HOMEPAGE E.

Frank Wijn

INDEX

Introduction to mass transfer operations 2

Introduction to trays, in general

Trays, in more detail

• Tray efficiency, in general

• Tray efficiency model

(including maldistribution, bypassing and mixing)

• Tray efficiency, flow regime and liquid height

Introduction

Tray types

Conventional (single- or multi-pass) Downcomer Trays.

Shell Calming-Section Trays.

Shell HiFi Trays.

UOP Multiple Downcomer Trays (MD-trays)

Tray model

General

A single contacting cell

One branch with one or more contacting cells

 Two branches with one or more ideally mixed cells

Multi-cell/Multi-branch model

Examples of 'large' column layouts

Conventional downcomer layout A

Conventional downcomer layout B

Shell Calming Section layout

Shell HiFi layout

UOP MD layout

• Summary of calculation results for 'Large' column examples.

Effect of Flow Regime and Liquid Height on Sieve Tray Efficiency

Operating range of trays

Hydraulics and Two Phase Flow regimes on trays

Tray pressure drop

Weeping, sealing & blowing: the lower operating limit(s)

Flooding, priming & entrainment: the upper operating limit(s)

Downcomers, Degassing and Bubble Coalescence.

Tray development; a bit of history

Gatekeeping on new activities on trays

Recently Developed Trays

Facilities for tray testing, development and trouble shooting

Links to tray related expertise, tray vendors, etc..

★
Advising Specialist (Petro-) Chemical Distillation & Oil Refining
Equipment
Trays for Absorption, Distillation, Stripping and Drying Columns

Tray efficiency, capacity, pressure drop, hydraulics and

gatekeeping
Introduction to mass transfer operations.

Absorption, Distillation, Stripping, Drying, Extraction are mass transfer operations.

These operations are widely used in various (petro-) chemical separation processes.
During such an operation either one (or more) component(s) in the vapour phase is
transferred to the liquid phase and/or from the liquid phase to the vapour phase. Some
examples are:

• Absorption is used for removal of gaseous hydrogen sulfide and/or carbon

dioxide and/or mercaptans from Natural Gas, Synthesis Gas, etc. (any process gas
stream, really) by dissolving them in a (reacting) liquid stream.
• Distillation is used in every oil refinery and also in many chemical manufacturing
plants in the separation and purification of the desired products. Distillation is the
most important separation technique, in general. Distillation columns are very
visible in the skyline of any refiny and many chemical plants.

• Stripping is the reverse of Absorption. Stripping of dissolved and contaminating

volatile organic components (VOC's) from (ground) water, is a good and actual
example.
• Drying (removal of water vapour) from Natural Gas by dehydrating liquids (i.e.
TriEthyleneGlycol) is an actual process example.
• Extraction is a liquid/liquid contacting operation, and is used in the petroleum
industry (to separate aromatics and aliphatic species) and in the pharmaceutical
industry (to recover penicillins).

These operations are usually performed in cylindrical columns. These columns come in a
wide range of sizes, their diameter ranges from 0.05 m (for a typical laboratory scale
column) to about 10 to 12 m (for the largest industrial columns) and their height ranges
from about 0.5 m to about 100 m (Yes!). The required contacting needed for the
 separation(s) is provided by filling these columns with packings and/or trays.
Packings (either structured or random) come in many types and sizes and so do
trays.

Commonly used tray types are:

Bubble cap trays,

Sieve tray,

Valve trays, etc.

Last Updated: 12 april 1998

HOMEPAGE E. Frank Wijn

Introduction to trays, in general.

Usually, trays are horizontal, flat, specially prefabricated metal sheets, which are placed
at a regular distance in a vertical cylindrical column.

Trays have two main parts:

1) the part where vapour (gas) and liquid are being contacted;
the contacting area and
2) the part where vapour and liquid are separated, after
having been contacted; the downcomer area.

Classification of trays is based on:

• type of plate used in the contacting area

• type and number of downcomers making up the downcomer area
• direction and path of the liquid flowing across the contacting area of the tray
• vapour (gas) flow direction through the (orifices in) the plate
• presence of baffles, packing or other additions to the contacting area to improve
the separation performance of the tray

Common plate types, for use in the contacting area:

• Bubble cap tray. Caps are mounted over risers fixed on the plate. These caps
come in a wide variety of sizes and shapes, round, square, rectangular (tunnel),
etc..
• Sieve trays come with different hole shapes (round, square, triangular, rectangular
(slots), star), various hole sizes (from ~2 mm to ~25 mm) and several punch
patterns (triangular, square, rectangular).
• Valve tray come in a variety of valve shapes (round, square, rectangular,
triangular), valve sizes, valve weights (light and heavy), orifice sizes and either as
fixed or floating valves.
• Combinations of these types are applied: a.o. plates with both sieve openings and
valves, as well as plates with both light and heavy valves.

Trays usually have one or more downcomers. Dual flow trays, i.e. trays without
downcomers, are used in exceptional (-ly fouling) services, only. Examples: the Stone &
Webster Ripple tray and the Shell Turbogrid tray.

Type and number of downcomers used mainly depends on the amount of downcomer
area required to handle the liquid flow. Single pass trays are trays with one downcomer
delivering the liquid from the next higher tray, a single bubbling area across which the
liquid passes to contact the vapour and one downcomer for the liquid to the next lower
tray.

This Nutter MVG-tray is an example of a single pass tray:

Trays with multiple downcomers and hence multiple liquid passes can have a number of
layout geometries. The downcomers may extend, in parallel, from wall to wall, as in

an example of a conventional two pass sieve tray:

The downcomers may be rotated 90 (or 180) degrees on successive trays,

as on the UOP MultiDowncomer tray:

The downcomer layout pattern determines the liquid flow path arrangement and liquid
flow direction in the contacting area of the trays.

Giving a preferential direction to the vapour flowing through the orifices in the plate will
induce the liquid to flow in the same direction. In this way, liquid flowrate and flow
direction, as well as liquid height, can be manipulated. In the past, this was exploited in
Exxon's Jet tray and is again being used on GHH's Perform Kontakt tray and BOC's
Expanded metal tray (developed by M.W. Biddulph at Nottingham University).

The presence of baffles, screen mesh or demister mats, loose or restrained dumped
packing and/or the addition of other devices in the contacting area can be benificial for
improving the contacting performance of the tray; viz. its separation efficiency. Various
examples of this (have been or still) are around;

• the Baffle tray, invented and developed by prof. Haselden at Leeds University,
U.K.
• the T-By baffle tray of UNI-FRAC Inc. (invented and developed by B.M. Parker
and T.J. Parker),
• the Duo-Sorb tray of Latoka Eng., Tulsa, Oklahoma, USA, (inventor H.O.
Ebeling),
• the Screen tray of Glitsch, Inc., Dallas, Texas, USA (invented and developed by
K.T. Chuang et c.s. at Univ. of Alberta, Edmonton, Alberta, Canada),
 • etc.

The most important parameter of a tray is its separation performance. The curve of
distillation tray efficiency versus vapour flowrate determines the operation of a trayed
distillation column. Commonly, this curve is actually obtained in a sufficiently large
experimental distillation column (such as those of Fractionation Research Inc, Stillwater,
OK, USA) with a binary testmixture. As an example:

The efficiency curve for Nutter Engineering's MVG tray

Four parameters are of importance in the design and operation of a trayed column:

• the level of the tray efficiency, in the normal operating range

• the vapour rate at the 'upper limit', i.e. the maximum vapour load
• the vapour rate at the 'lower limit', i.e. the minimum vapour load and
• the tray pressure drop (not shown).

Repeatedly, research and development efforts are being aimed at improving the
separation performance of trays, as well as their maximum allowable vapour rate.
This leads to a progressive improvement in the production capacity of trayed columns.
At the same time, tray pressure drop and minimum allowable vapour rate have to be
made as low as possible (to minimize energy consumption).

Last Updated: 2 june 2001

HOMEPAGE E. Frank Wijn

 Trays, in more detail.

In the pages offered here, the subject matter will be dealt with in more detail and
becomes more complicated. The matter presented in these pages represents the 'state of
the art' (... as perceived by Frank) and may reflect an uneven level of understanding of the
underlying phenomena. Usually, this is the current 'state of affairs'. From time to time,
these pages will be updated with information about new insights and developments.
Frank will be very grateful for any material being sent to him, that assists him in creating
a Web-site of value to any visitor, interested in this field of science and technology.
Please, do not hesitate to sent him your contributions, remarks, comments, etc..

• Tray efficiency, in general

• Tray efficiency model
(including maldistribution, bypassing and mixing)
• Tray efficiency, flow regime and liquid height
• Operating range
• Hydraulics, flow regimes
• Pressure drop
• Weeping, the lower operating limit(s)
• Flooding, the upper operating limit(s)
• Downcomer behaviour

• Tray technology, historical development

• Recent trends & activities on trays
• New trays, recent developments
• Facilities for testing and developing trays
• Courses on the design of trays and columns
• Links to tray expertise, vendors, etc.

Last Updated: 29 june 2001

 Tray Efficiency, in general

The separation performance of a tray is the basis of the performance of the column as a
whole. The primary function of, for instance, a distillation column is the separation of a
feed stream in (at least) one top product stream and one bottom product stream. The
quality of the separation performed by a column can be judged from the purity of the top
and bottom product streams. The specification of the impurity levels in the top and
bottom streams and the degree of recovery of pure products set the targets for a
successful operation of a distillation column.

Because of its importance, a vast literature exists on the separation performance of

distillation trays. Roughly, more than half of the more than 1600 publications in journals,
conferences, books, etc.deal with the mass transfer performance, i.e. the tray efficiency.

Earlier developed mathematical models for the separation behaviour of distillation

columns used the concept of equilibrium stages (theoretical trays) and incorporated the
tray efficiency, either as the Fenske Overall efficiency (Eo) or the Murphree vapour
efficiency (Emv). Newer computing models (like ChemSep™ and Aspen Ratefrac™)
use non-equilibrium stages i.e. the rate based approach, which is a better more general
approach, esp. for multicomponent systems. This approach uses information on flowrates,
mass transfer coefficients, interfacial area and liquid hold up, which can conveniently be
combined in one variable prozaically called; the NTU (Number of Transfer Units). The
rate based approach is more computationally intensive, however, and also requires much
more input information. At the moment, both approaches are in use, while the rate based
approach appears to be gaining ground on the traditional equilibrium stage method,
because its reliability is increasing.

In due course, typical tray efficiency values have become available from plant test runs
and various experimental programs (usually executed with well-defined binary
testsystems). Perry's Chemical Engineers Handbook (table 18-4) gives representative
plate efficiencies, Emv's, which all vary between 60 - 120 %. These typical efficiency
values depend to a more or lesser extend on the test systems used, the tray type and
geometry tested, the liquid submergence, etc., etc.. Picking the right value for a new
column design is still largely based on experience, but may be underpinned by one (or
more) of the following methods:

• data interpolation (from previous tests/experience and/or from published

literature).

• direct scale up from efficiencies measured in a 1) Oldershaw column; 2) pilot

plant column; 3) existing commercial column.
• empirical prediction methods (Drickamer & Bradford; O'Connell; Bakowski).
 • theoretical prediction methods (the procedure given in AIChE Bubble Tray
Design Manual and the Chan & Fair correlation).

When test data are available (or become available), the first two methods are preferred,
because of the 'proven' nature of the numbers involved and because the uncertainty in
these methods is the least.

Summarizing available prior experience, it has become evident that tray efficiency (or
NTU) can be influenced by:

• the specific component under consideration (this holds specially for multi-
comp. systems in which the efficiency can be different for each component,
because of different diffusivities, diffusional interactions, different stripping
factors, etc.).

• the vapour flowrate; usually increasing the flowrate increases the effective
mass transfer rate, while it decreases the contact time, at the same time. These
counteracting effects lead to a roughly constant efficiency value, for a tray in its
normal operating range. Upon approaching the lower operating limit a tray starts
weeping and looses efficiency. The tray also looses efficiency upon approaching
the upper operating limit, because of droplet entrainment.
• the liquid flowrate; usually increasing the fowrate increases the tray liquid
hold up and contact time and hence tray efficiency.
• the outlet weir height; higher weirs raise the liquid level, increase the
interfacial area and contact time giving improved efficiencies.
• the operating pressure; tray efficiency increases with pressure. This apparent
pressure effect may be a reflection of the increase in liquid hold up (either
because of relative larger liquid flowrates or relatively less weir lengths).
• the type of tray appears to be of secondary importance, as tray efficiencies of
valve and sieve tray are quite comparable, in their normal operating ranges.
• the free area (fractional hole area); efficiency increases with a reduction in free
area.
• the liquid flow path length; longer path lengths enhance efficiency, as liquid
mixing is less able to flatten out the axial liquid composition profile developing,
across the entire tray.
• the presence of non-uniform flow distributions for the liquid and/or vapour
flow reduce tray efficiency .
• the presence of dead water regions (stagnant zones) is detrimental to tray
efficiency, especially so on trays were these regions on successive trays "stack
up" (because they lie directly above each other) and allow part of the vapour flow
to effectively bypass contact with liquid across more trays.
• the surface tension and surface tension gradients; a lower surface tension
increases the interfacial area and hence efficiency. Surface tension gradients can
either lower or enhance efficiency depending on the sign (direction) of the
gradient. In the liquid continuous flow regime, a positive gradient can enhance
 efficiency (reduces bubble coalescence and increases interfacial area). A negative
gradient diminishes efficiency for the opposite reasons.
• the liquid viscosity; there is no conclusive evidence for the effect of liquid
viscosity on tray efficiency, in distillation. In absorption, this situation can be
different for components for which the mass transfer becomes controlled by liquid
phase mass transfer.

There are still a number of Unknowns remaining, i.e. topics on which consensus has
not yet been reached. For instance; the best way to correlate the vapour phase mass
transfer coefficient; the mechanisms determining the amount of interfacial area available
for mass transfer; the 'stubborn' use of the inadequate Francis' weir equation in the
calculation of liquid hold up on a tray operating in the spray regime; the estimation of the
contribution of the liquid phase mass transfer resistance; and last but not least the
influence of the two phase flow regime.

It will be appreciated from the above, that the estimation of distillation tray efficiency is
wrought with uncertainties and inaccuracies. Hence, confidence in the validity of
available data is important.

Because of the considerable commercial value attached to these efficiency numbers, the
proper numbers have been (and are) slow in becoming available, publicly. Either they are
considered to be proprietary know how or are kept confidental (as company-sponsored
FRI does), because of the high costs involved in obtaining them. This has slowed down
the scientific development of fundamental theories of mass transfer on trays (which could
have been of benefit to the companies involved, as well ... ).

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Multi-cel/Multi-branch Tray Efficiency Model

This 'page' gives a summary of the model. A more detailed description of the model
and its development can be found in the paper presented at the Studiedag Trends in
Destillatie, Utrecht, 26 october 1995 and in The Chemical Engineering Journal, 63
(1996), pp 167-180.

Contents:

• Introduction
• Tray types
o Conventional (single- or multi-pass) Downcomer Trays.
o Shell Calming-Section Trays.
o Shell HiFi Trays.
o UOP Multiple Downcomer Trays. (MD-trays)
• Tray model
o General
o A single contacting cell
o One branch with one or more contacting cells
o Two branches with one or more ideally mixed cells
o Multi-cell/Multi-branch model
o Examples of 'large' column layouts
ƒ Conventional downcomer layout A
ƒ Conventional downcomer layout B
ƒ Shell Calming Section layout
ƒ Shell HiFi layout
ƒ UOP MD layout
• Summary of calculation results for 'Large' column examples.

Move back one page

Introduction:

The two most important parameters in the design of a (distillation or absorption) column
are:
 - the maximum capacity (througput) of gas/vapour and liquid of a tray
and
- the tray efficiency (separation performance).

The maximum gas/vapour capacity (at the upper limit) dictates the size of the contacting
area and the maximum liquid capacity (at the upper limit) dictates the downcomer area,
which are needed in a specific column design. The total column area is obtained by
adding up the contacting and downcomer area and so the column diameter is determined.

The tray efficiency governs the number of actual trays needed to achieve the desired
product purity specification(s). With highly efficient trays one can install a lower number
of these highly efficient trays to achieve the separation desired.

Several tray efficiency definitions in use.

Three (different) efficiencies are useful, in particular:

- the overall efficiency (Eo) (the socalled: Fenske efficiency),

- the average tray efficiency (first defined by Murphree) (Emv) and
- the local (or point) efficiency (Ep).

Historically, tray efficiency values have been experimentally derived by performing test
distillations with binary test systems and from plant experience (plant testrun data) for a
wide variety of physical systems, tray types and tray dimensions.

The last few years, the modelling of tray efficiency was revisited, with the aim to gain a
better insight of the effect of tray-layout (especially the distribution of the downcomers
across the column area).

Top of File

Tray types:
 Conventional Downcomer Trays.

This is the most common type of tray layout, which can have a single pass for the liquid
over the tray or multiple passes. The example shown has two passes for the liquid
flowing over the tray.

Shell calming section

tray

Shell Calming-Section Trays.

The 'Calming Sections' are a special type of downcomers developed over the period of
approx. 1955 to 1965 by F.J. Zuiderweg and coworkers. The rectangularly shaped
'Calming Sections' are relatively narrow and short, because of this they can be distributed
more easily across the tray area. Typically, they are installed at some 0.5 to 2.5
CS's/sq.mtr, depending on liquid load.

Shell HiFi Tray

 Shell HiFi Trays.

This layout was developed for use at high liquid loadings (high flowparameters, hence:
HiFi) and the relevant patent mentions the name of A.D. Freije as its inventor (1973).
Several narrow and long downcomers run in parallel across one-half of the column
diameter. The other half of the column contains another set of parallel downcomers,
which are placed at intermediate positions with respect to the downcomers on the first
half. On subsequent trays, these tray halves will be mirror-images. As these 'HiFi'-
downcomers are placed in parallel, the flow path lengths are uniform, although they may
be short.

UOP Multiple Downcomer Trays (MD-trays).

This tray layout was originally developed by Union Carbide (W. Bruckert, 1968). Later
Union Carbide transferred its tray technolgy to UOP, which is actively developing and
marketing these trays, nowadays. More recently, they have introduced Enhanced
Capacity and Enhanced Efficiency versions of the original MD-layout (which is shown).
Note, the downcomers on subsequent trays are rotated by an angle of 90 degrees. This
downcomer rotation forces the liquid to flow across the tray in a 90 degree turn, as well.

Top of File

Tray model:

Any tray mass transfer model is a combination of a flow-structure with mass transfer rate
relations (usually specified by Nog's) and V/L-equilibrium relations. The favoured
'network' structure make use of 'unit'-cells. These 'unit'-cells are put in series in a branch
(or zone) and tow or more of these branches (zones) are put in parallel. This structure
provides several degrees of freedom in the specification of:

1) the number of contacting cells (stages) in series in each branch

(zone)
2) the distribution of liquid flow into the various branches, and
3) the distribution of vapour flow transversally across the branches
and
longitudinally in the branches.

The additional advantage of such a cellular/zonal flow structure is, that it is relatively
easy to create an effective and fast computer program. Because the relevant mass transfer
relations and equilibrium relations are included in an analytical description of the
individual cell, which can be called over and over again as a subroutine in the
programme.

General

A PC-based computer programme simulating the multi-branch/multi-cell model was built

up and validated in several steps. The first (1) step was the construction and validation of
the subroutine describing the basic contacting cell, followed by (2) extending to any
number of cells in a branch. By expansion (3) to two branches, it became possible to
check for the already known effects of bypassing and maldistribution, before doing
calculations (4) for an arbitrary number of branches with an arbitrary number of cells.

A single contacting cell

The contacting cell is modelled in similarity to the vapour/liquid flow commonly

observed around a vapour jet (or bubble-chain) issuing from a hole in a tray. On an
overall longitudinal flow of liquid across a tray, the vapour jets create a (up & down)
recirculating liquid flow in the dispersion on a tray. A high-velocity vapour-jet entrains
the liquid phase and contacting of vapour and liquid takes place cocurrently, locally.
After contact, the liquid returns to the tray floor and may recycle around the same
vapour-jet again, or flow to a nearby vapour-jet. In the direction of the overall liquid
flow, this may be the next one forward (or backward: resulting in locally retrograde
flow). Similarly, part of the liquid is allowed to recycle internally in a cell, to the cel-
inlet. However, to keep the model relatively simple, the possibility of forward and
backward (as well as sideways) inter-cellular transport will be excluded, for the moment.
The only liquid flow from cell to cell is the nominal liquid flow, that proceeds along the
chain of cells from inlet to outlet. By retaining the internal recycle, we will be able to
explicitly take into account the effect of liquid mixing in the vertical direction.
Although this approach to modeling the local mass transfer relations is unusual, it is not
new in itself. Berkovskii, Aleksandrov, Skoblo et Sheinman (1971) and Malafe'ev and
Malyusov (1971) have been pioneering this approach, in the development of cross- and
co-current contactors. Also, Standart, Bragg, Uddin El Yafi and Yaroson (1979) proposed
to use this concept in the modeling of point efficiency on distillation and absorption
cross-flow trays.

For a single contacting cell, a detrimental effect of the inernal liquid recycle in the cells
shows up, because the returning liquid has already been in contact with vapour. This
reduces the driving force for mass transfer and hence efficiency. With an increasingly
larger recycle ratio (Rrec), the liquid phase on the tray becomes increasingly better
mixed, vertically. It was expected that the point efficiency should become constant at a
sufficiently large recycle ratio. For Rrec > 20 this behaviour is observed. The limiting
value agrees with the analytical equation for a vertically ideally mixed liquid phase.

With no recycling of liquid (Rrec = 0), the tray is operating as one cocurrent contactor
with plug flow in both vapour and liquid phase. In this special case, the point efficiency
varies with the stripping-factor(S = K.V/L) and the number of overall vapour phase
transfer units (Nog). This additional influence of the stripping-factor is a distinctive
feature of the cocurrent contacting cell and has consequences for the calculation of the
efficiency of high stripping-factor components in a multicomponent mixture.

At intermediate recycle ratios (0 < Rrec < 20), the point efficiency lies (depending on
Rrec) between the upper plug flow limit, for the specified S, and the lower limit for
ideally mixed liquid.
One branch with one or more contacting cells

The operating principle of one branch with a series of ideally mixed cells (Rrec > 20)
across a tray is equivalent to the case choosen originally by Gautreaux & O'Connell
(1955). With the number of ideally mixed cells in series going to infinity, the solution
given by Lewis (1936) for the limit of liquid in plug flow is reproduced. For a low, finite
number of cells in series, the same efficiency values are calculated as given by the
Gautreaux & O'Connell equation.

Two branches with one or more ideally mixed cells

With two branches in parallel, it becomes possible to specify an arbitrary number of cells,
in each of the zones. Maldistribution of vapour and/or liquid flow can be included in a
two brnach system by increasing the flow in one zone by a specified amount and
decreasing the corresponding flow in the other zone by the same amount. In a similar
way, liquid bypassing is modelled by assuming that all vapour flows through one zone
and that a part of the liquid flow is diverted from this contacting zone to the second
branch, being the bypass. For the case with liquid bypassing a contacting branch with
liquid in plug flow though it, the same values as given by the equation of Savel'ev,
Sabitov & Nikolaev (1979) are retrieved. When about 40% of the liquid bypasses, the
tray efficiency drops to a value equal to or below that of the point efficiency and the
beneficial liquid cross-flow effect (staging) has been lost. Similarly, the effect of vapour
bypassing can be modelled. For the case of vapour bypassing a contacting branch with
liquid in plug flow the values of the equation of Strand (1963) are reproduced. Vapour
bypassing significantly deteriorates tray efficiency. The percentage vapour bypassing,
that can be allowed before tray efficiency reduces below the point efficiency is about 25-
35%.

Multi-cell/Multi-branch model

A multi-branch structure on a conventional single pass side-downcomer tray is pictured

below. The tray part between the inlet and outlet downcomer is characterized by two side
segments and a rectangular middle part. As it is assumed that there is no mixing between
branches and liquid is fed in only to the middle part, the side segments are starved of any
flowing liquid and hence develop into stagnant zones, through which part of the vapour
flow will be bypassing.
A (rectangular middle part) ten branch structure has been used to model the effect of
transversal liquid flow maldistribution on tray efficiency. Several liquid inlet flow
distributions have been studied (lineair, parabolic and Gaussian shaped). As these results
were essentially similar when presented in a generalised graph of an efficiency
enhancement ratio, it will suffice to show the results for the lineair inlet liquid flow
distribution. The degree of maldistribution is characterized by the parameter MDl,
defined as the relative standard deviation in the liquid flow rate. A good agreement
between calculations for a ten branch and a two branch model has been found. Also, a
good agreement was obtained with calculations done by Stichlmair & Ulbrich (in 1987).

Please be aware, that at a severe degree of maldistribution the tray efficiency can drop
below the value of the point efficiency; Emv/Ep < 1. Also important to observe, that a
substantial amount of liquid maldistribution (20% - 30%) can be accommodated, before
the tray efficiency deteriorates markedly.
From the above shown generalized graph of the efficiency enhancement ratio, a graph for
a specific stripping factor (S) can be extracted, as shown below. The graph demonstrates
the increasing difference between the tray averaged efficiency (Emv) and the local (point)
efficiency at increasing values of Ep, especially at low degrees of liquid maldistribution.
Examples of 'Large' column layouts:

Positioning of the downcomers dictates both the location of the feed-point and the
discharge-point of the liquid on a tray. Consequently, it governs the path taken by (parts
of) the liquid-flow in crossing the 'active area' of the tray and hence the amount of vapour
being seen by (parts of) the liquid flow during its crossing. The distribution of liquid flow
is assumed to be set by the amount of of downcomer outlet area of the next higher tray,
discharging liquid into the specific branches. The distribution of vapour flow,
transversally across the branches, is based on the amount of 'active area' included in the
liquid flow paths, which are assumed to sweep across the tray. The longitudinal
distribution of vapour flow in the branches is assumed to be uniform. The flow path
length of each branch may be different, because of a different number of contacting cells
in series. However, in all calculations that follow the length of a contacting cell is kept
the same, in this way putting the results on a comparable basis.

Conventional Downcomer Trays.Two conventional downcomer 4-pass layouts are

considered, because of the reversal in liquid flow direction on subsequent trays. Because
in both cases each half of the column cross-section is a mirror-image of the other half,
only one half of the tray needed to be modeled. This has been achieved by using 16
branches, which contained a total of 110 cells.
Layout (A):

Layout (B):

Shell Calming-Section Trays.

For the Shell Calming Section layout pattern, it was sufficient to model only one half of
the tray (15 branches, with a total of 103 cells). The colored cells identify the cells
recieving liquid discharging directly from the overlying Calming Sections, in the next
higher tray.

Shell HiFi Trays.

The Shell HiFi downcomer layout pattern could be modeled with one half of the tray, as
well.
UOP Multiple Downcomer Trays (MD-trays).

For the Multi-Downcomer layout pattern, commercialised by UOP, Inc., only one
quadrant needed to be modelled, because of the two-fold symmetry around two
centrelines, which intersect at 90 degrees. This symmetry arises from the rotation by 90
degrees of the inlet downcomers (gray lined) with respect to the outlet downcomers
(black lined) on the next tray. This rotation also forces the liquid to flow across the tray in
a 90 degree turn. (The differently colored cells help in identifying the liquid flow paths.)

Top of File

Summary of calculation results for 'Large' column examples.

The calculated (= predicted) results are tabulated and ranked below, at one specific
combination of stripping factor (S) and point efficiency (Ep). The reason for the ranking
of a particular layout pattern can be gleaned from three parameters listed in the table, viz.
the average flow path length (Nc,avg) and the maldistribution coefficients for the liquid
and the vapour flow distribution across the branches (MDl and MDv). The familiar
relationship between tray efficiency and flow path length can be recognised, however
being modified by maldistribution of either the liquid or the vapour flow (or both at the
same time). The examples show that depending on the number, type and size of
downcomers employed, different tray efficiencies will be obtained. As the design
procedures allow for a substantial degree of freedom in the design of trays, it may be
appreciated at this point, why the development of an efficiency calculation tool, enabling
the best choice of tray layout, was important.

Assumptions made:
- Ideally mixed vapour in between the trays (hence: Lewis Case I)
- Refreshment of liquid in stagnant zones
- S = K.V/L = 1.000
- Rrec = 1000 (hence: ideally mixed cells)
- Nog per cell = 1.000
- Ep, point efficiency = 0.632

Emv Emv/Ep Nc,avg MDl

MDv
Reference values:

Liquid in plug flow 0.882 1.395 infinite 0.00

0.00

Liquid ideally mixed 0.632 1.000 1.0 0.00

0.00

Examples of 'Large' column downcomer area distribution patterns:

Shell Calming Section tray 0.796 1.260 6.9 0.24

0.18

4-pass Conv. Downcomer tray (A) 0.784 1.240 3.4 0.24

0.25

4-pass Conv. Downcomer tray (B) 0.772 1.222 3.4 0.13

0.25

Shell HiFi tray 0.716 1.132 1.9 0.07

0.08

UOP MD-tray 0.665 1.052 2.1 0.24

0.46

General comparison of 'Large' column examples.

The ranking of layouts obtained at one specific combination of stripping factor and point
efficiency is valid for other combinations, as is shown conclusively in the next
generalised graph.

These calculations have the character of predictions, because of the simplifying

assumptions and the specific flow distributions used. They are also a worst case analysis,
because of the neglection of transversal exchange of material (mixing) between branches,
which could have transversally smoothed out the developing concentration differences
between the branches. So, experimental validation is needed, if not for evaluating the
layouts themselves, than for testing the validity of these predictions and their underlying
assumptions.

Last Updated: 12 april 1998

This 'page' gives a summary of the model. A more detailed description of the model
and its development can be found in the paper presented at the Studiedag Trends in
Destillatie, Utrecht, 26 october 1995 and in The Chemical Engineering Journal, 63
(1996), pp 167-180.

Top of File

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HOMEPAGE E. Frank Wijn

Effect of Flow Regime and Liquid Height on Sieve Tray Efficiency

This 'page' gives a summary of the more detailed paper presented at the Distillation
& Absorption Conference at Maastricht on 8-10 september 1997.

Move back one page

HOMEPAGE E. Frank Wijn

Introduction.

On the basis of test data, it has been proven to be surprisingly difficult to show that there
is a link between flow regime, liquid height on the tray and sieve tray efficiency. On
theoretical grounds, one expects that there is such a link, if only because the interfacial
area for mass transfer depends on the amount of liquid on a sieve tray, on the bubble and
drop sizes and on the gas and liquid volume fractions in the dispersion.

The paper by Porter, Yu Chambers and Zhang presented at the 1992 Distillation
Conference in Birmingham was the trigger to revisit this subject. Porter's group operates
a unique large scale test facility at the Aston University, so their data merit careful
consideration. Their large scale sieve tray efficiency data, obtained by the water cooling
technique, were especially welcome because they allowed comparison with other test
data available in the literature.

One reference of special importance was the paper by Prado and Fair presented at the
1987 Distillation Conference in Brighton. In that paper they describe a sieve tray
efficiency model, which was validated against efficiency test data, that were obtained by
humidifying air.

Apart from using the same test system both groups used an almost identical sieve tray
layout. So, one would expect similar sieve tray efficiencies at comparable operating
conditions, as well.
As the flow path length was quite short, Prado and Fair measured point efficiencies on
their sieve tray. Porter, Yu, cs. experimented on in a 8 ft. diameter column. They
measured the efficiency for the tray as a whole. They also measured flow fields on their
tray (by temperature sensors, static liquid height sensors, etc.) and were able to reduce the
efficiency for the entire tray to point efficiency values.

Some experimental data.

When the point efficiencies obtained by Prado & Fair and Porter, Yu, c.s. are plotted
versus the vapour load factor (which essentially is the vapour velocity times the square
root of the ratio of vapor density over liquid density) it is immediately evident that there
is a wide scatter in the data.

The highest point efficiency values found by Porter, Yu, c.s. agree with the values
obtained by Prado & Fair. The trend in the efficiency values of Porter, Yu, c.s. were
found to correlate with weir height and weir liquid load. The lowest efficiencies were
found when a tray was tested with a low weir height and was operated at low weir liquid
loads. This was pointing to an influence of the liquid height on the tray (submergence of
the holes). In the original paper these liquid height values were not provided. However,
Prof. Porter has been so kind to provide me with these data and I gratefully acknowledge
his permission to use these data.

This graph shows that the two data sets appear to 'line up' when plotted versus the liquid
height. To obtain a maximal efficiency at least some 30 mm of liquid is needed on the
tray. The meaning of the two different efficiency ranges was not immediately recognized,
however.

A possible reason for this came in view when the ranges of the test conditions of the two
groups was plotted in the flow regime plot published by Muller & Prince (1972). These
authors performed a classic study for a single 1/4" hole (6.35 mm diam.) with the (same)
air/water system. They varied the submergence systematically and obtained the gas rates
corresponding to changes in the observed hydraulic behavior at the hole. This graph
summarized their results.
As can be seen the data set of Porter, Yu, c.s. falls in the 'pulsating jet' regime and the
data set reported by Prado & Fair falls almost completely in the 'deformed bubbling'
regime. This result shows that each group operated preferentially in a different flow
regime.

That being the case, than what is causing the difference in efficiencies ?

The answer appeared to be a change in the mechanism producing the interfacial area on a
sieve tray. This can be shown by a graph reported by the Russian authors Ulyanov,
Rodionov, c.s. (1982). This figure provides the real key to the resolution of the efficiency
puzzle.
At low submergence, the interfacial area a increases with submerge and at high
submergence decreases again. At the lowest liquid height, the vapour tends to blow fairly
stable 'holes' or 'jets' in the dispersion, which give a small interfacial area. At increasing
liquid heights, these 'jets' increasingly tend to 'pulsate' (i.e. alternate between jetting and
bubble formation), which produces an increasing (time averaged) interfacial area. At a
sufficiently high liquid height, the steady formation of bubbles has reached its full
development. A still higher liquid heights, coalescence of the rising bubbles again
reduces the (volume averaged) interfacial area.

With this dependence of interfacial area on submergence it can now be estimated how
sieve tray efficiency should behave.

Some calculations.

As usual, it is assumed that the liquid phase is well mixed and that the vapour passes in
plug flow through the liquid. The point efficiency Ep is related to the number of overall
gas phase mass transfer units by: Ep = 1 - exp(-Nog). As the mass transfer rate during the
evaporation of water is completely gas side limited: Nog = Ng and the gas side number of
mass transfer units is given by: Ng = kg.al.Hl/Vg (kg is the gas side mass transfer
coefficient; al is the interfacial area per unit volume of liquid; Hl is the liquid height; Vg
is the gas velocity in the contacting area).

Two types of calculations for the number of mass transfer units Ng have been made:

• kg.al/Vg is kept constant and only the liquid height Hl is varied (dashed lines),

• kg/Vg is kept constant and both Hl and al are varied (fully drawn lines).
The expected dependency of efficiency on submergence becomes:

The difference between the two sets of three curves is caused by the extra effect of liquid
height on interfacial area. The fully drawn curves, including both the influence of both Hl
and al, agree with the earlier figure containing the test data of Prado and Fair and of
Porter, Yu, c.s.. This confirmed the effect of changes of liquid height and flow regime on
sieve tray efficiency.

Some more evidence.

Could more information be found in the literature in support of this conclusion ?

A French group of investigators (Nagy, Laurent, c.s.) also reported in 1986 that
interfacial area on a sieve tray declined with increasing dispersion height (which confirms
the trend at high submergence). They obtained their interfacial area data by measuring the
absorption rate of carbon dioxide in an aqueous alkanol amine solution (the Danckwerts
technique).
Another example was reported by Pohorecki who studied the absorption of carbon
dioxide in an aqueous and ionic solution of carbonate and bicarbonate using the same
technique. His data show the declining trend at high submergence, also.

But, what about the behavior at low liquid heights ? can that be confirmed ?
Two examples could be located which show this effect.

One example was reported in 1964 already, by a British group studying sieve tray
efficiency in distillation. In this particular example only the range at low liquid heights is
visible. The trend they found confirms the expected effect of submergence on efficiency.

The second example was brought up during a discussion session in the 1969 Distillation
Conference in Brighton by Shore. (Note, that he gives the liquid height in inches.) In his
rather unique study both submergence ranges are visible, already.

So, for both the low and the high submergence range, confirmation could be found in the
literature for the expected effect of liquid height on sieve tray efficiency.
Conclusions.

• Sieve tray efficiency depends on flow regime.

• Sieve tray efficiency depends on liquid height (submergence).

• The flow regime chart of Muller and Prince is useful.
• The data set of Porter, Yu, c.s. falls in the pulsating jet regime.
• The data set of Prado and Fair falls almost completely in the deformed bubbling
regime.

Last Updated: 12 april 1998

This 'page' gives a summary only of the more detailed paper presented at the
Distillation & Absorption Conference at Maastricht on 8-10 september 1997.

Move back one page

HOMEPAGE E. Frank Wijn

 Operating range of trays

The successful operation of a refinery and chemical plant depends on the operability of
its separation columns. The range of operation of the installed trays governs the
maximum and minimum gas and liquid loads a column can handle and consequently the
capacity and turndown capability of a plant.

In the illustrative operating diagram below, an operating line as well as an operating

point are indicated.

Stop... Don't hit that button

There's good shit down
below!!!!!

For both the vapour and the liquid flow rate, lower and
upper limits exist. Several different hydraulic mechanisms
control these limits. The operating point of a column
should be chosen by carefully considering these
limitations.

The physical mechanisms behind the most common limitations are:

• Entrainment flooding is caused by an excessive liquid flowrate generated by

droplets carried out of the gas-liquid dispersion on the tray and up to the next tray
by the gas stream. Preferentially, this occurs on trays operating in the spray
regime, when most of the liquid is present as droplets.

• Bed expansion flooding (priming) sets in when the gas-liquid layer on the tray
extends up to the next higher tray, i.e. the dispersion height becomes equal to the
tray spacing. This condition depends on the liquid height on a tray and its
expansion, by the gas flowing through it. Actually, this condition is equivalent to
entrainment flooding. As the dispersion level approaches the next higher tray,
drop entrainment rises rapidly, causing accumulation of liquid on the next tray
and increasing the liquid flow, for the downcomers to cope with.
• Choking (blockage) of the downcomer entrance can set a limit to the liquid
flowrate, when the inflowing gas-liquid dispersion needs more space, when a
large amount of gas is set free in the downcomer. The separated gas rises counter
 to the downflowing gas-containing liquid and can induce a (vapour) blocking
condition.
• Downcomer overflowing (flooding) sets a limit to the liquid handling capacity
of a downcomer. This happens when the height of the gas-liquid layer in a
downcomer exceeds the height of the downcomer. This condition is controlled by
the height of liquid in the downcomer and by the expansion caused by entrapped
gas bubbles. The liquid height is given by the condition that the hydrostatic liquid
head should be sufficient to make the liquid flow down against the tray pressure
drop.
• Weeping (raining) of liquid occurs when gas flowing through the perforations
in the tray floor is no longer able to counterbalance the hydrostatic head of liquid
on a tray and liquid starts leaking.
• Sealing (dumping) of the downcomers requires, that at least some liquid flows
through them. This condition is fullfilled when the gas-liquid layer on a tray is
sufficiently expanded by the gasflow to make liquid flow over the outlet weir.
• Blowing occurs when insufficient liquid flows through the downcomers to
prevent gas from bypassing up through them. This is detrimental to the separation
performance of a tray, because a significant part of the gas bypasses the
contacting area of the tray in this way. During the (re-) start up phase of a column,
this phenomenon needs special attention, as the penalty will be a severe
underperformance of the column (in both throughput and separation
performance).

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Hydraulics and Two Phase Flow regimes on trays

The full spectrum of all possible two phase flow regimes can be encountered on trays,
depending on flow conditions, system properties and tray geometry. In general, four two
phase flow regimes are recognised to classify the flow patterns in the contacting area:

• Bubbling; gas bubbles in liquid, < ~ 50 % gas content, wide bubble size
distribution.

• Foaming; gas bubbles in liquid, > ~ 65% gas content, fairly uniform bubble
sizes.
• Churn turbulent / Compounded regime; ~ 20 to 60 % liquid content, at
bottom gas in liquid and at top droplets in gas.
• Spray; droplets in gas, < ~ 25 % liquid content, wide drop size distribution.

The Bubbly Flow regime occurs at low(er) gas velocities and relatively thick liquid
layers (while bubbles or jets forming at the tray floor are fully submerged). The bottom
layer has the highest liquid content (may be 80-90 % liquid) and the formation of bubbles
or jets at the orifices can be seen clearly. The upper layer has a higher gas content and is
in constant chaotic motion. The bubble population has a wide size distribution; large
bubbles may be a few centimeter (~0.02 - 0.05 m) in diameter, while small bubbles
measure a few millimeter ( ~ 0.002 - 0.005 m). Chaotic (turbulent) motion causes large
bubbles to break up, while coalescence of colliding bubbles makes them grow. The wide
bubble size distribution is the result of a dynamic equilibrium between break up and
coalescence. The chaotic motion is caused by the rising bubbles and makes the surface
level fluctuate constantly, as well.

The bubbly flow regime can be found:

• in high pressure absorption or distillation; over whole operating range.

• in medium pressure distillation; from the lower limit up to somewhere in the

middel of the operating range.
• in subatmospheric or vacuum pressure distillation: only at operation near the
lower limit.

The bubbly liquid flows over the outlet weir into the downcomer entrance. On its way,
the liquid is degassing, as bubbles keep escaping from the liquid phase. The gas content
of the liquid falling into the surface level in the downcomer will depend on the rate of
degassing and its original gas content (on exit from the contacting zone). This degassing
process causes an upflow of gas out of the downcomer. This upflow of gas may interfere
with the downflow of liquid. Particularly, when the dispersion height becomes tall in
comparison to the horizontal width of the downcomer (the available distance for the
throw of liquid becomes too small) and choking of the entrance occurs.

The two phase dispersion sitting in the downcomer has two strata, usually. A clear
bottom layer with an upper bubbly layer. In the upper gas-liquid layer, bubbles come in
with the falling liquid and rise up and out of the dispersion, again. Thereby generating the
upflow of gas out of the downcomer.

The Foaming regime, also, occurs at low(er) gasvelocities and relatively thick liquid
layers (while bubbles or jets forming at the tray floor are fully submerged). The
difference with the Bubbly Flow regime is caused by inhibition of bubble coalescence, as
break up of liquid films in between the bubbles is retarded by the presence of surface
tension gradients, very fine solids, a high liquid viscosity, etc.. This leads to an increased
residence time of the gas in the dispersion, thus giving rise to a high gas content. The
degassing process is inhibited as well.

Commonly, the bottom layer is fairly clear. However, for fairly strongly foaming
systems, it can turn into a 'milky' emulsion containing many tiny bubbles ( < ~ 50 % gas
content).

The mobile foam layer on top contains a fair amount of motion. Its structure is that of a
typical foam with fairly stable liquid lamella in between bubbles, of various sizes. The
gas content can be quite high, up to 90 - 95 %.

The foamy liquid falling over the outlet weir into the downcomer entrance takes along a
large amount of gas, as a result of the inhibited bubble coalescence and hampered
degassing. A much larger amount of gas now becomes available from the degassing
process in the downcomer. This gives rise to a large(r) upflow of gas out of the
downcomer, which can interfere even more strongly with the volumetricly enlarged
incoming foamy liquid. Thus leading to earlier inducement of choking phenomena and
hence premature flooding of the downcomer.

The two phase dispersion sitting in the downcomer again has two strata, but their relative
contributions differ. The upper bubbly (foamy) layer is enlarged, as an enlarged gas flow
has to be separated, while the separation takes more time, as well. Very small bubbles
may not separate at all and can be 'carried under', with the liquid leaving the downcomer
exit.

The Churn turbulent regime (or Compounded regime) might also be called a
'sandwiched' regime, as it is intermediate to both the bubbly regime and the spray regime.
In the structure of this two phase dispersion, these regimes can still be identified: the
bubbly regime in the bottom layer and the spray regime in the upper layer. Often, an
approximately uniform intermediate layer is present, as well. The dispersion is very
chaotic and full of fluctuations of many size and frequencies scales. The 'bubble' size
distribution is even wider than in the bubbly flow regime, as now large clusters of
bubbles or even 'slugs' (large voids) can be seen rising at high velocity through the two
phase layers.

The churn turbulent regime can be found:

• in high pressure absorption or distillation; over the whole operating range.

• in medium pressure distillation; from somewhat above the lower operating limit
up to somewhere before the upper limit.
• in subatmospheric or vacuum pressure distillation: somewhat above the lower
limit.

The liquid flow passing into the downcomer entrance is composed of a part contributed
by drops thrown over (out of the spray layer) and a part contributed by a disintegrating
bubbly liquid flowing over (out of the liquid continuous bottom layer). Their relative
contributions vary, depending on the operating conditions and the specifics of the outlet
weir geometry. The entering liquid flow is composed of a vary wide range of 'drop' sizes
falling down and impacting on the surface level. Gas bubbles are generated on impact, in
addition to being taken along as the bubbly part in the liquid.

Again, the dispersion in the downcomer has a two layered structure, with a bubbly upper
layer and a fairly clear liquid layer at the bottom.

The Spray regime occurs at high(er) gasvelocities and relatively thin liquid layers (while
bubbles or jets forming at the tray floor are only partly submerged). The bottom layer has
the highest liquid content. The high velocity of the gas flowing out the perforations in the
tray floor atomizes the liquid and accelerates the droplets upward. The drops can acquire
substantial initial propagation velocities. Because the vapor transferred its kinetic energy
and momentum to the liquid phase, the result is a fairly chaotic (randomized) movement
of droplets, with a very wide distribution in sizes, ranging from several cm (~ 0.02 m) for
the largest drops, to a few hundred (~ 0.0001 m) micron for the smallest droplets. A
fraction of the very smallest droplets is dragged along by the gas flow and carried away
to the next higher tray, as droplet entrainment (or carry over).

The upper layer is primarily a droplet propagating layer, whose height depends on the
initial propagation velocity acquired by the drops and the shearing interaction between
the moving drops and upflowing gas. Ultimately, the dispersion height is constrained by
the spacing between successive trays, which is somewhere between 0.3 to 0.9 m high.
The liquid content of this upper layer typically ranges from ~1 - ~15 %.
The spray regime can be found:

• in high pressure absorption or distillation; only at operation near the upper

limit.

• in medium pressure distillation; from somewhere in the middle of the operating

range up to the upper limit.
• in subatmospheric or vacuum pressure distillation: starting somewhat above the
lower limit it occurs over the whole operating range.

The liquid passes into the downcomer entrance by splashing over the outlet weir. Hence,
a cloud of droplets rains down in and impacts on the liquid surface in the downcomer. On
impact some gas is entrained and bubbles forming from entrained gas are dragged down
for some distance, before they escape the impacting liquid flow and rise up and separate.
The upper layer in the downcomer will (again) be a bubbly dispersion, while the bottom
layer will be fairly clear liquid.

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Tray pressure drop

Typical tray pressure drops lie in the range of 250 - 1500 N/m*m (or 2.5 mbar - 15 mbar
or 25 - 150 mm Water Column, in whatever units one prefers).

Usually, the drop in pressure caused by gas flowing through a tray is small in
comparison to the system pressure. Except for vacuum columns, where it can become
quite substantial and the gas velocity in the perforations may become comparable to the
velocity of sound.

The tray pressure drop plays an important part in filling up the downcomers. To
compensate for the pressure drop, a liquid head builds up in the downcomers, to enable
the liquid to flow down against it. When the tray pressure drop becomes excessive with
respect to the height of the downcomers, flooding will be the result.

The tray pressure drop is composed of (at least) two (major) contributions:

1. a pressure drop caused by the gas flowing through the perforations in the tray
floor. This contribution depends on gas flow rate, fraction free area and the
pressure drop coefficient of the particular perforations (or valves) being used.
This pressure drop coefficient depends on relative hole thickness (i.e. the ratio of
tray thickness over hole diameter), hole shape and nearness of other holes (ratio of
hole pitch to hole diameter).
2. a pressure drop caused by the liquid present on the tray. This liquid hold up effect
primarily increases with an increase in outlet weir height, decreases with an
increase in gas flowrate and increases with an increase in liquid flowrate. To a
lesser extent, it depends on physical properties of the gas/liquid system.

The figure shows a typical tray pressure drop curve for a particular sieve tray. The tray
pressure drop is shown in comparison to the pressure drop caused by gas flowing through
a tray without liquid on it; the dry tray pressure
drop. The gas velocity in the holes is given here
as the hole loadfactor, which essentially is the
gas hole velocity times the square root of the
ratio of the gas over the liquid density.

The pressure drop of a valve tray differs from

the sieve tray pressure drop mainly, because the
gas flowing through the valves in the tray floor experiences a different flow resistance. A
valve tray dry pressure drop curve exhibits two horizontally shifted parabola's connected
by a horizontal plateau. Operating at this plateau, the force exerted on the valves by the
gas pressure drop overcomes the weight of the valves.

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Weeping, sealing & blowing: the lower operating limit(s)

For assessment of the lower operating limit, three ranges in gas flowrate should be
considered.

In the normal operating range, the gas flowrate exceeds the critical weep point gas
flowrate needed to stop liquid from leaking through the perforations in the tray. All liquid
goes in cross flow from the inlet to the outlet downcomer. The separation performance of
a tray is at its best in this range.

In the weeping range, the gas flowrate is below the critical weep point gas flowrate and
above the critical seal point gas flowrate. The liquid changes from flowing across the
tray and over the outlet weir into flowing countercurrently to the gas through the
perforations. The separation performance remains good (at the same level as in its normal
operating range) quite some way down in gasrate (and up in weep rate).

In the dumping range, the gas flowrate is below the critical seal point (or dump point)
gas flowrate. All gas and liquid flow countercurrently through the perforations. A tray
operating in its dumping range operates similar to a downcomerless tray, i.e. a dual flow
tray. The separation performance falls off significantly upon operation below the seal
point.
The hole loadfactor is essentially the gas hole velocity times the square root of the ratio
of gas density over liquid density.

The weep point gas flowrate (wp).

The weep point gas flowrate is the gas flowrate where the first leakage (raining) of
liquid occurs, because the gas flowing through the perforations is no longer able to
counterbalance the hydrostatic head of liquid on a tray. The perforations are no longer
exclusively used by the gas. The weep point gas flowrate (or weep point for short) is
fairly well defined and numerous studies have reported their values and several empirical
correlations are available in the open literature. As a rule of thumb for sieve trays, the gas
flowrate expressed as loadfactor (= capacity factor) based on hole area usually lies
between 0.30 - 0.40 m/s. Recommended can be the article by M.J. Lockett and S. Banik,
"Weeping from Sieve Trays", Ind. Eng. Chem. Process Des. Dev., 25(1986)25, pp 561-
569.

The seal point (or dump point) gas flowrate (sp).

The seal point (or dump point) gas flowrate is the gas flowrate where the gas-liquid
layer on a tray is just expanded enough to make some liquid flow over the outlet weir
into the downcomer. At a lower gasrate, the dispersion height is less than the weir height
and all liquid is leaking away through the perforations (countercurrent to the gas).
Although the seal point gas flowrate (or seal point for short) is fairly well defined, only a
few studies are available in the open literature. Recommended for sieve trays can be the
article by R.G.H. Prince and B.K.C. Chan, Trans. Instn Chem. Engrs, 43(1965), pp. T49-
T55.
Upward blowing of gas through the downcomers occurs when insufficient liquid flows
through them. This is detrimental to the separation performance of a tray, because a
significant part of the gas bypasses the contacting area of the tray in this way. During the
(re-) start up phase of a column, this phenomenon needs special attention, as in this mode
of operation, the penalty can be a severe underperformance of the column (in both
throughput and separation performance). A proper tray design can take care of this
potentially deliterious situation.

Frank has reworked the subject of the lower operating limits of sieve and valve trays.
He reported his findings in an article in the "Chemical Engineering Journal", 70 (1998),
pp. 143-155.

Last Updated: 03 februari 1999

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HOMEPAGE E. Frank Wijn

Flooding, priming & entrainment: the upper operating limit(s)

The common mechanisms governing the upper limit of operation of a tray are:

Entrainment flooding is caused by an excessive liquid flowrate generated by droplets

carried out of the gas-liquid dispersion by the gas stream and up to the next higher tray.
Preferentially, this occurs on trays operating in the spray regime, when most of the liquid
is already present as droplets.

Bed expansion flooding (priming) sets in when the gas-liquid layer on the tray extends
up to the next tray, i.e. the dispersion height becomes equal to the tray spacing. This
condition depends on the liquid height on a tray and the expansion by the gas flowing
through it. Actually, this condition is 'the twin brother' of entrainment flooding. As the
dispersion level approaches the next higher tray, droplet entrainment rises rapidly as well,
causing an increasing liquid level on the next tray, thereby increasing the liquid flow
through the downcomers.

Choking (blockage) at the downcomer entrance (or in the dwoncomer itself) can set a
limit to the liquid flowrate, as the inflowing gas-liquid dispersion needs more space, in
the case when a large amount of gas has to be set free in the downcomer. The separated
gas rises counter to the downflowing gas-containing liquid and can induce a (vapour
b)locking condition.

Downcomer overflowing (flooding) sets a limit to the liquid handling capacity of a

downcomer. This happens when the height of the gas-liquid layer in a downcomer
exceeds the height of the downcomer. This condition is controlled by the height of liquid
in the downcomer and by the expansion caused by entrapped gas bubbles. The liquid
height is given by the condition that the hydrostatic liquid head should be sufficient to
make the liquid flow down against the tray pressure drop.

A downcomer velocity limitation occurs when the liquid downflow velocity becomes
large enough to stop gas bubbles from rising out of the liquid level and begins to drags
them down (resulting in carry under of these gas bubbles). The downcomer gets filled up
with a 'frothy' mass, which takes up more space than pure liquid.

The open literature on upper operating limits (or maximum capacities) of sieve and valve
trays is fairly large and several decades old already. Because of its importance, the
subject has been reviewed several times and their results have become text book material
as empirical design rules and design graphs, see: Henry Z. Kister, Distillation Design,
McGraw-Hill,Inc., 1992, and Section 18. Liquid-gas Systems by J.R. Fair in Perry's
Chemical Engineers Handbook, McGraw-Hill, Inc..
Rather unfortunately, a generic and unifying hydrodynamic theory of two phase flow of
gas and liquid on trays has not (yet) been developed. Consequently, a theory predicting
the upper operating limit(s) is still lacking and we will have to make the best use we can
of past experience, as embodied by these empirical design rules and graphs as given.

Usually, the maximum column loadfactor for a conventional sieve tray lies in the range
of 0.06 to 0.09 m/s, depending on tray spacing (typically 0.6 m), liquid load, weir height,
etc.. For conventional valve trays, the maximum capacity is about the same or somewhat
lower. Specially developed high capacity trays can handle substantially more, however.

The column loadfactor is essentially the column gas velocity times the square root of the
ratio of gas density over liquid density.

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Downcomers, Degassing and Bubble Coalescence.

The prime function of downcomers is to transport the liquid flow from one tray to the
next lower tray, without giving rise to (premature) flooding, caused by filling up of the
downcomers, by a vapour/liquid mixture.

The downcomer area needed to achieve this function, usually, occupies about 5 % to 20
% of the column cross section. For relatively very low liquid rates, it may be even less,
for instance in vacuum distillation columns and in columns using glycol in dehydrating
natural gas. On the other hand, heavily liquid loaded columns in some high pressure
distillations (demethanizers, deethanizers, etc.) and in high pressure acid gas treating, the
downcomer area may occupy up to 50 % of the column area.

Specification of the downcomer area needed in a particular application is still in an

empirical state and is based largely on past experience and test data. Scientific research
can contribute greatly by clarifying the mechanisms involved and give rise to this
unsatisfactory situation.

Factors to consider depend on the governing mechanism. The usual mechanism involves
filling up of the downcomer with a vapour/liquid mixture, to accomodate the pressure
difference between trays by building up an equivalent static presure. The height between
trays (tray spacing) sets the maximum height for a downcomer before it overflows onto
the next upper tray. The hole area (free area) in the tray controls (part of ) the tray
pressure drop. Increasing the hole area will reduce pressure drop and so delay the point of
overflowing of the downcomer. Gas captured in the liquid (the gas fraction) causes the
gas/liquid mixture to occupy more space, than needed for pure liquid, in this way
lowering the point of flooding. Proper degassing will help in delaying flooding of the
downcomers.

The rate of disengagement of gas (vapour) from the liquid may become constraining in
reaching the desired maximum liquid throughput.

Based on occasional observations with various gas containing liquid flows, the
awareness has grown, that the maximum liquid handling capacity of vertical tubular
downcomers can be affected greatly by the presence of either bubbles or slugs. Slugs are
'large bubbles' or large clusters of small bubbles. When slugs are formed, the maximum
liquid handling capacity of a downcomer can be lower (contrary to what one would
expect on basis of an increased velocity of rise), in comparison to the maximum liquid
throughput when bubbles are present. This is caused by temporary blockage of the
downflowing liquid by these fast upward rising slugs.
 Whether slugs will be formed out of the bubbles depends on a balance between
coalescence and break up, either an increase in the rate of break up or a decrease in the
rate of coalescence (growth) will reduce or prevent slug formation. Likewise, foam-
formation results from a reduction in the rate of bubble coalescence as caused by
stabilization of films between the bubbles. So, the foamability of a system can be an
important parameter in the transition from bubbly to slug flow. Consequently, the
maximum liquid handling capacity of downcomers depends on the foamability of the
vapour/liquid system being handled.

It is proposed, that the effect of the gas content in the liquid and the foamability be
studied more systematically in order to improve the understanding of the actual two
phase behaviour in downcomers and to arrive at a better predictability of the maximum
liquid handling capacity of downcomers.

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Tray development; a bit of history.

The beginning of the development of modern tray technology for absorption and
distillation coincides with the beginning of the development of chemical engineering, in
general. The flowering of chemical engineering started after Word War I, when the
(petro-) chemical manufacturing and oil refining industry in the U.S. of America found
themselves in the position of having to become selfsufficient in the products, which they
had been importing from Europe before. Moreover, they had to supply Europe as well, as
their chemical industry was knocked out (for some time). The Massachusets Institute of
Technology (MIT), Cambridge, Mass., USA played a pioneering and leading role in the
early developments, which shaped chemical engineering, as we know it.

Examination of the statistics of publications in scientific journals reveals, that four

phases can be discerned in the development of tray technology. Broadly:

• A initial period, from about 1920 upto 1945 (including World War II). This
period saw the development of the basics of process equipment development and
the necessary supporting tools (o.a. mathematical). This was accompanied by
characterisation and study of process equipment as it has been and was being
used. Attempts at rationalisation of proven practices were made and led to the first
design outlines/procedures. The Bubble Cap tray was THE commonly tray, in this
period. Interestingly, a quote to an early reference (of W.A. Peters, 1922) shows
that the sieve tray was already known, since the inception of modern tray
technology.

• A period of expansion and growth from 1945 till about 1965. Investments in
rebuilding the (petro-) chemical industry and oil refineries in Europe, Japan etc.
after World War II created lots of opportunities to incorporate the latest
developments in tray technology. This coincided with the (re-)establishment
chemical engineering departments at European universities and elsewhere, thus
intensiving the scientific and technical research and development. The sieve tray
began its revival, because of its simplicity, its cost effectiveness, it was better
researched, this time. A number of other tray types were developed, as well:
Turbogrid tray, Ripple tray, Kittel tray, valve tray, dual flow tray, shower deck
tray, baffle tray, etc..
• A consolidation phase from 1965 till about 1985. Industrial tray development
largely faded away. Collaborative (FRI) and (sponsored) scientific research
became dominant, in this period.
 • From about 1985 till now, interesting new or improved tray types were
introduced. The tray hardware industry was stirring to live again ! (because of
competitive pressures). High efficiency tray and high capacity tray developments.

Timeline of some key concepts, events and developments:

1915 A.D. Little (MIT) introduced the concept of Unit Operations.

1922 W.K. Lewis (MIT), Ind. Eng. Chem., 14, 492- . He defined the "overall plate
efficiency" as the ratio of the number of equilibrium contacts to the number of actual
plates, required for a given purpose.

1922 W.A. Peters, Ind. Eng. Chem., 14, 476-479. "Peters was the first to report any
values for the allowable rate to which a column may be run without priming. For two
bubble cap columns he reports the allowable vapor rate as 1.3 feet per second. For sieve
plate columns running weak alcohol/water or acetic acid/water mixtures he gives the
allowable vapor rate as 3 feet per second".

1924 W.K. Lewis and W.G. Whitman (MIT), Ind. Eng. Chem., 16(12), 1215-1239. The
first Absorption Symposium tested the usefulness of the theory of two-film resistance to
mass transfer at an interface (with thermodynamic equilibrium at the interface).

1925 E.V. Murphree, Ind. Eng. Chem., 17, (7), 747-750, Rectifying column calculations
with particular reference to N component Mixtures. Defined the tray efficiency, which is
now know as the Murphree tray efficiency; Emv.

1925 W.L. McCabe and E.W. Thiele (MIT), Ind. Eng. Chem., 17, 605- . Presentation of
graphical method for computing the number of equilibrium plates required in a
fractionating column for binary mixtures.

1928 W.K. Lewis, H.D. Wilde (MIT), Trans. Am. Inst. Chem. Eng., 21, 9-126. They
introduced the concept: key component.

1930 Kremser, Natl. Petroleum News, 43, May 21. Equations for computing the inlet and
outlet concentrations for absorbers and strippers, depending on absorption or stripping
factor and number of equilibrium plates.

1934 M. Souders and G.G. Brown, Ind. Eng. Chem., 26, 98-103. Design of fractionating
Columns.I. Entrainment and Capacity. Presentation of equation for maximum allowable
vapour velocity of bubble cap trays (with the constant C, which is now known as the
capacity factor).

1934 Perry, J.H., McGraw-Hill Book Co., Inc., N.Y., appearance of the first edition of
the Chemical Engineers Handbook.
1935 R. Higbie, Trans. Am. Inst. Chem. Engrs., 31, 365- . Penetration theory leads to use
of the contact time in the calculation of the mass transfer coefficients in the two-film
theory.

1936 W.K. Lewis (MIT), Rectification of binary mixtures. Plate Efficiency of Bubble
Cap Columns, Ind. Engng. Chem., 28, (1), 399. He gave the first quantitative
relationships to calculate the tray efficiency for the three cases, now known as the 'Lewis'
cases.

1941 C.F. Oldershaw (Shell Dev. Co., USA) published his experience with a
standardized perforated plate lab scale distillation column, Ind. Eng. Chem., Anal.
Edition, 13(4), 265-268.

1943 H.G. Drickamer and J.R. Bradford, Trans. Am. Inst. Chem. Eng., 39, 319-360.
Overall Plate Efficiency of Commercial Hydrocarbon Fractionating Columns as a
function of viscosity. This paper contains the presentation of a 'classic' empirical tray
efficiency correlation.

1943 Flexitray (Koch Eng Comp.) in first commercial use. By 1953 in more than 350
installations. (Petr. Refiner, 33(1954)05, 199-201).

1951 P.V. Danckwerts, Ind. Eng. Chem., 43(6), 1460-1467. Significance of liquid-film
coefficients in gas-absorption. Introduction of the surface renewal concept.

1952 D.S. Arnold, C.A. Planck, E.M. Schoenborn, Chem. Eng. Progr., 48(12), 633-642.
Performance of perforated plate distillation columns. Pioneering study on sieve trays.

1952 Shell Oil, USA published their Turbogrid tray. (Petr. Refiner, 31(11), 105-108).

1952 Founding of Fractionation Research, Inc.. (FRI) at Alhambra, Calif., USA, by 43

companies. (Petr. Refiner, 34(01), 154).

1953 P.V. Danckwerts, residence time distribution and mixing concepts, Chem. Eng. Sci.,
2(1), 1-13.

1954 Operations of Fractionation Research, Inc.. (FRI) start.

1955 M.F. Gautreaux & H.E. O'Connell publish a liquid mixing model based on series of
well mixed pools, Chem. Engng Progr., 51(5), 232-237. Introduction of the flow path
length effect on tray efficiency.

1956 The Ripple tray was made public by Stone & Webster Eng. Co., Chem. Eng. Progr.,
52(12), 503-508.

1956 P.H. Calderbank, Gas-Liquid Contacting on Plates, Trans. Instn. Chem. Engrs.,
34,79-90. Pioneering study on the fundamentals of mass transfer on trays.
1958 Publication of the A.I.Ch.E. Bubble Tray Design Manual.

1958 F.J. Zuiderweg and A. Harmens recognized the importance of surface tension
gradients; discovery of the Marangoni effect, Chem. Eng. Sci., 9(2/3), 89-102.

1959 Introduction of the Glitsch 'Ballast' tray (a type with two valves in a cage) with a 9 :
1 turn down ratio.

1961 Fair published his (by now classic) flooding correlation (i.e. maximum allowable
vapour velocity). Petro/Chem Engineer, Sept., 211-218.

1962 P.E. Barker & M.F. Self, important study on liquid mixing (eddy diffusivity) effects
on a sieve plate, Chem. Eng. Sci., 17, 541-553.

1969 Shell (F.J. Zuiderweg et al.) published the existence of their Calming Section trays

1972 R.L. Bell (FRI), Residence time and fluid mixing on commercial scale sieve trays,
AIChE Journal, 18(3), 491- 497 and 498-505. Large nonuniformities were found.

1974 Solari & Bell, calculated the effect of nonuniform velocity distributions on tray
efficiency, AIChE Journal, 20(4), 688-695.

1976 R. Krishna started to propagate the use of the Maxwell-Stephan equations for heat
and mass transfer in multicomponent mixtures (in effect starting 'rate based' modelling).

1984 H. Chan & J.R. Fair's tray efficiency correlation published in Ind. Eng. Chem.
Process Des. Dev., 23, 814-819.

1986 M.J. Lockett published his book: Distillation Tray Fundamentals, (now out of
print), Cambridge University Press, ISBN 0 521 32106 9.

1992 Henry Z. Kister published his book: Distillation Design, McGraw Hill, Inc., ISBN 0
07 034909 6

1993 R. Taylor and R. Krishna published their book: Multicomponent Mass Transfer,
John Wiley & Sons, Inc., ISBN 0 471 57417 7.

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Gatekeeping on new activities on trays.

Periodically, the status of technical developments has been assessed in the field of trays
for distillation and absorption columns. The most recent study (covering the period from
1984 to 1994) was done, May 1994.

The most striking points of this study were:

• The last 5 years have seen a moderate 20% activity growth in the development of
tray technology (compared to previous 5 yr. periods). The USA and China have
been largely responsible for this increase.
• The technical areas of growth were:
1. cryogenic distillation (o.a. Air Products, BOC Group, Hitachi, Union
Carbide/Praxair)
2. catalytic distillation (reactive distillation) (Chemical Research &
Licensing)
3. higher efficiency trays (with packing or baffles on the tray)
4. higher capacity trays (Glitsch: Nye tray, UOP: Enhanced MD tray)

The ten groups most active in patenting:

1. Moscow Chemical Equipment Inst., Rusland (48 patents in 10 yr.)

2. Krasnodar Polytechnic Rusland (11 patents)
3. Hitachi Japan (11 patents)
4. Oil Equipment Design Bureau Rusland (10 patents)
5. Kazan Kirov Chem. Techn. Rusland ( 9 patents)
6. Glitsch Inc. USA ( 9 patents)
7. Beloruss Kirov Tech. Inst. Rusland ( 7 patents)
8. Gas Processing Research Inst. Rusland ( 7 patents)
9. NOVO-UFA Oil Refinery Rusland ( 6 patents)
10. Gorki Polytechnic Rusland ( 6 patents)

The ten groups most active in publishing:

1. University of Nottingham U.K. (26 publications in 10

yr.)
2. University of Texas, Austin USA (19 publ's)
3. Tianjin University China (17 publ's)
4. Okayama University Japan (10 publ's)
5. Union Carbide / Praxair USA (10 publ's)
6. Univ. of Alberta (K.T. Chuang) Canada (10 publ's)
7. East China Univ., Shanghai China (10 publ's)
 8. Aston Univ., Birmingham U.K. (10 publ's)
9. Air Products USA ( 9 publ's)
10. UMIST, Manchester U.K. ( 9 publ's)

The most creative/productive individuals were:

1. Biddulph. M.W. Univ. of Nottingham (28 publ's &

pat's.)
2. Chekhov, O.S. Moscow Inst. Chem. Machine Building (27)
3. Solomakha, G.P. Moscow Inst. Chem. Machine Building (21)
4. Tarasov, V.A. Moscow Inst. Chem. Machine Building (17)
5. Skrynnik, Yu.N. Moscow Inst. Chem. Machine Building (14)
6. Fair, J.R. Univ. of Texas, Austin (14)
7. Lockett, M.J. Praxair (Union Carbide) (14)
8. Chuang, K.T. Univ. of Alberta (13)
9. Resetarits, M.R. UOP (Union Carbide) (11)
10. Mel'nikov, V.S. Moscow Gubkin (10)

Last Updated: 12 april 1998

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HOMEPAGE E. Frank Wijn

Recently Developed Trays.

The most recently introduced trays and improvements of existing tray hardware
are presented here in reversed chronological order (newest at the top and oldest at
the bottom). From this list it can be seen, that the last decade has seen interesting
developments in tray technology. This has resulted from the competitive activities of
the hardware manufacturers, who created a livelier tray market place by doing so.

(Please note: this listing is indicative and not exhaustive.)

The Triton™ high capacity tray was introduced during 1997 by

Norton Chemical Process Products Corp., Akron, Ohio, USA.

This tray combines a patented downcomer design (to increase the

tray's contacting area) with a tray deck provided with proprietary
fixed valves; the Provalve™. Up to 45% greater capacity than conventional valve trays
are claimed, while maintaining or improving mass transfer efficiency. This tray appears
to be Norton's answer the competition.

A unique collaboration of Nutter Engineering and UOP resulted in the introduction of

the new VG-MD tray, during the DA '97 Conference, september 1997 at Maastricht.

This tray is a further extension of the line of Multiple Downcomer trays of UOP. The
VG-MD tray combines Multiple Downcomers with MVG tray decks (replacing the usual
sieve tray decks). This tray type combines the characteristics of MD-trays and MVG-
trays: a higher capacity and an improved turn down.

The Bi-FRAC ™ high capacity tray was introduced by Koch

Engineering during 1997.

A tray design engineered to minimize entrainment (by

reducing gasvelocity) is the key to the claimed capacity
increase of up to 30 % over conventional valve and sieve
trays, without efficiency loss. (Another of Koch's competitors to Glitsch's Nye tray).

The tray floor is provided with fixed, open valves, which are positioned to create a bi-
directional flow pattern, hence Bi-Frac.

The Vortex (Downcomer) Tray™ was (re-) introduced by Sulzer Metawa B.V. (Tiel,
the Netherlands) during the ACHEMA at Frankfurt, FRG in 1997. This is an improved
high capacity tray enabling capacity increases up to 30% over
conventional trays, without losing separation efficiency (may even gain in
efficiency). The special feature of this tray is in the design of the
downcomers, which are cylindrical/conical with a seal pan underneath and
three baffles on top. These baffles are placed in such a way, that liquid
entering the downcomer induces a swirling motion on the liquid reservoir
in the downcomer. This improves vapour disengagement and hence higher
downcomer velocities are allowable. Reportedly, the first industrial
applications have confirmed the advantages of this tray design and more
applications are under study and being installed.

The (US 5453222) patent application for the SUPERFRAC® tray of Glitsch Inc. was
filed in 1994. The new tray was described in a paper presented at the DA '97 conference,
september 1997, Maastricht. The basic concept of the improvement embodied in this tray
is to transfer downcomer bottom area into contacting area by truncating the downcomer,
so that contacting area can be maximized. Due to this increase vapour velocity decreases
and entrainment will decrease, as well. The SUPERFRAC® tray is specifically
developed for large diameter (> 2m.) columns, which are heavily loaded with liquid.

In 1994 Stahl GmbH (Mannheim/Viernheim, Germany) introduced during the

ACHEMA their new Dualflex™ tray. This is a high performance tray based on a
downcomerless ('dual flow') tray provided with specially adapted Varioflex™ valves.
The main advantage of this tray is, that in situations where dual flow trays have to be
applied, it gives a better tray efficiency and an improved turn down capability. See,
Chemie-Ingenieur-Technik, 69(1997)5, 649-650.

In a 1994 paper presented in November at the Annual Meeting of the A.I.Ch.E. at San
Francisco, British Oxygen Company and the University of Nottingham unveiled their
newly developed Very High Capacity Expanded Metal tray. This tray was described in
the European Patent 635,292, filed 12/07/94, which mentioned J.T. Lavin as inventor.
This tray comprises a special combination of two Expanded Metal sheets (with different
geometric specifications) which are 'piggy backed' together and make up the tray floor in
the contacting area. In between these 'dual-sheet' tray floors of two successive trays, they
mount, at a distance of some 0.15 m from the lower tray, a third Expanded Metal sheet,
which collects and separates droplet entrainment and redirects this entrainment to a
downcomer. The overal tray spacing could be made as low as 0.2 m. In a rectangular
three tray air/water simulator (3inch wide and 4 ft flow path), they achieved a maximum
air flowrate which was almost twice the maximum air flow rate expected for
conventional sieve trays. Moreover, these trays showed a relative low pressure drop and
also the separation performance (tray efficiency) was satisfactory, i.e. comparable to
convential sieve trays. Apparently, this new tray type is under active development, at
BOC.

In March 1994 Koch Engineering Comp. Inc announced at the I.Chem.E.

"Debottlenecking Seminar" in London, that they had applied for the first time, with
success, a new type of high capacity tray, specifically developed for heavily liquid loaded
systems: the Ultra-Frac tray. For this tray, they claim at least a 20% capacity advantage,
in comparison to the original UOP Multiple Downcomer trays, while this is combined
with a comparable tray efficiency. They were enabled to do so, because several years
earlier they had acquired access to Russian tray technology on cocurrent vortex tube
contactors, as developed around 1978-1984 by Y.N. Lebedev, V.I. Sheinman and c.s..
See, paper by Kulov and Lebedev at 1992 Distillation and Absorption Conference at
Birmingham. This Russian technology in combination with Koch's expertise and an
additional development effort lead to their Ultra-Frac tray.

The MVG tray was introduced late in 1993 by Nutter

Engineering. It was the logical further extension of the V-grid
tray, which had two versions the SVG (Small V-Grid) and the
LVG (Large V-Grid) to a still smaller size of the fixed-valve
'grid'; the Mini V-Grid. All three sizes are shown in the picture.
Use of this type of plate in the contacting area of trays gives a
10 to 20% increase in the upper limit (when this limit is
contrained by an excessive entrainment rate) and an (initially
unexpected) decrease in the lower operating limit. Thus,
increasing the turn down ratio to a value of about four, which makes this tray
intermediate in flexibility, in comparison to convential sieve trays (with a turndown ratio
of ~2.5 to 3.0) and valve trays ( 5 and more). Since, their introduction Nutter Engineering
has been selling these tray decks successfully, for a wide range of applications.

In 1992 the second commercial application of UOP's Enhanced Capacity Multiple

Downcomer Trays came in use in a 18 ft. C3-splitter of Chevron Chemical Company at
Port Arthur Texas, USA. UOP had been developing this tray technology since 1989. The
first application had been in 1990 in a de-ethanizer owned by ÖMV Aktiengesellschaft, at
Schwechat, Austria. The results of this application were described in a paper presented at
the IChemE Distillation and Absorption Conference, Birmingham in 1992. UOP's EMCD
tray is a further development of their succesful MD-tray. Since the MD-tray was patented
in 1968 by the Linde division of Union Carbide, by 1992 they had been installed in
nearly 400 distillation columns, all over the world. The newly developed EMCD trays
have a 20% increased upper limit for the vapour flow in comparison to the original MD
trays. This may not sound spectacular, but financially can be highly attractive, because of
the increased productivity of these large scale columns.

The Duo-Sorb tray is an example of a high efficiency and high turndown tray invented
in 1992 by H.O. Ebeling (U.S. 5,116,393, US 5,514,305). The tray consists of a tray floor
having a multitude of small bubble caps with random packing dumped on it and a special
outlet downcomer construction. This tray was originally developed by Latoka
Engineering, Tulsa, OK, USA, for a particular range of applications, viz. dehydration of
natural with glycol at high pressure. The advantage these trays resides in the far greater
tray contacting area (greater efficiency), which lowers the number of actual trays, to be
installed in a new column design (or existing columns might be shortened). This results
in a reduced column height, which saves significantly on the cost of a high pressure
column shell. Latoka Eng. has been selling these trays for some time now, successfully.

1989 and 1990 saw the patenting of a tray, which combined a sieve tray (or a tray floor
made of closely spaced trapezoidal rods; a screen) with packing (or demistermat) on it.
(U.S. Patent 4,842,778, 1989; European Patent 381,388, filed 26-01-90). This type of tray
was invented and developed by G.K. Chen and K.T. Chuang at the University of Alberta,
Edmonton, Canada. Glitsch Inc. acquired a license to this tray and initially they have
been actively trying to sell it, under their trademark SCREEN TRAY. The much
improved efficiency of this type of tray could have made it attractive for many
applications. Lately, this type of tray type is no longer actively promoted and has quitely
disappeared from Glitsch's advertisements, for reasons which have not been made public.

The T-By tray was patented by B.M. Parker and T.J. Parker in 1988 (US 4,762,651) and
test results were published in 1992. The T-By tray is a modified sieve tray using a system
of transversal weirs with side baffles to approximate liquid plug flow across the tray.
Vapour/liquid contact is modified by patterning the vapour sieve holes in parallel to the
weirs and with the intermediate weirs form circulation cells promoting liquid mixing and
stabilizing the froth bed in the contacting area. The test results showed that the tray
performance (efficiency) could be improved in significant ways. Additional studies were
proposed, for further improvements.
 In 1988 the Nye tray was applied for the first time in two deisobutanisers of different
diameter. Glitsch Inc. acquired in 1991 the worldwide rights for the design, marketing
and manufacturing of these trays, which were invented by James O. Nye. Since their
introduction, Glitsch Inc. has had great success in selling these higher capacity trays. As
of May 1994, they had been installed in 119 distillation columns, already. In existing
columns that were modified, by installing these trays, Glitsch claims that capacity
increases of 20% have been achieved by the 'Nye-tray' effect. The Nye tray achieves this
increase in vapour capacity by shortening of the downcomers, in effect lifting them 2 to 3
inches up from the tray floor. In a way, trading off downcomer area for the enlargement
of the contacting area. This reduces the vapour velocity in the contacting area and allows
larger volumetric flowrates before entrainment or bed expansion limitations set in, again.
This strategy works for both sieve tray decks and valve tray decks.

Already in 1982 William R. Trutna described in the U.S. Patent

4,361,469 (filed Feb. 17, 1981) a new device based on a combination of
a special contacting tray with cocurrent upflow of vapour and liquid and
a special liquid separation section. Relative to sieve trays this Trutna
tray would have a 30 to 100% greater vapour handling capacity. A
development programme has been in place at SRP, University of Texas,
Austin, USA from 1989 onwards. From a SRP-publication in the
Chemical Engineering Progress, June 1996, pp. 42-48, it can be learned,
that this type of tray can double the upper limit of the vapour throughput
and maintain the same separation efficiency. However, these trays are
more complex to make and more expensive to install. In debottlenecking
applications, the Trutna trays require about three to four times the
replacement cost of convential sieve trays, which make them about as
expensive as structured packing. By 1998, Jaeger Products., Inc. had started to offer this
tray under the trademark of CoFlo Tray.

Last Updated: 18 december 1998

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HOMEPAGE E. Frank Wijn

Facilities for tray testing, development and trouble shooting

Experimental facilities for research, development ot trouble shooting work on absorption

and distillation trays (have been) are available at various locations around the world.

Facilities owned by hardware vendors:

• Glitsch Inc., Dallas, TX, USA has a 3 ft (0.9 m) diameter tray simulator, operated
with the air/water or the air/Isopar M test system. They own also a 7 ft (2.1 m)
diameter air/water simulator.
• Koch Engineering, Wichita, USA operate a 0.20 m reactive distillation pilot plant
and command a 3 ft. (0.9 m) diameter air/water simulator. Their colleagues at
Bergamo, Italy have a 8 ft (2.4 m) diameter air/water simulator to their disposal.
• Norton Chemical Process Products Corporation, USA has the following test
columns at their Chamberlain Laboratory research facility near Akron, Ohio: 1) a
10 ft (3.0 m) diameter air/water simulator; 2) a 2.5 ft (0.76 m) diameter column
that is used for air/water simulation, gas absorption and direct-contact heat
transfer tests; 3) a 1.27 ft (0.39 m) diameter distillation column, which can be
operated from high vacuum to 24 psia (1.7 bara) and; 4) a 1.27 m (0.39 m)
diameter distillation column which can be operated from high vacuum to 300 psia
(22 bara).
• Nutter Engineering, Tulsa, OK, USA has a 4 ft (1.2 m) diameter simulator,
operating on air/water or air/Isopar M.
• Mitsui Engineering and Shipbuiding Co., Tamano, Japan has a 0.5 m diameter
distillation column, capable of doing total reflux testruns with the cyclohexane/n-
heptane system, at near atmospheric pressures.
• Shell Research and Technology Center, Amsterdam, the Netherlands
decommissioned their large 2.5 m diameter distillation column in 1996. They
operate a 0.45 m distillation column, capable of operation at sub- and super-
atmospheric pressures and have several air/water test setups.
• Sulzer Chemtech, Winterthur, Switzerland has a 1.0 m diameter distillation
column, operating on the chlorobenzene/ethylbenzene test system, at amospheric
pressures and below. They have a number of smaller scale test columns.
• UOP Process Equipment, Tonawanda, NY, USA has a square 2 ft by 2 ft (0.6 m x
0.6 m) air/water simulator and a large 8 ft (2.4 m) diameter simulator.

Facilities at Universities:

• Aston University, Birmingham, U.K. has a 4 ft (1.2 m) and a 8 ft (2.4 m) diameter

air/water simulator.
 • Laboratory for Equipment for the Chemical Process Industry at Delft Technical
University, the Netherlands operate a 0.45 m distillation column and a large scale
0.8 m by 1.4 m air/water test rig, which has been used in the past for tray research
and is currently being exploited for structured packing studies.
• Separations Research Program, University of Texas, Austin, U.S.A. has a range of
experimental facilities: 1, 2, 3, 4 inch Oldershaw distillation lab-scale columns; 2,
4, 6 inch small-scale distillation columns; a 16.8 inch (0.426 m) diameter
air/water stripper and a 16.8 inch (0.426 m) diameter distillation/extraction pilot
plant.
• UMIST, Manchester, U.K. has a large 2 ft (0.6 m) diameter distillation column
(which is unique for a university) operating with a methanol/water test system.
• University of Alberta, Dept. of Chemical Engineering, Edmonton, Canada has 6
inch (0.15 m) diameter air/water column; a 6 inch (0.15 m) distillation column; a
0.30 m air/water column; a 0.30 m distillation column and a 0.60 m diameter
air/water column
• University of Nottingham, U.K. has a 6 ft (1.8 m) diameter air/water simulator for
hydraulic testing and acetone stripping experiments. There is also a three tray 3
inch wide and 4 ft long air/water simulator and a one tray 6 inch diameter
distillaton column, operating on methanol/water.
• University of Wroclaw, Poland has a 1 m diameter air/water simulator.

Other facilities:

• Fractionation Research Inc., Stillwater, OK, USA operates a 4 ft (1.2 m)

diameter low pressure column, which has a topsection of 8 ft (2.4 m) diameter.
Also, they have a 4 ft (1.2 m) diameter high pressure column. Various
hydrocarbon mixtures or hydrocarbon/water mixtures have been and can be used
by them, as test systems.

Last Updated: 18 december 1998

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HOMEPAGE E. Frank Wijn

Links to tray related expertise, tray vendors, etc..

Tray expertise can being found at:

• Aston University, CEAC Separations Research Group, Birmingham, UK (B.

Davies, J.P. Fletcher)
• FRI, Fractionation Research Inc., Stillwater, Oklahoma, USA
• University of Texas at Austin, Separations Research Program, Texas, USA (J.
Fair, F. Seibert)
• Technical University of München, Lehrstuhl für Fluidverfahrenstechnik,
Boltzmannstr. 15, D-85747 Garching (near München), Germany (J. Stichlmair)
• Technical University of Wroclaw, Poland (A. Koziol, J. Kuzniar)
• Tianjin University, National Engineering Research Center of Distillation
Technology, Research Center of Packed tower and Tower Internals, Tianjin,
NanKai district, China (K.T. Yu)
• Tokyo Institute of Technology, Dept. Chem. Eng., Kosuge Laboratory, Tokyo,
152-8552, Japan (Ass. Prof. KOSUGE, Hitoshi)
• University of Alberta, Dept. of Chemical and Materials Engineering, Edmonton,
Alberta, Canada (K.T. Chuang)

Tray hardware vendors:

• ACS Separations & Mass Transfer Products, 14211 Industry Road, Houston, TX
77053, USA, tel. 800 321 0077.
• AMT International, Inc., 100 N.Central Expressway suite #1100, Richardson, TX
75080, USA, tel. 972 783 1770 and 972 783 1924.
• APV Invensys, Solutions and Services, 395 Fillmore Ave. , Tonawanda, NY
14150, USA.
• Distall, Brunel House, Third Avenue, The Pensnett Estate, Kingswinford, West
Midlands, U.K., OY6 7XE, tel. +(44)-(0) 1384 277 776.
• Eta Process Plant, Ltd., The Levels, Brereton, Rugeley, Staffordshire, WS15 1RD,
U.K., tel. +(44)-(0) 1889 576 501. (Note: in 2000 acquired by Norton and in 2002
acquired by Koch-Glitsch).
• Finepac™ Structures PVT Ltd., M.I.D.C., Bhosari, Pune, 411 026 India, tel. +91
20 712 5014, +91 20 712 3025, +91 20 711 9517.
• GESIP GmbH (Gesellschaft fuer Informations und Prozesstechnik), Rudower
Chaussee 29, 12489 Berlin, Germany, tel. +(49)-(0) 3067 892 160.
• Jaeger Products Inc., 1611 Peachleaf, Houston, TX 77039, USA, tel. 281 449
9500 or 800 678 0345. Special high capacity tray: the CoFlo tray.
 • Kansai Chemical Engineering Co. Ltd., 2-chome, 9-7 Minaminanamatsu-cho,
Amagasaki-city, Hyogo pref., Japan, tel (06) 6419 7121. Special dual flow tray:
the LIFT Tray.
• Koch-Glitsch, Inc., 4111 E. 37th Street North, Wichita, Kansas 67220, USA tel.
316 828 5110 and 4900 Singleton Blvd., Dallas, TX 75212, USA, tel. 214 583
3000.
• Kühni AG, Gewerbestrasse 28, Allschwil 2, CH-4123 Switzerland, tel. +41 (0)61
486 37 37. Special (fixed valve) tray type: the Kühni Slit Tray.
• Julius Montz GmbH, Postbox 530, D-40705 Hilden, Germany, tel. +(49)-(0) 2103
8904. Special tray types: Thormann tray, Tunnel trays, KSG-trays.
• Raschig GmbH, Mundenheimer Strasse 100, D-67061 Ludwigshafen, Germany,
tel. +49 621 5618 648, email: mailto:MSchultes@raschig.de(allied with UOP
LLC).
• Rauschert Verfahrenstechnik GmbH, Paul-Rauschert-Str. 6, D-96349
Steinwiesen, Germany, tel. +49 92 62 77736, email: i.wagner@rauschert.de.
• Saint Gobain NorPro Corporation, P.O. Box 350, Akron, OH 44309-0350 USA,
tel. 330-673-5860 (formerly Norton Chemical Process Products Corporation)
(Note: The tray hardware activties were acquired by Koch-Glitsch in 2002).
• Nutter Engineering, Tulsa, OK 74170 USA, tel. 918-446-6672 (Note: acquired by
Sulzer Chemtech).
• Stahl Apparate- und Gerätebau GmbH, D 6806 Viernheim, Germany, tel. 0 62 04
20 25 (Note: acquired by Koch-Glitsch).
• Metawa Tray, 4000 HA Tiel, the Netherlands, tel. 0344 636 600, (Note: acquired
by Sulzer Chemtech).
• Sulzer Chemtech, P.O. Box 65, CH-8404 Winterthur, Switzerland, tel. +41 52
262 6008.
• Tray Hardware, Inc., P.O. Box 62004, Houston, TX 77205-2026, USA, tel. 281-
540-8990.
• UOP LLC, Engineered Products Group, Tonawanda, NY 14151-0044 USA, tel.
716-879-2444 (Multi-Downcomer trays).
• Consult also the Buyer's Guide on the PetroPages

Distillation and Absorption Processes:

• Distillation, an introduction by M.T. Tham of the University of Newcastle upon

Tyne (UK).
• Reactive Distillation research and design, Moscow, Russia (L.A. Serafimov; Y.
Pisarenko, D. Efromov)
• Andrew Sloley's Distillation page, Houston, Texas, USA.
• Lurgi A.G., Lurgi-allee 5, 60295 Frankfurt am Main, BRD, tel. +(49)-(0)
6958080, distillation and gas sweetening processes.
• AQUAlibrium, John J. Carrol's web page on phase equilibria in Natural
Gas/Water systems, P.O. Box 52163, Edmonton, Alberta, Canada, T6G 2T5.
 • Optimized Gas Treating, Inc., 15638 Whitewater Lane, Houston, Texas, USA
developed the ProTreat™ simulator, a mass and heat transfer rate based software
tool for simulating the removal of H2S, CO2 and mercaptans from industrial
gases by absorption into amine-based aqueous solvents.
• WelChem GmbH, Weyarner Strasze 1, D-83629 Wattersdorf (near München),
Germany, tel. +49 80 20 90 83 70.

Column Scanning Services:

• Canadian Tower Scanning, P.O. Box 347, Corunna, Ontario, N0N 1G0, Canada.
• Gamma Surveys, 16511 Space Center Drive, #400, Houston, TX 77058, USA.
Site with animated GIF's of tray action.
• Scanning Technologies Inc., 13524 - 117 Street, Edmonton, Alberta, T5E 5K6,
Canada.
• Synetix Tracerco-services, Synetix Process Diagnostics, P.O. Box 1, Belasis Hall
Technology, Billingham, Cleveland, TS231LB, UK.
• Nuclear Scanning Services, Inc., 2437 Bay Area Blvd., Suite 297, Houston, TX
77058, USA.
• Tru-Tec Services Inc., Division of Koch Engineering, 11005 W. Fairmont Pkwy.,
LaPorte, TX 77571, USA.

Other interesting/useful links:

• American Petroleum Institute, USA.

• ePTQ, Petroleum Technology Quaterly, Hopesay, Craven Arms, SY7 8HD, U.K.
• Gas Processors Association, USA.
• TU Delft, Laboratory for Chemical Engineering Process Equipment.
• Oil and Gas Journal, USA.
• ProcessAssociates of America, Newtown, PA 18940, USA, confederation of
consultants.
• Refining On-Line, USA.
• Shell Global Solutions, Amsterdam.
• Shell Research & Technology Center, Amsterdam.

Last Updated: 4 februari 2004

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 Introduction to Distillation
By M.T. Than

Introduction

Types of Columns

Basic Equipment and Operation

Column Internals

Reboilers

Distillation Principles

Vapour Liquid Equilibria

Distillation Column Design

Effects of the Number of Trays or Stages

Factors Affecting Operation

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 [visit our sponsor]

an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction Distillation is defined as:

Types of Columns
a process in which a liquid or vapour mixture of two or more substances is separated into its
Basic Equipment component fractions of desired purity, by the application and removal of heat.
and Operation
Column Internals Distillation is based on the fact that the vapour of a boiling mixture will be richer in
the components that have lower boiling points.
Reboilers
Distillation Therefore, when this vapour is cooled and condensed, the condensate will contain
Principles more volatile components. At the same time, the original mixture will contain more
Vapour Liquid of the less volatile material.
Equilibria
Distillation Column Distillation columns are designed to achieve this separation efficiently.
Design
Although many people have a fair idea what “distillation” means, the important
Effects of the
aspects that seem to be missed from the manufacturing point of view are that:
Number of Trays or
Stages
Factors Affecting
Operation
distillation is the most common separation technique
Crossword
Other Resources it consumes enormous amounts of energy, both in terms of cooling and heating requirements
Copyright
it can contribute to more than 50% of plant operating costs
Information
Visit our sponsor The best way to reduce operating costs of existing units, is to improve their efficiency and operation via process
optimisation and control. To achieve this improvement, a thorough understanding of distillation principles and how
distillation systems are designed is essential.

The purpose of this set of notes is to expose you to the terminology used in distillation practice and to give a very
basic introduction to:
types of columns

basic distillation equipment and operation

column internals

reboilers

distillation principles

vapour liquid equilibria

distillation column design and

the factors that affect distillation column operation

Next
 [visit our sponsor]

an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction TYPES OF DISTILLATION COLUMNS

Types of Columns
There are many types of distillation columns, each designed to perform specific types of separations, and each
Basic Equipment design differs in terms of complexity.
and Operation
Column Internals Batch and Continuous Columns
Reboilers One way of classifying distillation column type is to look at how they are operated. Thus we have:
Distillation batch and
Principles
Vapour Liquid continuous columns.
Equilibria
Distillation Column Batch Columns
Design In batch operation, the feed to the column is introduced batch-wise. That is, the column is charged with a
Effects of the 'batch' and then the distillation process is carried out. When the desired task is achieved, a next batch of
Number of Trays or feed is introduced.
Stages
Factors Affecting Continuous Columns
Operation
Crossword In contrast, continuous columns process a continuous feed stream. No interruptions occur unless there is a
problem with the column or surrounding process units. They are capable of handling high throughputs and
Other Resources are the most common of the two types. We shall concentrate only on this class of columns.
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Types of Continuous Columns
Continuous columns can be further classified according to:

the nature of the feed that they are processing,

binary column - feed contains only two components

multi-component column - feed contains more than two components

the number of product streams they have

multi-product column - column has more than two product streams

where the extra feed exits when it is used to help with the separation,
extractive distillation - where the extra feed appears in the bottom product stream

azeotropic distillation - where the extra feed appears at the top product stream

the type of column internals

tray column - where trays of various designs are used to hold up the liquid to provide better contact
between vapour and liquid, hence better separation
packed column - where instead of trays, 'packings' are used to enhance contact between vapour and liquid

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 [visit our sponsor]

an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction BASIC DISTILLATION EQUIPMENT AND OPERATION

Types of Columns
Main Components of Distillation Columns
Basic Equipment
and Operation Distillation columns are made up of several components, each of which is used either to tranfer heat energy or
Column Internals enhance materail transfer. A typical distillation contains several major components:
Reboilers a vertical shell where the separation of liquid components is carried out
Distillation
Principles column internals such as trays/plates and/or packings which are used to enhance component separations
Vapour Liquid a reboiler to provide the necessary vaporisation for the distillation process
Equilibria
Distillation Column a condenser to cool and condense the vapour leaving the top of the column
Design
a reflux drum to hold the condensed vapour from the top of the column so that liquid (reflux) can be
Effects of the
recycled back to the column
Number of Trays or
Stages The vertical shell houses the column internals and together with the condenser and reboiler, constitute a
Factors Affecting distillation column. A schematic of a typical distillation unit with a single feed and two product streams is shown
Operation below:
Crossword
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Basic Operation and Terminology

The liquid mixture that is to be processed is known as the feed and this is introduced usually somewhere near the
middle of the column to a tray known as the feed tray. The feed tray divides the column into a top (enriching or
rectification) section and a bottom (stripping) section. The feed flows down the column where it is collected at the
bottom in the reboiler.

Heat is supplied to the reboiler to generate vapour. The source of heat input can be
any suitable fluid, although in most chemical plants this is normally steam. In
refineries, the heating source may be the output streams of other columns. The
vapour raised in the reboiler is re-introduced into the unit at the bottom of the
column. The liquid removed from the reboiler is known as the bottoms product or
simply, bottoms.

The vapour moves up the column,

and as it exits the top of the unit, it is
cooled by a condenser. The condensed liquid is stored in a holding
vessel known as the reflux drum. Some of this liquid is recycled back
to the top of the column and this is called the reflux. The condensed
liquid that is removed from the system is known as the distillate or
top product.

Thus, there are internal flows of vapour and liquid within the column
as well as external flows of feeds and product streams, into and out of the column.

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an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction COLUMN INTERNALS

Types of Columns
Trays and Plates
Basic Equipment
and Operation The terms "trays" and "plates" are used interchangeably. There are many types of tray designs, but the most
Column Internals common ones are :
Reboilers
Bubble cap trays
Distillation A bubble cap tray has riser or chimney fitted over each hole, and a cap that
Principles covers the riser. The cap is mounted so that there is a
Vapour Liquid space between riser and cap to allow the passage of
Equilibria vapour. Vapour rises through the chimney and is directed
downward by the cap, finally discharging through slots in
Distillation Column the cap, and finally bubbling through the liquid on the tray.
Design
Effects of the
Number of Trays or Valve trays
Stages In valve trays, perforations are covered by liftable caps. Vapour flows lifts
the caps, thus self creating a flow area for the passage of vapour. The lifting
Factors Affecting
cap directs the vapour to flow horizontally into the liquid, thus providing
Operation
better mixing than is possible in sieve trays.
Crossword
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Information
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Sieve trays are simply metal plates with holes in them. Vapour passes
straight upward through the liquid on the plate. The arrangement, number
and size of the holes are design parameters.

Because of their efficiency, wide operating range, ease of maintenance and cost factors, sieve and valve trays have
replaced the once highly thought of bubble cap trays in many applications.

Liquid and Vapour Flows in a Tray Column

The next few figures show the direction of vapour and liquid flow across a tray, and across a column.

Each tray has 2 conduits, one on each side, called ‘downcomers’. Liquid falls through the downcomers by gravity
from one tray to the one below it. The flow across each plate is shown in the above
diagram on the right.

A weir on the tray ensures that there is always some liquid (holdup) on the tray and is
designed such that the the holdup is at a suitable height, e.g. such that the bubble caps
are covered by liquid.
Being lighter, vapour flows up the column and is forced to pass through the liquid, via the
openings on each tray. The area allowed for the passage of vapour on each tray is called
the active tray area.

The picture on the left is a photograph of a section of a

pilot scale column equiped with bubble capped trays. The tops of the 4 bubble caps
on the tray can just be seen. The down- comer in this case is a pipe, and is shown
on the right. The frothing of the liquid on the active tray area is due to both
passage of vapour from the tray below as well as boiling.

As the hotter vapour passes through the liquid on the tray above, it transfers heat
to the liquid. In doing so, some of the vapour condenses adding to the liquid on
the tray. The condensate, however, is richer in the less volatile components than is
in the vapour. Additionally, because of the heat input from the vapour, the liquid
on the tray boils, generating more vapour. This vapour, which moves up to the
next tray in the column, is richer in the more volatile components. This continuous
contacting between vapour and liquid occurs on each tray in the column and brings
about the separation between low boiling point components and those with higher
boiling points.

Tray Designs
A tray essentially acts as a mini-column, each accomplishing a fraction of the separation task. From this we can
deduce that the more trays there are, the better the degree of separation and that overall separation efficiency will
depend significantly on the design of the tray. Trays are designed to maximise vapour-liquid contact by
considering the

liquid distribution and

vapour distribution

on the tray. This is because better vapour-liquid contact means better separation at each tray, translating to better
column performance. Less trays will be required to achieve the same degree of separation. Attendant benefits
include less energy usage and lower construction costs.

There is a clear trend to improve separations by supplementing the use of trays by additions of packings.

Packings
Packings are passive devices that are designed to increase the interfacial area for vapour-liquid contact. The
following pictures show 3 different types of packings.

These strangely shaped pieces are supposed to impart good vapour-liquid contact when a particular type is placed
together in numbers, without causing excessive pressure-drop across a packed section. This is important because
a high pressure drop would mean that more energy is required to drive the vapour up the distillation column.

Packings versus Trays

A tray column that is facing throughput problems may be de-bottlenecked by replacing a section of trays with
packings. This is because:
packings provide extra inter-facial area for liquid-vapour contact

efficiency of separation is increased for the same column height

packed columns are shorter than trayed columns

Packed columns are called continuous-contact columns while trayed columns are called staged-contact columns
because of the manner in which vapour and liquid are contacted.

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 [visit our sponsor]

an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction COLUMN REBOILERS

Types of Columns There are a number of designs of reboilers. It is beyond the scope of this set of introductory notes to delve into
Basic Equipment their design principles. However, they can be regarded as heat-exchangers that are required to transfer enough
and Operation energy to bring the liquid at the bottom of the column to boiling boint. The following are examples of typical
Column Internals reboiler types.

Reboilers
Distillation
Principles
Vapour Liquid
Equilibria
Distillation Column
Design
Effects of the
Number of Trays or
Stages
Factors Affecting
Operation
Crossword
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Copyright Photo courtesy of Brian Kennedy
Information
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A novel development in reboiler design is the self-cleaning shell-and-tube heat exchangers by Klarex Technology
for applications where heat exchange surfaces are prone to fouling by the process fluid. Particles are introduced
into the process stream and these produce a scouring action on the heat exchange surfaces. An example is shown
in the diagram on the left. [Click on it to find out more]

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 [visit our sponsor]

an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction DISTILLATION PRINCIPLES

Types of Columns
Separation of components from a liquid mixture via distillation depends on the differences in boiling points of the
Basic Equipment individual components. Also, depending on the concentrations of the components present, the liquid mixture will
and Operation have different boiling point characteristics. Therefore, distillation processes depends on the vapour pressure
Column Internals characteristics of liquid mixtures.
Reboilers Vapour Pressure and Boiling
Distillation
Principles The vapour pressure of a liquid at a particular temperature is the equilibrium pressure exerted by molecules
leaving and entering the liquid surface. Here are some important points regarding vapour pressure:
Vapour Liquid
Equilibria energy input raises vapour pressure
Distillation Column
Design vapour pressure is related to boiling
Effects of the a liquid is said to ‘boil’ when its vapour pressure equals the surrounding pressure
Number of Trays or
Stages the ease with which a liquid boils depends on its volatility
Factors Affecting
Operation liquids with high vapour pressures (volatile liquids) will boil at lower temperatures
Crossword the vapour pressure and hence the boiling point of a liquid mixture depends on the relative amounts of the
Other Resources components in the mixture
Copyright distillation occurs because of the differences in the volatility of the components in the liquid mixture
Information
Visit our sponsor The Boiling Point Diagram
The boiling point diagram shows how the equilibrium compositions of the components in a liquid mixture vary with
temperature at a fixed pressure. Consider an example of a liquid mixture containing 2 components (A and B) - a
binary mixture. This has the following boiling point diagram.

The boiling point of A is that at which the mole fraction of A is 1.

The boiling point of B is that at which the mole fraction of A is 0. In
this example, A is the more volatile component and therefore has a
lower boiling point than B. The upper curve in the diagram is called
the dew-point curve while the lower one is called the bubble-point
curve.

The dew-point is the temperature at which the saturated

vapour starts to condense.

The bubble-point is the temperature at which the liquid starts

to boil.

The region above the dew-point curve shows the equilibrium

composition of the superheated vapour while the region below the
bubble-point curve shows the equilibrium composition of the
subcooled liquid.

For example, when a subcooled liquid with mole fraction of A=0.4

(point A) is heated, its concentration remains constant until it reaches the bubble-point (point B), when it starts to
boil. The vapours evolved during the boiling has the equilibrium composition given by point C, approximately 0.8
mole fraction A. This is approximately 50% richer in A than the original liquid.

This difference between liquid and vapour compositions is the basis for distillation operations.

Relative Volatility
Relative volatility is a measure of the differences in volatility between 2 components, and hence their boiling
points. It indicates how easy or difficult a particular separation will be. The relative volatility of component ‘i’ with
respect to component ‘j’ is defined as

yi = mole fraction of component ‘i’ in the vapour

xi = mole fraction of component ‘i’ in the liquid

 [visit our sponsor]

an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction VAPOUR LIQUID EQUILIBRIA

Types of Columns
Distillation columns are designed based on the boiling point properties of the components in the mixtures being
Basic Equipment separated. Thus the sizes, particularly the height, of distillation columns are determined by the vapour liquid
and Operation equilibrium (VLE) data for the mixtures.
Column Internals
Vapour-Liquid-Equilibrium (VLE) Curves
Reboilers
Distillation Constant pressure VLE data is obtained from boiling point diagrams. VLE
Principles data of binary mixtures is often presented as a plot, as shown in the figure
on the right. The VLE plot expresses the bubble-point and the dew-point of
Vapour Liquid a binary mixture at constant pressure. The curved line is called the
Equilibria equilibrium line and describes the compositions of the liquid and vapour in
Distillation Column equilibrium at some fixed pressure.
Design
Effects of the This particular VLE plot shows a binary mixture that has a uniform vapour-liquid
Number of Trays or equilibrium that is relatively easy to separate. The next two VLE plots below on the
Stages other hand, shows non-ideal systems which will present more difficult separations. We
can tell from the shapes of the curves and this will be explained further later on.
Factors Affecting
Operation
Crossword
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The most intriguing VLE curves are generated by azeotropic systems. An azeotrope is a liquid mixture which when
vaporised, produces the same composition as the liquid. The two VLE plots below, show two different azeotropic
systems, one with a minimum boiling point and one with a maximum boiling point. In both plots, the equilibrium
curves cross the diagonal lines, and this are azeotropic points where the azeotropes occur. In other words
azeotropic systems give rise to VLE plots where the equilibrium curves crosses the diagonals.

Note the shapes of the respective equilibrium lines in relation to the diagonal lines that bisect the VLE plots.
Both plots are however, obtained from homogenous azeotropic systems. An azeotrope that contains one liquid
phase in contact with vapour is called a homogenous azeotrope. A homogenous azeotrope cannot be separated by
conventional distillation. However, vacumn distillation may be used as the lower pressures can shift the azeotropic
point.Alternatively, an additional substance may added to shift the azeotropic point to a more ‘favourable’ position.
When this additional component appears in appreciable amounts at the top of the column, the operation is
called azeotropic distillation.
When the additional component appears mostly at the bottom of the column, the operation is called
extractive distillation
The VLE curve on the left is also generated by an azeotropic system, in this
case a heterogenous azeotrope. Heterogenous azeotropes can be identified
by the ‘flat’ portion on the equilibrium diagram.

They may be separated in 2 distillation columns since these substances

usually form two liquid phases with widely differing compositions. The
phases may be separated using settling tanks under appropriate conditions.

Next, we will look at how VLE plots/data are used to design distillation columns.

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an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction DISTILLATION COLUMN DESIGN

Types of Columns
As mentioned, distillation columns are designed using VLE data for the mixtures to be separated. The vapour-
Basic Equipment liquid equilibrium characteristics (indicated by the shape of the equilibrium curve) of the mixture will determine the
and Operation number of stages, and hence the number of trays, required for the separation. This is illustrated clearly by
Column Internals applying the McCabe-Thiele method to design a binary column.
Reboilers McCABE-THIELE DESIGN METHOD
Distillation
Principles The McCabe-Thiele approach is a graphical one, and uses the VLE plot to determine the theoretical number of
stages required to effect the separation of a binary mixture. It assumes constant molar overflow and this implies
Vapour Liquid that:
Equilibria
Distillation Column molal heats of vaporisation of the components are roughly the same
Design
heat effects (heats of solution, heat losses to and from column, etc.) are negligible
Effects of the
Number of Trays or for every mole of vapour condensed, 1 mole of liquid is vaporised
Stages
Factors Affecting The design procedure is simple. Given the VLE diagram of the binary mixture, operating lines are drawn first.
Operation
Operating lines define the mass balance relationships between the liquid and vapour phases in the column.
Crossword
Other Resources There is one operating line for the bottom (stripping) section of the column, and on for the top (rectification
or enriching) section of the column.
Copyright
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Operating Line for the Rectification Section

The operating line for the rectification section is constructed as follows. First the desired top product
composition is located on the VLE diagram, and a vertical line produced until it intersects the diagonal line
that splits the VLE plot in half. A line with slope R/(R+1) is then drawn from this instersection point as
shown in the diagram below.

R is the ratio of reflux flow (L) to distillate flow (D) and is called the reflux ratio and is a measure of how
much of the material going up the top of the column is returned back to the column as reflux.
Operating Line for the Stripping Section
The operating line for the stripping section is constructed in a similar manner. However, the starting point is
the desired bottom product composition. A vertical line is drawn from this point to the diagonal line, and a
line of slope Ls/Vs is drawn as illustrated in the diagram below.
Ls is the liquid rate down the stripping section of the column, while Vs is the vapour rate up the stripping
section of the column. Thus the slope of the operating line for the stripping section is a ratio between the
liquid and vapour flows in that part of the column.
Equilibrium and Operating Lines
The McCabe-Thiele method assumes that the liquid on a tray and the vapour above it are in equilibrium. How
this is related to the VLE plot and the operating lines is depicted graphically in the diagram on the right.

A magnified section of the operating line for the stripping section is shown in relation to the corresponding
n'th stage in the column. L's are the liquid flows while V's are the vapour flows. x and y denote liquid and
vapour compositions and the subscripts denote the origin of the flows or compositions. That is 'n-1' will
mean from the stage below stage 'n' while 'n+1' will mean from the stage above stage 'n'. The liquid in
stage 'n' and the vapour above it are in equilibrium, therefore, xn and yn lie on the equilibrium line. Since the
vapour is carried to the tray above without changing composition, this is depicted as a horizontal line on the
VLE plot. Its intersection with the operating line will give the composition of the liquid on tray 'n+1' as the
operating line defines the material balance on the trays. The composition of the vapour above the 'n+1' tray
is obtained from the intersection of the vertical line from this point to the equilibrium line.
Number of Stages and Trays
Doing the graphical construction repeatedly will give rise to a number of 'corner' sections, and each section
will be equivalent to a stage of the distillation. This is the basis of sizing distillation columns using the
McCabe-Thiele graphical design methodology as shown in the following example.

Given the operating lines for both stripping and rectification

sections, the graphical construction described above was
applied. This particular example shows that 7 theoretical
stages are required to achieve the desired separation. The
required number of trays (as opposed to stages) is one less
than the number of stages since the graphical construction
includes the contribution of the reboiler in carrying out the
separation.

The actual number of trays required is given by the formula:

(number of theoretical trays)/(tray efficiency)

 Typical values for tray efficiency ranges from 0.5 to 0.7 and depends on a number of factors, such as the
type of trays being used, and internal liquid and vapour flow conditions. Sometimes, additional trays are
added (up to 10%) to accomodate the possibility that the column may be under-designed.
The Feed Line (q-line)
The diagram above also shows that the binary feed should be introduced at the 4'th stage. However, if the
feed composition is such that it does not coincide with the intersection of the operating lines, this means that
the feed is not a saturated liquid. The condition of the feed can be deduced by the slope of the feed line or q-
line. The q-line is that drawn between the intersection of the operating lines, and where the feed composition
lies on the diagonal line.

Depending on the state of the feed, the feed lines will have
different slopes. For example,

q = 0 (saturated vapour)
q = 1 (saturated liquid)
0 < q < 1 (mix of liquid and vapour)
q > 1 (subcooled liquid)
q < 0 (superheated vapour)

The q-lines for the various feed conditions are shown in the
diagram on the left.

Using Operating Lines and the Feed Line in McCabe-Thiele Design

If we have information about the condition of the feed mixture, then we can construct the q-line and use it in
the McCabe-Thiele design. However, excluding the equilibrium line, only two other pairs of lines can be used
in the McCabe-Thiele procedure. These are:

feed-line and rectification section operating line

feed-line and stripping section operating line
stripping and rectification operating lines

This is because these pairs of lines determine the third.

[see Flash tutorial on Distillation Basics written by Jon Lee]

OVERALL COLUMN DESIGN

Determining the number of stages required for the desired degree of separation and the location of the feed tray is
merely the first steps in producing an overall distillation column design. Other things that need to be considered
are tray spacings; column diameter; internal configurations; heating and cooling duties. All of these can lead to
conflicting design parameters. Thus, distillation column design is often an iterative procedure. If the conflicts are
not resolved at the design stage, then the column will not perform well in practice. The next set of notes will
discuss the factors that can affect distillation column performance.

Previous Next
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an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction EFFECTS OF THE NUMBER OF TRAYS OR STAGES

Types of Columns
Here we will expand on the design of columns by looking briefly at the effects of
Basic Equipment
and Operation the number of trays, and
Column Internals
the position of the feed tray, and
Reboilers
Distillation on the performances of distillation columns.
Principles
Effects of the Number of Trays
Vapour Liquid
Equilibria It can be deduced from the previous section on distillation column design that the number of trays will
influence the degree of separation. This is illustrated by the following example.
Distillation Column
Design
Consider as a base case, a 10 stage column. The feed is a binary mixture that has a composition of 0.5 mole
Effects of the fraction in terms of the more volatile component, and introduced at stage 5. The steady-state terminal
Number of Trays or compositions of about 0.65 at the top (stage 1) and 0.1 at the bottom (stage 10) are shown below:
Stages
Factors Affecting
Operation
Crossword
Other Resources
Copyright
Information
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Composition Profile: 10 stages, feed at stage 5

Suppose we decrease the number of stages to 8, and keep the feed at the middle stage, i.e. stage 4. The
resulting composition profile is:

Composition Profile: 8 stages, feed at stage 4

We can see that the top composition has decreased while the bottom composition has increased. That is, the
separation is poorer.

Now, if we increase the number of stages to 12, and again introduce the feed at mid-column, i.e. stage 6,
the composition profile we get is:
 Composition Profile: 12 stages, feed at stage 6

Again, the composition has changed. This time the distillate is much richer in the more volatile component,
while the bottoms has less, indicating better separation.

Thus, increasing the number of stages will improve separation.

Effect of Feed Tray Position
Here we look at how the position of the feed tray affects separation efficiency. Suppose we have a 20 stage
column, again separating a binary mixture that has a composition of 0.5 mole fraction in terms of the more
volatile component. The terminal compositions obtained when the feed is introduced at stages 5, 10 and 15
(at fixed reflux and reboil rates) are shown in the following plots.

Composition profile: 20 stages, feed at stage 5

Composition profile: 20 stages, feed at stage 10

Composition profile: 20 stages, feed at stage 15

 [Click on green button to see animated display of how the composition
profiles change with feed stage position]

As the feed stage is moved lower down the column, the top composition becomes less rich in the more
volatile component while the bottoms contains more of the more volatile component. However, the changes
in top composition is not as marked as the bottoms composition.

The preceding examples illustrate what can happen if the position of the feed tray is shifted for this particular
system. They should not be used to generalise to other distillation systems, as the effects are not
straightforward.

Previous Next
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an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction FACTORS AFFECTING DISTILLATION COLUMN OPERATION

Types of Columns
The performance of a distillation column is determined by many factors, for example:
Basic Equipment
and Operation feed conditions
Column Internals
state of feed
Reboilers composition of feed
Distillation trace elements that can severely affect the VLE of liquid mixtures
Principles
internal liquid and fluid flow conditions
Vapour Liquid
Equilibria state of trays (packings)
Distillation Column
Design weather conditions
Effects of the
Some of these will be discussed below to give an idea of the complexity of the distillation process.
Number of Trays or
Stages Feed Conditions
Factors Affecting The state of the feed mixture and feed composition affects the operating lines and hence the number of
Operation stages required for separation. It also affects the location of feed tray. During operation, if the deviations
Crossword from design specifications are excessive, then the column may no longer be able handle the separation task.
Other Resources To overcome the problems associated with the feed, some column are designed to have multiple feed points
when the feed is expected to containing varying amounts of components.
Copyright
Information Reflux Conditions
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As the reflux ratio is increased, the gradient of operating line
for the rectification section moves towards a maximum value of
1. Physically, what this means is that more and more liquid
that is rich in the more volatile components are being recycled
back into the column. Separation then becomes better and
thus less trays are needed to achieve the same degree of
separation. Minimum trays are required under total reflux
conditions, i.e. there is no withdrawal of distillate.

On the other hand, as reflux is decreased, the operating line

for the rectification section moves towards the equilibrium line.
The ‘pinch’ between operating and equilibrium lines becomes
more pronounced and more and more trays are required.This
is easy to verify using the McCabe-Thiele method.

The limiting condition occurs at minimum reflux ration, when

an infinite number of trays will be required to effect separation.
Most columns are designed to operate between 1.2 to 1.5
times the minimum reflux ratio because this is approximately
the region of minimum operating costs (more reflux means
higher reboiler duty).
Vapour Flow Conditions
Adverse vapour flow conditions can cause

foaming
entrainment
weeping/dumping
flooding

Foaming
Foaming refers to the expansion of liquid due to passage of vapour or gas. Although it provides high
interfacial liquid-vapour contact, excessive foaming often leads to liquid buildup on trays. In some
cases, foaming may be so bad that the foam mixes with liquid on the tray above. Whether foaming
will occur depends primarily on physical properties of the liquid mixtures, but is sometimes due to
tray designs and condition. Whatever the cause, separation efficiency is always reduced.
Entrainment
Entrainment refers to the liquid carried by vapour up to the tray above and is again caused by high
vapour flow rates. It is detrimental because tray efficiency is reduced: lower volatile material is
carried to a plate holding liquid of higher volatility. It could also contaminate high purity distillate.
Excessive entrainment can lead to flooding.
Weeping/Dumping
 This phenomenon is caused by low vapour flow. The pressure exerted by the vapour is insufficient to
hold up the liquid on the tray. Therefore, liquid starts to leak through perforations. Excessive weeping
will lead to dumping. That is the liquid on all trays will crash (dump) through to the base of the
column (via a domino effect) and the column will have to be re-started. Weeping is indicated by a
sharp pressure drop in the column and reduced separation efficiency.
Flooding
Flooding is brought about by excessive vapour flow, causing liquid to be entrained in the vapour up
the column. The increased pressure from excessive vapour also backs up the liquid in the
downcomer, causing an increase in liquid holdup on the plate above. Depending on the degree of
flooding, the maximum capacity of the column may be severely reduced. Flooding is detected by
sharp increases in column differential pressure and significant decrease in separation efficiency.
Column Diameter
Most of the above factors that affect column operation is due to vapour flow conditions: either excessive or
too low. Vapour flow velocity is dependent on column diameter. Weeping determines the minimum vapour
flow required while flooding determines the maximum vapour flow allowed, hence column capacity. Thus, if
the column diameter is not sized properly, the column will not perform well. Not only will operational
problems occur, the desired separation duties may not be achieved.
State of Trays and Packings
Remember that the actual number of trays required for a particular separation duty is determined by the
efficiency of the plate, and the packings if packings are used. Thus, any factors that cause a decrease in tray
efficiency will also change the performance of the column. Tray efficiencies are affected by fouling, wear and
tear and corrosion, and the rates at which these occur depends on the properties of the liquids being
processed. Thus appropriate materials should be specified for tray construction.
Weather Conditions
Most distillation columns are open to the atmosphere. Although many of the columns are insulated, changing
weather conditions can still affect column operation. Thus the reboiler must be appropriately sized to ensure
that enough vapour can be generated during cold and windy spells and that it can be turned down
sufficiently during hot seasons. The same applies to condensors.
These are some of the more important factors that can cause poor distillation column performance. Other factors
include changing operating conditions and throughputs, brought about by changes in upstream conditions and
changes in the demand for the products. All these factors, including the associated control system, should be
considered at the design stages because once a column is built and installed, nothing much can be done to rectify
the situation without incurring significant costs. The control of distillation columns is a field in its own right, but
that's another story.

Previous
 [visit our sponsor]

an introduction
© Copyright 1997-2003 by M.T. Tham

Introduction OTHER RESOURCES

Types of Columns
General Articles
Basic Equipment
and Operation
Andrew Sloley's Distillation page contains links to information about the history of distillation processes as
Column Internals well as a glossary of terms.
Reboilers Distillation - a Microsoft Concise Encyclopaedia article
Distillation Columns (introductory notes from Cornell Uni.)
Distillation Energy Conservation in Distillation - from cheresources.com
Principles
Vapour Liquid Extractive Distillation - a good detailed description of what the method is all about, including the separation of
Equilibria mixtures that form azeotropes.
Distillation Column
Design Reactive Distillation
Effects of the
Number of Trays or Reactive Distillation - an introduction by T.Mashue and H.S. Fogler (UMich)
Stages Website devoted to reactive distillation
Factors Affecting
Design
Operation
Crossword McCabe-Thiele
Other Resources Distillation column design using the McCabe-Thiele method (Stage 1 and above)
On-line McCabe-Thiele calculation of number of theoretical plates (Stage 1 and above)
Copyright
Calculation of a rectification column - an example from Uni. Koeln, Germany
Information
On-line distillation design for separating a single feed stream containing up to 20 components from databank
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of 480 components.
Calculating distillation tray efficiencies has plenty of illustrative photos

Internals

Distillation trays, with some nifty photographs (by tray specialist E. Frank Wijn)
Introduction to trays, in general
Trays, in more detail
Packings
Segmented Packed Columns - see what these guys are capable of
Case studies of the use of packings to improve column performance

Process Control

Distillation column control design using steady state models - describes the use of steady models to come up
with a practicable distillation column control strategy (by Hyprotech)
Pressure control of distillation columns by Andrew Sloley

Other Links

Distillation resources from miningco.com

Petroleum Topics: Refining of Petroleum

Refining of Petroleum

Petroleum is a complex mixture of organic liquids

called crude oil and natural gas, which occurs
naturally in the ground and was formed millions of
years ago. Crude oil varies from oilfield to oilfield
in colour and composition, from a pale yellow low
viscosity liquid to heavy black 'treacle' consistencies.

Crude oil and natural gas are extracted from the ground, on land or under the oceans, by
sinking an oil well and are then transported by pipeline and/or ship to refineries where
their components are processed into refined products. Crude oil and natural gas are of
little use in their raw state; their value lies in what is created from them: fuels, lubricating
oils, waxes, asphalt, petrochemicals and pipeline quality natural gas.

An oil refinery is an organised and coordinated arrangement of manufacturing processes

designed to produce physical and chemical changes in crude oil to convert it into
everyday products like petrol, diesel, lubricating oil, fuel oil and bitumen.

As crude oil comes from the well it contains a mixture of hydrocarbon compounds and
relatively small quantities of other materials such as oxygen, nitrogen, sulphur, salt and
water. In the refinery, most of these non - hydrocarbon substances are removed and the
oil is broken down into its various components, and blended into useful products.

Natural gas from the well, while principally methane, contains quantities of other
hydrocarbons - ethane, propane, butane, pentane and also carbon dioxide and water.
These components are separated from the methane at a gas fractionation plant.

Petroleum hydrocarbon structures

Petroleum consists of three main hydrocarbon groups:

Paraffins
★

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Petroleum Topics: Refining of Petroleum

These consist of straight or branched carbon rings saturated with hydrogen atoms, the
simplest of which is methane (CH4) the main ingredient of natural gas. Others in this
group include ethane (C2H6), and propane (C3H8).

Hydrocarbons

With very few carbon atoms (C1 to C4) are light in density and are gases under normal
atmospheric pressure. Chemically paraffins are very stable compounds.

Naphthenes

Naphthenes consist of carbon rings, sometimes with side chains, saturated with hydrogen
atoms. Naphthenes are chemically stable, they occur naturally in crude oil and have
properties similar to paraffins.

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Petroleum Topics: Refining of Petroleum

Aromatics

aromatic hydrocarbons are compounds that contain a ring of six carbon atoms with
alternating double and single bonds and six attached hydrogen atoms. This type of
structure is known as a benzene ring. They occur naturally in crude oil, and can also be
created by the refining process.

The more carbon atoms a hydrocarbon molecule has, the "heavier" it is (the higher is its
molecular weight) and the higher is its the boiling point.

Small quantities of a crude oil may be composed of compounds containing oxygen,

nitrogen, sulphur and metals. Sulphur content ranges from traces to more than 5 per cent.
If a crude oil contains appreciable quantities of sulphur it is called a sour crude; if it
contains little or no sulphur it is called a sweet crude.

The refining process

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Petroleum Topics: Refining of Petroleum

Every refinery begins with the separation of crude oil into different fractions by
distillation.

The fractions are further treated to convert

them into mixtures of more useful saleable
products by various methods such as
cracking, reforming, alkylation,
polymerisation and isomerisation. These
mixtures of new compounds are then
separated using methods such as fractionation
and solvent extraction. Impurities are
removed by various methods, e.g.
dehydration, desalting, sulphur removal and
hydrotreating.

Refinery processes have developed in

response to changing market demands for
certain products. With the advent of the
internal combustion engine the main task of
refineries became the production of petrol.
The quantities of petrol available from
distillation alone was insufficient to satisfy
consumer demand. Refineries began to look for ways to produce more and better quality
petrol. Two types of processes have been developed:

● breaking down large, heavy hydrocarbon molecules

● reshaping or rebuilding hydrocarbon molecules.

Distillation (Fractionation)

Because crude oil is a mixture of hydrocarbons with different boiling temperatures, it can
be separated by distillation into groups of hydrocarbons that boil between two specified
boiling points. Two types of distillation are performed: atmospheric and vacuum.

Atmospheric distillation takes place in a distilling column at or near atmospheric

pressure. The crude oil is heated to 350 - 400oC and the vapour and liquid are piped into
the distilling column. The liquid falls to the bottom and the vapour rises, passing through
a series of perforated trays (sieve trays). Heavier hydrocarbons condense more quickly
and settle on lower trays and lighter hydrocarbons remain as a vapour longer and
condense on higher trays.

Liquid fractions are drawn from the trays and removed. In this way the light gases,

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Petroleum Topics: Refining of Petroleum

methane, ethane, propane and butane pass out the top of the column, petrol is formed in
the top trays, kerosene and gas oils in the middle, and fuel oils at the bottom. Residue
drawn of the bottom may be burned as fuel, processed into lubricating oils, waxes and
bitumen or used as feedstock for cracking units.

To recover additional heavy distillates from this residue, it may be piped to a second
distillation column where the process is repeated under vacuum, called vacuum
distillation. This allows heavy hydrocarbons with boiling points of 450oC and higher to
be separated without them partly cracking into unwanted products such as coke and gas.

The heavy distillates recovered by vacuum distillation can be converted into lubricating
oils by a variety of processes. The most common of these is called solvent extraction. In
one version of this process the heavy distillate is washed with a liquid which does not
dissolve in it but which dissolves (and so extracts) the non-lubricating oil components
out of it. Another version uses a liquid which does not dissolve in it but which causes the
non-lubricating oil components to precipitate (as an extract) from it. Other processes
exist which remove impurities by adsorption onto a highly porous solid or which remove
any waxes that may be present by causing them to crystallise and precipitate out.

Reforming

Reforming is a process which uses heat, pressure and a catalyst (usually containing
platinum) to bring about chemical reactions which upgrade naphthas into high octane
petrol and petrochemical feedstock. The naphthas are hydrocarbon mixtures containing
many paraffins and naphthenes. In Australia, this naphtha feedstock comes from the

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Petroleum Topics: Refining of Petroleum

crudes oil distillation or catalytic cracking processes, but overseas it also comes from
thermal cracking and hydrocracking processes. Reforming converts a portion of these
compounds to isoparaffins and aromatics, which are used to blend higher octane petrol.

● paraffins are converted to isoparaffins

● paraffins are converted to naphthenes
● naphthenes are converted to aromatics

e.g.

catalyst
heptane -> toluene + hydrogen
C7H16 -> C7H8 + 4H2
catalyst
cyclohexane -> benzene + hydrogen
C6H12 -> C6H6 + 3H2

Cracking

Cracking processes break down heavier hydrocarbon molecules (high boiling point oils)
into lighter products such as petrol and diesel. These processes include catalytic cracking,
thermal cracking and hydrocracking.

e.g.

A typical reaction:

catalyst
C16H34 - C8H18 + C8H16
>

Catalytic cracking is used to convert heavy hydrocarbon fractions obtained by vacuum

distillation into a mixture of more useful products such as petrol and light fuel oil. In this
process, the feedstock undergoes a chemical breakdown, under controlled heat (450 -
500oC) and pressure, in the presence of a catalyst - a substance which promotes the
reaction without itself being chemically changed. Small pellets of silica - alumina or
silica - magnesia have proved to be the most effective catalysts.

The cracking reaction yields petrol, LPG, unsaturated olefin compounds, cracked gas

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Petroleum Topics: Refining of Petroleum

oils, a liquid residue called cycle oil, light gases and a solid coke residue. Cycle oil is
recycled to cause further breakdown and the coke, which forms a layer on the catalyst, is
removed by burning. The other products are passed through a fractionator to be separated
and separately processed.

Fluid catalytic cracking uses a catalyst in the form of a very fine powder which flows
like a liquid when agitated by steam, air or vapour. Feedstock entering the process
immediately meets a stream of very hot catalyst and vaporises. The resulting vapours
keep the catalyst fluidised as it passes into the reactor, where the cracking takes place and
where it is fluidised by the hydrocarbon vapour. The catalyst next passes to a steam
stripping section where most of the volatile hydrocarbons are removed. It then passes to a
regenerator vessel where it is fluidised by a mixture of air and the products of
combustion which are produced as the coke on the catalyst is burnt off. The catalyst then
flows back to the reactor. The catalyst thus undergoes a continuous circulation between
the reactor, stripper and regenerator sections.

The catalyst is usually a mixture of aluminium oxide and silica. Most recently, the
introduction of synthetic zeolite catalysts has allowed much shorter reaction times and
improved yields and octane numbers of the cracked gasolines.

Thermal cracking uses heat to break down the residue from vacuum distillation. The
lighter elements produced from this process can be made into distillate fuels and petrol.
Cracked gases are converted to petrol blending components by alkylation or
polymerisation. Naphtha is upgraded to high quality petrol by reforming. Gas oil can be
used as diesel fuel or can be converted to petrol by hydrocracking. The heavy residue is
converted into residual oil or coke which is used in the manufacture of electrodes,
graphite and carbides.

This process is the oldest technology and is not used in Australia.

Hydrocracking can increase the yield of petrol components, as well as being used to
produce light distillates. It produces no residues, only light oils. Hydrocracking is
catalytic cracking in the presence of hydrogen. The extra hydrogen saturates, or
hydrogenates, the chemical bonds of the cracked hydrocarbons and creates isomers with
the desired characteristics. Hydrocracking is also a treating process, because the
hydrogen combines with contaminants such as sulphur and nitrogen, allowing them to be
removed.

Gas oil feed is mixed with hydrogen, heated, and sent to a reactor vessel with a fixed bed
catalyst, where cracking and hydrogenation take place. Products are sent to a fractionator
to be separated. The hydrogen is recycled. Residue from this reaction is mixed again with
hydrogen, reheated, and sent to a second reactor for further cracking under higher

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Petroleum Topics: Refining of Petroleum

temperatures and pressures.

In addition to cracked naphtha for making petrol, hydrocracking yields light gases useful
for refinery fuel, or alkylation as well as components for high quality fuel oils, lube oils
and petrochemical feedstocks.

Following the cracking processes it is necessary to build or rearrange some of the lighter
hydrocarbon molecules into high quality petrol or jet fuel blending components or into
petrochemicals. The former can be achieved by several chemical process such as
alkylation and isomerisation.

Alkylation

Olefins such as propylene and butylene are produced by catalytic and thermal cracking.
Alkylation refers to the chemical bonding of these light molecules with isobutane to form
larger branched-chain molecules (isoparaffins) that make high octane petrol.

Olefins and isobutane are mixed with an acid catalyst and cooled. They react to form
alkylate, plus some normal butane, isobutane and propane. The resulting liquid is
neutralised and separated in a series of distillation columns. Isobutane is recycled as feed
and butane and propane sold as liquid petroleum gas (LPG).

e.g.

catalyst
isobutane + butylene -> isooctane
C4H10 + C4H8 -> C8H18

Isomerisation

Isomerisation refers to chemical rearrangement of straight-chain hydrocarbons

(paraffins), so that they contain branches attached to the main chain (isoparaffins). This
is done for two reasons:

● they create extra isobutane feed for alkylation

● they improve the octane of straight run pentanes and hexanes and hence make
them into better petrol blending components.

Isomerisation is achieved by mixing normal butane with a little hydrogen and chloride

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Petroleum Topics: Refining of Petroleum

and allowed to react in the presence of a catalyst to form isobutane, plus a small amount
of normal butane and some lighter gases. Products are separated in a fractionator. The
lighter gases are used as refinery fuel and the butane recycled as feed.

Pentanes and hexanes are the lighter components of petrol. Isomerisation can be used to
improve petrol quality by converting these hydrocarbons to higher octane isomers. The
process is the same as for butane isomerisation.

Polymerisation

Under pressure and temperature, over an acidic catalyst, light unsaturated hydrocarbon
molecules react and combine with each other to form larger hydrocarbon molecules.
Such process can be used to react butenes (olefin molecules with four carbon atoms) with
iso-butane (branched paraffin molecules, or isoparaffins, with four carbon atoms) to
obtain a high octane olefinic petrol blending component called polymer gasoline.

Hydrotreating and sulphur plants

A number of contaminants are found in crude oil. As the fractions travel through the
refinery processing units, these impurities can damage the equipment, the catalysts and
the quality of the products. There are also legal limits on the contents of some impurities,
like sulphur, in products.

Hydrotreating is one way of removing many of the contaminants from many of the
intermediate or final products. In the hydrotreating process, the entering feedstock is
mixed with hydrogen and heated to 300 - 380oC. The oil combined with the hydrogen
then enters a reactor loaded with a catalyst which promotes several reactions:

● hydrogen combines with sulphur to form hydrogen sulphide (H2S)

● nitrogen compounds are converted to ammonia

● any metals contained in the oil are deposited on the catalyst

● some of the olefins, aromatics or naphthenes become saturated with hydrogen to

become paraffins and some cracking takes place, causing the creation of some
methane, ethane, propane and butanes.

Sulphur recovery plants

The hydrogen sulphide created from hydrotreating is a toxic gas that needs further

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Petroleum Topics: Refining of Petroleum

treatment. The usual process involves two steps:

● the removal of the hydrogen sulphide gas from the hydrocarbon stream

● the conversion of hydrogen sulphide to elemental sulphur, a non-toxic and useful

chemical.

Solvent extraction, using a solution of diethanolamine (DEA) dissolved in water, is

applied to separate the hydrogen sulphide gas from the process stream. The hydrocarbon
gas stream containing the hydrogen sulphide is bubbled through a solution of
diethanolamine solution (DEA) under high pressure, such that the hydrogen sulphide gas
dissolves in the DEA. The DEA and hydrogen mixture is the heated at a low pressure and
the dissolved hydrogen sulphide is released as a concentrated gas stream which is sent to
another plant for conversion into sulphur.

Conversion of the concentrated hydrogen sulphide gas into sulphur occurs in two stages.

1. Combustion of part of the H2S stream in a furnace, producing sulphur dioxide

(SO2) water (H2O) and sulphur (S).

2H2S + 2O2 - SO2 + S + 2H2O

>
2. Reaction of the remainder of the H2S with the combustion products in the
presence of a catalyst. The H2S reacts with the SO2 to form sulphur.

2H2S + 2O2 - 3S + 2H2O

>

As the reaction products are cooled the sulphur drops out of the reaction vessel in a
molten state. Sulphur can be stored and shipped in either a molten or solid state.

Click here to view a flow chart of a refinery .

Refineries and the environment

Air, water and land can all be affected by refinery operations. Refineries are well aware
of their responsibility to the community and employ a variety of processes to safeguard
the environment.

The processes described below are those used by the Shell refinery at Geelong in
Victoria, but all refineries employ similar techniques in managing the environmental
aspects of refining.

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Petroleum Topics: Refining of Petroleum

Air

Preserving air quality around a refinery involves controlling the following emissions:

● sulphur oxides
● hydrocarbon vapours
● smoke
● smells

Sulphur enters the refinery in crude oil feed. Gippsland and most other Australian crude
oils have a low sulphur content but other crude's may contain up to 5 per cent sulphur. To
deal with this refineries incorporate a sulphur recovery unit which operates on the
principles described above.

Many of the products used in a refinery produce hydrocarbon vapours. The escape of
vapours to atmosphere are prevented by various means. Floating roofs are installed in
tanks to prevent evaporation and so that there is no space for vapour to gather in the
tanks. Where floating roofs cannot be used, the vapours from the tanks are collected in a
vapour recovery system and absorbed back into the product stream. In addition, pumps
and valves are routinely checked for vapour emissions and repaired if a leakage is found.

Smoke is formed when the burning mixture contains insufficient oxygen or is not
sufficiently mixed. Modern furnace control systems prevent this from happening during
normal operation.

Smells are the most difficult emission to control and the easiest to detect. Refinery smells
are generally associated with compounds containing sulphur, where even tiny losses are
sufficient to cause a noticeable odour.

Water

Aqueous effluent's consist of cooling water, surface water and process water.

The majority of the water discharged from the refinery has been used for cooling the
various process streams. The cooling water does not actually come into contact with the
process material and so has very little contamination. The cooling water passes through
large "interceptors" which separate any oil from minute leaks etc., prior to discharge. The
cooling water system at Geelong Refinery is a once-through system with no recirculation.

Rainwater falling on the refinery site must be treated before discharge to ensure no oily
material washed off process equipment leaves the refinery. This is done first by passing

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Petroleum Topics: Refining of Petroleum

the water through smaller "plant oil catchers", which each treat rainwater from separate
areas on the site, and then all the streams pass to large "interceptors" similar to those
used for cooling water. The rainwater from the production areas is further treated in a
Dissolved Air Flotation (DAF) unit. This unit cleans the water by using a flocculation
agent to collect any remaining particles or oil droplets and floating the resulting flock to
the surface with millions of tiny air bubbles. At the surface the flock is skimmed off and
the clean water discharged.

Process water has actually come into contact with the process streams and so can contain
significant contamination. This water is treated in the "sour water treater" where the
contaminants (mostly ammonia and hydrogen sulphide) are removed and then recovered
or destroyed in a downstream plant. The process water, when treated in this way, can be
reused in parts of the refinery and discharged through the process area rainwater
treatment system and the DAF unit.

Any treated process water that is not reused is discharged as Trade Waste to the sewerage
system. This trade waste also includes the effluent from the refinery sewage treatment
plant and a portion of treated water from the DAF unit.

As most refineries import and export many feed materials and products by ship, the
refinery and harbour authorities are prepared for spillage from the ship or pier. In the
event of such a spill, equipment is always on standby at the refinery and it is supported
by the facilities of the Australian Marine Oil Spill Centre at Geelong, Victoria.

Land

The refinery safeguards the land environment by ensuring the appropriate disposal of all
wastes.

Within the refinery, all hydrocarbon wastes are recycled through the refinery slops
system. This system consists of a network of collection pipes and a series of dewatering
tanks. The recovered hydrocarbon is reprocessed through the distillation units.

Wastes that cannot be reprocessed are either recycled to manufacturers (e.g. some spent
catalysts can be reprocessed), disposed of in EPA-approved facilities off-site, or
chemically treated on-site to form inert materials which can be disposed to land-fill
within the refinery.

Waste movements within the refinery require a "Process liquid, Sludge and Solid waste
disposal permit". Wastes that go off-site must have an EPA "Waste Transport Permit".

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 Process Consulting Services
Index

PCS Students Guide to Refining

PCS Hydroprocess Options

PCS Increasing FCC propylene recovery

PCS newsletter vol. 1.1

PCS newsletter vol. 1.2

PCS newsletter vol. 1.3

PCS newsletter vol. 1.4

PCS newsletter vol. 2.1

PCS newsletter vol. 2.2

PCS Revamping Crude Units ★

 The high speed and low cost of
computing permits sophisticated
computer technologies to be
applied in a petroleum refinery.
What have not changed are the
fundamental workings of a
petroleum refinery. Refinery units
are made of process equipment
configured to meet certain
objectives. Ultimately, it is the
process flow scheme, equipment
performance, and operating model
(whether sophisticated or not) that
determine maximum profitability.
No advanced technologies will
overcome process and equipment
shortcomings.

Low graphic / printable site version

Flash movie EXPERIENCED PEOPLE WORKING

THE FUNDAMENTALS WILL MAXIMIZE SUSTAINABLE PROFIT IMPROVEMENTS.
 Shared Content in cooperation with Refining Technology Online

Students' Guide to Refining

Crude Distillation

Distillation is the first step in the processing of crude oil and it takes place in a tall steel
tower called a fractionation column. The inside of the column is divided at intervals by
horizontal trays. The column is kept very hot at the bottom (the column is insulated) but
as different hydrocarbons boil at different temperatures, the temperature gradually
reduces towards the top, so that each tray is a little cooler than the one below.

The crude needs to be heated up before entering the fractionation column and this is done
at first in a series of heat exchangers where heat is taken from other process streams
which require cooling before being sent to rundown. Heat is also exchanged against
condensing streams from the main column. Typically, the crude will be heated up in this
way upto a temperature of 200 - 280 0C, before entering a furnace.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

As the raw crude oil arriving contains quite a bit of water and salt, it is normally sent for
salt removing first, in a piece of equipment called a desalter. Upstream the desalter, the
crude is mixed with a water stream, typically about 4 - 6% on feed. Intense mixing takes
place over a mixing valve and (optionally) as static mixer. The desalter, a large liquid full
vessel, uses an electric field to separate the crude from the water droplets. It operates best
at 120 - 150 0C, hence it is conveniently placed somewhere in the middle of the preheat
train.

Part of the salts contained in the crude oil, particularly magnesium chloride, are
hydrolysable at temperatures above 120 0C. Upon hydrolysis, the chlorides get converted
into hydrochloric acid, which will find its way to the distillation column's overhead
where it will corrode the overhead condensers. A good performing desalter can remove
about 90% of the salt in raw crude.

Downstream the desalter, crude is further heated up with heat exchangers, and starts
vaporising, which will increase the system pressure drop. At about 170 -200 0C, the
crude will enter a 'pre-flashvessel', operating at about 2 - 5 barg, where the vapours are
separated from the remaining liquid. Vapours are directly sent to the fractionation
column, and by doing so, the hydraulic load on the remainder of the crude preheat train
and furnace is reduced (smaller piping and pumps).

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Just upstream the preflash vessel, a small caustic stream is mixed with the crude, in order
to neutralise any hydrochloric acid formed by hydrolysis. The sodium chloride formed
will leave the fractionation column via the bottom residue stream. The dosing rate of
caustic is adjusted based on chloride measurements in the overhead vessel (typically 10 -
20 ppm).

At about 200 - 280 0C the crude enters the furnace where it is heated up further to about
330 -370 0C. The furnace outlet stream is sent directly to the fractionation column. Here,
it is separated into a number of fractions, each having a particular boiling range.

At 350 0C, and about 1 barg, most of the fractions in the crude oil vapourise and rise up
the column through perforations in the trays, losing heat as they rise. When each fraction
reaches the tray where the temperature is just below its own boiling point, it condenses
and changes back into liquid phase. A continuous liquid phase is flowing by gravity
through 'downcomers' from tray to tray downwards. In this way, the different fractions
are gradually separated from each other on the trays of the fractionation column. The
heaviest fractions condense on the lower trays and the lighter fractions condense on the
trays higher up in the column. At different elevations in the column, with special trays
called draw-off trays, fractions can be drawn out on gravity through pipes, for further
processing in the refinery.

At top of the column, vapours leave through a pipe and are routed to an overhead
condenser, typically cooled by air fin-fans. At the outlet of the overhead condensers, at
temperature about 40 0C, a mixture of gas, and liquid naphtha exists, which is falling into
an overhead accumulator. Gases are routed to a compressor for further recovery of LPG
(C3/C4), while the liquids (gasoline) are pumped to a hydrotreater unit for sulfur
removal.

A fractionation column needs a flow of condensing liquid downwards in order to provide

a driving force for separation between light and heavy fractions. At the top of the column
this liquid flow is provided by pumping a stream back from the overhead accumulator
into the column. Unfortunately, a lot of the heat provided by the furnace to vaporise
hydrocarbons is lost against ambient air in the overhead fin-fan coolers. A clever way of
preventing this heat lost of condensing hydrocarbons is done via the circulating refluxes
of the column. In a circulating reflux, a hot side draw-off from the column is pumped
through a series of heat exchangers (against crude for instance), where the stream is
cooled down. The cool stream is sent back into the column at a higher elevation, where it
is been brought in contact with hotter rising vapours. This provides an internal
condensing mechanism inside the column, in a similar way as the top reflux does which
is sent back from the overhead accumulator. The main objective of a circulating reflux
therefore is to recover heat from condensing vapours. A fractionating column will have
several (typically three) of such refluxes, each providing sufficient liquid flow down the
corresponding section of the column. An additional advantage of having circulating

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

refluxes is that it will reduce the vapour load when going upwards in the column. This
provided the opportunity to have a smaller column diameter for top sections of the tower.
Such a reduction in diameter is called a 'swage'.

The lightest side draw-off from the fractionating column is a fraction called kerosene,
boiling in the range 160 - 280 0C, which falls down through a pipe into a smaller column
called 'side-stripper'. The purpose of the side stripper is to remove very light
hydrocarbons by using steam injection or an external heater called 'reboiler'. The
stripping steam rate, or reboiled duty is controlled such as to meet the flashpoint
specification of the product. Similarly to the atmospheric column, the side stripper has
fractionating trays for providing contact between vapour and liquid. The vapours
produced from the top of the side stripper are routed back via pipe into the fractionating
column.

The second and third (optional) side draw-offs from the main fractionating column are
gasoil fractions, boiling in the range 200 - 400 0C, which are ultimately used for blending
the final diesel product. Similar as with the kerosene product, the gasoil fractions (light
and heavy gasoil) are first sent to a side stripper before being routed to further treating
units.

At the bottom of the fractionation column a heavy, brown/black coloured fraction called
residue is drawn off. In order to strip all light hydrocarbons from this fraction properly,
the bottom section of the column is equipped with a set of stripping trays, which are
operated by injecting some stripping steam (1 - 3% on bottom product) into the bottom of
the column.

Hydrotreating

The objective of the Hydrotreating prococess is to remove suplur as well as other

unwanted compunds, e.g. unsaturated hydrocarbons, nitrogen from refinery process
streams.

Until the end of World War 2, there was little incentive for the oil industry to pay
significant attention to improving product quality by hydrogen treatment. However, soon
after the war the production of high sulphur crudes increased significantly, which gave a
more stringent demand on the product blending flexibility of refineries, and the
marketing specifications for the products became tighter, largely due to environmental
considerations. Furthermore, the catalyst used in the Platforming process can only handle
sulfur in the very low ppm level, so hydrotreating of naphtha became a must. The
necessity for hydrotreating of middle distillates (kerosene/gasoil) originates from
pressure to reduce sulfur emissions into the environment. Overall, this situation resulted
in an increased necessity for high sulphur removal capability in many refineries.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

As catalytic reforming gives hydrogen as a byproduct, it gave additional momentum to

the development of sulphur removal process by hydrogen treatment. In this treatment, the
sulphur compounds are removed by converting them into hydrogen sulphide by reaction
with hydrogen in the presence of a catalyst. This results in high liquid product yields,
since only sulphur is removed. Furthermore, the hydrogen sulphide produced can be
easily removed from the product gas stream, for example by an amine wash. In this way,
hydrogen sulphide is recovered as a higly concentrated stream and can be further
converted into elemental sulphur via the "Claus" process.

Hydrodesulphursiation has been extensively used commercially for treating naphtha as

feedstock for catalytic reformers to meet the very stringent sulphuir specification of less
than 1 ppm wt to protect the platinum catalyst. It has also been widely used for removal
of sulphur compounds from kerosine and gasoils to make them suitable as blending
components. In cases where products are from catalytic or thermal crackers, hydrogen
treatment is used to improve product quality specifications like colour, smoke point,
cetane index, etc.

For Hydrotreating, two basic processes are applied, the liquid phase (or trickle flow)
process for kerosine and heavier straight-run and cracked distillates up to vacumn gas oil
and the vapour phase process for light straight-run and cracked fractions.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Both processes use the same basic configuration: the feedstock is mixed with hydrogen-
rich make up gas and recycle gas. The mixture is heated by heat exchange with reactor
effluent and by a furnace and enters a reactor loaded with catalyst. In the reactor, the
sulphur amd nitrogen compounds present in the feedstock are converted into hydrogen
sulphide and ammonia respectively. The olefins present are saturated with hydrogen to
become di-olefins and part of the aromatics will be hydrogenated. If all aromatics needs
to be hydrogenated, a higher pressure is needed in the reactor compared to the
conventional operating mode.

The reactor operates at temperatures in the range of 300-380 0C and at a pressure of 10-
20 bar for naphta and kero, as compared with 30-50 bar for gasoil, with excess hydrogen
supplied. The temperature should not exceed 380 0C, as above this temperature cracking
reactions can occur, which deteriorates the colour of the final product. The reaction
products leave the reactor and, after having been cooled to a low temperature, typically
40-50 0C, enter a liquid/gas separation stage. The hydrogen-rich gas from the high
pressure separation is recycled to combine with the feedstock, and the low pressure off-
gas stream rich in hydrogen sulphide is sent to a gas-treating unit, where hydrogen
sulphide is removed. The clean gas is then suitable as fuel for the refinery furnaces. The
liquid stream is the product from hydotreating. It is normally sent to a stripping column
where H2S and other undesirable components are removed, and finally, in cases where
steam is used for stripping, the product is sent to a vacumn drier for removal of water.
Some refiners use a salt dryer in stead of a vacuum drier to remove the water.
The catalyst used is normally cobalt, molybdenum and nickel finely distributed on
alumina extrudates. It slowly becomes choked by coke and must be renewed at regular
intervals (typically 2-3 years). It can be regenerated (by burning off the coke) and reused
typically once or twice before the breakdown of the support's porous structure
unacceptably reduces its activity. Catayst regeneration is, nowadays, mainly carried out
ex- situ by specialised firms. Other catalysts have also been developed for applications
where denitrification is the predominant reaction required or where high stauration of
olefins is necessary.

A more recent development is the application of Hydrotreating for pretreatment of

feedstcok for the catalytic cracking process. By utilisation of a suitable hydrogenation-
promoting catalyst for conversion of aromatics and nitrogen in potential feedstocks, and
selection of severe operating conditions, hydrogen is taken up by the aromatic molecules.
The increased hydrogen content of the feedstock obtained by this treatment leads to
significant conversion advantages in subsequent catalytic cracking, and higher yield of
light products can be achieved.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Hydrotreatment can also be used for kerosine smoke point improvement (SPI). It closely
resembles the conventional Hydrotreating Process however an aromatic hydrogenation
catalyst consisting of noble metals on a special carrier is used. The reactor operates at
pressure range of 50-70 bar and temperatures of 260-320 0C. To restrict temperature rise
due to the highly exothermic aromatics conversion reactions, quench oil is applied
between the catalysts beds. The catalyst used is very sensitive to traces of sulphur and
nitrogen in the feedstock and therefore pretreatment is normally applied in a conventional
hydrotreater before kerosine is introduced into the SPI unit. The main objective of Smoke
Point Improvement is improvement in burning characteristics as the kerosine aromatics
are converted to naphthenes.

Hydrotreatment is also used for production of feedstocks for isomersiation unit from
pyrolysis gasoline (pygas) which is one of the byproducts of steam cracking of
hydrocarbon fractions such as naphtha and gasoil.

A hyrotreater and a hydrodesulphuriser are basically the same process but a hydotreater
termed is used for treating kerosene or lighter feedstock, while a hydodesulhuriser mainly
refers to gasoil treating. The hydrotreatment process is used in every major refinery and
is therefore also termed as the work horse of the refinery as it is the hydrotreater unit that
ensures several significant product quality specifications. In most countries the Diesel
produced is hydrodesulhurised before its sold. Sulphur specifications are getting more
and more stringent. In Asia, countries such as Thailand, Singapore and Hong Kong
already have a 0.05%S specification and large hydrodesulphurisation units are required to
meet such specs.

The by-products obtained from HDT/HDS are light ends formed from a small amounts of
cracking and these products are used in the refinery fuelgas pool. The other main by-
product is Hydrogen Sulphide which is oxidized to sulphur and sold to the chemical
industry for further processing.

In combination with temperature, the pressure level (or rather the partial pressure of
hydrogen) generally determines the types of components that can be removed and also
determines the working life of the catalyst. At higher (partial) pressures, the
desulphurisation process is 'easier', however, the unit becomes more expensive for
instance due to larger compressors and heavier reactors. Also, at higher pressure, the
hydrogen consumption of the unit increases, which can be a signficant cost factor for the
refinery. The minimum pressure required typically goes up with the required severity of
the unit, i.e. the heavier the feedstock, or the lower levels of sulphur in product required.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Platforming

Motor gasoline (Mogas) production starts with the distillation of crude oil. One of the
products out of that process is a fraction of low octane gasoline, normally referred to as
naphtha, typically boiling in the range 100 - 160 0C. Other gasoline fractions are
produced as a result of secondary processes like catalytic cracking, isomerisation,
alkylation and platforming. Petrol is then produced by blending a variety of these
gasoline components of different qualities to meet a series of product specifications.

One very important property of Mogas is the octane number, which influences
"knocking" or "pinking" behaviour in the engine of cars. Traditionally lead compounds
have been added to petrol to improve the octane number. Over the past years, in many
countries legislation has been implemented aimed at reducing the emission of lead from
exhausts of motor vechiles and this, calls for other means of raising the octane number.

The role of a platformer is to pave the way for this by a process which reforms the
molecules in low octane naphtha to produce a high octane gasoline component. This is
achieved by employing a catalyst with platinum as its active compound; hence the name
Platformer. For many refinery catalyst applications, a promoter is used, and in the
platforming process, it is a chloride promoter which stimulates the 'acidity' of the catalyst
and thereby the isomerisation reactions. Often, a bimetallic catalyst is used, i.e. in
addition to the platinum, a second metal, for instance Rhenium is present on the catalyst.
The main advantage is a higher stability under reforming conditions. The disadvantage is
that the catalyst becomes more sensitive towards poisons, process upsets and more
susceptible to non-optimum regenerations.

Chemistry:
By: Hardeep Hundal, original editing by Jeroen Buren
Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

The main reactions of platforming process are as follows:

• Dehydrogenation of naphthenes, yielding aromatics and hydrogen
• Dehydro-isomerisation of alkyl cyclopentanes to aromatic and hydrogen
• Isomerisation of paraffins and aromatics
• Dehydrocyclisation of paraffins to aromatics and hydrogen
• Hydrocracking of paraffins and naphthenes to ligher, saturated paraffins at the
expense of hydrogen

The process literally re-shapes the molecules of the feed in a reaction in the presence of a
platinum catalyst. Normally it is the hydrocarbon in the C6-C10 parafins that get
converted to aromatics.

The above reactions takes place concurrently and to a large extent also sequentially. The
majority of these reactions, involve the conversion of paraffins and naphtenes and result
in an increase in octane number and a nett production of hydrogen. Characteristic of the
total effect of these reactions is the high endothermicity, which requires the continuous
supply of process heat to maintain reaction temperature in the catalyst beds. That is why
the process is typically done in four reactors in series with furnaces in between, in order
to remain sufficiently high reactor temperatures.

The reactions takes place at the surface of the catalyst and are very much dependent,
amongst other factors, on the right combination of interactions between platinum, its
modifiers or activators, the halogen and the catalyst carrier. During operating life of the
catalyst, the absolute and relative reaction rates are influenced negatively by disturbing
factors like gradual coke deposition, poisons and deterioration of physical characteristic
of the catalyst (surface area decline).

The process of platforming:

The feedstock of the platformer is drawn from the refinery's distillation units. This is first
treated by passing the feedstock together with hydrogen over a catayst, in a process called
'hydrotreating, to convert the sulphur and nitrogen compunds to hydrogen sulphide and
ammonia, in order to prevent poisoning of the expensive platformer catalyst. After
hydrotreating, ,the reactor effluent moves on through a stabiliser column to remove the
gases formed (hydrogen sulphide, ammonia and fuel gas). In a second column, the C5
and some of the C6 is removed in a separate fraction called 'tops'. The reason to remove
C5/C6 is that this component will crack in the platformer to produce fuel gas, while C6
gets converted into benzene, which can only be allowed in limited amount into the mogas
because of its toxicity. From the bottom of the splitter column, the naphtha stream is
produced, which is the feed for the Platforming section.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

At the heart of the Platformer process are the four reactors, each linked to furnaces to
sustain a suffiently high reaction temperature, about 500 0C at the inlet of the reactors.

Over time, coke will build up on the catalyst surface area, which reduces the catalyst
activity. The catalyst can be easily regenerated however, by burning the coke off with air.
After coke burning, the catalyst needs to be reconditioned by a combined treatment of air
and HCl under high temperature. This regeneration step is called 'oxy-chlorination'. After
this step the catalyst is dryed with hot nitrogen and subsequently brought in its active
condition by reducing the surface with hot hydrogen. The refinery will therefore regulary
have to take out one of the reactors to undergo this regeneration process. This type of
process is therefore called semi-regen platforming.

During the regeneration process, the refinery will suffer production loss, which is the
reason why UOP developed a major process enhancement by making the regeneration
possible continuously, in a Continous Catalytic Reformer, CCR. In the CCR unit, the
reactors are cleverly stacked, so that the catalyst can flow under gravity. From the bottom
of the reactor stack, the 'spent' catalyst is 'lifted' by nitrogen to the top of the regenerator
stack. In the regenerator, the above mentioned different steps, coke burning,
oxychlorination and drying are done in different sections, segregated via a complex
system of valves, purge-flows and screens. From the bottom of the regenerator stack,
catalyst is lifted by hydrogen to the top of the reactor stack, in a special area called the
reduction zone. In the reduction zone, the catalyst passes a heat exchanger in which it is
heated up against hot feed. Under hot conditions it is brought in contact with hydrogen,
which performs a reduction of the catalyst surface, thereby restoring its activity. In such a
continuous regeneration process, a constant catalyst activity can be maintained without
unit shutdown for a typical runlength of 3 - 6 years. After 300 - 400 cycles of
reaction/regeneration, the surface area of the catalyst will have dropped to a level (120 -
130 m2/g) that it becomes more difficult to maintain catalyst activity and at such a time
normally the catalyst will be replaced by a fresh batch. The batch of 'spent' catalyst is
then sent for platinum reclaim to recover the valuable precious metals.

For economic reasons, the design capacities of Platformer units vary from 1000 - 4500
t/d; operating pressures can vary over a wide range, units with from 3.5 barg up to 30
barg can be found, whereby the latest generation CCR's are typically at the lower
pressure range. A lower pressure enhances the endothermic reactions, which gives less
cracking reactions and thereby a higher liquid yield. However, at a lower reactor
operating pressure, the hydrogen partial pressure will be lower as well, which favours
coke formation. The reason why semi regen platformers will not operate at a too low
pressure, otherwise the cycle length between regenerations becomes to short. A second
disadvantage of operating at a lower pressure is that a larger compressor will be required
to boost the pressure of the hydrogen up to the normal pressure of the hydrogen system
(about 20 barg). Typical design reformate octane numbers are in the 95-104 range. The
reactor temperature is in a region of 450-530 0C.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

At the outlet of the last reactor the product is still well above 400 0C. It is cooled down
against cold feed in massive heat exchanger, either a so called 'Texas Tower' or a
Packinox plate-pack heat exchanger. The special design of those heat exchangers ensures
that minimum heat loss occurs in order to minimise the fuel consumption of the furnaces.
After passing the feed/effluent exchanger, the reaction products are cooled in air/water
coolers and routed to a product separator, where the hydrogen is the main gaseous
product. Part of the hydrogen produced is recycled back (via a compressor) to the feed, in
order to maintain a high enough hydrogen partial pressure in the reactors. The remainder
of the gases are compressed and brought in contact again with the liquid from the product
separator. This is step is called 'recontacting' and is done in order to recover as much as
possible hydrocarbons from the hydrogen produced. The reactor product, now in liquid
form, goes on to the platformer stabiliser which removes Liquid Petroleum Gas ( LPG)
and other gases to leave a liquid high octane gasoline component called platformate,
ready for blending into the refinery mogas pool. Summarising, the Platformer unit
produces about 85% liquid platformate, 10% hydrogen and 5% LPG.

The Continuous Catalytic Reforming unit or better known as CCR Platformer is licensed
by UOP, Universal Oil Products, based in USA. More recently, other technology vendors
have copied the concept, one of the main competitors for UOP in this field is IFP from
France.

Main Equipment in a CCR Platformer:

A CCR typically contains a feed/effluent heat exchanger (Texas Tower or Packinox), 4
furnaces, 4 reactors, a regenerator, overhead recontacting section, net gas compressor,
recycle gas compressor and a stabiliser column.

Isomerization

The isomerisation process involves the transformation of one molecular structure into
another (isomer) whose component atoms are the same but arranged in a different
geometrical structure. Since isomers may differ greatly in physical and chemical
properties, isomerisation offers the possibility of converting less desirable compounds
into isomers with desirable properties, in particular to convert n-paraffins into iso-
paraffins, thereby increasing the octane of the hydrocarbon stream. The main fields of
application of isomerisation are:
• ISOMERISATION of normal butane into isobutane
• ISOMERISATION of pentanes and hexanes into higher- branched isomers

Since branched isomers have a higher antiknock quality than the corresponding linear
paraffins, this form of isomerisation is important for the production of motor fuels.

By: Hardeep Hundal, original editing by Jeroen Buren

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In addition to the above applications, isomerisation is applied for the conversion of ortho-
xylene and meta- xylene into para- xylene, used for the manufacturing of polester fibres.

Isomerisation of low molecular weight paraffins has been commercially applied for many
years. After extensive laboratory work had been carried out during the 1930s, World War
2 prompted the development of the laboratory processes into full- scale commercial units
in order to meet the demand for isobutane necessary for the manufacture of large amounts
of alkylate. While the first butane isomerisation unit went on stream in late 1941, by the
end of the war nearly 40 butane isomerisation units were in operation in the USA and the
Caribbean. Two pentane and two light naphtha isomerisation units also came on stream
towards the end of the war to provide an additional source of blending aviation gasoline.

Though butane isomerisation has maintained its importance, present day interest
isomerisation is specially focussed on the upgrading of fractions containing C5 Pentane
and C6 Hexane for use as motor gasoline components. This application has been
prompted by the world drive to remove the lead additives gradually from motor gasoline
in order to reduce air pollution. The octane loss caused by the removal or reduction of
lead antiknock additives can be compensated for by isomerisation of pentane/hexane
paraffin fraction of the light gasoline fractions.

Isomerisation technology has also improved substantially due to the hard work of many
technologist. In order to achieve the low temperature necessary to obtain an acceptable
yield of isomers, the Catalyst systems used in the early units were based on aluminium
chloride in some form. These catalyst systems, however, had the drawback of being
highly corrosive and difficult to handle. In recent years, catalyst of a different type have
come in use. These are solid catalysts consisting of a support having an acidic carrier and
a hydrogenation function, frequently a noble metal. Modern isomerisation units utilise

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these dual- function catalysts and operate in the vapour phase and the presence of
hydogen. For these reasons, these process are called hydro- isomerisation processes.

The first hydro- isomerisation unit was introduced nin 1953 by UOP, followed in 1965 by
the first BP one, while in 1970 the first Shell hydro-isomerisation (HYSOMER) unit was
started up. At present the following hydro-isomerisation processes are commercially
available:
• UOP BUTAMER for butane isomerisation
• UOP PENEX for pentane/hexane isomerisation
• BP C4 isomerisation for butane isomerisation
• BP C5/C6 isomerisation for pentane/hexane isomerisation
• SHELL Hysomer for pentane/hexane isomerisation

All these processes takes place in the vapour phase on a fixed bed catalyst containing
platinum on a solid carrier.

As an example, the Shell Hysomer process will be briefly described. The liquid feedstock
is pentane/hexane from light naphtha. naphtha splitters are widely used to split light
naphtha, heavy naphtha and also LPG. The light naphtha (C5/C6) is combined with the
recycle gas/ fresh gas mixture. The resultant combined reactor feed is routed to a feed/
effluent heat exchanger, where it is heated and completely vaporised by the effluent of
the reactor. The vapourised combined reactor feed is further heated to the desired reactor
inlet temperature in the reactor charge heater. The hot charge enters the Hysomer reactor
at the top and flows downwards through the catalyst bed, where a portion of normal and
mono- branched paraffins is converted into higher branched (high octane) components.
Temperature rise from the heat of reaction release is controlled by a cold quench gas
injection into the reactor. Reactor effluent is cooled and subsequently separated in the
product separator into two streams: a liquid product (isomerate) and a recycle gas stream
returning to the reactor via the recycle gas compressor.

The catalyst is a dual function catalyst consisting of platinum on a zeolite basis, highly
stable and regenerable.

Temperatures and pressure vary in a range of 230 - 285 0C and 13-30 bar, C5/C6 content
in product relative to that in feed is 97% or better, and octane upgrading ranges between
8 and 10 points, depending on feedstock quality. The Hysomer process can be integrated
with catalytic reformer, resulting in substantial equipment savings, or with iso-normal
separation processes which allows for a complete conversion of pentane/hexane mixtures
into isoparaffin mixtures. An interesting application in this field is the total isomerisation
process (TIP) in which the isomerisation is completely integrated with a Union Carbide
molecular sieve separation process or the naphtha IsoSiv Process by UOP.

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Highlights of TIP
The following are some of the highlights of the TIP process:
A. TIP has been in commercial operation since 1975
B. UOP manufacturers both the zeolite isomerisation catalyst and the IsoSiv Grade
Molecular Sieve adsorbent.
C. UOP's zeolite catalyst will tolerate sulfur and/or water upsets, the effects of which
are usually reversible, either with time or by in situ regeneration (which
minimises any down time).
D. The expected life of the catalyst and adsorbent is 10 years or more.
F. The combination of zeolite isomerisation and IsoSiv molecular separation is
possible because each station has similar operating conditions of temperature,
pressure and environment. This eliminates the need for a second compressor,
intermediate stabilisation and the costs associated with cooling, purifying and
reheating the recylce normal paraffins.
G. TIP and IsoSiv separation permits maximum flexibility in changing the C5/C6
ratio and iso/normal ratio of the feed.

Conclusion:
Nowadays many refiners are looking into the isomerisation processes to add potential
extra value and complimentary to the platforming process. Directly both the platforming
and isomerisation process work hand in hand in several ways. C5 paraffins tend to crack
away in the platformer, but give high upgrading in the isomerisation unit. C6 components
convert nicely to benzene in the platformer, but nowadays the specs on aromatics and
benzene are tightening, which makes conversion of these components to C6 isomers
preferred. Furthermore, benzene is hydrogenated in the isomerisation unit. By adjusting
the cutpoint between the light and heavy naphtha, i.e. the cutpoint between the feed to the
isomerisation feed and the platformer feed, the refiner has the flexibility to control the
benzene content of its gasoline pool.

Hydrocracking

Introduction
The need for gasoline of a higher quality than that obtained by catalytic cracking led to
the development of the hydrocracking process. The history of the process goes back to
the later 1920s when a plant for the commercial hydrogenation of brown coal was
commissioned at Leuna in Germany. Tungsten sulphide was used as a catalyst in this
one-stage unit , in which high reaction pressures, 200-300 bar, were applied. The catalyst
displayed a very high hydrogenation activity: the aromatic feedstock, coal and heavy
fractions of oil, containing sulphur, nitrogen and oxygen, was virtually completely
converted into paraffins and isoparaffins. The result of the Leuna plant - loss of octane
number from aromatic hydrogenation of impurities in the feedstock, notably nitrogen
compounds, followed by a hydrocracking step. In 1939, ICI developed the second-stage

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catalyst for a plant that contributed largely to Britain's supply of aviation gasoline in the
subsequent years.

During World War II, two stage processes were applied on a limited scale in Germany,
Britain and USA. In Britain, feedstock were creosote from coal tar and gas oil from
petroleum. In the USA, Standard Oil of New Jersey operated a plant at Baton Rouge,
producing gasoline from a Venezuelan kerosine/light gasoil fraction. Operating
conditions in those units were comparable: approximate reaction temperature 400 0C and
reaction pressures of 200-300 bar.

After the war, commercial hydrocracking was stopped because the process was too
expensive. Hydocracking research, however, continued intensively. By the end of the
1950s, the process had become economic, for which a number of reasons are identified.

The development of improved catalyst made it possible to operate the process at

considerably lower pressure, about 70-150 bar.

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This in turn resulted in a reduction in equipment wall thickness, whereas simultaneously,

advances were made in mechanical engineering, especially in the field of reactor design
and heat transfer. These factors, together with the availability of relatively low cost
hydrogen from steam reforming process, brought hydrocracking back on the refinery
scene. The first units of the second generation were built in USA to meet the demand for
conversion of surplus fuel oil in the gasoline-oriented refineries.

Now, hdyrocracking is well established process from many licensors.

Basis for the Choice of Conversion Route

Refiners are continuously faced with trends towards increased conversion, better product
qualities and more rapidly changing product patterns. Various processes are available that
can meet the requirements to a greater or lesser degree: coking, visbreaking/thermal
cracking, catalytic cracking and hydrocracking.

The type of process applied and the complexity of refineries in various parts of the world
are determined to a greater extent by the product distribution required. As a consequence,
the relatively importance of the above process in traditionally fuel-oil dominated
refineries such as those in Western Europe will be quite different from those of gasoline-
oriented refineries in, for instance, the USA.

An important aspect of the coking, thermal cracking and catalytic cracking process is that
they operate at low pressures. This gives advantages in the fields of capital, matallurgy
and engineering.

A particular feature of the hydrocracking process, as compared with its alternatives, is its
flexibility with respect to product outturn and high quality of its products. In the areas
where quantitative imbalance exists of lighter products, middle distillates and fuel,
hydrocracking is a most suitable process for correction. Moreover, the hydrocracker does
not yield any coke or pitch byproduct: the entire feedstock is converted into the required
product range, an important consideration in a situation of limited crude oil availability.
The development of the low-pressure catalytic reforming process, which produces
relatively cheap, high quality hydrogen, has continued substantially to the economic
viability of hydrocracking. On the whole, hydrocracking can handle a wider range of
feedstock than catalytic cracking, although the latter process has seen some recent
catalyst developments which narrowed the gap. There are also examples where
hydrocracking is complementary rather than alternative to the other conversion process;
an example, cycle oils, which cannot be recycled to extinction in the catalytic cracker,
can be processed in the hydrocracker.

Notwithstanding many extensive comparisons between the various processes, the

experince shows the generalisation with respect to the optimum conversion route still
cannot be made.

By: Hardeep Hundal, original editing by Jeroen Buren

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Process Description
All hydrocracking process are characterised by the fact that in a catalytic operation under
relatively high hydrogen pressure a heavy oil fraction is treated to give products of lower
molecular weight.

Hydrocracking covers widely different fuels, ranging from C3/C4 production from
naphta, on the other hand, to luboil manufacture from deasphalted oils, on the other.
Most hydrocrackers use fixed beds of catalyst with downflow of reactants. The H-Oil
process developed by Hydrocarbon Research Corp and Cities Service R & D employs an
ebullient bed reactor in which the beds of particulate catalyst are maintained in an
ebullient or fluidised condition in upflowing reactants.

When the processing severity in a hydrocracker is increased, the first reaction occuring
leads to saturation of any olefinic material present in feedstock. Next comes the reaction
of desulphurisation, denitrogenation and de-oxygenation. These reactions constitute
treating steps during which in most cases, only limited cracking takes place. When the
severity is increased further, hydrocracking reaction is initiated. They proceed at various
rates, with the formation of intermediate products (eg. saturation of aromatics), which are
subsequently cracked into lighter products.

Process Configuration
When the treating step is combined with the cracking reaction to occur in one reactor, the
process is called a SINGLE-STAGE PROCESS.

SINGLE-STAGE PROCESS: In this simplest of the hydrocracker configuration, the lay

out of the reactor section generally resembles that of hydrotreating unit. This
configuration will find application in cases where only moderate degree of conversion
(say 60% or less) is required. It may also be considered if full conversion, but with a
limited reduction in molecular weight, is aimed at. An example is the production of
middle distillates from heavy distillate oils. The catalyst used in a single-stage process
comprises a hydrogenation function in combination with a strong cracking function. The
hydogenation function is provided by sulphided metals such as cobalt, molybdenum and
nickel. An acidic support, usually alumina, attends to the cracking function. Nitrogen
compounds and ammonia produced by hydrogenation interfere with acidic activity of the
catalyst. In the cases of high/full conversion is required, the reaction temperatures and
run lenghts of interest in commercial operation can no longer be adhered to. It becomes
necessary to switch to a multi-stage process, in which the cracking reaction mainly takes
place in an added reactor. With regard to the adverse effect of ammonia and nitrogen
compounds on catalyst activity, two versions of the multi stage hydrocracker have been
developed: the TWO STAGE HYDROCRACKER and SERIES FLOW
HYDROCRACKER.

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In the first type, the undesirable compounds are removed from the unconverted
hydrocarbons before the latter are charged to the cracking reactor. This type is called the
TWO STAGE PROCESS. The other variety is ofen referred to as SERIES FLOW
HYDROCRACKER. This type uses a catalyst with an increased tolerance towards
nitrogen, both as ammonia and in organic form.

TWO STAGE CONFIGURATION: Fresh feed is preheated by heat exchange with

effluent from the first reactor. It is combined with part of a not fresh gas/recycle gas
mixture and passes through a first reactor for desulphurisation/denitrogenation step.
These reactions, as well as those of hydrocracking, which occurs to a limited extent in the
first reactor, are exothermic. The catalyst inventory is therefore divided among a number
of fixed beds. Reaction temperatures are controlled by introducing part of the recycle gas
as a quench medium between beds. The ensuing liquid is fractionated to remove the
product made in the first reactor. Unconverted, material, with a low nitrogen content and
free of ammonia, is taken as a bottom stream from the fractionation section. After, heat
exchange with reactor effluent and mixing with heated recycle gas, it is sent to the second
reactor. Here most of the hydrocracking reactions occur. Strongly acidic catalyst with a
relatively low hydrogenation activity (metal sulphides on, for example, amorphous silica-
alumina) are usually applied. As in the first reactor, the exothermicity of the process is
controlled by using recycle gas as quench medium the catalyst beds. Effluent from the
second reactor is cooled and joins first stage effluent for separation from recycle gas and
fractionation. The part of the second reactor feed that has remained unconverted is
recycled to the reactor. Feedstock is thereby totally converted to the product boiling
range.

SERIES FLOW CONFIGURATION: The principal difference is the elimination of first

stage cooling and gas/liquid separation and the interstage ammonia removal step. The
effluent from the first stage is mixed with more recycle gas and routed direct to the inlet
of the second reactor. In contrast with the amorphous catalyst of the two-stage process,
the second reactor in series flow generally has a zeolitic catalyst, based on crystalline
silica-alumina. AS in the two stage process, material not converted to the product boiling
range is recycled from the fractionation section.

Conclusion
Both two stage and series flow hydrocracking are flexible process: they may yield, in one
mode of operation, only naphtha and lighter products and, in a different mode, only
gasoil and lighter products. In the naphtha mode, both configurations have comparable
yield patterns. In modes for heavier products, kerosine and gasoil, the two stage process
is more selective because product made in the first reactor is removed from the second
reactor feed, In series flow operation this product is partly overcracked into lighter
products in the second reactor.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Catalytic Cracking

Introduction
Already in the 30's it was found that when heavy oil fractions are heated over clay type
materials, cracking reactions occur, which lead to significant yields of lighter
hydrocarbons. While the search was going on for suitable cracking catalysts based on
natural clays, some companies concentrated their efforts on the development of synthetic
catalyst. This resulted in the synthetic amorphous silica-alumina catalyst, which was
commonly used until 1960, when it was slightly modified by incorporation of some
crystalline material (zeolite catalyst). When the success of the Houdry fixed bed process
was announced in the late 1930s, the companies that had developed the synthetic catalyst
decided to try to develop a process using finely powdered catalyst. Subsequent work
finally led to the development of the fluidised bed catalytic cracking (FCC) process,
which has become the most important catalytic cracking process.

Originally, the finely powdered catalyst was obtained by grinding the catalyst material,
but nowadays, it is produced by spray-drying a slurry of silica gel and aluminium
hydroxide in a stream of hot flue gases. Under the right conditions, the catalyst is
obtained in the form of small spheres with particles in the range of 1-50 microns.

When heavy oil fractions are passed in gas phase through a bed of powdered catalyst at a
suitable velocity (0.1-0.7m/s), the catalyst and the gas form a system that behaves like
liquid, i.e. it can flow from one vessel to another under the influence of a hydrostatic
pressure. If the gas velocity is too low, the powder does not fluidise and it behaves like a
solid. If velocity is too high, the powder will just be carried away with the gas. When the
catalyst is properly fluidised, it can be continously transported from a reactor vessel,
where the carcking reactions take place and where it is fluidised by the hydrocarbon
vapour, to a regenerator vessel, where it is fluidised by the air and the products of
combustion, and then back to the reactor. In this way the proces is truly continous.

The first FCC unit went on stream in Standard Oil of New Jersey's refinery in Baton
Rounge, Louisiana in May 1942. Since that time, many companies have developed their
own FCC process and there are numerous varieties in unit configuration.

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FCC Process Configuration:

Hot feed, together with some steam, is introduced at the bottom of the riser via special
distribution nozzles. Here it meets a stream of hot regenerated catalyst from the
regenerator flowing down the inclined regenerator standpipe. The oil is heated and
vaporised by the hot catalyst and the cracking reactions commence. The vapour, initially
formed by vaporisation and successively by cracking, carries the catalyst up the riser at
10-20 m/s in a dilute phase. At the outlet of the riser the catalyst and hydrocarbons are
quickly separated in a special device. The catalyst (now partly deactivated by deposited
coke) and the vapour then enter the reactor. The vapour passes overhead via cyclone
separator for removal of entrained catalyst before it enters the fractionator and further
downstream equipment for product separation. The catalyst then descends into the
stripper where entrained hydrocarbons are removed by injection of steam, before it flows
via the inclined stripper standpipe into the fluidised catalyst bed in the regenerator.

Air is supplied to the regenerator by an air blower and distributed throughout the catalyst
bed. The coke deposited is burnt off and the regenerated catalyst passes down the
regenerator standpipe to the bottom of the riser, where it joins the fresh feed and the cycle
recommences.

The flue gas (the combustion products) leaving the regenerator catalyst bed entrains
catalyst particles. In particular, it entrains "fines", a fine dust formed by mechanical
rubbing of catalyst particles taking place in the catalyst bed. Before leaving the
regenerator, the flue gas therefore passes through cyclone separators where the bulk of
this entrained catalyst is collected and returned to the catalyst bed.

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Normally modern FCC is driven by an expansion turbine to mimimise energy

consumption. In this expansion turbine, the current of flue gas at a pressure of about 2
barg drives a wheel by striking impellers fitted on this wheel. The power is then
transferred to the air blower via a common shaft. This system is usually referred to as a
"power recovery system". To reduce the wear caused by the impact of catalyst particles
on the impellers (erosion), the flue gas must be virtually free of catalyst particles. The
flue gas is therefore passed through a vessel containing a whole battery of small, highly
efficient cyclone separators, where the remaining catalyst fines are collected for disposal.

Before being disposed of via a stack, the flue gas is passed through a waste heat boiler,
where its remaining heat is recovered by steam generation.

In the version of the FCC process described here, the heat released by burning the coke in
the regenerator is just sufficient to supply the heat required for the riser to heat up,
vaporise and crack the hydrocarbon feed. The units where this balance occurs are called "
heat balanced" units. Some feeds caused excessive amounts of coke to be deposited on
the catalyst, i.e. much more than is required for burning in the regenerator and to have a
"heat balanced" unit. In such cases, heat must be removed from the regenerator, e.g. by
passing water through coils in the regenerator bed to generate steam. Some feeds cause so
little coke to be deposited on the catalyst that heat has to be supplied to the system. This
is done by preheating the hydrocarbon feed in a furnace before contacting it with the
catalyst.

Main Characteristics
• A special device in the bottom of the riser to enhance contacting of catalyst and
hydrocarbon feed.
• The cracking takes place during a short time (2-4 seconds) in a riser ("short-contact
time riser") at high temperatures ( 500-540 0C at riser outlet).
• The catalyst used is so active that a special device for quick separation of catalyst and
hydrocarbons at the outlet of the riser is required to avoid undesirable cracking after
the mixture has left the riser. Since, no cracking in thereactor is required, the reactor
no longer functions as a reactor; it merely serves as a holding vessel for cyclones.
• The regenerator takes place at 680-720 0C. With the use of special catalysts, all the
carbon monoxide (CO) in the flue gas is combusted to carbon dioxide (CO2) in the
regenerator.
• Modern FCC includes a power recovery system for driving the air blower.

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Equipment in FCC
• Large storage vessels for catalyst (fresh and equilibrium)
• Regenerator
• Reactor
• Main Fractionator
• Product Work Up section (several distillation columns in series
• Product treating facilities

Feedstock & Yield

Before the introduction of residues, vacumn distillates were used as feedstock to load the
Catalytic Cracker fully. These days, even residues are used to load the cracker. The term
used for this type of configuation is Long Residue Catalytic Cracking Complex. The only
modification or addition needed are a residue desalter and a bigger and more heat
resistent reactor.

The yield pattern of an FCC unit is typically as follows:

Product % weight on fresh feed

C3 & C4 15
Gasoline 40-50
Heavy Gas Oil 10
Coke 5

Conclusion
The FCC Unit can a real margin improver for many refineries. It is able to convert the
residues into high value products like LPG , Butylene, Propylene and Mogas together
with Gasoil. The FCC is also a start for chemical production (poly propylene). Many
FCC's have 2 modes: a Mogas mode and a Gasoil mode and FCC's can be adapted to
cater for the 2 modes depending on favourabale economic conditions. The only
disavantage of an FCC is that the products produced need to be treated (sulfur removal)
to be on specification. Normally Residue FCCs act together with Residue
Hydroconversion Processes and Hydrocrackers in order to minimise the product quality
give away and get a yield pattern that better matches the market specifications. Via
product blending, expensive treating steps can be avoided and the units prepare excellent
feedstock for eachother: desulfurised residue or hydrowax is excellent FCC feed, while
the FCC cycle oils are excellent Hydrocracker feed.

In the near future, many refiners will phase the challenge how to desulfurise cat cracked
gasoline without destroying its octane value. Catalytic destillation appears to be one of
the most promising candidate processes for that purpose.

By: Hardeep Hundal, original editing by Jeroen Buren

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 Shared Content in cooperation with Refining Technology Online

Thermal Cracking

Introduction
Thermal cracking is the oldest and, in a way, the simplest cracking process. It basically
aims at the reduction of molecular size by application of heat without any additional
sophistication such as catalyst or hydrogen. At a temperature level of 450-500 C, the
larger hydrocarbon molecules become unstable and tend to break spontaneously into
smaller molecules of all possible sizes and types. By varying the time, temperature and
pressure under which a particular feedstock remains under cracking conditions, the
desired degree of cracking (conversion) can be controlled. Temperature and time
(residence time) are important process variables pressure plays a secondary role.

Obviously, the cracking conditions to be applied and the amount and type of cracked
products will depend largely on the type of feedstock. In practice, the feedstock for
thermal cracking is a mixture of complex heavy hydrocarbon molecules left over from
atmospheric and/or vacuum distillation of crude. The nature of these heavy, high
molecular weight fractions is extremely complex and much fundamental research has
been carried out on their behaviour under thermal cracking conditions. However, a
complete and satisfactory explanation of these reactions that take place cannot be given,
except for relatively simple and well-defined types of products. For instance, long chain
paraffinic hydrocarbon molecules break down into a number of smaller ones by rupture
of a carbon-to-carbon bond (the smaller molecules so formed may break down further).
When this occurs, the number of hydrogen atoms present in the parent molecule is
insufficient to provide the full complement for each carbon atom, so that olefins or
"unsaturated" compounds are formed. The rupturing can take place in many ways,
usually a free radical mechanism for the bond rupture is assumed.

However, paraffinic hydrocarbons are usually only a small part of the heavy petroleum
residues, the rest being cyclic hydrocarbons, either aromatic or naphthenic in character.
In these, the rupture takes place in the paraffinic side-chain and not in the ring. Other side
reactions also take place. In particular, the condensation and polymerisation reactions of
olefins and of the aromatics are of considerable practical importance, since they can lead
to undesirable product properties, such as an increase in the sludge or tar content. Hence,
in practice, it is very difficult to assess the crackability of various feedstocks without
plant trials. The final products consist of gas, light hydrocarbons in the gasoline and
gasoil range and heavier products. By selection of the type of unit, feedstock and
operating conditions, the yields and quality of the various products can, within limits be
controlled to meet market requirements.

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The maximum conversion that can be obtained will be determined by the quality of the
bottom product of the thermal cracker, thermally cracked residue. This stream is normally
routed to the fuel oil blending pool. When the cracking has taken place at a too high
severity, the fuel can become 'unstable' upon blending with diluent streams (see below).
Normally, the refinery scheduler will assess what the maximum severity is that the
thermal cracking unit can operate on, without impacting on the stability of the refinery
fuel blending pool.

When thermal cracking was introduced in the refineries some 80 years ago, its main
purpose was the production of gasoline. The units were relatively small (even applying
batch processing), were inefficient and had a very high fuel consumption. However, in
the twenties and thirties a tremendous increase in thermal cracking capacity took place,
largely in the version of the famous DUBBS process, invented by UOP. Nevertheless,
thermal cracking lost ground quickly to catalytic cracking (which produces gasoline of
higher octane number) for processing heavy distillates with the onset of the latter process
during World War II. Since then and up to the present day, thermal cracking has mostly
been applied for other purposes : cracking long residue to middle distillates (gasoil), short
residue for viscosity reduction (visbreaking), short residue to produce bitumen, wax to
olefins for the manufacture of chemicals, naphtha to ethylene gas (also for the
manufacturing of chemicals), selected feedstocks to coke for use as fuel or for the
manufacture of electrodes.

In modern oil refineries there are three major applications of the thermal cracking
process:
• VISBREAKING
• THERMAL GASOIL PRODUCTION
• COKING

Visbreaking
Visbreaking (i.e. viscosity reduction or breaking) is an important application of thermal
cracking because it reduces the viscosity of residue substantially, thereby lessening the
diluent requirements and the amount of fuel oil produced in a refinery. The feed, after
appropriate preheat, is sent to a furnace for heating to the cracking temperature, at about
450-460 degrees C. The cracking takes place to a small extent in the furnace and largely
in a soaker (reaction chamber) just downstream of the furnace. At the soaker outlet, the
temperature is lower than at the furnace outlet (soaker inlet) because the cracking
reactions are endothermic. The products are quenched at the soaker outlet to stop the
cracking reaction (to prevent excessive coke formation). After that, the products enter the
fractionator at a temperature level of 300- 400 degrees C and from here onward the
processing is similar to any normal distillation process. The products are separated into
gas, gasoline, kero, gasoil and residue. The residue so obtained has a lower viscosity that
the feed (visbreaking), which leads to a lower diluent requirement to make the fuel on
specification for viscosity. The up-flow soaker provides for a prolonged residence time

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and therefore permits a lower cracking temperature than if the soaker was not used. This
is advantageous as regards cost in furnace and fuel. Modern soakers are equipped with
internals so as to reduce back mixing- effects , thus maximising the viscosity reduction.
Since only one cracking stage is involved, this layout is also named one-stage cracking.
The cracking temperature applied is about 440-450 degree C at a pressure of 5-10 barg in
the soaker. The fractionator can be operated at 2-5 barg, depending on furnace
constraints, condenser constraints and fuel cost.

Thermal Gasoil Production

This is a more elaborate and sophisticated application of thermal cracking as compared
with visbreaking. Its aim is not only to reduce viscosity of the feedstock but also to
produce and recover a maximum amount of gasoil. Altogether, it can mean that the
viscosity of residue (excluding gasoil) run down from the unit is higher than that of the
feed.

In the typical lay out is the first part of the unit quite similar to a visbreaking unit. The
visbroken residue is vacuum-flashed to recover heavy distillates, which are then sent
back to a thermal cracking stage, together with heavy distillate recovered from the
fractionator, in a second furnace under more severe cracking conditions ( temperature
500 degrees C; pressure 20-25 barg) . More severe conditions are necessary because the
feedstock has a smaller molecular size and is therefore more difficult to crack than the
larger residue molecules in the first stage. This layout is referred to as tow-stage cracking.

Delayed Coking
This is an even more severe thermal cracking application than the previous one. The goal
is to make a maximum of cracking products - distillates - whereby the heavy residue
becomes so impoverished in hydrogen that it forms coke. The term "delayed" is intended
to indicate that the coke formation does not take place in the furnace (which would lead
to a plant shutdown) but in the large coke drums after the furnace. These drums are
filled/emptied batch-wise (once every 24 hours), though all the rest of the plant operates
continuously. A plant usually has two coke drums, which have adequate capacity for one
day's coke production (500-1500 m2). The process conditions in the coke drum are 450-
500 degrees C and 20 - 30 bar. Only one coke drum is on-line; the other is off line, being
emptied or standing by. Only the vapour passes from the top of the coke drums to the
fractionator, where the products are separated into the desired fractions. The residue
remains in the coke drum to crack further until only the coke is left. Often the heaviest
part of the fractionator products is recycled to feed.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Product Quality
Thermally cracked products - distillates - are not suitable for commercial use as produced
in other units; they require further refinement or treatment in order to improve their
quality, particularly sulfur and olefins content. Formerly, wet treating processes, for
example treatment with caustic or an other extraction medium, were applied to remove or
"sweeten" the smelly sulfur products, but nowadays the catalytic hydrotreating is
employed almost without exception, both for gasoline and for gas oil range products. Of
course, the gases too have to be desulphurised before being used as fuel gas within the
refinery.

The residual products from thermal cracking are normally not treated any further, except
for coke, which may be calcined if the specifications require it to be treated. The cracked
residue is normally disposed of as refinery or commercial fuel. Here a very important
aspect of the process is the stability of the cracked residues or of the final fuels after
blending with suitable diluents. Residue contains asphaltenes,which are colloidally
dispersed uniformly in the oil in a natural way. In the cracking process, the character of
the asphaltenes as well as of the oil changes, and if the cracking is too severe the natural
balance of the colloidal system can be affected to the extent that part of the asphaltenes
precipitates in the equipment or in the storage tanks, forming sludge. If the sludge
formation is excessive, i.e. above a certain specified limit, the product (fuel) is considered
to be unstable.

Plant Operations/Decoking
A practical aspect of operation of thermal cracking units is that, in spite of good design
and operating practice, furnaces, and sometimes also other equipment, gradually coke up,
so that the unit has to be shut down and decoked. Furnaces can be decoked by "
turbining" (using special rotary tools to remove coke from inside furnace pipes) or by
steam-air decoking process. In the latter case, the coke is burnt off in a carefully
controlled decoking process in which air and steam are passed through the tubes at
elevated temperatures. Air serves to burn coke, where as the steam serves to keep the
burning temperatures low so that they do not exceed the maximum tolerable
temperatures.

More recently, a new decoking method using studded 'pigs' propelled with water, is
getting more popular. The plastic pigs have a size slightly smaller than the tube inside
diameter and are equipped with metal studs. When the pigs are pumped through the
furnace pipes, they move around in a rotating fashion, thus scraping the cokes from the
inside of the furnace tubes.

Other coked equipment is usually cleaned by hydrojetting techniques. Owing to these

unavoidable stops for decoking, the on-stream time i.e. on stream days per annum, for
thermal cracking units is slightly shorter than for most other oil processes.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Vacuum Distillation

Introduction
To recover additional distillates from long residue, distillation at reduced pressure and
high temperature has to be applied. This vacuum distillation process has become an
important chain in maximising the upgrading of crude oil. As distillates, vacuum gas oil,
lubricating oils and/or conversion feedstocks are generally produced. The residue from
vacuum distillation - short residue - can be used as feedstock for further upgrading, as
bitumen feedstock or as fuel component. The technology of vacuum distillation has
developed considerably in recent decades. The main objectives have been to maximise
the recovery of valuable distillates and to reduce the energy consumption of the units.

At the place where the heated feed is introduced in the vacuum column - called the flash
zone - the temperature should be high and the pressure as low as possible to obtain
maximum distillate yield. The flash temperature is restricted to about 420 0C, however,
in view of the cracking tendency of high-molecular-weight hydrocarbons. Vacuum is
maintained with vacuum ejectors and lately also with liquid ring pumps. Lowest
achievable vacuum in the flash zone is in the order of 10 mbar.

In the older type high vacuum units the required low hydrocarbon partial pressure in the
flash zone could not be achieved without the use of "lifting" steam. The steam acts in a
similar manner as the stripping steam of crude distillation units. This type of units is
called "wet" units. One of the latest developments in vacuum distillation has been the
deep vacuum flashers, in which no steam is required. These "dry" units operate at very
low flash zone pressures and low pressure drops over the column internals. For that
reason the conventional reflux sections with fractionation trays have been replaced by
low pressure- drop spray sections. Cooled reflux is sprayed via a number of specially
designed spray nozzles in the column countercurrent to the up-flowing vapour. This spray
of small droplets comes into close contact with the hot vapour, resulting in good heat and
mass transfer between the liquid and vapour phase.

To achieve low energy consumption, heat from the circulating refluxes and rundown
streams is used to heat up the long residue feed. Surplus heat is used to produce medium
and/or low-pressure steam or is exported to another process unit (via heat integration).
The direct fuel consumption of a modern high-vacuum unit is approximately 1% on
intake, depending on the quality of the feed. The steam consumption of the dry high-
vacuum units is significantly lower than that of the "wet" units. They have become net
producers of steam instead of steam consumers.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

Three types of high-vacuum units for long residue upgrading have been developed for
commercial application:
• FEED PREPARATION UNITS
• LUBOIL HIGH- VACUUM UNITS
• HIGH - VACUUM UNITS FOR BITUMEN PRODUCTION

Feed Preparation Units

These units make a major contribution to deep conversion upgrading ("cutting deep in the
barrel"). They produce distillate feedstocks for further upgrading in catalytic crackers,
hydrocrackers and thermal crackers. To obtain an optimum waxy distillate quality a wash
oil section is installed between feed flash zone and waxy distillate draw-off. The wash oil
produced is used as fuel component or recycled to feed. The flashed residue (short
residue) is cooled by heat exchange against long residue feed. A slipstream of this cooled
short residue is returned to the bottom of the high-vacuum column as quench to minimise
cracking (maintain low bottom temperature).

Luboil High-Vacuum Units

Luboil high vacuum units are specifically designed to produce high-quality distillate
fractions for luboil manufacturing. Special precautions are therefore taken to prevent
thermal degradation of the distillates produced. The units are of the "wet" type. Normally,
three sharply fractionated distillates are produced (spindle oil, light machine oil and
medium machine oil). Cutpoints between those fractions are typically controlled on their
viscosity quality. Spindle oil and light machine oil are subsequently steam- stripped in
dedicated strippers. The distillates are further processed to produce lubricating base oil.
Short residue is normally used as feedstock for the solvent de-asphalting process to
produce deasphalted oil, an intermediate for bright stock manufacturing.

High-Vacuum Units for Bitumen Production

Special vacuum flashers have been designed to produce straight-run bitumen and/or
feedstocks for bitumen blowing. In principle, these units are designed on the same basis
as the previously discussed feed preparation units, which may also be used to provide
feedstocks for bitumen manufacturing.

Vacuum Distillation

Asphaltic bitumen, normally called "bitumen" is obtained by vacuum distillation or

vacuum flashing of an atmospheric residue. This is " straight run" bitumen. An
alternative method of bitumen production is by precipitation from residual fractions by
propane or butane- solvent deasphalting.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 Shared Content in cooperation with Refining Technology Online

The bitumen thus obtained has properties which derive from the type of crude oil
processed and from the mode of operation in the vacuum unit or in the solvent
deasphalting unit. The grade of the bitumen depends on the amount of volatile material
that remains in the product: the smaller the amount of volatiles, the harder the residual
bitumen.

In most cases, the refinery bitumen production by straight run vacuum distillation does
not meet the market product quality requirements. Authorities and industrial users have
formulated a variety of bitumen grades with often stringent quality specifications, such as
narrow ranges for penetration and softening point. These special grades are manufactured
by blowing air through the hot liquid bitumen in a BITUMEN BLOWING UNIT. What
type of reactions take place when a certain bitumen is blown to grade? Bitumen may be
regarded as colloidal system of highly condensed aromatic particles (asphaltenes)
suspended in a continuous oil phase. By blowing, the asphaltenes are partially
dehydrogenated (oxidised) and form larger chains of asphaltenic molecules via
polymerisation and condensation mechanism. Blowing will yield a harder and more
brittle bitumen (lower penetration, higher softening point), not by stripping off lighter
components but changing the asphaltenes phase of the bitumen. The bitumen blowing
process is not always successful: a too soft feedstock cannot be blown to an on-
specification harder grade.

The blowing process is carried out continuously in a blowing column. The liquid level in
the blowing column is kept constant by means of an internal draw-off pipe. This makes it
possible to set the air-to-feed ratio (and thus the product quality) by controlling both air
supply and feed supply rate. The feed to the blowing unit (at approximately 210 0C),
enters the column just below the liquid level and flows downward in the column and then
upward through the draw-off pipe. Air is blown through the molten mass (280-300 0C)
via an air distributor in the bottom of the column. The bitumen and air flow are
countercurrent, so that air low in oxygen meets the fresh feed first. This, together with the
mixing effect of the air bubbles jetting through the molten mass, will minimise the
temperature effects of the exothermic oxidation reactions: local overheating and cracking
of bituminous material. The blown bitumen is withdrawn continuously from the surge
vessel under level control and pumped to storage through feed/product heat exchangers.

By: Hardeep Hundal, original editing by Jeroen Buren

Adapted for The Chemical Engineers’ Resource Page by Chris Haslego
 S P E C I A L F O C U S O N

HYDROPROCESSING
OPTIONS

Pushing Plant Limits.

Test Runs, Plant Expectations
and Performance Confidence
By SCOTT W. GOLDEN, Process Consulting Services Inc., Houston, TX

R
eliable minimum-capital- any revamp, yet it is
investment revamps require rarely performed. Often
pushing the plant to the it is considered an
limit of the existing equip- unnecessary cost. To-
ment. While various small equipment day, there is an overrid-
design errors are tolerable at moderate ing belief that ease-of-
equipment capacity, they are not when use computer models
the unit is pushed.1,2 Process design can be used by inexpe-
for reliable low-cost revamps must rienced personnel; and
focus on the following areas: that all data can be rec-
• Test-run: quantifying existing equip- onciled by the appro-
ment performance priate software algo-
• Alternative flow schemes: maximiz- rithms. Unfortunately,
ing existing equipment utilization refinery equipment
• Reliable revamps: detail equipment may not operate per
design. the model assumptions
All phases of a revamp process or the presumed origi-
design must be addressed to meet nal equipment design.
both the low cost and reliability objec- Installed equipment
tives. 3,4 Shortcutting this process performance must be
results in either higher investment measured. For in-
costs due to unnecessary equipment stance, knowing that
modifications or unreliable revamp the FCC wet gas com-
designs being installed because exist- pressor is limiting the
ing equipment problems were not throughput or unit
identified. This article will address the conversion is not the
importance of test-runs in the design same as knowing what
of refinery revamps. FCC units are causes the limitation.
used to highlight the importance of Accurate operating Photo 1. Revamps Push Equipment Limits
the test-run in quantifying equipment data is required to
performance. (Photo 1). quantify existing equipment perfor- While management information
A well-planned comprehensive mance and pinpoint specific causes of systems contain volumes of data, often
test-run is the most important step in equipment operating problems. this data is not complete or worse, it

1 March/April 1999
F O C U S O N : HYDROPROCESSING OPTIONS

may not be accurate. Pushing refinery

unit operations to the major equip-
ment limits requires a different PRELIMINARY REVAMP REFINERY DEFINED
PROCESS OBJECTIVE OBJECTIVES
approach to a revamp design. Reliable
low-cost revamps must address the
myriad of potential detail equipment REFINERY STAFF +
design considerations that cause units FIELD REVIEW
REVAMP ENGINEERS
not to perform. Revamp process
design must quantify existing equip-
ment performance and eliminate bot- COMPREHENSIVE REFINERY STAFF +
TEST RUN REVAMP ENGINEERS
tlenecks that prevent maximum
equipment utilization. Ultimately,
revamps must be both cost effective TEST RUN
REVAMP ENGINEERS
and reliable. DATA EVALUATION

Conventional Project
Process Design EXISTING EQUIPMENT
REVAMP ENGINEER
REVIEW
A conventional project process design
is executed by a series of activities that
are often performed by different DEVELOPMENT OF ALTERNATIVE
REVAMP ENGINEER
groups of individuals.5 First a heat- REVAMP FLOW SCHEME
and-material balance based on prede-
termined revamp objectives is com-
REFINERY STAFF +
pleted. Existing equipment evaluation FINAL OBJECTIVES REVAMP ENGINEERS
then follows. Generally, all this work is
completed in the design office. If the
calculated existing equipment capacity DESIGN PACKAGE REVAMP ENGINEERS
is insufficient, the equipment is
replaced or paralleled. In this
approach, the office-based calculation COST ESTIMATE
COST ESTIMATORS
of the plant performance sets the & SCHEDULE
future performance expectation. Engi-
neering calculations are performed
based on presumed equipment perfor- Figure 1. Revamp Approach
mance, vendor supplied estimates of
performance, industry averages or the ty margin required to adequately cover ever, in this era of computers and
specific E&C’s design standards. the errors in the conventional project’s advanced process control equipment,
Revamped process equipment perfor- approach is often greater than the real a test-run generally means gathering
mance is often a small difference changes required. data from the computer and sending it
between two large numbers (future to the engineering contractor’s process
requirements– current capabilities). Reliable Revamps- designers. Our definition of test-run is
The conventional project’s process Minimum Investment an accurate set of composition, pres-
design determines the future require- By contrast, revamp process design sure and temperature data on all
ments based on the criteria previously focuses on measuring the current process streams that impact the unit
mentioned. In meeting the revamp equipment performance with a com- heat and material balance including
objective, equipment design assump- prehensive well-planned and executed data needed to quantify the major
tions often force unnecessary equip- test run (Figure 1). Gathering specific process equipment performance.
ment modifications or underestimate test-run data and the subsequent Generally, revamp requirements are an
existing equipment capacity. To pre- analysis is the basis for an accurate incremental change to current mea-
vent revamp failures, engineering and bench-mark of current equipment sured (not calculated) performance
construction firms add in safety mar- performance. The term test-run has and not an average or presumed per-
gins to the changes required. The safe- come to mean different things. How- formance. The existing equipment

2 WORLD REFINING March/April 1999

 F O C U S O N : HYDROPROCESSING OPTIONS

estimates based on a similar service in

another unit. Often the estimates used
for design are so low, that they make
the installed exchanger surface area
unworkable. Additionally, calculated
exchanger tube and shell-side pressure
drop assumes the exchanger is clean.
Fouling can dramatically increase
exchanger pressure drop.
Many FCC catalytic sections are
being revamped to increase conversion
and/or product selectivity. 7 These
revamps make more C3-C12 boiling
range products from the reactor. The
additional lower boiling products must
be fractionated in the gas plant
columns. Figure 2 shows the depro-
panizer in the FCC gas plant. The
overhead product from this column is
processed in an ethylene plant C3 split-
ter to produce chemical grade propy-
lene and propane for sales. The FCC
Figure 2. Depropanizer Kettle Reboiler depropanizer overhead product C 4
content is an important control para-
performance must be quantified. This basis. This approach is shown for sev- meter. The depropanizer column is
minimizes revamp errors and cost, as eral FCC product recovery revamps. reboiled with heat from the FCC main
well as identifies equipment perfor- Case study examples highlight the fractionator slurry pumparound.
mance problems. Higher confidence in importance of test-run data when Increased reactor C3-C12 production
your current equipment performance revamping for reliability and mini- always increases gas plant reboiler heat
allows the revamp engineer to push mum investment cost. requirements. The proposed depro-
the equipment limits further.6 This Performing a comprehensive panizer heat and material balance, for
minimizes investment for any given test-run is a necessity. The time spent the revamp, shows higher reboiler duty
revamp objective. determining real plant limitations to meet the C3 recovery and overhead
directly relates to the revamp capital product C4 composition target.
Test Runs-Why Are cost. Test-run data simplify engineer- The conventional project’s
They Needed? ing calculations and help avoid calcu- approach rigorously models the TEMA
An accurate bench-mark requires a lation errors. Ease-of-use engineering AKT kettle with one of the several com-
plant test-run. Revamping requires a software calculations do not necessari- mercially available shell-and-tube
knowledge of the absolute magnitude ly represent reality. For instance, sys- exchanger models. The exchanger
of the equipment capabilities still tem hydraulic calculations involve equipment specialist uses the specific
unused. The standard information equivalent lengths of pipe, friction fac- E&C’s design criteria for this service
available to plant owners from process tors, exchanger pressure drop algo- which is a total shell and tube-side foul-
control systems and reports is trend rithms, and oil physical and transport ing factor of 0.005 Ft2-hr-°F/Btu. The
data, which often is not adequate to properties. In theory, a heat exchanger rigorous exchanger model indicates the
identify deficiencies. Hidden offsets in performance can be rigorously calcu- existing depropanizer reboiler will meet
trend data do not affect period-to-peri- lated by a computer program. Howev- future duty requirements. This conclu-
od operating performance compar- er, the reality for many petroleum sion is based on assumed fouling fac-
isons. However, hidden offsets can refinery exchanger services is that the tors and office-based calculations.
have a dramatic impact on revamp actual fouling factor value controls the However, the measured total fouling
analysis. The surest way to have a suc- service heat transfer coefficient. It is factor calculated with accurate field
cessful, low investment revamp is to the largest resistance to heat transfer. measured data averages 0.009. Based
carefully define your current operating User input fouling factors are, at best, on the measured equipment perfor-

March/April 1999 WORLD REFINING 3

F O C U S O N : HYDROPROCESSING OPTIONS

a plot is the objective of the PI and this

is how it is used by the experienced
operators. These trend lines are great
for showing if today’s operation is bet-
ter or worse than yesterday’s, last
week’s or the same month last year.
Biases and offsets, built into the trend
line for an entire time period, do not
matter in making comparisons
between time periods. In looking at
absolute limits of equipment perfor-
mance being pushed in a revamp,
absolute values are critical.
Second, advanced plant infor-
mation systems all have means to rec-
oncile plant data. Whatever the
method, the assumption is that some
of the data is incorrect and must be
reconciled. Reconciled data simply
means the incorrect numbers are cal-
culated from some relationship. It may
be a simple material balance (flow
Figure 3. Kettle Reboiler Flooding in=flow out) or the heat and material
balance (HM&B). However, each
mance, the exchanger is either under- tems and the numbers they report. increasing level of sophistication
surfaced or has high fouling. An Management information systems (material balance is less complex than
exchanger’s performance one week after (MIS) gather numbers from plant the HM&B) must use more plant mea-
startup is not an indicator of average instrumentation systems, store the sured data. Again, the under-lying
operation and should not be the basis numbers and manipulate the numbers assumption of data reconciliation is
for a revamp service exchanger rating. into a variety of reports. At best, statis- that computer models can correct any
tical methods may be used to fill in the piece of data given enough informa-
Data, Numbers and Test Runs missing numbers or to attempt to cor- tion. The need to reconcile the data is
The proper place to start a revamp, rect obvious errors. caused by numerous problems from
whatever the objective, is with a well Using only the MIS (or plant instrument failures to incorrect instru-
planned and executed test-run of the information, PI) data to perform a ment location in the field vs. the flow
operating unit. Test-run planning revamp requires several assumptions diagram. The test-run is a structured
identifies specific data to be gathered for this approach to be successful. effort to gather comprehensive compo-
so that equipment performance can be Three assumptions are of particular sition, temperature and pressure data
quantified. The specific data is identi- importance. First, that nearly all of the around each major piece of equip-
fied on a “flag sheet.” “Why perform a information coming in is correct. Sec- ment. 9 If data reconciliation is
test run, when we have data from our ond, that statistical methods will cor- required, it is applied locally where it
control system?” is the inevitable ques- rectly identify the faulty data and fix it has a reasonable chance of working.
tion that arises from management. based on a relationship with the cor- Plant-wide data reconciliation is still
This question is natural. After all, hav- rect data. And last, that all the only a dream.
ing spent what is most likely millions required data to quantify equipment Third, often the data required to
of dollars for an advanced manage- performance has been gathered. These quantify equipment performance is not
ment information system, manage- assumptions often fail. First, plant pre- routinely gathered by the MIS systems.
ment expects that the numbers from ventive maintenance to get instru- Returning to our FCC depropanizer
the system will be usable for any ments to read to a correct absolute shown in Figure 2, a properly planned
desired purpose. This question, while reading is rarely done in a refinery.8 It and executed test-run must gather the
natural, arises from misconceptions is generally not necessary for day-to- required data to quantify the
about management information sys- day operations. Having a trend line on depropanizer reboiler performance.

4 WORLD REFINING March/April 1999

 F O C U S O N : HYDROPROCESSING OPTIONS

Rigorous heat exchanger models

use the exchanger mechanical configu-
ration and the process fluid physical
and transport properties to rate the
heat exchanger. Heat exchanger pro-
grams “rate” the exchanger based on
user input fouling factors and process
conditions. If the data is not accurate
the exchanger rating will not be cor-
rect. While errors may not matter
when a new plant is built and the REACTOR
EFFLUENT
exchanger is significantly over-sur-
faced, low cost revamps must fully uti-
lize the existing equipment perfor-
mance. When determining if a piece of
equipment can be pushed, these data
errors can lead to wrong conclusions.

Computer Models
Engineering computer models are use-
ful tools, however, they are not a
panacea. As an example, rigorous heat
exchanger models require measured Figure 4. Slurry Pumparound Hydraulics
pressure and temperature data to be
properly calibrated. The two major vice duty without having to replace will limit the exchanger service duty
problems with rigorous heat exchang- the exchanger. Figure 3 shows the and must be corrected or the exchang-
er models are the user-input (guessed) depropanizer kettle reboiler system er service duty cannot be increased
fouling factors and calculated pressure. with some relative equipment dimen- beyond current performance.
The exchanger fouling factor must be sions and the test-run measured pres- Returning to the example shown
calculated from plant flowrates and sure and temperature data. Kettle in Figure 2, increasing FCC reactor
temperatures. Often the fouling factor reboiler performance is affected by the conversion requires increased depro-
is defined by the user based on an shell and tube-side process flows. Data panizer reboiler duty. A conventional
“industry” average, which is not accu- analysis and properly interpreting the project’s process design approach has
rate for the specific service. Addition- measured data is important. The mea- concluded that the reboiler surface
ally, the calculated tube-side pressure sured pressure drop from the column area is adequate for the revamp condi-
drop from the exchanger models are to the reboiler is 2 psi. The kettle tions. Figure 4 shows the relative loca-
not correct unless it has been calibrat- reboiler baffle elevation difference is tion of the depropanizer reboiler in the
ed with field measured pressure drop. only 7'-6" to the reboiler vapor return. slurry pumparound circuit and some
This is especially true for fouling ser- Distillation column kettle reboiler hydraulic data. Increasing the depro-
vices. FCC slurry pumparound is a design requires the height of liquid in panizer reboiler service duty by 20%
high boiling condensed aromatics the bottom of the column equal to the requires increased slurry circulation
stream containing coke and catalyst pressure drop from the column to the rate to the exchanger. Increased slurry
fines. The coke and catalyst content reboiler. Assuming the feed to the circulation rate will impact the
vary depending on numerous factors. reboiler has a density of 0.68, the mea- pumparound system hydraulics. The
Catalyst and coke foul heat exchang- sured column to reboiler pressure existing hydraulic limitations will dic-
ers. The fouling tendency is unit spe- drop requires 7' of liquid static head to tate whether higher slurry flowrates
cific. Numbers should never be taken overcome the 2 psi pressure drop. are possible. Test-run measured pres-
as accurate data and the data must be The measured pressure drop implies sure drop and valve position data help
complete. the liquid level in the bottom of the quickly identify opportunities. Re-
In our earlier example, a mini- column is near the vapor return noz- vamp process design uses this data in
mum-cost revamp objective would be zle. This affects both the reboiler and conjunction with the calculated
to increase depropanizer reboiler ser- column performance. This problem exchanger fouling factor to determine

March/April 1999 WORLD REFINING 5

F O C U S O N : HYDROPROCESSING OPTIONS

valve changes to increase the tube

velocity and reduce fouling. Decreas-
ing the number of tube passes is not a
reliable solution. Whether pump mod-
ifications are feasible is situation spe-
cific. However, decreasing the number
of tube passes aggravates an existing
problem. Reliable minimum-cost
revamp designs are not determined by
rote application of rules. There is only
one given in a revamp and that is
“there are no rules, only the specifics
of the situation.”

Case Studies—Use Of
Test Run Data
In recent years, FCC unit converter
section revamps have increased con-
version through higher riser feed tem-
peratures, higher catalyst/oil ratio, and
improved converter section hardware
and catalyst technology. These con-
Figure 5. Main Fractionator Pumparounds verter section revamps have affected
major FCC product recovery equip-
future performance. The following tube erosion. FCC exchanger designs ment. A product recovery revamp to
potential modifications will allow the (hydraulics permitting) using slurry recover the lighter reactor products
slurry pumparound flowrate to be pumparound should be operated at 8 will need to evaluate the integrated
increased: ft/sec or higher to minimize the foul- main fractionator, wet gas compressor
• Decrease tube side passes from 4 to 2 ing rate and the fouling factor. and gas plant performance. Three
• Increase slurry pumparound control Process flow sheet models and parts of the product recovery section
valve size rigorous equipment evaluation pro- that are significantly impacted are the
• Pump impeller size or turbine speed grams are necessary tools for revamp slurry pumparound system, gas plant
Prior to determining the most process design. They should never be reboilers and main fractionator heat
cost-effective and reliable modifica- used without comprehensive plant balance. These case studies will high-
tion, the specific exchanger service test-run data. Rigorous exchanger light examples where test-run data was
must be considered. Previous operat- models can be used to quantify tube used to minimize capital investment,
ing history on the exchanger indicates velocity, fouling factors and exchanger identify under-performing equipment
a fouling rate that requires the surface to meet future conditions. In and implement reliable revamps in
exchanger be taken out of service after the depropanizer reboiler case, the FCC product recovery section.
12 months. FCC slurry pumparound conventional project’s approach rec-
service has relatively high fouling ommended decreasing the number of Slurry Pumparound Steam
rates. Fouling service exchangers are tube passes to eliminate the slurry Generators-Example
affected by the process fluid, tube side pumparound hydraulic limitations. In Increasing conversion typically in-
velocity, fouling rate and fouling fac- fact, lowering tube velocity will creases the slurry pumparound heat
tor. These are related. Higher exchang- increase the rate of fouling and reduce removal requirements. Figure 5
er tube velocity lowers the fouling rate exchanger run length. The depro- shows the four main fractionator
and the fouling factor. Conversely, panizer reboiler measured fouling fac- pumparound heat removal systems.
lower tube velocity will increase the tor of 0.009 is above what can be tol- Figure 6 shows the slurry pump-
fouling rate. Industry “standards” state erated if future service duty requirements around system. The slurry pump-
the velocity in this service should be are to be achieved. The only reliable around system removes riser super-
5-8 ft/sec, however, some refiners have modifications to the depropanizer heat with two identical parallel steam
operated at 13 ft/sec with minimal reboiler system is pump and/or control generators (third is used as a spare)

6 WORLD REFINING March/April 1999

 F O C U S O N : HYDROPROCESSING OPTIONS

and the debutanizer reboiler. The total

slurry pumparound system flowrate
goes to the two steam generators with
a portion of the steam generator outlet
used to reboil the debutanizer. The
remainder is by-passed around the
debutanizer reboiler. A converter sec-
tion revamp will have the following
impact on the slurry pumparound sys-
tem:
• Increase total slurry pumparound
duty
• Increase the debutanizer reboiler
duty-more gasoline and alky feed
Slurry pumparound steam gen-
erators can produce 150-600 psig
steam. The steam pressure is a func-
tion of refinery steam system balance
and the investment cost. In this case
steam is produced at 250 psig from
two parallel kettle (TEMA AKT) steam
generators. Typical industry total foul-
ing factors for this service vary from Figure 6. Slurry Pumparound Exchangers
0.003-0.006. These fouling factors
result in a service U-value between 75- The implications of the test-run mea- lated U-value is correct, let’s use the
130 Btu/hr-ft2-°F. Applying this design sured data is important when deter- test-run data to check for consistency.
criteria to our system results in a cal- mining what equipment should be First, the steam and the boiler feed
culated steam generator surface area modified. meter showed 88,000 lb/hr and
equal to twice that installed. The exist- Prior to assuming that the calcu- 92,000 lb/hr, respectively. Assuming a
ing steam generator rating shows the
steam generators under-surfaced,
when they appear to be operating at
much higher service U-values.
Figure 7 is a schematic of the
kettle steam generator. The boiler feed
water rate, steam rate, and the specific
test-run data collected is indicated.
The measured temperatures and
flowrates are used to model the steam
generator. The calculated service U-
value is 240 Btu/hr-ft2-°F. Therefore,
the exchanger was exceeding the
design heat duty based on industry
standard service U-values by almost
100%. The steam generator slurry
pumparound inlet flow resulted in a
tube velocity of 13 ft/sec. The
exchangers had been operating for
several years under these conditions.
Every 12-18 months an exchanger is
taken out of service for cleaning and
the spare exchanger put in service. Figure 7. Slurry Steam Generator

March/April 1999 WORLD REFINING 7

F O C U S O N : HYDROPROCESSING OPTIONS

not follow design averages, it nonethe-

less is reality. Minimizing capital
investment in this FCC would involve
maintaining the high slurry steam gen-
erator tube velocity. Without accurate
test-run data this conclusion would
not be possible. If one wants rote
design practices they should stick to
grassroot projects and avoid the
vagaries of revamps.

Debutanizer Reboiler
Debutanizer reboiler duty increases
when the FCC unit conversion
increases. The duty increase is a func-
tion of reactor yield shift and the
debutanizer fractionation objectives of
alky feed recovery, alky feed C5 olefin
content, and gasoline Rvp. Generally,
the debutanizer reboiler duty increases
between 30–70% when conversion is
increased. Accurately measuring the
Figure 8. Debutanizer Reboilers debutanizer service U-value is impor-
tant. Figure 8 shows the measured
reasonable blow-down of 2% the and valve Cv curve, the steam pro- temperature and pressure data for the
steam/water data are consistent. Based duction was calculated at approxi- reboiler. The low tube-side pressure
on the measured service U-value and mately 90,000 lb/hr. Additionally the drop indicates low velocity (start-of-
theoretical calculations of shell and exchanger tube-side calculated pres- run). Low velocity will increase the
tube-side heat transfer coefficients a sure drop based on the metered slurry fouling rate and increase the average
total fouling factor of 0.0005 is calcu- pumparound flowrate was 24 psi. The fouling factor over the exchanger run
lated. This is essentially a clean measured pressure drop was 22 psi. length. This exchanger has a history of
exchanger. Flowmeters can be wrong, Therefore, the metered slurry severe fouling. Debutanizer reboiler
although both the slurry and pumparound flowrate also appear to (using slurry) fouling can be high on
water/steam flows would all need to be correct. both the shell and tube-sides,
be in error. The test-run data and subse- although high shell-side fouling usual-
Comprehensive test-run data quent data shows high tube velocity ly results from using 700°F slurry.
includes information to reconcile data results in very high service U-values. Table 1 shows the calculated
errors. The test-run data included con- The high slurry tube velocity keeps debutanizer reboiler fouling factors for
trol valve pressure drop and % open. the exchanger tubes clean. Revamping three different FCC units using slurry
Control valve data can be used to this system by adding an additional as reboiler heat. The numbers vary
identify major flowrate errors. Using steam generator in parallel will require significantly. Variation is caused by
both the boiler feed water and steam higher slurry pumparound flow rate, slurry oil physical properties (API
control valve pressure drop, % open, otherwise the existing exchanger tube Gravity=-3.0 to 4.0), catalyst fines,
Table 1. Debutanizer Reboiler
velocity will decrease. Decreasing tube coke fines, tube-side velocity and
(Fouling Factor) velocity to 8 ft/sec lowers the calculat- exchanger design.
ed service U-value clean to 200 Revamping the converter section
Unit # Fouling Factor Btu/hr-ft 2-°F. A qualitative revamp increases debutanizer reboiler duty sig-
(Ft2-hr-°F/Btu)
consideration must be made to deter- nificantly. Therefore, the slurry
1 0.006 mine the impact of lower tube velocity pumparound flowrate must be
2 0.016 on the existing exchanger fouling rate increased to meet higher reboiler duty.
3 0.009 and the fouling factor. While operating This debutanizer reboiler has a calcu-
at a tube velocity above 13 ft/sec does lated fouling factor of 0.016. Rigorous

8 WORLD REFINING March/April 1999

 F O C U S O N : HYDROPROCESSING OPTIONS

exchanger calculations indicate the

tube velocity is approximately
4 ft/sec.10,11 Ultimately, the revamp
engineer will have to make a decision
about the relationship between fouling
factor and tube velocity. Reducing the
overall fouling factor from 0.016 (2
months operation) to 0.009 significant-
ly affects the service U-value. It is pos-
sible to meet future reboiler duty for
this service by increasing tube velocity
and decreasing the average fouling fac-
tor to 0.009. Increasing slurry
pumparound flowrate would increase
the exchanger tube velocity, which will
reduce the fouling factor. Minimum-
cost revamps require engineering deci-
sions that are not part of any ease-of-
use computer model.

Stripper (De-ethanizer Reboiler)

Gas plant stripper columns remove C2
and H 2 S prior to feeding the Figure 9. Deethanizer (Stripper) Reboilers
depropanizer or the debutanizer. The
reboiler duty largely tracks the % gaso- to each reboiler the above design assumptions are cor-
line, and to a lesser extent, the alky • 50% of the LCO pumparound total rect. Test-run measured temperature
feed production. Increasing gasoline circulation rate flow to each reboiler data should be used to identify once-
production by 20% will increase the The plant piping and column through and parallel reboiler operating
stripper reboiler duty by 20-35%. internal design will dictate whether performance. Evaluating the measured
Large changes in alky feed production
also increase the stripper duty. Figure
9 shows a stripper reboiler system
using FCC main fractionator LCO
pumparound (Figure 5) as reboiler
heat. This is a relatively common gas
plant/main fractionator heat integra-
tion scheme. Often, the stripper
reboiler performance is limited by
LCO pumparound heat availability or
poor stripper reboiler draw/piping sys-
tem design.
Figure 10 shows the reboiler
system and test-run measured stream
temperatures. The unit is designed as a
once-through reboiler system. The
stripper column bottom tray liquid
feeds two identical parallel reboilers.
The following assumptions are made
when this system was designed: –TEMPERATURE, °F
• 100% of bottom tray liquid feeds the
two reboilers
• 50% of the bottom tray liquid flow Figure 10. Deethanizer Reboiler Temperatures

March/April 1999 WORLD REFINING 9

F O C U S O N : HYDROPROCESSING OPTIONS

ucts. If bottom tray liquid by-passes

the reboiler, then the reboiler return
temperature must be higher to meet
the column bottom product C2 control
objective. In this case (more often than
not), a valve tray is used to draw liq-
uid to the reboiler system. Active trays
should never be used to draw liquid to
a once-through reboiler because they
will leak. Figure 11 represents the
actual operation of this reboiler sys-
tem. The bottom product temperature
is 12°F colder than the reboiler vapor
return (measured at the column). This
indicates that some of the bottom tray
liquid leaks through the tray deck.
The leakage rate can be estimat-
–TEMPERATURE, °F ed with a computer model. Replacing
the active tray with a seal-welded col-
lector tray will ensure all the liquid
Figure 11. Reboiler Draw System goes to the once-through reboiler.
The reboiler design assumes
test-run temperatures shows that all bottom tray liquid and bottom prod- equal process stream flows to both
the above assumptions are incorrect. uct. Using a once-through reboiler reboilers. Evaluating the measured
A once-through thermosyphon maximizes LMTD. FCC, saturate gas temperatures shows the tube and
reboiler is essentially one theoretical and delayed coker strippers (de-etha- shell-side process flows are not equal,
stage of fractionation. These reboilers nizers) all have 40-60°F temperature although, the exact imbalance cannot
are designed to take advantage of large differences between the bottom tray be determined. The two reboiler outlet
temperature differences between the liquid and the column bottom prod- vapor stream temperatures vary by 5°F
and the LCO pumparound outlet tem-
peratures are not equal. This reboiler
system was designed improperly,
resulting in under-utilized exchanger
surface area. A revamp should consid-
er opportunities to fully utilize existing
equipment.

FCC Main Fractionator

Heat Balance
Increasing reactor conversion shifts
product yields from LCO and decant
oil to gasoline and lighter product.
Reducing the LCO and decant oil prod-
uct yields, shifts the main fractionator
heat balance. Figure 5 shows an FCC
main fractionator pumparound system.
When gasoline yield is increased, the
reflux ratio in the gasoline/LCO frac-
tionation section must be maintained.
The top pumparound duty must be
increased to maintain gasoline/LCO
Figure 12. Main Fractionator Gasoline/LCO Fractionation separation. Concurrently, the LCO

10 WORLD REFINING March/April 1999

 F O C U S O N : HYDROPROCESSING OPTIONS

pumparound duty decreases because

more column internal vapor must flow
to the top pumparound. At the same
time, the LCO pumparound duty
decreases, the gas plant stripper reboiler
duty increases. In this case, the LCO
pumparound supplies the stripper
reboiler heat, then the reactor conver-
sion increase creates a fundamental
main fractionator/gas plant heat integra-
tion system design problem. Minimum-
cost revamps must address the integrat-
ed system affects of shifting yields.
Figures 12 and 13 show a poten-
tial revamp that addresses both the
increased main fractionator top
pumparound heat removal and the
increased stripper reboiler duty
requirements.12 Assuming the stripper
reboiler design uses two parallel reboil-
ers, then one can be converted to top Figure 13. Deethanizer Reboiler Revamp
pumparound service. The feasibility of
this will be unit specific, however, this rate, the flux limitation often prevents circulation rate increases LMTD but
specific modification has been imple- taking advantage of the higher decreases the service U-value. Heat flux
mented successfully. The top exchanger LMTD. Once the exchanger limitations result in no additional strip-
pumparound draw temperature is typi- flux limit is reached, the service U- per reboiler duty.
cally 325°F vs. an LCO pumparound value decreases as the LMTD is Figures 12 and 13 modifica-
draw temperature of 440°F. Therefore, increased. Higher LCO pumparound tions address both the main fractiona-
the top pumparound exchanger LMTD
is less than the LCO pumparound
exchanger. The use of series reboilers
has advantages that are not apparent
unless the reboiler system is thorough-
ly evaluated.
Using top pumparound heat for
one of the two parallel reboilers
requires a thorough evaluation of the
reboiler system. Thermosyphon service
heat transfer coefficient is affected by
both a heat flux limitation and the
shell-side percent vaporization. The
heat transfer coefficient drops when
vaporization increases beyond
25–30%. Stripper reboilers using one
heat service (Figure 9) have a vaporiza-
tion rate between 40-50%. Therefore,
the inherent stripper reboiler service
U-value is low due to high percent
vaporization. Also, the heat flux in
these services is limited to about
13,000-15,000 Btu/hr-ft2. Depending
on the LCO pumparound circulation Photo 2. Maximize Equipment Utilization

March/April 1999 WORLD REFINING 11

F O C U S O N : HYDROPROCESSING OPTIONS

tor heat balance problem and an cumventing the limitations. Test-runs Capital Approach,” Fuel Technology
inherently poor reboiler system are the starting point of a revamp, not and Management, March/April
design. All of the bottom tray liquid is theoretical calculations. ■ 1996, pp.37-45.
withdrawn with a seal-welded collec- 7. Golden, S. W., “Approaching the
tor tray to ensure maximum top Acknowledgment Revamp Project,” Hydrocarbon Tech-
pumparound exchanger LMTD. The Prepared for Presentation at the nology Quarterly, Autumn 1995,
top pumparound reboiler supplies Process Optimization Conference Gulf pp.47-55.
40–60% of the total stripper column Publishing Company, Houston, Texas, 8. Sloley, A. W., “Avoid Problems Dur-
reboiler duty. The top pumparound 7-10 April 1997 ing Distillation Column Startup,”
reboiler return feeds one side of the Chemical Engineering Progress,
column bottom, which is partitioned References July 1996, pp.30-39.
with a baffle. The liquid from the top 1. Lieberman, N. P., and Lieberman, 9. Golden, S. W., “Temperature, Pres-
pumparound reboiler return-side of E.T., “Design and Installation Pit- sure Measurements Solve Column
the baffle feeds the LCO pumparound falls Appear in Vacuum Column Operating Problems,” Oil and Gas
reboiler. This series reboiler arrange- Retrofit,” Oil and Gas Journal, Aug. Journal, Dec. 25, 1995.
ment reduces vaporization to less 21, 1991, pp. 75-79. 10.Sloley, A. W. and Martin, G. R.,
than 30% in both reboilers, therefore 2. Lieberman, N.P., and Lieberman, E. “Effectively Design and Simulate
the service U-values go up. The result T., “Inadequate Inspection Cause of Thermosyphon Reboiler Systems:
is maximum service U-value and min- Vacuum Tower Revamp Failure,” Part 1,”Hydrocarbon Processing, June
imum exchanger surface area. An Oil and Gas Journal, Dec. 14, 1992, 1995, pp.67-78.
additional benefit is improved energy pp. 33-35. 11.Sloley, A. W. and Martin, G. R.,
recovery. The main fractionator heat 3. Bloch, H. P., “How the Best Petro- “Effectively Design and Simulate
balance problem is corrected by chemical Petrochemical Companies Thermosyphon Reboiler Systems:
increasing top pumparound duty and will Achieve Reliability,” Hydrocarbon Part 2,”Hydrocarbon Processing, July
decreasing LCO pumparound duty Processing, July 1996, pp. 83-86. 1995, pp.101-110.
for the stripper reboiler. 4. Golden, S. W., “Revamping FCC’s- 12.Golden, S.W., Sloley, A. W., and
Process and Reliability”, Petroleum Fleming, P. B., “Revamping FCC
Equipment Costs Technology Quarterly, Summer Unit MainFractionator Energy Sys-
Test-run data and analysis are a vital 1996, pp.85-93. tems,”Hydrocarbon Processing, Nov.
part of the revamp process. Revamp 5. Golden, S. W., “Minimize Capital 1993, pp.43-50.
process design quantifies equipment Investment for Refinery Revamps,” 13.Golden, S. W., Schmidt, K. D., and
performance, identifies existing unit Hydrocarbon Processing, Jan. 1997, Martin, G. R., “Field Data, New
equipment bottlenecks and economi- pp. 103-112. Design Correct Faulty FCC Tower
cally eliminates the bottleneck. 13 6. Kowalczyk, D., and Golden, S. W., Revamp,” Oil and Gas Journal, May
Understanding how the integrated sys- “FCC Optimization–A Minimum 31,1993, pp.54-60.
tem works and modifying the system
maximizes use of existing equipment.
Process design controls the majority of The author
capital investment, not project execu- Scott W. Golden is a chemical engineer with
tion and procurement. Project execution Process Consulting Services Inc., Houston, TX.
and procurement are important. Never- His work includes field troubleshooting and
theless, these functions change a small applying fundamental chemical engineering
project (<20–30 MM$) installed cost by skills to improve refinery profitability. The com-
only 10–15%. However, process design pany provides revamp process design, optimiza-
determines what equipment must be tion and troubleshooting to the refinery indus-
modified. It is minimizing the pump, try. Mr. Golden was previously a refinery
process engineer and distillation system trou-
piping, heat exchanger and tower mod-
bleshooter. He has a BS in chemical engineering
ifications that have the biggest impact from the University of Maine. He has authored
on a revamp installed cost. (Photo 2). and co-authored more than 75 technical papers
Revamp process design must focus on concerning refinery unit troubleshooting, design
measuring existing equipment perfor- and simulation.
mance and conceptual aspects of cir-

Copyright © Hart Publications • 4545 Post Oak Place, Suite 210 •Houston, TX 77027 • 713/993-9320
 Process Technology
Processing and Operations

Information Management
UPDATES U P D A T E O N

Terminals & Storage

Increase FCC Propylene

Production and Recovery
SCOTT GOLDEN, Process Consulting Services, Inc., Houston, TX, DENNIS KOWALCZYK and
BOB COMPAGNA, Refining Process Services, Inc., Cheswick, PA

F
CC unit propylene production is dependent on the con-
verter, gas compression, and product recovery sections Flue
gas
operation, process flow scheme and equipment design.
Increasing reactor propylene production requires changes to the
reactor operation, catalyst system and/or reactor hardware
Reactor
( F i g u re 1). Once the propylene is produced, it must be Regenerator
processed through the wet gas compressor and gas plant (Figure
2) and ultimately recovered as liquid product. Fundamental
converter and product recovery considerations impacting propy- Steam
1 , 2
lene production and re c o v e ry are re v i e w e d . H o w e v e r,
increasing propylene liquid barrels requires knowing the specific
process and equipment limitations.
Converter section, gas compression and product recovery
Air
performance are interdependent. Improved reactor hardware
and catalyst systems have reduced the dry gas yield, reduced Oil feed
wet gas rate, increased gasoline yield and increased percent
Air blower
propylene recovery. FCC propylene recovery varies from
below 75% to above 98%, depending on reactor yield, gas Figure 1. Converter section impacts % propylene recovery.
plant design and operating conditions. Some general rules
concerning propylene recovery are: FCC Converter
• For a fixed reactor dry gas yield, the reactor gasoline/propy- FCC converter performance can be optimized to increase the
lene ratio affects propylene recovery. Increasing the ratio propylene yield, while minimizing the effect on the existing wet
i n c reases propylene re c o v e ry. Decreasing this ratio gas compressor and gas plant systems. As reactor propylene yield
decreases propylene recovery. is increased, it is essential that the dry gas yield be minimized to
• Decreasing reactor dry gas yield (C2’s and lighter) at con- keep the wet gas rate at a minimum value. Otherwise, potential
stant reactor gasoline/propylene ratio increases the % reactor yield changes may require high capital investment on the
propylene recovery. wet gas compressor and gas plant. All the reactor incremental
• Increasing the gasoline/propylene ratio in the primary propylene and dry gas make flow through the wet gas com-
absorber increases propylene recovery. p ressor to the gas plant. When the total reactor C3 y i e l d s
• Decreasing absorber gas and liquid feed temperatures increase, the dry gas rate must be reduced to minimize compres-
increases propylene recovery. sion system changes and gas plant capital spending.At some
• Decreasing the absorber operating temperature increases point, the gas plant will require a water chilling or refrigeration
propylene recovery. system to recover incremental propylene.
• Increasing the absorber operating pressure increases propy- R e a c t o r / regenerator operation cannot be separated.
lene recovery. However, for the sake of simplifying the discussions, the major
 U P D AT E S O N P R O C E S S T E C H N O L O G Y

This requires maximum C/O ratio and

minimum feed temperature to limit dry
Main
Fractionator Sponge
absorber
gas production. Minimum dry gas produc-
Vapor
tion reduces the wet gas rate and permits
Cooling more propylene to be compressed. Less
water
Cooling
water
dry gas also makes it easier to recover the
Cooling Primary
Comp. Wet gas
water absorber propylene in the primary absorber.
driver comp

Catalyst Management
Catalyst properties have a large impact on
Overhead Interstage High
receiver receiver pressure
receiver
the propylene yield. An increase in equi-
Stripper librium catalyst activity will result in
higher propylene yield. When it is desir-
Cooling able to reduce the degree of gasoline over-
water
Sour
water
cracking, increasing catalyst acti vity
should be considered. Higher catalyst
Debutanizer
feed activity allows optimizing reaction temper-
a t u re to increase propylene yield and
maintain a relatively high gasoline yield.
Figure 2. Wet gas compressor/gas plant. Reactor temperature should be favored
over increasing catalyst activity because
focus will be on the reactor system performance. Here, converter delta propylene yield will be 50% to 100% higher per unit of
section issues affecting propylene yield will include: increased conversion by using temperature.
• Riser outlet temperature/catalyst-to-oil ratio. Increasing propylene yield at the expense of gasoline yield
• Catalyst management. can be achieved by lowering the rare earth content and/or zeolite
• Reactor equipment. unit cell size (UCS), while increasing catalyst zeolite content to
maintain equivalent activity. Catalyst addition rate can also be
Temperature and C/O Ratio increased. In the lab, when adjusted to constant activity level, an
Higher riser outlet temperature increases propylene yield. This octane type catalyst (low UCS and rare earth) will produce 30%
will generally make more propylene, dry gas and increase gaso- more propylene and C4 olefins than a gasoline type catalyst (high
line production up to the point of over-cracking. In a pilot unit UCS and rare earth). However, the gasoline yield will potentially
operated at constant catalyst-to-oil ratio (C/O), propylene yield drop by more than 5 vol% on an absolute basis.
increased from 4.8 wt% to 6.0 wt% as riser outlet temperature Equipment limits should drive any catalyst system optimiza-
was increased from 970ºF to 1,000ºF. In a commercial FCC unit, tion or change. Reducing rare earth or zeolite UCS increases dry
raising reaction temperature increases both propylene and gaso- gas yield, increases the wet gas rate and raises propylene yield.
line proportionately to the point of gasoline over-cracking. Dry These yield shifts may also make propylene recovery more diffi-
gas rate increases will generally limit reactor temperature due to cult. Higher gas rate may require a reduction in conversion or
wet gas compressor constraints. lower feed rate, neither of which is typically desirable from an
At some temperature, the rate of gasoline cracking will optimization standpoint. Often, to maximize propylene against
exceed the rate of gasoline production and a “gasoline over- gas compression constraints, the optimum catalyst formulation
cracking” condition will be reached. Beyond this temperature
(typically 985ºF or above), gasoline yield will remain relatively Higher Propylene
Base Case
constant with increasing conversion or possibly even decrease. At Conversion Production
temperatures above the gasoline over-cracking point, the amount
of propylene produced per unit of conversion will increase as 3.1
Dry Gas 2.9 2.2
gasoline is cracked. Gasoline over-cracking does not reduce
refinery gasoline yields assuming all the C4 olefins are alkylated C3 1.3 1.1 1.5
and there are no limitations on wet gas compression, gas plant, or C3 = 3.7 4.1 6.4
alkylation unit capacity. In one unit, the riser temperature was
C4 S 9.0 9.5 14.4
increased from 995°F to 1,025°F, conversion increased by 3.8
vol% fresh feed (FF), gasoline yield remained flat, and propylene C5 to 430 ºF 48.4 52.2 44.4
yield increased by over 1.1 vol% FF.
430 to 650ºF 20.9 18.0 16.8
Specific equipment limitations impact optimization and
potential revamps. When the wet gas compressor capacity is lim- Decant 8.8 7.8 8.3
iting, achieving the highest possible conversion while minimizing
dry gas yields maximizes the potential propylene yield increases. Table 1. Unit Yields (wt % feed).
 U P D AT E S O N P R O C E S S T E C H N O L O G Y

will produce the lowest dry gas yield per unit of incre-
mental conversion. Base Case Higher Propylene
ZSM-5 additive can be used to increase propylene yield at the Conversion Production
expense of gasoline. However, use of ZSM-5 offers a better overall
yield selectivity profile because it does not produce additional dry Total C3 7.9 7.4 11.0
gas. ZSM-5 makes propylene by cracking the C6 to C12 gasoline minus gas
yield
molecules almost exclusively to light olefins. Results of numerous
pilot plant and commercial studies show that adding ZSM-5 addi- Delta from
tives at levels of up to 6% by weight of catalyst can increase propy- base -6% +39%
lene yield by 80% to 100%. There are a number of different
formulations of ZSM-5 additives and most of these will favor Table 2. C3 minus yields (wt % feed).
propylene production relative to C4 olefins.
feed nozzle changes only. Lower dry gas and higher gasoline yield
Hardware Upgrades can increase propylene recovery; therefore, liquid C3 recovery
FCC hardware changes and the addition of advanced reactor tech- can increase even at lower reactor yields.
nology can have an impact on the yields of both propylene and Advanced riser termination devices have the most profound
gasoline. The changes made in the past decade include improved impact on propylene yield. There are many different commercial
radial feed injection nozzles, advanced riser termination and riser termination designs available for license. The common jus-
reactor quench systems. These equipment changes allow opti- tification is the desire to reduce the residence time of the hydro-
mization of reactor temperature and catalyst system performance carbon vapor in the reactor vessel. This reduces thermal cracking
to maximize production and recovery of propylene. that occurs after the hydrocarbon is separated from the catalyst.
Converting from axial or older style radial nozzles to the In one commercial study designed to determine the potential
most advanced types of FCC feed nozzles reduces dry gas yield impact of installing an advanced riser termination device, the
and increases gasoline yield. This occurs because improved feed reaction mix sampling (RMS) technique was used to collect
dispersion and vaporization results in an increase in catalytic simultaneous samples from the riser outlet and the inlet to the
cracking and a reduction in the thermal cracking reactions that reactor overhead line.3 The results showed that after the reaction
produce dry gas and loss of gasoline yield. Better mixing of oil vapor mixture exited the riser, conversion increased by 4.0
with catalyst reduces delta coke and increases C/O ratio, thereby vol%FF, gasoline yield dropped by 1.0 vol%, and propylene yield
increasing both conversion and the degree of hydrogen transfer. increased by 1.4 vol%.4
The total amount of C3 + C4 increases, but the olefinicity of this A concern with the new riser termination devices is that the
stream is reduced. The net impact on propylene yield is a func- incremental conversion and post-riser propylene yield will be
tion of many complex factors, but in general, total FCC propy- lost. It is often necessary to adjust catalyst formulation, increase
lene production will be about the same or drop slightly assuming catalyst activity and/or increase reactor temperature to regain lost
conversion after installing advanced riser term i n a t i o n s
Primary devices. These adjustments can often recover most if not all
Main absorber
Fractionator bottom of the loss in propylene yield. In the end, shifts in propylene
Vapor
Stripper yield will depend upon the degree to which any post-riser
vapor
Cooling conversion loss is recovered. At equivalent conversion, post-
water
Cooling revamp propylene yield may even increase slightly.
water
Cooling Quench liquid injection at the end of the riser lowers the
water
Comp.
driver
Wetgas
comp
reaction mixture temperature and is sometimes viewed as an
alternative to advanced riser termination. Lower reaction mix-
Primary
absorber ture temperature reduces the degree of gasoline yield loss to
gas
thermal cracking. This results in reduced production of light
Overhead Interstage High
receiver pressure olefins and dry gas caused by thermal mechanisms, and will
receiver receiver
potentially reduce unit propylene yield. Again, adjustments in
catalyst activity and reactor temperature can be used to offset
this reduction.
Sour
water Case Studies
stripper
Stripper
feed
The basic yield structures for three types of commercial FCC
Primary
operations shown in Table 1 provide a range of FCC yields
absorber for quantifying the operational and design issues affecting
liquid
wet gas compression and propylene recovery. The base case
Figure 3. Wet gas compressor system. represents yields from a unit using axial feed nozzles and
 U P D AT E S O N P R O C E S S T E C H N O L O G Y

conventional riser disengaging technology.

The high conversion case incorporates the impact of advanced Base Case Higher Propylene
riser termination and feed nozzles and increased catalyst activity to Conversion Production
improve propylene yield. More gasoline and less total wet gas are
produced in the high conversion case, which will affect both gas Total C3s and 7.9 7.4 11.0
compression and propylene recovery. The third case represents a dry gas yield
commercial resid cracking operation optimized for maximum
propylene production. This is accomplished by adding ZSM-5 C5 to 430 ºF 48.4 44.4
52.2
additive to the circulating inventory and operating at relatively
higher reaction temperature. This results in a significantly lower
gasoline yield than the other cases. Ratio Gasoline 6.1 7.1 4.0
to C3 lighter
Wet Gas Compression
The wet gas compressor system (Figure 3) moves mainly C3- and Table 3. Gasoline C 3 minus ratio.
a smaller amount of heavier gases from the main column over-
head receiver to the gas plant. The wet gas rate is a function of the Eq. 1 Hp= T1 (1545/MW) Zavg [(n/n-1) (P2/P1) n/n-1 -1)]
dry gas yield, mixed C3 production, and the overhead receiver
operating conditions (Table 2). Where:
The base case has a total mixed C3’s and dry gas yield of Hp= polytropic head in feet
7.9%. The high conversion case has a lower value of 7.4% and T1= inlet gas temperature in absolute units, ºR
the high propylene case has 11.0%. The propylene production MW= gas molecular weight
case results in an increase in the total C3- yield of almost 40% over Zavg= average gas compressibility
the base case. The wet gas rate will track the C3- yields unless oper- P2= discharge pressure, psia
ating conditions are changed. P1= suction pressure, psia
The liquid in the main column overhead receiver absorbs n= gas constant
some of the propylene and most of the heavier hydrocarbons.
Table 3 looks at the ratio of gasoline/total C3- gas to show how Eq. 2 (n/n-1) = (k/k-1) np
gasoline yield impacts the wet gas rate. The high conversion case
produces less C3- gas and has more gasoline to absorb the propy- Where:
lene. Alternately, the high propylene case over-cracks gasoline to k= heat capacity ratio of the gas, Cp/Cv
C3 and lighter; therefore, its wet gas rate will be even higher than np= polytropic efficiency
the total %C3- gas increase because there is less gasoline to
absorb it. Polytropic efficiency is a function of the compressor design
and varies with the gas flowrate. The compressor manufacturer
Performance provides an efficiency curve.
Maximizing wet gas compressor performance is essential when
increasing propylene production. Figure 3 shows a typical wet
gas compressor system. Compressor performance curves
determine the capacity and head developed. Total com-
pressor system performance is determined by the low-stage
performance curve (before inter-stage cooler), inter-con-
denser performance and high stage performance curve
(after inter-stage cooler).
A centrifugal compressor operating at a fixed speed
develops a fixed differential head (not pressure) for any
given inlet flowrate. The compressor performance curve is
based on actual inlet gas flowrate, not standard units as
reported by refinery flowmeters. Refinery metered flowrates
must be corrected to inlet conditions to evaluate com-
pressor performance. The head term is reported as poly-
tropic head. Equation 1 shows the components of the
polytropic head term:

Figure 4. Increased capacity: lower polytropic head.

 U P D AT E S O N P R O C E S S T E C H N O L O G Y

compressor high-stage discharge pressure is approxi-

mately 245 psia. Raising or lowering discharge pressure
by 5 psi has a small effect on the compression ratio.
Alternately, the suction pressure on many wet gas com-
pressors is in the range of 24 psia. Increasing this
pressure to 29 psia decreases the compression ratio by
a significant amount and lowers polytropic head.
Lower head increases the compressor inlet capacity.
Therefore, the main column overhead receiver oper-
ating pre s s u re has a large impact on compre s s o r
capacity.
The centrifugal compressor moves a compressible
fluid. The gas density is a function of pressure, temper-
ature, and gas molecular weight. Reactor performance
sets gas molecular weight. Temperature changes have
only a small impact on gas density. Pressure changes
will significantly increase the gas density. Increasing
Figure 5. Overhead receiver temperature: wet gas rate. compressor suction pressure from 18.7 psia (4 psig) to
20.7 psia (6 psig) will increase gas density and reduce
the gas volume by 10.6%.
Gas density is calculated with Eq. 3. Wet gas molec-
ular weight is largely set by the reactor composition. As
the gas molecular weight increases, the gas density
increases for a given pressure. In the base case, the wet gas
has a high percentage of dry gas. The high conversion
case wet gas has more C3 mix than dry gas. Dry gas has a
molecular weight of 21-23. Propylene/propane mixtures
have a molecular weight of 43.5.
As the converter reduces the dry gas rate yield, the
wet gas density increases, therefore, driver power con-
sumption increases. The gas density can be calculated
from the ideal gas law:
Eq. 3 Gas density = P (MW)/RT
Thus, the gas density and the compressor inlet volume
determine compressor mass flowrate, which controls
unit conversion and/or charge rate.
Figure 6. Overhead receiver pressure: wet gas rate.
Driver Power
Polytropic Head Compressor driver power requirements can limit the compressor
Compressor curves show that for a fixed polytropic head, the wet maximum flowrate. When the drivers are limited, the turbine
gas inlet flowrate is fixed. A reduction in polytropic head steam rate and speed, or the motor amps are at maximum.
increases the compressor capacity, while an increase in polytropic Compressor driver energy consumption is a function of the mass
head reduces capacity. Therefore, process changes that decrease flow, compressor polytropic head, compressor efficiency, and gear
head will increase compressor inlet gas volume capacity. efficiency. Compressor shaft horsepower (SHP) is shown in Eq. 4:
Alternately, reactor yield changes that reduce wet gas molecular Eq. 4 Compressor SHP= (m) Hp/[(np) 33000] 1.02
weight decrease compressor inlet gas capacity.
Compressor capacity is also affected by compressor suction Where:
and discharge pressure. This is shown as compression ratio P2/P1. SHP = Shaft horsepower
High-stage compressor discharge pressure P2 is set by the sponge m= Mass flowrate of gas
absorber pressure controller and the system pressure drop. Hp = Polytropic head
Increasing the compression ratio increases polytropic head, which np = Polytropic efficiency
reduces the compressor capacity. Conversely, Figure 4 shows the 1.02 = Assume 2% gear losses.
effect of a 1,500 foot polytropic head reduction on capacity. Small Compressor horsepower limitations affect the capacity and/or
suction and discharge pressure changes have significantly different conversion of some units. Main column overhead receiver tem-
impacts on polytropic head and capacity. For instance, the wet gas perature and pressure affect driver energy consumption. As dry
 U P D AT E S O N P R O C E S S T E C H N O L O G Y

propylene is recovered in the primary absorber.

Compressed gas
Inter-stage water Sponge
Most of the mixed C3’s and some C4’s enter the
Inter-stage oil absorber absorber in the feed gas from the high-pressure
receiver and a much smaller amount in the feed
liquid (Figure 7). Gasoline from the main
column and possibly recycled debutanized
Cooling Primary
water absorber gasoline are used in the absorber to recover
propylene.Typically, the gas leaving the primary
absorber feeds a sponge absorber. The sponge
absorber recovers very little C3’s except when
High the propylene recovery is low (<80%). The flow
pressure
receiver
Stripper
scheme in Figure 7 has a high- pre s s u re
receiver, absorber inter-cooler and uses only
gasoline from the main column in the primary
Cooling
absorber.
Sour water Reactor yields affect propylene recovery by
water
impacting the amount of gas feed entering the
Main Debutanizer absorber from the high-pre s s u re re c e i v e r.
column feed
liquid Lower dry gas yields increase the high-pressure
receiver C3’s condensation, which lowers the
Figure 7. Primary absorber high pressure receiver. gas feed rate to the absorber. Less absorber gas
flow reduces the amount of propylene that
gas rate drops and more propylene is compressed, the driver must be absorbed in the primary absorber. Reducing gas feed
energy consumption increases. rate decreases propylene absorption re q u i rements in the
absorber and lowers operating temperature, which improves the
Receiver Operation percent propylene recovery.
The wet gas rate, for a given reactor yield, is set by the over- Dry gas flowing through the absorber is analogous to steam in
head receiver operating pressure and temperature. Increasing a hydrotreater H2S stripper. Gasoline is trying to absorb the propy-
o v e rhead receiver pre s s u re or decreasing temperature will lene while the dry gas is trying to strip it. The higher the dry gas
decrease the wet gas rate. This increases available wet gas com- yield, the higher the propylene losses for a given absorber/stripper
pressor capacity by:
• Increasing the amount of condensation in the overhead Higher Propylene
Base Case
receiver. Conversion Production
• Increasing the gas density which lowers the volume of
wet gas. Dry Gas 2.9 2.2 3.1
• Decreasing the compressor polytropic head increases the
amount of gas that can be compressed. Total C3s 5.0 5.2 7.9
Main fractionator pressure and temperature can be optimized
through operating changes or by revamping. Figures 5 and 6 show Total C3s and 7.9 7.4 11.0
the effect of inlet pressure and temperature on wet gas rate for one Dry Gas Yield
unit. A low-capital revamp may involve replacing the 4-tube row
fin-fan bundles with 6-row bundles. The 6-tube row bundles will C5 to 430 ºF 48.4 52.2 44.4
have less than half the pressure loss of the 4-tube rows. In one
instance, this increased compressor capacity by over 20% by Absorber L/V
increasing receiver pressure by 2 psi and reducing temperature by Liquid/vapor 3.5 4.9 3.0
10ºF. In another instance, structured packing was used to reduce ratio
main fractionator pressure drop by 5 psi and the column heat bal-
ance was adjusted, thereby reducing the wet gas rate by over Relative gas
30%.5,6 Minimizing wet gas rate by optimizing main column over- feed rate to Moderate Low High
head receiver temperature and pressure is essential when increasing absorber
reactor propylene yields.
Relative
absorber C3 Moderate High Low
Propylene Recovery recovery
Gas plant process flow scheme, equipment design and operating
conditions impact propylene re c o v e ry. The majority of the Table 4. Absorber C3 recovery.
 U P D AT E S O N P R O C E S S T E C H N O L O G Y

high-pressure receiver/absorber equip-

ment control temperature. Absorber
operating pre s s u re should be maxi-
mized up to the compressor or equip-
ment pre s s u re limits. High-pre s s u re
receiver and absorber operating tem-
peratures should be minimized. Gas
plant revamps can include pre-satu-
rator or inter-coolers to minimize
absorber operating temperatures.
The high-pressure receiver oper-
ating temperature sets the absorber gas
feed rate, the primary absorber operating
temperature, and determines propylene
recovery assuming all other variables are
fixed. Figure 8 shows the impact of the
high-pressure receiver temperature on
Figure 8. High pressure receiver temperature. propylene recovery for the three reactor
yield cases, assuming a fixed absorber
operation and equipment design. Table 4 shows the high conver- gasoline temperature, pressure, and inter-cooler duty. As the high-
sion case having low dry gas. The high conversion case is also the pressure receiver operating temperature increases, the propylene
easiest to maintain high percent propylene recovery. Additionally, recovery drops in all cases. The high conversion yield case also
many refiners process external feed streams through the FCC gas shows the highest propylene recovery at any temperature due to
plant. Some examples are listed below: low reactor dry gas production.
• Delayed coker gas Figure 8 shows that propylene recovery is higher for the
• Hydrotreater gas base and high conversion cases. The high propylene yield case
• Reformer stabilizer gas. has relatively low propylene recovery. This is caused by high dry
Many of these feed streams have a low molecular weight gas yield and a high amount of propylene that must be con-
(hydrogen, methane, ethylene, ethane, CO2, H2S, etc.) and these densed and absorbed. Maximizing recovery for the high dry
streams can reduce propylene liquid barrel recovery. In some gas/high propylene yield requires much lower operating tem-
instances, removing these streams increases the total liquid peratures than can be achieved with cooling water tempera-
propylene yield. tures. Additional cooling with a chilled water or refrigeration
Gasoline flow rate in the absorber/high-pressure receiver system is required to boost the propylene recovery. Figure 9
system affects propylene recovery. Higher reactor
gasoline yield decreases the absorber gas feed rate.
The absorber gas feed rate rises with increasing Compressed gas
reactor dry gas yield. Increasing reactor C3’s at Inter-stage water Sponge
Inter-stage oil absorber
constant dry gas yield also raises the absorber gas Cooling
feed rate. Table 4 shows the reactor yields and the water

relative amount of gas flowing to the absorber. Chilled

water
The primary absorber liquid/vapor (L/V) ratio
Cooling Saturator
also affects the propylene recovery. The value water receiver

shown in Table 4 is not the true L/V in the Chilled

water
absorber, but the ratio of reactor moles of gasoline Sour
water
and dry gas assuming 100% recovery of propy-
lene. While not a theoretically accurate value, it High
pressure
receiver
helps show the difficulty of propylene recovery in Stripper
the high propylene yield case. Chilled
water

Operating Changes Primary

absorber
T h e re are numerous FCC gas plant designs. Sour
water
F i g u re 2 shows the base case system design.
Debutanizer
Increasing high-pressure receiver and absorber Main column feed
operating pressure and decreasing temperature overhead liquid

improves propylene recovery. Wet gas compressor

performance limits operating pressure and the Figure 9. Minimizing operating temperature maximize propylene recovery.
 U P D AT E S O N P R O C E S S T E C H N O L O G Y

Minimum Capital Approach,” Fuel

Technology & Management,
March/April 1996, pp. 37-45.
5. Golden, S. and S. Fulton,
“Low-Cost Methods to Improve
FCCU Energy Efficiency,”
Petroleum Technology Quarterly,
Summer 2000, pp. 95-103.
6. Golden, S.W, et al., “FCC Main
No Intercooler
Fractionator Revamps,”
Intercooler
Hydrocarbon Processing, March
1993, pp. 77-81.
7. Campagna, R. J., D. Kowalczyk, J.
Wilcox, “FCC Operating Variable
and Feed Quality Effects,” 4th FCC
Forum, May 2-5, 2000.

The Authors
Scott W. Golden is a chemical engineer
Figure 10. Propylene recovery with chilling. with Process Consulting Services, Inc.,
Houston, TX. He specializes in identi-
shows a modified process flow scheme using a pre-saturator fying refinery profit making opportunities and specifying min-
drum and chilling to improve C3 recovery by lowering oper- imum capital cost solutions. Mr. Golden was previously a
ating temperatures well below those achievable with cooling refinery process engineer and distillation system troubleshooter.
water. While not a standard feature, pre-saturators are installed He has a BS in chemical engineering from the University of
on some FCC gas plants. Figure 10 shows the impact of the Maine. He has authored more than 75 technical papers on
chilling system on propylene recovery for the high propylene refinery unit troubleshooting, design and simulation. Process
case. The curve shows recovery with and without an absorber Consulting Services provides revamp, optimization, and trou-
inter-cooler. The inter-cooler lowers the absorber operating bleshooting services to the worldwide refining industry.
temperature and typically increases recovery by about 3%.
Dennis C. Kowalczyk is director of operations and one of the
Conclusions principals of Refining Process Services. He oversees FCC refinery
Reactor yields, wet gas compressor and gas plant performance are technical services and various FCC catalyst evaluation programs.
integrally linked. Increasing reactor propylene yield and the per- In addition, he is responsible for monitoring the technical and
cent propylene recovery requires a concurrent evaluation of con- economic trends, which will affect the petroleum re f i n i n g
verter7, wet gas compressor system and gas plant operation. As industry and for developing new technical seminars to meet the
reactor dry gas production rises, the wet gas rate increases and refiner’s evolving needs. Prior to the startup of Refining Process
the propylene recovery drops at the same operating condition. Services, Mr. Kowalczyk spent 5 years with the Gulf Oil Corp.,
Maximizing equipment perf o rmance becomes incre a s i n g l y where he was involved in hydrotreating, catalytic reforming and
i m p o rtant as incremental propylene production targets are FCC projects and was responsible for providing technical sup-
increased; otherwise, revamp capital investment will be high. port for Gulf ’s fluid catalytic cracking units. He holds a BS
Optimizing or revamping the FCC unit to increase propylene degree in chemical engineering from West Virginia University
liquid barrel production requires a thorough evaluation of the
complete system that affects propylene recovery. Robert J. Campagna is the director of Technical Services and one of
the principals of Refining Process Services. He is currently involved
References in catalyst evaluation studies, refining industry technical service and
1. Campagna, R. J., et al., “New Optimization Tools and training program presentations. Mr. Campagna was previously
Technologies for the FCCU,” 1993 NPRA Spring National employed by Filtrol Corp., where he provided technical and mar-
Meeting, March 19-21, 1993. keting support for Filtrol’s fluid catalytic cracking catalysts. He also
2. Golden, S. W., “Approaching the Revamp,” Hydrocarbon spent 10 years with the Gulf Oil Corp., where he contributed to the
Technology International, Autumn 1995, pp. 47-55. areas of hydrotreating, catalytic reforming and fluid catalytic cracking
3. Campagna, R. J.,et al., “FCC Reactor Effluent Sampling: A technical service. Mr. Campagna is an independent consultant in the
Valuable Tool for FCC Unit Optimization,” 1989 NPRA field of fluid catalytic cracking and has presented numerous technical
Spring National Meeting, March 19-21, 1989. seminars throughout the world. He holds BS and MS degrees in
4. Kowalczyk, D. and S. Golden, “FCCU Optimization: A chemical engineering from the University of Pittsburg h .
Volume 1, Issue 1

■ OPERATIONS IMPROVEMENT AND FIELD TROUBLESHOOTING

■ PROCESS EQUIPMENT MODIFICATIONS
■ PROCESS UNIT REVAMPS

Process Consulting Services offers a range of engineering

services that provide you with sustainable refinery
improvements.

Even the best-run refinery has opportunities to improve

its bottom line. While sophisticated computer modeling

and theoretical knowledge may improve profitability,

many times, they don’t. When refinery experience shows

computer models alone do not yield reliable and sus-

tainable performance improvements, a different approach

is required. This approach assumes nothing; it measures

actual performance in the field and doesn’t make office-

based assumptions. Ultimately, it is process unit design

and actual equipment performance that determine

maximum unit profitability, not the ideals of a computer

model. Practical and field-proven experience must sup-

plement the use of computer models to identify, capture,

and maintain profit improvements. ■

THE BOTTOM LINE

Process engineering services that
identify improvement opportu-
nities, eliminate bottlenecks, and
correct performance problems
improve the bottom line.
 PROCESS CONSULTING SERVICES

Operations improvement ASSUME NOTHING

and field troubleshooting Operations improvement and

quickly identify the opportu- field troubleshooting require

nities. Sometimes no capital accurate data. Knowing what

investment is required to data is meaningful and where to

improve the bottom line. get the data requires experience.

Other times, process equip- Nothing should be assumed.

ment modifications are Office-based assumptions

needed. Major equipment about equipment performance

design must consider oper- are often wrong.

ability, flexibility, and relia-

bility issues. Overall process
unit performance depends
on integrated equipment
systems (process design) and OPERATIONS ically reduce profits. Other
OPERATIONS IMPROVE-
actual equipment operation. IMPROVEMENT times, exaggerated claims
MENTS CAN INVOLVE:
A process unit revamp may AND FIELD of equipment performance
be required to meet refinery TROUBLESHOOTING cause the bottlenecks. What- • Crude/Vacuum Units

management’s goals. Mini- ever the case, there is usually • FCC Product Recovery:

mizing capital cost while Opportunities exist to one feature in common: the Main Fractionator

producing reliable and improve operations and exact root cause can be hard Through Gas Plant

measurable results is our circumvent bottlenecks to diagnose. • Coker Product Recovery:

objective. Process Consulting with very low or no capital. The first step to improve Coker Main Fractionator

Services uses an incremental Identifying operations operations or correct a prob- Through Gas Plant

approach that allows engi- improvement areas and lem is to find out, in detail, • Saturate Gas Plant

neering cost-control and field troubleshooting require how the unit is operating. • Refinery Distillation

flexibility when identifying more than sophisticated This requires accurate field-
and capturing operation computer modeling tools. measured data and then
improvements. ■ Sometimes, simple things analysis of the data by
like non-ideal equipment practical and experienced
piping arrangements dramat- engineers.
PERFORMANCE
What is the
real potential?
Finally discovering

the real cause of a

REALITY CHECK
problem shows the Equipment models, such as heat
exchangers, are built with
way to its definitive
underlying assumptions about

solution. how equipment should perform

in a given service. Refinery
equipment must be designed to
minimize fouling, erosion,
corrosion, and coking, not just
meet one set of process loads.

Volume 1, Issue 1
 ESSENTIAL KNOW-HOW
Once accurate field data is

For troubleshooters, gathered, fundamental

equipment operating principles
the most useful are used to evaluate actual
performance. The difference
tools are often the
between actual and potential
simplest. performance determines the
improvement opportunities.
Equipment design know-how
is essential.

FIELD At Process Consulting Each type of major process OPERATIONS

TROUBLESHOOTING Services, troubleshooting equipment requires different IMPROVEMENT
A troubleshooter has several engineers are experts in data to evaluate performance. Once a unit performance
tools at his disposal. The knowing what data is needed Data requirements are based limitation is identified, every
most useful tools are often to identify opportunities and on fundamental equipment effort is made to make oper-
the simplest: thermocouples, diagnose existing problems. operating principles. Iden- ating changes that circum-
pressure gauges, and field This experience is proven in tifying problems with heat vent or mitigate the bottle-
observations. Whatever the field, not behind a desk exchangers, distillation neck. Sometimes, simply
field equipment the trouble- working on theoretical com- columns, vacuum ejectors, changing part of the unit
shooter uses, it should puter calculations. fired heaters, vacuum unit operation can provide signif-
provide reliable, accurate, transfer lines, and other icant improvements. Other
and easy-to-interpret clues refinery equipment requires times, low-cost process flow
to help find the opportunities knowing how to apply scheme changes can be made
or the problems. Data from fundamental equipment without having to shut the
the control room and PI sys- operating principles. unit down. At times, a unit
tems may be useful. Sophis- shutdown is required to
ticated tools, such as gamma fix the existing equipment.
scans, can provide useful While this is an undesirable
information. However, it is situation, experiencing
often field measurements of an unscheduled shutdown
temperature, pressure, and and then failing to fix the
composition, and field obser- real problem is the worst
FIELD MEASUREMENTS
vations, that indicate where outcome.
ENSURE RESULTS
the opportunity or the trou- The key to finding an
Computer models make simpli-
ble really lies. opportunity or solving a
fied assumptions about equip-
Tools are useful only when problem, big or small, is
ment operation. Sometimes
they are in the hands of finding the exact equipment
model results do not represent
someone who knows where, or system limitation. ■
reality. Field measurements are
when, and how to use them.
needed to calibrate computer
models to be sure that models
match reality.
 FIELD-PROVEN EXPERIENCE
Simple process flow scheme
changes can eliminate bottle-
necks and problems. Knowing
what changes to make and how
to design the equipment to
accomplish the process objectives
takes field-proven experience,
not book knowledge alone.

PROCESS than on theoretical consider- ment, Process Consulting coker gas plant heat integra-
EQUIPMENT ations alone, modifications Services engineers provide tion systems.
MODIFICATIONS work the first time. increased capacity and ener- Fired heater modifications
Presently, numerous pro- gy efficiency, improved relia- to reduce the coking rate,
Refinery process equipment prietary distillation technol- bility, and improved product increase firing rate, improve
can under perform because ogies exist with varying quality in many services, efficiency, or remove hydrau-
the process operation has operating characteristics and including: atmospheric lic bottlenecks require an
changed, the equipment performance claims. Some crude, vacuum crude, lube understanding of localized
has suffered damage, or the technologies have bona fide vacuum, FCC main fraction- heat-flux variations. Once
original design may have benefits while others make ator, and delayed coker main this is determined, necessary
been flawed. Implementing exaggerated claims. Correct- fractionator columns. changes to coil layout, tube
changes can make sustain- ing operating limitations Practical, low-cost modi- size transition locations,
able improvements in the and design problems takes fications to heat exchanger transfer lines, burner layout,
refinery’s operating perfor- knowledge of what works equipment require consider- and flame pattern can be
mance. and what does not work in ation of all proximate made. Such modifications
When modifications are the real world. upstream and downstream only succeed, however, if
based on fundamental, prac- By making modifications equipment. To increase heat they are based on practical
tical equipment knowledge to existing distillation equip- recovery in a crude unit knowledge. Process Consult-
and a firm grasp of chemical preheat train with minimum ing Services has successfully
engineering principles rather investment, for instance, revamped crude and vacuum
modifications must start heaters.
with process flow scheme
changes involving not just
Minimizing capital
exchangers, but other equip-
MINOR DETAILS –
ment as well. investment takes
MAJOR IMPACT
Process Consulting
Equipment reliability, flexibility, knowledge of what
Services engineers modify
and operability depend on
exchanger system services, works in the real
design details. Small design
including: crude preheat
details, which cost very little, can
trains, FCCU gas plant heat world.
have a large impact on unit
integration, and delayed
profitability.
 PROCESS CONSULTING SERVICES

CAPITAL COST
MAIN SPONGE
FRACTIONATOR
VAPOR COOLING
ABSORBER
Full utilization of existing
WATER

COOLING CHILLED equipment minimizes capital

WATER WATER
COOLING
WATER SATURATOR
expenditure when revamping
COOLING RECEIVER
WATER WATER DRAW
COMP.
DRIVER
WET GAS
COMP. TRAY a unit. The revamp engineer’s
CHILLED
WATER
SOUR SOUR creativity and experience
WATER WATER

HIGH
determine capital cost, not rote
OVERHEAD INTERSTAGE
RECEIVER RECEIVER PRESSURE
RECEIVER
project management structure
STRIPPER
CHILLED
WATER
30
and hierarchy.
PRIMARY
SOUR ABSORBER
WATER
STRIPPER

DEBUTANIZER
FEED
STRIPPER FEED
PREHEATER

Vacuum system and liquid to improve vacuum system PROCESS UNIT a process design package
ring pump performance performance. REVAMPS follows in which all major
problems are leading causes Improper design of equipment design details are
of low product yield and vacuum transfer lines can Implementing a revamp that established based on prac-
poor vacuum gas oil quality. lower vacuum gas oil cut- works the first time and tical engineering know-how,
Vacuum system limits can point, increase coking in the keeps investment to a mini- not vendor proposals.
almost always be traced to heater, and cause high mum is the goal of every Process unit revamps
damaged, poorly sized, or residue entrainment in the refiner. Rather than commit begin with a field survey to
badly piped steam ejectors vacuum column, all of which immediately to a full-scale quickly determine opportu-
and liquid ring pumps. When cause product yield and project, this goal is best nities. Then, a comprehen-
proper consideration is given quality to suffer. Proper reached by conducting pro- sive test run is conducted to
to vacuum system gas loads, design must take into account cess engineering in stages, establish a benchmark, not
as well as equipment opera- pressure drop, critical mass each offering the chance to just of the overall refinery
tion, modifications can correct velocity, and two-phase flow continue or halt work depend- unit in question, but also of
design flaws, fix equipment regime. Process Consulting ing on the results of the all major equipment within
damage, and augment or Services engineers have preceding step. the unit. Feed and product
compensate for undersized corrected existing transfer Determining revamp feasi- streams are analyzed in the
equipment. Process Consult- line problems and designed bility via a field survey comes laboratory. Pressure, temper-
ing Services engineers have new vacuum transfer lines. ■ first. The field survey identi- ature, and composition data
modified process operation fies major bottlenecks, which are gathered throughout the
and ejector system equipment PROCESS EQUIPMENT define project costs and unit to identify all bottle-
MODIFICATIONS MAY potential benefits. Establish- necks that affect the oper-
Modifications INVOLVE: ing a unit benchmark with a ating revenues. Alternative
comprehensive test run comes process flow schemes are
• Distillation
succeed only if they next. At this stage, bottle- evaluated using rigorous
• Heat Exchangers and
necks are defined in much detailed models that have
are based on practi- Exchanger Networks
more detail. Then, alternate been calibrated with test run
• Fired Heaters
cal knowledge. process flow schemes are data. Actual constructability
• Hydraulic Systems
evaluated in the conceptual is determined with extreme
• Vacuum Ejector Systems
process design stage. Finally, care.
• Vacuum Transfer Lines
 PROCESS CONSULTING SERVICES

EXPERIENCE REQUIRED
The final report from
Revamps start by measuring
Process Consulting Services
existing unit performance with a
provides yield summaries,
field survey. Current performance
flow sheet simulation models,
determines whether low-capital
process flow diagrams,
changes can make significant
equipment layout drawings,
improvements. Knowing how
and equipment lists and
major equipment bottlenecks can
specifications — in short,
be eliminated, whether through
all the information a refiner
equipment modifications, process
needs to make a thorough
flow scheme changes, or both,
review and select a detailed
requires experience.
engineering contractor,
equipment vendors, and
a construction contractor.
A workable process design
based on actual field experi- MAN-HOURS OR RESULTS?
ence, not on suppositions and
hopes, is the end result. ■ Operating limitations and problems in a refinery can cause
millions of dollars in losses and needless expenses. What
PROCESS UNIT REVAMPS is needed is a quick and certain fix. That is why, with the sole HOUSTON OFFICE:

COMPLETED INCLUDE: exception of field troubleshooting and field installation 3400 Bissonnet, Suite 130

inspection, Process Consulting Services works only on a Houston, Texas 77005

• Crude/Vacuum Units
lump-sum basis. We focus on results, not billable man-hours Tel: (713) 665-7046
• Lube Vacuum Units
that run up engineering costs and excuses and result only Fax: (713) 665-7246
• FCC Product Recovery:
Main Fractionator in patching, not solving problems. Cost-plus contracts are
notorious in this respect. The hours devoted to any given DALLAS OFFICE:
Through Gas Plant
job by Process Consulting Services engineers may be long 1901 Central Drive, Suite 811
• Coker Product Recovery:
or short, but what you pay for is not the expenditure Bedford, Texas 76021
Coker Main Fractionator
of hours — it is the final, satisfactory, and guaranteed Tel: (817) 358-1566
Through Gas Plant
completion of a necessary task. ■ Fax: (817) 358-2040
• Saturate Gas Plant
• All Refinery Distillation,
E-MAIL:
Absorption, and Extraction
info@revamps.com
Services

WEBSITE:
www.revamps.com

Process Consulting
©1999
Services engineers Process Consulting Services
All rights reserved.
focus on results, not Printed in the U.S.A.

billable man-hours.

Volume 1, Issue 1
Volume 1, Issue 2

■ CRUDE/VACUUM UNITS ■ FCC PRODUCT RECOVERY: MAIN

FRACTIONATOR THROUGH GAS PLANT ■ COKER PRODUCT
RECOVERY: COKER MAIN FRACTIONATOR THROUGH GAS PLANT
■ SATURATE GAS PLANT ■ REFINERY DISTILLATION

A practical hands-on approach can help identify

operations improvement opportunities, unit bottlenecks,
and equipment limitations that are reducing the bottom
line. But just what is involved in such an approach?

Capitalizing on opportunities to improve operations and

circumvent bottlenecks with very low or no capital can

enhance the bottom line. Operating problems, even in the

best-run refineries, are an unfortunate fact of life. Oper-

ations improvement and field troubleshooting require

more than sophisticated computer modeling tools. Some-

times, simple things like non-ideal equipment piping

arrangements dramatically reduce profits. Other times,

exaggerated claims of equipment performance cause the

bottlenecks. Whatever reduces the bottom line, however,

there is one feature in common: the root cause can be

hard to diagnose exactly. ■

MEASUREMENTS
AND KNOW-HOW
Exploiting the difference between
measured and potential equip-
ment performance leads
to sustainable improvements.
 PROCESS CONSULTING SERVICES

FINDING OPPORTUNITIES
Trying to improve oper-
Process equipment performance
ations can be frustrating.
depends on proximate upstream
Instrumentation may indi-
and downstream equipment
cate normal operation in
design. Computer models make
spite of the fact that capacity,
underlying assumptions about
product quality, and/or
the installed equipment that
production rates suffer.
may not match reality. Field data
Operating guidelines may
gathering, field observations,
tell you that you are doing
and an understanding of equip-
everything right, yet obeying
ment interaction help identify
them to the letter does not
the opportunities.
help. Running computer
simulations seems to offer
no better insight either.
Problems occur because
a refinery operates in a harsh FINDING experienced people to are experts in knowing what
environment: equipment OPPORTUNITIES; perform the tests correctly, data is needed to identify
is prone to coking, fouling, IDENTIFYING and personnel experienced opportunities and diagnose
plugging, corrosion, erosion, PROBLEMS with the specific process existing problems. This
and upsets that damage equipment being evaluated experience is proven in the
equipment. The list goes Many different types of tools to interpret the data. Often, field, not behind a desk
on and on. Not only are can be used to find opportu- it is field measurements working on theoretical com-
refineries operating longer, nities or identify problems. of temperature, pressure, puter calculations.
operating severity has Some are as sophisticated as and composition, and field Each type of major process
increased, feedstock changes gamma ray scans, computer observations that indicate equipment has different data
are more prevalent, and most aided topography, neutron where the opportunity requirements to evaluate
refinery equipment is older. scatters that identify distil- or the trouble really lies. performance. Trouble-
Hence, problems occur lation column problems, and It is important to note: shooters must thoroughly
because equipment is being computer simulations. How- tools are useful only in understand equipment
pushed closer to its inherent ever, the most useful tools the hands of someone who operating principles. Iden-
limit. are often the simplest: ther- knows where, when, and tifying problems with heat
The first step to improve mocouples, pressure gauges, how to use them. At Process exchangers, distillation
operations or correct a and field observations. Consulting Services, column internals, vacuum
DO YOU KNOW?
problem is to identify oper- Whatever field equipment troubleshooting engineers ejectors, fired heaters,
Measuring
ating variable changes that the troubleshooter uses, it vacuum unit transfer lines,
equipment
may account for the prob- should provide reliable, ac- and other refinery equip-
performance.
lem’s symptoms. If none are curate, and easy-to-interpret For troubleshooters, ment requires knowing how
What is required?
found, then it is necessary clues to help find the oppor- to apply fundamental equip-
to quickly find the root cause tunities or the problems. the most useful ment operating principles. ■
of the problem — in other Data from the control room
tools are often the
words, troubleshoot. ■ and PI systems may be use-
ful. Sophisticated tools can simplest.
also provide information that
may be helpful. However,
sophisticated tools require

Volume 1, Issue 2
 FIELD MEASUREMENTS
OPERATIONS
Actual pressure, temperature,
IMPROVEMENTS
and composition profiles are
AND FIELD
measured in the field. Assump-
TROUBLESHOOTING
tions are never made! Simple
field measurement techniques,
Once the root cause of a unit
such as pressure gauge calibra-
performance limitation is
tion, are essential. Knowing
identified, every effort is
what data is required, where
made to implement operating
to get the data in the field,
changes that circumvent or
and how to interpret the data
mitigate the bottleneck.
requires experience.
Sometimes, simply changing
part of the unit operation can
provide significant improve-
ments. Other times, low-cost
process flow scheme changes ing the equipment shortfalls Finding opportunities nities and many potential
can be made without having that cause the unit to under demands a thorough under- problems. Presently, several
to shut the unit down. At perform. standing of: 1) integrated proprietary distillation tech-
times, a unit shutdown is To make a long story short: equipment performance, nologies exist with varying
required to fix the existing the key to finding an oppor- as well as an understanding operating characteristics and
equipment. While this is an tunity or solving a problem, of 2) individual equipment performance claims. Some
undesirable situation, it is big or small, is finding out operation. Process Consult- technologies have bona fide
often better to shut down exactly what is the problem. ing Services engineers have benefits while others make
and fix the problem. Experi- That is a specialty of Process experience with all types of exaggerated claims. Finding
encing an unscheduled shut Consulting Services engi- refinery process equipment. ■ opportunities or correcting
down and then failing to neers. ■ design and operating prob-
fix the real problem is the EQUIPMENT INCLUDES: lems with little or no capital
worst outcome. PROCESS UNIT • Distillation Columns
investment takes not only
On operations improve- OPERATIONS • Heat Exchangers and
textbook learning but also
ment and troubleshooting AND EQUIPMENT Exchanger Networks
knowledge of what works
jobs, Process Consulting TROUBLESHOOTING • Fired Heaters
and what does not work in
Services engineers gather the real world.
• Vacuum Ejector Systems
necessary field data, make Refinery processing begins
• Vacuum Transfer Lines
field observations of actual with crude units that sepa-
equipment layouts, and im- rate crude oil into products
DISTILLATION The key to finding
plement operating changes for further processing. Crude
to test the unit response to units and other refinery
Refineries handle heavy oil an opportunity and
specific disturbances. Chem- process units have numerous
streams that feature wide
ical engineering fundamen- types of process equipment
boiling range hydrocarbon
solving a problem
tals and equipment operating with complex interaction
mixtures with varying prop- is finding out exactly
principles are used, rather between the equipment.
erties. Therefore, the design
than hard-to-explain com- what is the problem.
of distillation columns can be
puter-modeling technology.
complex, providing numer-
Problems and solutions are
ous improvement opportu-
explained in terms of real
measurements, clearly show-
 PROCESS CONSULTING SERVICES

PROCESS CONSULTING SERVICES

Engineers from Process FIELD-PROVEN EXPERIENCE

ENGINEERS HAVE IDENTIFIED Consulting Services have Results are what one should pay
IMPROVEMENT OPPORTUNITIES extensive distillation expe- for, not lengthy studies that
AND PROBLEMS INCLUDING:
rience with commercially detail all the activities performed.
DISTILLATION available, proven technol- Activities do not make money
• High diesel product cloud point
ogies, including emerging for the refiner. Process Consulting
• Black atmospheric gas oil
• High microcarbon and metals proprietary high capacity Services engineers have per-
in the vacuum gas oil trays. Operations improve- formed operations improvements
• Low gas oil yield
• Fouled or plugged trays
ments include increased and field troubleshooting in

• Damaged internals capacity and energy effi- more than a hundred refineries
• Premature column flooding ciency, improved reliability, worldwide.
• Off-color LCO
• High FCC gasoline RVP
and improved product
• Poor gasoline/LCO fractionation quality in many different
• Coking — vacuum wash, FCC distillation services. ■
slurry, and delayed coker wash
sections 6–8 years. Vacuum gas oil
• Inability to withdraw sidecut EXCHANGERS yield targets are higher, oil is
products from heavy oils columns
AND EXCHANGER heavier, and firing rates have
EXCHANGER SYSTEMS NETWORKS increased. Controlling local-
• Crude heat exchanger networks
• FCC energy utilization
ized oil film temperatures and
• Reboiler fouling
Exchanger design consider- oil residence time is more HOUSTON OFFICE:
• Reboiler piping problems ations — such as surface critical today than ever. On
• Condenser: inerts venting 3400 Bissonnet, Suite 130
area, shell and tube side top of the increased process
• FCCU slurry steam generators Houston, Texas 77005
• Reboilers heat transfer coefficient, and severity, low NOX burners Tel: (713) 665-7046
• Reboiler draw problems pinch — are important theo- have a longer, colder flame Fax: (713) 665-7246
FIRED HEATERS
retical concerns. However, that further complicates heat
• Rapid heater tube coking evaluation of heat exchange flux variations throughout
• Low convection section duty DALLAS OFFICE:
equipment requires consider- the radiant section. ■
• Air preheat system limitations 1901 Central Drive, Suite 811
• Heat flux distribution
ation of many practical
Bedford, Texas 76021
• High oil residence time issues such as actual fouling, VACUUM EJECTORS
• High localized oil peak film Tel: (817) 358-1566
corrosion, erosion, localized AND LIQUID
temperature Fax: (817) 358-2040
• Coil steam injection optimization water condensation, trace RING PUMPS
feed components, piping,
VACUUM EJECTORS E-MAIL:
AND LIQUID RING PUMPS
and all proximate upstream Vacuum system performance info@revamps.com
• Undersized ejectors and downstream equipment problems are a leading cause
• Ejector nozzle erosion designs. ■
• “Breaking” ejectors
of low product yield and poor WEBSITE:
• Ejector nozzle plugging vacuum gas oil quality. ■ www.revamps.com
• Inter-condenser dip leg plugging FIRED HEATERS
• High piping pressure drop
• Inter-condenser fouling
VACUUM ©1999
• Pressure control piping Crude and vacuum unit TRANSFER LINES Process Consulting Services
• Liquid ring pump strainer systems heater reliability is becoming
All rights reserved.
VACUUM TRANSFER LINES
more important as targeted Moving oil from the heater Printed in the U.S.A.
• Sonic velocity heater run-lengths increase to the column is often con-
• Two-phase flow regime from 3–4 years to as high as
• High pressure drop
sidered so simple that inade-
• Multi-pass heater outlet quate attention is given to
pressure imbalances the design of transfer lines. ■

Volume 1, Issue 2
Volume 1, Issue 3

■ DISTILLATION COLUMNS ■ HEAT EXCHANGERS AND EXCHANGER

NETWORKS ■ FIRED HEATERS ■ VACUUM EJECTOR SYSTEMS AND
LIQUID RING PUMPS ■ VACUUM TRANSFER LINES

Correcting design flaws and debottlenecking operating

equipment while holding capital investment within
reasonable bounds takes not only textbook learning,
but also knowledge of what works and what doesn’t
work in the real world.

Existing refinery equipment can under perform because

the process operation has changed, the equipment has

suffered damage, or the original design may have been

flawed. Whatever the cause, to make the corrective

modifications, the difference between actual and potential

performance must be exploited. When modifications are

based on fundamental, practical equipment knowledge

and a firm grasp of chemical engineering principles rather

than on theoretical considerations alone, necessary capital

investment can be held to a minimum. ■

THE KEY
Understanding process equipment
principles and how integrated
equipment performs is key to
minimizing capital expenditure.
 PROCESS CONSULTING SERVICES

DISTILLATION PRACTICAL

COLUMNS CONSIDERATIONS
Exchanger network design must

With refinery streams consider practical implications;

composed of wide boiling otherwise, actual performance

range hydrocarbon mixtures will not match theory. Fouling,

with varying properties, corrosion, and erosion are

the design of distillation common. Many times, these

columns can be complex. factors determine equipment

Presently, numerous propri- performance.

etary distillation technol-

ogies exist with varying
operating characteristics and
performance claims. Some
technologies have bona fide
benefits while others make Process Consulting EXCHANGERS tightly linked. To increase
exaggerated claims. Services offers a full range AND EXCHANGER heat recovery in a crude unit
Correcting design and of distillation services from NETWORKS preheat train with minimum
operating limitations while field troubleshooting and investment, for instance,
holding capital investment equipment design through Exchanger design consider- evaluation of atmospheric
within reasonable bounds start-up. ■ ations, such as surface area, and vacuum crude distilla-
takes not only textbook LMTD, and pinch, are impor- tion operation is essential.
learning, but also knowledge PROCESS CONSULTING tant theoretical consider- Modification must start with
of what works and what SERVICES ENGINEERS HAVE ations. Exchanger network process flow scheme changes
doesn’t work in the real MADE MODIFICATIONS design also must consider involving not just exchang-
world. TO THESE DISTILLATION practical implications; other- ers, but other equipment
Engineers from Process SYSTEMS: wise, actual performance as well. The same compre-
REALITY CHECK Consulting Services have will not match predictions. hensive approach applies
• Atmospheric Crude
Often, real-world these practical capabilities. Fouling, erosion, and to other refinery units. ■
• Vacuum Crude
equipment They have field-proven corrosion are common.
• Preflash Crude
performance does experience with all commer- Many times, these factors PROCESS CONSULTING
• Preflash Vacuum
not match theory. cially available technologies, determine equipment SERVICES ENGINEERS HAVE
• Lube Vacuum
including emerging propri- performance. MODIFIED EXCHANGER
• FCC Main Fractionator
etary high capacity trays. Practical, low-cost modi- SYSTEMS INCLUDING:
• Delayed Coker
By making corrective fications to heat exchanger
Main Fractionator • Crude Preheat Trains
modifications to new equip- equipment require consider-
• HF Alky Isostripper • FCCU Gas Plant Heat
ment during the design ation of all proximate
• Sulfuric Acid Integration
stage or to existing equip- upstream and downstream
Alkylation Columns • Delayed Coker Gas Plant
ment, Process Consulting equipment. Exchangers are
Heat Integration
Services engineers provide often connected to distilla-
• FCCU Slurry
increased capacity and tion equipment, and the
Steam Generators
energy efficiency, improved performance of the column
• Reboilers
reliability, and improved equipment can affect ex-
• Fin-Fan and Shell-and-Tube
product quality. changer utilization. Objec-
Condensers
tives of column fractionation
and heat exchange are

Volume 1, Issue 3
FIRED HEATERS PERFORMANCE = PROFITABILITY
Deepcut vacuum unit profitability

Heater reliability is becom- depends on steam-jet ejector

ing more important as target- and liquid ring pump operating

ed unit run-lengths continue performance. Series and parallel

to be increased from three equipment interaction, as well

or four years to as high as as individual component perfor-

six to eight years. Yet, when mance, determine vacuum unit

oil film temperatures and operating pressure. Ultimately,

oil residence time are not vacuum unit operating pressure

properly controlled, many controls product yield and quality.

atmospheric crude and crude

vacuum heaters must be
decoked, at least every two
years. Fired heater perfor-
mance depends on practical Once this is understood, VACUUM
PROCESS CONSULTING
considerations, such as coil necessary changes to coil EJECTORS AND
SERVICES ENGINEERS HAVE
layout, tube size transition layout, tube-size transition LIQUID
MODIFIED:
locations, burner flame locations, transfer lines, RING PUMPS
pattern, burner type, and burner layout, and flame • Ejector Nozzles

transfer line design. patterns can be made. Such Vacuum system performance • Steam Systems

When making modifica- modifications succeed, how- problems are a leading cause • Inter-Condenser Design

tions to reduce coking rate, ever, only if they are based of low product yield and poor • Inter-Condenser Dip Legs

increase firing rate, improve on practical knowledge. ■ vacuum gas oil quality. Such • Pressure Control

efficiency, and remove hy- problems can often be traced • Liquid Ring Pump Systems

draulic bottlenecks, actual PROCESS CONSULTING to damaged, poorly sized, or

heater performance must SERVICES ENGINEERS HAVE badly piped steam ejectors
be evaluated and measured. MODIFIED ATMOSPHERIC or liquid ring pumps. When
Controlling localized heat- CRUDE AND CRUDE proper consideration is given
flux, oil film temperature, VACUUM HEATERS TO to vacuum system gas loads,
and oil residence time (rather IMPROVE PERFORMANCE. as well as equipment oper-
than the average parameters) ation, modifications can
is essential to improving correct design flaws, improve
overall reliability. pressure control, fix equip- Modifications succeed
ment damage, and augment
or compensate for undersized only if they are
equipment. ■
based on practical

HOT SPOTS knowledge.

Localized coking causes
hot spots. Process oil film
temperature and oil residence
time must be controlled.
Otherwise, unscheduled
decokings will result.
 PROCESS CONSULTING SERVICES

UNDER VACUUM
VACUUM
Transfer lines should move oil
TRANSFER LINES
without causing high pressure
drop, an unstable flow regime,
Moving oil from the heater
or a choke flow condition any-
to the column is often
where in the line. Otherwise,
considered so simple that
low product yield and poor gas
inadequate attention is
oil quality can result.
given to the design of the
transfer lines. Yet, improper
design can lower vacuum
gas oil cutpoint, increase
coking in the heater, and
cause high residue entrain-
ment in the vacuum column.
As a result, product yield
and quality suffer. Proper
design must take into
account pressure drop,
critical mass velocity, and MAN-HOURS OR RESULTS?
two-phase flow regime. ■
Operating limitations and problems in a refinery can cause HOUSTON OFFICE:
PROCESS CONSULTING millions of dollars in losses and needless expenses. What 3400 Bissonnet, Suite 130
SERVICES ENGINEERS HAVE is needed is a quick and certain fix. That is why, with the Houston, Texas 77005
REDESIGNED EXISTING sole exception of field troubleshooting and field installation Tel: (713) 665-7046
TRANSFER LINES AND inspection, Process Consulting Services works only on a Fax: (713) 665-7246
DESIGNED NEW VACUUM lump-sum basis. We focus on results, not billable man-hours
TRANSFER LINES TO that run up engineering costs and excuses and result only DALLAS OFFICE:
INCREASE PRODUCT YIELD in patching, not solving problems. Cost-plus contracts are 1901 Central Drive, Suite 811
AND QUALITY. notorious in this respect. The hours devoted to any given Bedford, Texas 76021
job by Process Consulting Services engineers may be long Tel: (817) 358-1566
or short, but what you pay for is not the expenditure Fax: (817) 358-2040
of hours — it is the final, satisfactory, and guaranteed
completion of a necessary task. ■ E-MAIL:
info@revamps.com

Process Consulting
WEBSITE:

Services engineers www.revamps.com

focus on results, not ©1999

Process Consulting Services
billable man-hours.
All rights reserved.
Printed in the U.S.A.

Volume 1, Issue 3
Volume 1, Issue 4

■ REVAMP FEASIBILITY
■ UNIT BENCHMARKING
■ CONCEPTUAL PROCESS DESIGN
■ PROCESS DESIGN PACKAGE

A four-stage approach, starting with Revamp Feasibility

and culminating with a Process Design Package,
minimizes capital expenditure, controls engineering costs,
and maximizes return on investment.

Implementing a revamp that works the first time and

keeps investment to a minimum is the goal of every

refiner. Rather than commit immediately to a full-scale

project, you can best reach this goal by conducting process

engineering in stages, each offering the chance to continue

or halt work depending on results of the preceding step.

Determining revamp feasibility via a field survey comes

first. The field survey identifies major bottlenecks, which

define project costs and potential benefits. Establishing

a unit benchmark in a comprehensive test run comes next.

In this stage, bottlenecks are defined in much more detail.

Then, alternative process flow schemes are evaluated

in the conceptual process design stage. Finally, a process

design package follows in which all major equipment

design details are established based on practical engineer-

ing know-how, not vendor proposals. ■

DETERMINING OPPORTUNITIES
Improvement areas must
be identified before they can be
exploited. Some opportunities
are apparent, while others
require extensive field experi-
ence to uncover.
 PROCESS CONSULTING SERVICES

PROCESS FLOW
SCHEME CHANGES

Minimizing total capital Revamps must fully utilize

expenditure while imple- existing equipment; otherwise,

menting a reliable revamp capital costs are not minimized.

that has a quick return is Finding improvement opportu-

every refiner’s goal. A four- nities in the revamp feasibility

stage engineering approach and unit benchmarking stages

that focuses the appropriate is where a minimal capital-cost

resources at the right time revamp should begin.

helps develop a revamp that

can be funded and controls
engineering costs.
Revamp process engineer-
data analysis and computer costs are not justified. If work
ing begins by identifying
improvement opportunities
REVAMP modeling are performed — stops, engineering costs are

and their associated “rough”

FEASIBILITY — determines major equipment minimal. Revamp feasibility

costs. If costs are justified by

Finding constraints. Major equipment avoids expensive engineering

the benefits, then engineer-

Opportunities constraints define gross studies that do not material-
investment levels. Potential ize into funded revamps. ■
ing progresses. Opportuni-
Revamp feasibility estab- benefits are also determined.
ties and costs are defined in
lishes the direction of the Whether it is a crude unit, PROCESS ENGINEERING
more detail as engineering
revamp and determines FCC, or delayed coker, a unit OF A REVAMP SHOULD
proceeds. Each stage offers
the modifications that offer field survey is necessary to BE CONDUCTED IN FOUR
the chance to continue work
the quickest return while get a handle on revamp costs STAGES:
or halt it based on the results
controlling engineering and benefits.
of the previous stage. This • Revamp Feasibility –
costs. A field survey identi- Once feasibility work is
way, engineering costs are Unit Field Survey
fies opportunities and major complete, process engineer-
controlled while progressing • Unit Benchmarking –
constraints that require ing can continue to the next
toward a funded revamp. ■ Comprehensive Test Run
capital outlays. stage, revamp objectives can
• Conceptual Process
A field survey in which be redefined based on avail-
FULL UTILIZATION Design – Evaluating
select measurements are able capital, or the work can
There is no E&C Alternative Process
gathered — and subsequent be stopped because capital
“design safety Flow Schemes
factor” when • Process Design Package –
operating at the NORMALLY
CLOSED
Major Equipment Design
OFF GAS OFF GAS
limits.
MF OVHDS
NORMALLY
Practical engineering CLOSED
CREATIVITY, NOT ROUTINE

know-how keeps PROJECT SKILLS

NEW LINE EJECTOR

Discovering an underutilized
SYSTEM
refiners’ investments LT NAP
crude unit compressor will
materially lower capital cost
to a minimum. VACUUM
COLUMN versus rote project methods
that result in purchasing
a new compressor.

Volume 1, Issue 4
UNIT MAKING EQUIPMENT CHANGES

BENCHMARKING — Full utilization of existing equip-

Identifying ment requires the practical appli-

Bottlenecks cation of fundamental equip-

ment knowledge. This requires

Every process unit has an understanding of how the

indicators of overall perfor- actual physical equipment design

mance. These indicators, affects performance. This know-

together with major equip- how comes from experience,

ment and system perfor- not computer models. Exploiting

mance, determine the poten- the differences between current

tial return on investment and potential equipment perfor-

of a revamp. For instance, mance minimizes costs.

product cutpoint, in conjunc-

ing requires correct feed must be tempered with prac-
tion with process and equip-
characterization, complete tical knowledge. For example,
ment design, can be used
heat and material balances, a fired heater model assumes
to determine opportunities
and internal stream data. that individual pass heat flux
for improving product yield
Feed and product streams and flow rates are equal, but
and quality.
must be analyzed in the the actual heater design may
To accurately assess unit
laboratory. Accurate unit not match this ideal. Exten-
opportunities, unit limita-
profiles of temperature, sive hands-on field experi-
tions must be clearly under-
pressure, and composition ence is needed for modeling
stood. Unit benchmarking
must be created from actual to accurately reflect reality.
identifies unit bottlenecks COMPOSITION, PRESSURE,
measurements in the field, Once the test run is com-
at the ground level. Nothing AND TEMPERATURE PROFILES
not taken from PI systems pleted, priorities can be
is taken for granted and Practical considerations such as
alone. An on-site review of established for revamp objec-
no assumptions are made. examining a black atmospheric
piping must be conducted. tives. Alternative revamp
Equipment performance gas oil (AGO) sample can iden-
However powerful computer schemes can then be consid-
is determined by hands-on tify a low-capital improvement
modeling may be, during ered and investment strategy
measurements and field opportunity.
equipment evaluation, it defined more closely. ■
observations. Test run data
analysis develops even more
detailed information: Does
a crude charge system have
9
8
hydraulic limitations? Is the PROCESS OPERABILITY
FCC wet gas compressor Process design has a major
running at maximum capa- impact on unit operability. Some
FEED

city? Are there other major process flow schemes are inher-
system limits? ently difficult to operate. Others STEAM
LOOP SEAL

Computer modeling are fundamentally easier to STEAM

is one tool that provides operate. An example of a low

answers. But accurate model- capital-cost, high return-on-
AGO PRODUCT
investment process design
modification is achieving direct REDUCED
CRUDE

control of wash oil flow rate.

CONCEPTUAL EQUIPMENT DESIGN DETAIL

PROCESS DESIGN — FCC main fractionator draw nozzle

Selecting a Low sizes are a seemingly small design

Capital Process detail. But, when nozzles are under-

Flow Scheme sized, pumparound heat removal

is reduced. Reduced heat removal

Brute force initiatives to lowers capacity and/or impairs

parallel or replace existing fractionation. Sometimes, low-cost

equipment can needlessly eat modifications, such as increased

up huge amounts of money nozzle sizes, can improve profits.

and time. A revamp carried Many times, sophisticated refinery-

out by Process Consulting wide computer studies will not find

Services’ four-stage method these problems.

can minimize such expend-

or eliminate operating of the air blower. Changing revamp costs and ensuring
itures by determining the
constraints. any one variable can drasti- long-term unit reliability.
best process flow scheme for
Successfully altering a cally alter overall unit perfor- Once the process flow scheme
the situation. Often, existing
process flow scheme requires mance. is set, no project management
equipment can be utilized
a thorough understanding of Developing a minimum- function will materially
more efficiently in a different
integrated unit performance. cost conceptual process decrease investment costs. ■
process flow scheme. Such
Each unit has unique equip- design demands not only a
equipment may be found to
ment interactions deter- firm grasp of chemical engi-
have been operating at less-
mined by the particular pro- neering fundamentals but
than-capacity or its perfor-
cess design. In crude units, extensive experience with
mance has been limited by
for example, heat integra- refinery units, different types
the process flow scheme.
tion, heater outlet tempera- of equipment, as well as hard
Rigorous, detailed models
ture, and distillation column practical knowledge of what
that have been calibrated
heat and material balance are works and what does not
with test run data, show the
all tightly linked. In an FCC, work in the real world.
limits of the existing process
reactor to wet gas compres- Conceptual design, in
equipment and support eval-
sor pressure profile materi- short, is the single most
uation of new process flow MEASURING
ally affects the performance important step in minimizing
schemes that can circumvent PERFORMANCE
Field measure-
VENT GAS
OFF GAS
ments eliminate
WIDE OPEN
guesswork and
assumptions.
EJECTOR
NAPH SYSTEM VACUUM
CRUDE COLUMN
ATMOSPHERIC
COLUMN STEAM
KEROSENE
LVGO
DESALTER
STEAM
LT DIESEL
HVGO
STEAM
THEORY VERSUS PRACTICE
FLASH
DRUM HVY DIESEL
Constructability must be considered during the
STEAM
STEAM process engineering of a revamp, not during
QUENCH

FUEL
GAS
VTB
detailed design. Equipment location, pipe rack
FUEL GAS
congestion, and new piping lengths are a few
AGO

of these considerations.
 PROCESS CONSULTING SERVICES

TYPICAL REVAMPS PERFORMED BY PROCESS CONSULTING SERVICES

FCC PRODUCT RECOVERY — USA REVAMP CONSTRUCTABILITY

The main fractionator-through-gas plant battery limits capac- Conceptual process design

ity was increased from 43 to 55 MBPD. Work included a evaluates alternate process

process design package with all major equipment design flow schemes. “Minimum cost”

specifications. The process flow scheme was modified to meet flow schemes are established

processing objectives while minimizing capital investment. based on total installed cost.
Constructability must be consid-
CRUDE/VACUUM UNIT — AUSTRALIA ered in conceptual process
Crude unit capacity was increased from 280 to 350 m3/hr.
designs.
Middle distillate yield and throughput increased without
additional heater firing. Process flow scheme and major
equipment were modified to improve jet fuel flash point,
diesel cold flow properties, and total middle distillate yield.
The existing heat exchanger network was modified to better residue was reduced. The process flow scheme was modified
utilize surface area. to improve overall unit operability and reliability. A process
design package included pump, distillation, and exchanger
DELAYED COKER GAS PLANT — USA
equipment modifications.
Gas plant modifications included heat integration, distillation
equipment design, a modified debutanizer condenser system, SATURATE GAS PLANT — USA
and deethanizer water removal to increase reliability and Gas plant capacity was increased, fractionation was im-
stabilize operability during drum switches. proved, and C3+ recovery was increased. Work included a
process design package with all major equipment design
FCC PRODUCT RECOVERY — SOUTH AFRICA
specifications.
Main fractionator-through-gas plant battery limits capacity
is being increased from 130 to 170 m3/hr. Work included CRUDE/VACUUM UNIT — SOUTH AFRICA
a process design package with all major equipment design Crude unit capacity is being increased from 580 to 720 m3/hr.
specifications. The process flow scheme was modified to meet Process flow scheme modifications included preflash crude
processing objectives while minimizing capital investment. and preflash vacuum column additions to permit reuse of
existing atmospheric and vacuum columns. Middle distillate
CRUDE UNIT — LATIN AMERICA
yield and throughput increased without additional heater
Total middle distillate yield was increased while diesel cold
firing. ■
flow properties were improved, and marine diesel carbon

OFF GAS VENT GAS

NEW
DRUM

RELIABLE LOW- NAPH NAPH

EJECTOR
CAPITAL REVAMPS SYSTEM
ATMOSPHERIC
CRUDE COLUMN
Existing unit and equipment STEAM
KEROSENE
LVGO
performance must be measured DESALTER 1

9 STEAM
VACUUM
and analyzed. Computer models LT DIESEL COLUMN HVGO
NEW
FLASH COLUMN STEAM
are important tools. However, DRUM HVY DIESEL

without accurate data on what

STEAM STEAM
is really happening on the unit, QUENCH

FUEL VTB
computer model results may be GAS
FUEL GAS

wrong and may result in making AGO

changes that do not work.

 PROCESS CONSULTING SERVICES

PROCESS DESIGN REVAMPING REQUIRES

PACKAGE CREATIVITY AND EXPERIENCE

Finding low-capital revamp

The process design package opportunities is an iterative

specifies, in detail, what process. Unit performance

equipment changes are depends on equipment inter-

necessary for the revamp to action. Changing the process

be implemented successfully. flow scheme can circumvent

All equipment modifications bottlenecks. There is no standard

are specified on the basis of crude unit or standard set of

actual, measured equipment system bottlenecks. Each case

performance. is unique and the best solution

The final report from will be one of a kind.

Process Consulting Services

provides heat and material
balances, yield summaries, MAN-HOURS OR RESULTS?
flow sheet simulation models,
process flow diagrams, lay- Operating limitations and problems in a refinery can cause
out drawings covering modi- millions of dollars in losses and needless expenses. What
fications to heaters, tower is needed is a quick and certain fix. That is why, with the HOUSTON OFFICE:
internals, vessels, and all sole exception of field troubleshooting and field installation 3400 Bissonnet, Suite 130
other needed equipment, and inspection, Process Consulting Services works only on a Houston, Texas 77005
equipment lists and specifi- lump-sum basis. We focus on results, not billable man-hours Tel: (713) 665-7046
cations. Vendor claims are that run up engineering costs and excuses and result only Fax: (713) 665-7246
evaluated against the record in patching, not solving problems. Cost-plus contracts are
of proven field experience. notorious in this respect. The hours devoted to any given DALLAS OFFICE:
A work schedule and job by Process Consulting Services engineers may be long or 1901 Central Drive, Suite 811
estimate of costs complete short, but what you pay for is not the expenditure of hours Bedford, Texas 76021
the report. The information — it is the final, satisfactory, and guaranteed completion Tel: (817) 358-1566
furnished allows engineer- of a necessary task. ■ Fax: (817) 358-2040
ing, operating, and mainte-
nance departments to make E-MAIL:
a thorough review and select info@revamps.com
detail engineering contrac-
tors, equipment vendors, A workable process WEBSITE:
and construction contractors. www.revamps.com
design based on
A workable process design
based on actual field exper- field experience is ©1999
ience, not on suppositions Process Consulting Services
and hopes, is the end the end product. All rights reserved.
product. ■ Printed in the U.S.A.

Volume 1, Issue 4
Volume 2, Issue 1

VACUUM UNIT COKING

■ FIRED HEATER
■ VACUUM COLUMN

Deepcut vacuum units operate at high temperatures.

Meeting profitability objectives depends on unit
reliability. Meeting 4-6 year runs has been a problem.

Revamping vacuum units to increase vacuum gas oil

yield can improve refinery profits. Profits depend on
both short-term yields and long-term reliability. Operating
temperatures can approach or exceed 800º F (427º C).
Shutdowns due to coking are common on deepcut
vacuum units. Coking is caused by both temperature
and oil residence time. Temperature is often considered
the primary cause of coking. However, oil residence time
must be controlled to reduce the rate of coking. Meeting
a 4-6 year run on deepcut vacuum units requires oil
residence time control in the heater and in the column.
Failure to control residence time results in unscheduled
shutdowns to remove the coke. ■

TECHNICAL PAPERS

“Heat flux imbalances in fired heaters cause operating

problems,” Martin, G. R., Hydrocarbon Processing,

May 1998, pgs. 103-109.

“Troubleshooting vacuum unit revamps,” Golden, S. W.,

WASH ZONE COKING
Wash zones coke because the Petroleum Technology Quarterly, Summer 1998, pgs. 107-113.
packing dries out. “Improved Flow Sheet Topology for Vacuum Distillation

Simulation,” Golden, S. W., et. al, Canadian Chemical

Engineering Conference, October 2-5, 1994.

 PROCESS CONSULTING SERVICES

FIRED HEATER OIL RESIDENCE TIME CONTROL

VACUUM VACUUM
Coil steam injection is essential to
UNIT UNIT
Monitoring vacuum heater FEED FEED minimize oil residence time. Other
operation is the key to meeting factors are tube layout, tube size,
run length objectives. HEATER HEATER number of passes, etc. Fundamen-
PASS PASS
Historically, average radiant OUTLETS OUTLETS
tal heater design considerations
section heat flux has been STEAM STEAM
determine the rate of coking. High
STEAM STEAM
used to monitor operating
rates of coking cause hotspots.
severity. Typically 10,000
btu/hr-ft2 (31.7 kw/m2) is
the upper limit. However,
average heat flux does not
represent localized conditions.
Some heaters coke at 9,000 STEAM STEAM

btu/hr-ft2 (28.5 kw/m2) or

less, while others operate at
11,000 btu/hr-ft2 (34.9 kw/
m2) or higher. FIRED HEATER OPERATION coked in less than 18 months.
The three factors that AND DESIGN PRINCIPLES: Low wash oil flow rate is
determine coking are oil • Equal heat flux in each pass the primary cause of coking.
residence time, localized heat • High mass velocity at the The middle of the packed
flux, and oil mass flux rates. inlet to the radiant section bed cokes because it dries
HOUSTON OFFICE:
Low radiant section oil • Coil steam injection to out. In fact, the bottom and
3400 Bissonnet, Suite 130
residence time is less than control oil residence time the top of the packed beds
• Good burner operation Houston, Texas 77005
10 seconds. However, dry are clean. There is no coke!
Tel: (713) 665-7046
heaters can operate with Several refineries have shut
more than 60 seconds VACUUM COLUMN down for turnaround and Fax: (713) 665-7246

residence time. Minimizing failed to check the packing.

residence time is essential. Deepcut vacuum unit After start-up, wash section DALLAS OFFICE:
Localized heat flux is revamps increase HVGO pressure drop of 20-30 1901 Central Drive, Suite 811
EQUIPMENT determined by burner layout product yield. These mmHg (2.7-4.0 kPa) were Bedford, Texas 76021
DESIGN and design and coil layout. revamps also increase measured. These units had to Tel: (817) 358-1566
For instance, minimizing operating temperature and be shutdown to replace the
Equipment design Fax: (817) 358-2040
excess air maximizes vaporize heavier gas oil. packing. ■
requires low oil
localized heat flux. The last HVGO product microcarbon
residence time. E-MAIL:
major factor is oil mass flux and metals content increase VACUUM COLUMN
rate. Higher mass flux rate info@revamps.com
as yield increases. Often OPERATION AND DESIGN
reduces the rate of coking. vacuum column wash PRINCIPLES:
Ideally, mass flux at the section efficiency must be • Adequate wash oil WEBSITE:
radiant section inlet should increased to meet HVGO flow rate www.revamps.com
be 450 lb/sec-ft2 (2200 kg/ product specifications. • Minimum packing
sec-m2). Ultimately peak oil Higher operating tempera- efficiency to meet HVGO ©1999
film temperature and oil ture and higher wash zone product quality Process Consulting Services
residence time determine efficiencies have dramati- • Good overflash collector
All rights reserved.
the rate of coke formation. cally increased the incidence vapor distribution
Printed in the U.S.A.
Operating changes that of wash zone coking. • Good vapor horn design
decrease peak oil film temp- Approximately 75% of • Low residence time
erature and oil residence time deepcut vacuum columns stripping trays
reduce the rate of coking. ■ have coked. Many have

Volume 2, Issue 1
 Volume 1, Issue 6

FCC GASOLINE FRACTIONATION

■ GASOLINE FRACTIONATION AND SULFUR DISTRIBUTION
■ HEAT INTEGRATION
■ GASOLINE ENDPOINT REDUCTION

FCC gasoline must be desulfurized and the endpoint

REDUCED CONDENSER reduced to meet Federal Tier II regulations. Economics
TEMPERATURE
will require gasoline fractionation prior to treating.
SALT
FORMATION
FCC unit product recovery section operation and
design will be changed by the Federal Tier II regula-
MORE
EFFICIENT USE tions. Reduced gasoline endpoint and gasoline fraction-
OF LOW
TEMPERATURE
NEW HEAT
ation requirements will lower main fractionator top
STRIPPER
temperature, increase the plugging potential on the
HEAVY MORE
NAPHTHA EFFICIENT USE
main fractionator internals, and require a large amount
PRODUCT OF LOW
TEMPERATURE of high temperature heat to run the gasoline splitter
HEAT
reboiler. Gasoline fractionation requires either high-pres-
sure steam or a fired heater for reboiler heat. A 50,000
GASOLINE
SPLITTER
LCO bpd (7,950 m3/day) FCC will require approximately
PRODUCT
70,000 lb/hr (31,800 kg/hr) of 600 psig (41.3 barg) steam
or 70 mmbtu/hr (17.65 mmkcal/hr) of fuel gas to a
600 PSIG
STEAM
FEED
(GASOLINE
fired heater for gasoline fractionation. Additional fired
SPLITTER)
heater duty or high-pressure steam generation from a
DECANT
PRODUCT
boiler increases NOx emissions.
FCC unit energy efficiency improvements will be
Minimizing Capital driven by the need to maximize high-pressure steam
Gasoline fractionation will generation. Current FCC equipment design and heat
be required to minimize integration will determine the most economical solution.
octane-barrel losses.
Dealing with the salt in the top of the fractionator will
Low-capital revamps will
require process and equipment changes. Minimizing
require creative solutions.
capital expenditure while implementing reliable modifi-
cations will be a challenge. ■
 PROCESS CONSULTING SERVICES

GASOLINE Gasoline Fractionation

FRACTIONATION Light, medium, and heavy cat

AND SULFUR naphtha contain different

amounts and species of sulfur.

Additional FCC gasoline Producing heavy naphtha from

fractionation will be the main fractionator is possible
required for the various in some units. In others, it’s
sulfur removal processes. not! Knowing what’s practical is
Currently, less than 50 % of important. Some changes will
refiners have gasoline split-
reduce reliability. Reducing unit
ters and many of these split-
reliability lowers profits.
ters do a poor job of frac-
tionation. The specific treat-
ing process will determine
fractionation requirements.
Fractionation will require
high temperature reboiler
utilization. Typically 100-250 in the main fractionator. At
heat, a high temperature
psig (6.9-17.3 barg) steam is some temperature, ammoni-
heat source, and additional
generated from the slurry um chloride and other salts
equipment. Existing equip-
pumparound. These steam begin to plug the internals
ment potential will deter-
temperatures aren't high in the fractionator. The frac-
mine the low-capital route. HOUSTON OFFICE:
enough for the gasoline tionator floods and gasoline
In many cases, existing 3400 Bissonnet, Suite 130
splitter reboiler. Improving endpoint increases. Salt for-
equipment modifications
high temperature heat use, mation depends on a num- Houston, Texas 77005
can be made that reduce
while balancing ber of variables, however, Tel: (713) 665-7046
new equipment costs. There
gasoline/LCO fractionation it is formed in some units Fax: (713) 665-7246
is no one best solution and
requirements is essential. at temperatures between
each case will require inno-
Low-capital FCC energy 235-250°F (112-121°C).
vative thinking to minimize DALLAS OFFICE:
efficiency changes require Once the salts are formed;
capital spending. ■ 1901 Central Drive, Suite 811
an understanding of main they must be removed
Bedford, Texas 76021
fractionator design and fun- either by on-line water
Tel: (817) 358-1566
HEAT INTEGRATION damental low temperature washing or a shutdown.
energy use in the gas plant. Reducing gasoline endpoint Fax: (817) 358-2040

It is practical heat integration will require operating

What changes Finding the heat for
schemes and distillation and/or equipment changes E-MAIL:
are required? gasoline fractionation, with-
know-how, not theoretical to deal with the salts. ■ info@revamps.com
out increasing refinery fuel
pinch analysis that solve
gas consumption and emis-
this problem and minimize SOLUTION WEBSITE:
sions, will be a challenge.
capital spending. ■
One possible route will be www.revamps.com

changes to the main frac- Each refiner's solution

GASOLINE ENDPOINT
tionator/gas plant heat inte- will need to be customized . © 2000
gration scheme. Many FCC
REDUCTION Supplying off-the-shelf solu- Process Consulting Services
unit main fractionators have tions will cost more money. All rights reserved.
two or three pumparounds. Federal regulations are Process Consulting Services
Printed in the U.S.A.
The available temperature also forcing refiners to can evaluate your specific
levels from two or three reduce gasoline endpoint. needs, find innovative solu-
pumapounds simply do Lowering gasoline endpoint tions, and minimize capital
not permit effective energy reduces the top temperature spending. ■

Volume 1, Issue 6
 Reprinted from: June 2000 issue, p. 45-56. Used with permission.

Revamping crude unit increases

reliability and operability
By improving the operation of pan attributes the success of this revamp to the relia-
bility gains for nonrotating equipment.
nonrotating equipment, a refiner
minimizes losses and raises Unit description. Fig. 1 is a process flow diagram of
process profitability Refpan’s crude unit before the revamp. The refinery pro-
cesses blends of Oriente, Arabian Heavy, Cano Limon,
Leona, Maya and Isthmus crude oils. During the unit
G. R. Martin, Process Consulting Services Inc., performance test, the atmospheric bottoms (ATB) yield
Bedford, Texas; E. Luque, Refinería de Panamá, was 56% of crude. The column’s lowest side cut is a heavy
S.A., Texaco, Inc., Colon, Panamá; and diesel product (HDO) used for power-company gas tur-
R. Rodriquez, Texaco, Inc., Bellaire, Texas bine and marine diesel sale. Light diesel (LDO) product
is used for road diesel and power generation. The HDO
mproving reliability increases profitability. When product contained 75% LDO boiling range material. ATB

I refinery margins are low, equipment reliability, or

lack of, can be the difference between making a
profit or losing money. Maintenance and rotating
is feed to the visbreaker unit, asphalt unit, and fuel-oil
blendstock. Twenty percent of the ATB was recoverable
LDO and HDO boiling-range material.
equipment engineers are aggressively improving Optimal crude unit fractionation depends on the
rotating equipment reliability. Rotating equipment refinery configuration and product market. Refpan’s
reliability programs are well established and have light-diesel product has higher value than heavy diesel.
been successful in many cases. Yet, the same attention The local market can absorb all the light diesel pro-
is not applied to nonrotating equipment. duced. HDO and LDO boiling range material in the
ATB are downgraded to fuel oil. The
Yardsticks. Rotating equipment reli- heavy-diesel product market is lim-
ability programs have successfully ited by local marine diesel sales.
increased the mean time between fail- Therefore, LDO product yield and
ures (MTBF) for rotating equipment. fractionation should be maximized.
MTBF is a yardstick for rotating Underperforming HDO boiling range material in the
equipment reliability; it is a measure ATB should be controlled based on
of time between failures and is easy marine diesel sales.
to monitor. But, how do you measure
equipment can
the reliability of a distillation system Minimum capital-cost revamp.
or a preheat train? Reliability— lead to lower unit Revamping an existing process unit is
applied to nonrotating equipment—is a four-stage approach. Process engi-
a measure of its performance relative neering identifies processing options
to its intended design performance. If
capacity, reduce from field-survey results, and ulti-
equipment does not perform as well mately captures those opportunities
as design, then it is unreliable. This product yields and with operating, process and equip-
reduces profits. ment changes. Using a four-stage
Refinería de Panamá (Refpan)
revamped its crude unit in the fall of
generate poor approach progressively moves the
refiner toward a funded revamp that
1996 to increase middle-distillate will produce real results while con-
yield and improve crude-oil process- product quality—all trolling engineering costs. The four
ing flexibility. Ultimately, the revamp engineering stages are:
increased middle-distillate yield by
factors will lower • Stage 1: Feasibility—f ield
10% volume on crude, improved both survey
marine- and road-diesel quality, and • Stage 2: Unit benchmarking—
enhanced crude-oil processing flexi- profitability. comprehensive performance test
bility. This project had a simple pay- • Stage 3: Conceptual design—
out of less than three months. Ref- eliminating bottlenecks
HYDROCARBON PROCESSING / JUNE 2000
Fig. 1. Pre-revamp crude unit PFD.

• Stage 4: Process design package—major equip- when evaluating equipment. A thorough understand-
ment design. ing of the process, as well as equipment operation and
Feasibility and benchmarking identify reliability design, is essential when identifying the root cause of
improvement areas. underperforming equipment. Wet feed, high liquid
level and pressure surges are examples of process-spe-
Determine reliability—unit benchmarking. Rotat- cific conditions that can impact a unit’s profitability.
ing equipment reliability programs focus on identify- A minimum capital-cost revamp must first measure
ing compressors and pumps with repetitive failures, equipment performance.1 This identifies unreliable
determining the root cause of the failure and elimi- process and equipment designs. Successful revamps
nating it. This same process has not been rigorously exploit the difference between actual and potential
applied to other process equipment and equipment sys- equipment performance to minimize capital expendi-
tems. During unit benchmarking, nonrotating equip- tures. Otherwise, more capital is spent to achieve
ment reliability is determined. revamp objectives. In some cases, the real unit limits
Nonrotating equipment reliability measures actual are never identified. These revamps fail to achieve an
equipment performance compared to its intended acceptable return-on-investment.
design performance. If equipment is not functioning as Unit benchmarking—comprehensive perfor-
intended, then it is not reliable. Underperforming mance test. Unit benchmarking establishes actual per-
equipment can lead to lower unit capacity, reduce prod- formance. Benchmarking is an expensive, time consuming
uct yields and generate poor product quality—all factors task. Often, this stage is skipped due to a perception that
will lower profitability. computer models alone can do this job. Field measure-
Identify process and equipment underperfor- ments are an integral part of benchmarking. Fieldwork
mance. Revamp engineers must identify the root cause consists of measuring temperature, pressure and compo-
of underperforming equipment before it can be fixed. sition profiles. These measurements are then used to cal-
While the fix for a reliability problem is often straight ibrate process and equipment models, and they can be used
forward, diagnosing the root cause can be difficult and directly to infer equipment conditions.2 For instance, at
time consuming. Reliability problems stem from pro- Refpan, field measurements confirmed that the wash- and
cess and mechanical design flaws. Potential design stripping-section trays were damaged. Fig. 2 shows the
flaws are numerous, and attention to detail is essential field-measured pressure drop. Trays require pressure drop
HYDROCARBON PROCESSING / JUNE 2000
Fig. 2. Pressure survey of the crude column. Fig. 3. Crude-column stripping and wash section.

to work properly. When there is no pressure drop across a not designed properly. These stresses are due either
tray, then it is either damaged or poorly designed. The Ref- to normal startup conditions or inadequate operat-
pan crude-unit benchmarking identified that these areas ing procedures. Whatever the case, column areas that
were affecting reliability and profitability: are susceptible to damage must be mechanically
• Practical process considerations designed to handle the startup, shutdown and abnor-
• Crude-column fractionation mal conditions.
• Equipment design. Wet stripping steam and high liquid level can dam-
age trays. Wet stripping steam introduces water to the
Process design reliability—root-cause analysis. bottom of the crude column. The water contacts hot oil
Process design does impact unit profitability. The and expands rapidly. As the water vaporizes into steam,
revamp conceptual design establishes a reliable pro- a pressure surge is created that exceeds the column
cess flow scheme, and the design package provides internal’s mechanical strength. High liquid level also
essential equipment details. While it is always possible damages crude-column internals due to the energy of
to make operating changes to minimize the effects of the liquid caused by steam or transfer-line vapor veloc-
poor process and equipment design, design f laws ity. While wet stripping steam and high liquid level can
inevitably affect profits. knock out trays during normal operation, they are more
Crude-unit profitability is driven by product yields. likely to cause damage during startup.
Crude-unit product yields are dependent on heat input, The stripping-section and wash-section tray (Fig. 3)
heat removal, operating pressure and fractionation. performance affects LDO and HDO product yields and
Heat input into crude charge consists of crude- unit profitability. The following symptom and root-
exchanger-network heat recovery and fired-heater duty. cause analysis relationship was valid:
Pumparound and condenser equipment design and
operation determine heat removal. Crude-unit operat- Symptom: Lower than expected LDO and
ing pressure is largely determined by the condensing HDO product yields
system capacity, fouling and corrosion. Root cause: Mechanical stress encountered
Fractionation depends on the process design and during startup, shutdown and upset conditions.
equipment performance. Problems in any of these areas
can affect unit capacity and product yields and qual- Damaged stripping trays alone resulted in a loss of
ity. Improving crude-unit profitability requires reliable 2,200 bpd of heavy- and light-diesel product to fuel oil.
process and equipment design. Stripping- and wash-section tray damage is very com-
Practical process considerations. Revamps solely mon (Fig. 4). Standard mechanical design trays (0.2
based on computer models often fail to meet their profit psi uplift) had been installed in the wash and stripping
objectives. Not addressing practical considerations can sections. Standard strength trays work fine for stan-
result in unsuccessful and unreliable revamps.3,4 Some dard conditions. However, stripping- and wash-section
practical considerations are: trays are not always exposed to standard conditions.
• Startup and normal unit upsets They must be designed to stay intact during startup
• Corrosion and fouling and normal unit upsets. The stripping-section trays
• Operability. were replaced with a heavy-duty tray design.
Refpan’s original crude-unit design highlights some Crude-column overflash is the vaporized oil that
practical considerations that should be addressed in returns to the flash zone as liquid. Minimum column
revamp projects. overflash maximizes the HDO-product yield. Prior to
the revamp, Refpan controlled the wash-oil rate by
Startup and normal unit upsets. Crude-unit start- changing the HDO-product yield. These systems are
ups and normal unit upsets can cause severe mechan- difficult to operate and cause either low yield or poor
ical stress on the column internals.5 Startups can be HDO quality.6 The revamp used a total-draw tray with
particularly damaging to column internals if they are flow-controlled reflux to a packed bed. This design per-
HYDROCARBON PROCESSING / JUNE 2000
Fig. 4. Stripping and wash section tray damage. Fig. 6. Fouled overhead exchanger bundle.

The crude-column operating pressure increases as

the condensers foul. Large increases in the column oper-
ating pressure from start-of-run to end-of-run are a com-
mon cause of distillate-yield loss. The column operat-
ing pressure determines the amount of vapor generated
at the heater outlet temperature for any given crude-
oil mix and temperature. The heater outlet oil vapor-
ization strongly affects product yields and economics.
Refpan’s overhead condenser system uses three par-
allel and three series exchangers (Fig. 5). The first two
exchangers, on each parallel train, exchange heat
against cold crude oil from storage. The last exchanger
is a water-cooled exchanger with seawater as the cool-
ing water supply. In some cases, these designs work.
However, in other systems, the exchanger fouling and
corrosion is so severe that it reduces unit reliability
and actually reduces profits due to product yield losses.
Condenser fouling has been a chronic problem at
Refpan. Historically, the crude versus overhead-
exchanger fouling has reduced heat transfer, and thus,
Fig. 5. Crude-column overhead condenser system. loaded the downstream cooling water exchangers. Load-
ing the water-cooled exchangers caused severe water-
mited minimum overflash operation and maximum side fouling, further reducing condenser capacity. The
LDO and HDO yield. The wash-section mechanical overhead receiver temperature increases, which raises
design was upgraded. the compressor-gas load. Higher compressor-gas load
raises operating pressure and reduces LDO and HDO
Corrosion and fouling. Crude-column top trays and product yields.
overhead condensers are susceptible to high levels of Crude-unit overhead systems that exchange heat
general corrosion and under-deposit corrosion and foul- against crude commonly experience corrosion and foul-
ing.7 Thus, neutralizing and filming chemicals are used ing (Fig. 6). The root cause is insufficient water at the
to reduce corrosion rates and extend condenser life. point of salt deposition to dissolve the salt. The only
Corrosion of overhead condensers affects the condenser effective way to remove these salts is by injecting water
system performance, which affects the column operat- in front of the exchangers. Whenever column overhead
ing pressure. Column top-tray corrosion decreases frac- vapor is exchanged against crude oil, the purpose is
tionation between naphtha and jet fuel. Top-tray cor- heat recovery. However, injecting sufficient wash water
rosion is a function of reflux rate, temperature and to remove these salts lowers the temperature to the
water content. first exchanger by 70°F, which reduces the driving force
Reduced condenser heat removal increases column for heat exchange. Proper water washing of crude-col-
operating pressure and decreases light- and heavy- umn overhead systems is difficult.
diesel product yields. When light- and heavy-diesel The parallel exchangers installed at Refpan can be
yields decrease, profits are lowered. Root-cause analy- periodically isolated for cleaning. Cleaning overcomes
sis led to these factors: the inherent reliability problems with these condenser-
system designs. Living with the consequences of this
Symptom: High column operating pres- system design is purely a business decision. The bene-
sure/low LDO/HDO product yield fit of cleaning the exchangers is lower, average operat-
Root cause: Overhead condenser system ing pressure. Reliable operation at lower pressure
design. increases distillate yield and improves unit profitability.
HYDROCARBON PROCESSING / JUNE 2000
Fig. 7. LDO side-stripper draw.

Operability—normal operation and startup. Pro-

cess flow scheme and equipment design affect oper-
ability. Unit operability is a measure of the unit’s capa-
bility to function during startup, through normal unit
disturbances, and the inevitable bounces associated
with crude switches. Good unit operability comes from Fig. 8. Crude-column pumparound/product draws.
practical know-how of what works and what does not.
Some units are difficult to startup, while others are
designed with startup and operability factored into the shift the heat balance to maintain stable operation;
process design. therefore, maximum LDO yield could never be
During startup, being able to inventory the achieved. Fig. 7 shows a previous modification to the
pumparound-draw trays and maintain liquid on the LDO-stripper and product-draw system; a flow con-
tray is essential to dryout (remove water) and reach troller was installed to feed the stripper. This design
stable operation. Long, drawn-out startups lead to unit was used to overcome the dry out problem. LDO prod-
upsets and equipment damage. Process and equipment uct was drawn to storage on level control. This con-
designs that make startup difficult and normal oper- trol system treated the symptom, but not the funda-
ation sensitive to routine disturbances leads to equip- mental problem.
ment damage.
When the pumparounds and products are with- Crude-column fractionation. Heat balance and
drawn from different trays, startup and normal oper- fractionation efficiency control the crude-column frac-
ation is more difficult. During startup, the feed tem- tionation. Heat removal is controlled by the condenser
perature to the column is slowly increased to remove and two pumparounds. The two pumparounds were
water. The column heat input is not constant; there- designed to exchange heat against crude oil. Both
fore, heat removal from the pumparounds must be location and the design of the distillation equipment
continuously adjusted. During normal operation, heat- determine operating flexibility and reliability. The
balance changes associated with crude switches can type and design of the internals control fractionation
effectively dry-out the product draws above the efficiency.
pumparound because of high heat removal. Once the Pumparound locations—column heat balance.
internal liquid flow from the draw tray reaches zero, The LDO-product end point control was poor because
then the product side-stripper loses level. Thus, the reflux below the LDO draw could not be maintained
stripper-bottoms pump will cavitate. throughout crude switches, charge rate changes or nor-
The Refpan crude-column pumparound and product mal unit upsets. The diesel-pumparound pump and
draw-tray locations caused reliability problems. The exchanger surface area sizes were large. Diesel-
following symptoms and root cause were identified: pumparound heat removal reduced the reflux from the
LDO-product draw tray. Reduced reflux downgraded
Symptom: Loss of LDO stripper level light diesel to heavy diesel. Ultimately, the diesel
Root cause: Pumparound flow scheme and pumparound was taken out of service because it
column heat balance. reduced LDO yield. This reduced total column heat
removal capacity, which limited the fired-heater duty,
With the diesel pumparound located between the and lowered the light- and heavy-diesel product yields.
light- and heavy-diesel products, it was possible to Root-cause analysis shows how incorrect pump-
dryout the LDO-stripper draw tray. Varying crude around location will result in operability problems from
slates and charge rates requires constantly changing poor control of internal reflux. The following cause/effect
heat removal in the diesel pumparound; otherwise, relationship was occurring with this existing design:
the LDO-stripper draw can dryout. Once this occurs,
the LDO stripper loses level, and the stripper-bottoms Symptom: Poor LDO quality and yield control
pump will cavitate. The operators had to constantly Root cause: Incorrect pumparound location.
HYDROCARBON PROCESSING / JUNE 2000
Fig. 9. Pumparound draw nozzle and piping.

Fig. 8 shows the original atmospheric-column flow

scheme. The pumparounds were located between the
product draws. When pumparounds are located in the
middle of fractionating sections, it is usually because Fig. 10. Crude-column draw tray.
the designer is trying to increase the pumparound-draw
temperature and provide higher exchanger LMTD val- or small draw-off nozzle sizes cause pump NPSH prob-
ues. This design can improve energy efficiency; how- lems.8 Pumparound or product-draw piping should not
ever, it also increases the system’s complexity and be reduced until 6–10 ft below the draw-nozzle eleva-
makes operations more difficult. tion. This permits pressure buildup of static head. If
Pumparound location has significant impact on col- the pump suction and draw nozzle are sized properly,
umn internal liquid and vapor rates. Increasing diesel the line will always be liquid full to the draw-off loca-
pumparound heat removal at constant heat input low- tion. An undersized draw nozzle can lead to a pump
ers the vapor rate leaving the pumparound section. NPSH problem.
This, in turn, reduces the internal reflux flowing from Review of the system showed that the suction line
the LDO-product draw tray. Reduced reflux increases to the pump had been designed improperly. The draw
the LDO-product endpoint. When the internal reflux nozzle was a 4-in. nozzle that increased in line size to an
from the LDO-draw tray reaches zero, there will be 8-in. line (Fig. 9). Ultimately, the hydraulic limit in the
insufficient liquid flowing to the LDO stripper to main- system was the 4-in. line. As the operators increased
tain stable operation. the pumparound flowrate to improve heat removal, the
liquid level in the draw-off piping decreased, and the
Equipment design. Ultimately, equipment design available NPSH dropped below the required NPSH.
determines how the unit operates. Computer models Thus, the pump would cavitate.
represent theoretical ideals; yet, they often do not reflect Column operating stability was poor due to erratic
the actual design or performance of the equipment. heat removal. Low heat removal causes more vapor to
Column draw-nozzle size. Field data and obser- flow up the column, which increases the condenser load
vations are used to identify actual equipment perfor- and raises the overhead receiver temperature. High
mance. During the performance test, product draw overhead temperature increases the gas load to the off-
rates and maintaining stable pumparound system per- gas compressor. Higher compressor load increases com-
formance were two major areas limiting this unit’s per- pressor-suction pressure. Higher column operating
formance. These problems led to this cause and effect pressure increases the atmospheric tower bottoms
relationship: (ATB) yield. The high ATB product yield reduced heavy-
and light-diesel yield. Pump problems are not appar-
Symptom: Yield limitations/pump cavitation ent unless the pump suction is measured. Basic equip-
Root cause: Column draw nozzle undersized. ment design must be correct; otherwise significant prof-
itability losses will result from unreliable equipment.
Pumparound and side-stripper hydraulic limitations Column draw-tray design. The crude column
are relatively common problems. Measuring the pres- draw-tray designs did affect operability. Refpan’s design
sure at either the pumparound-pump suction or the side- for the stripper-feed and pumparound-draw streams
stripper level control valve identifies these limitations. used sumps from active trays. These problems led to
These lines should be full of liquid. Therefore, the pres- this cause-and-effect relationship:
sure should equal column pressure at the tray where
the liquid draw is located plus the static head of liquid at Symptom: Yield limitations/pump cavitation
the point where the pressure is measured. If the line is Root cause: Column draw-tray design.
not full, then the pressure will be lower than expected.
Refpan’s diesel pumparound had experienced chronic Fig. 10 shows a typical valve tray with a sump
problems. Generally, column internal draw-tray design designed to draw liquid either to a side-stripper or a
HYDROCARBON PROCESSING / JUNE 2000
Fig. 11. Revamped crude unit simplified PFD.

pumparound pump. Withdrawing liquid from the tray to remove air. Part of the steam is condensed and accu-
sump requires that the liquid crosses the tray—feed- mulates inside the column. Circulating the
ing the sump—and flows into the downcomer. For an pumparounds helps remove water from the exchang-
active tray-draw sump to work properly, the leakage ers, piping and column internals. Startup procedures
rate through the tray must be less than the internal call for pump switches to drain water from the idle
reflux rate to the tray below. Once the tray leakage pump and all low points throughout the pumparound
rate is higher than the internal reflux rate, the exter- system. Once the column is hot, any water that enters
nal draw from the column will not be full. the column will vaporize violently. The resulting pres-
Product- and pumparound-draw trays should use sure surge can damage the column internals.
a seal-welded collector tray wherever possible. This Fractionation efficiency—LDO/HDO section.
ensures that liquid entering the tray can be with- LDO-product yield and fractionation between LDO and
drawn. All valve, sieve or bubblecap trays leak. The HDO product should be maximized. Root-cause analy-
quantity of leakage is a function of tray design and sis shows that the existing LDO/HDO section had poor
the process-vapor rate to the tray. Trays are designed fractionating efficiency. Seven fractionating trays were
as a series of panels that fit together at a metal-to- getting less than one theoretical stage of efficiency. The
metal seal. Vapor flow through the tray deck’s sieve, following cause/effect relationship was occurring with
valve or bubblecap holes determines the dry-tray pres- the existing design:
sure drop. The hole area and type of hole set how much
leakage occurs through the trays. Understanding Symptom: Poor LDO quality and yield control
vapor flowrate variation through a tray deck helps Root cause: Low tray efficiency.
identify why—under certain conditions—pump oper-
ating and reliability problems are more common. The LDO/HDO fractionation section experienced
Startup conditions, pumparound location, heat bal- large changes in vapor and liquid rates from the top to
ance and specific tray equipment design all affect the bottom trays. The liquid rate decreased by 50% from
draw rate from the column. the top tray to the bottom. The vapor rate also increases
Pumparound draws from active trays significantly by 30% from the bottom tray to the top. A single-tray
increase the difficulty of starting up a unit. During design used throughout the LDO/HDO section will
startup, the fractionator must be purged with steam cause low efficiency or reduced capacity in the column.
HYDROCARBON PROCESSING / JUNE 2000
The revamped LDO/HDO section trays used two dif-
Gary R. Martin is a chemical engineer for
ferent tray designs to handle the process requirements. Process Consulting Services Inc., Bedford,
Often, small equipment design changes result in sig- Texas. His responsibilities include revamps
nificant profit improvements; details are important. and troubleshooting of refinery processes. He
specializes in improving refining profitability by
troubleshooting, optimization and revamping
Increased profitability—improved reliability. of refinery units. He previously worked as a
Improving crude unit profitability with low capital-cost refinery process engineer and distillation sys-
revamps requires modifying only the required systems tem troubleshooter. He holds a BS degree in
and equipment. 9 –11 Revamping a unit should address chemical engineering from Oklahoma State
University. He is the author of more than 40 revamp and trou-
the practical process considerations that determine bleshooting technical papers.
reliability, operability and flexibility (Fig. 11). MTBF
should not only be a consideration for rotating equip- Elias Luque is the manager of the technical
ment, but for nonrotating equipment, as well. The unit services department for Refinería Panamá
S.A., a Texaco refinery located at Colón,
reliability, operability, flexibility and ultimately prof- Republic of Panamá. Mr. Luque has 33 years
itability of the modifications have been proven by three of experience in the refining industry in the
years of stable operation.  fields of analytical chemistry, process engi-
neering, refining operations, industrial safety,
LITERATURE CITED energy conservation, project management
1 Golden, S. W., “Minimum investment revamps: process design and achieving attractive and environmental protection. He holds a
economics,” 1997 NPRA Annual Meeting, March 16–18, 1997.
2 Martin, G. R., “Heat-flux imbalances in fired heaters cause operating problems,” Hydro-
BS degree in chemical engineering from the
carbon Processing, May 1998, pp. 103–109.
University of Florida and an MBA from the Universidad Santa
3 Barletta, T., “Practical consideration for crude unit revamps,” Petroleum Technology María La Antigua at Panamá.
Quarterly, Autumn 1998, pp. 93–103.
4 Barletta, T., “Revamping crude units,” Hydrocarbon Processing, February 1998, pp. 51–57.
5 Martin, G. R., S. W. Golden, and B. Fleming, “Distillation column inspection,” Third Inter-
Reynaldo A. Rodriguez is a process special-
national Conference & Exhibition on Improving Reliability in Petroleum Refineries and ist at Texaco Inc., General Engineering Depart-
Chemical Plants, Houston, Texas, Nov. 15–17, 1998. ment (GED) in Bellaire, Texas. His 20 years of
6 Golden, S. W., “ Prevent preflash drum foaming,” Hydrocarbon Processing, May 1997, pp.
industrial process engineering experience
141–153.
7 Strong, R., “Field history with low-salting crude unit neutralizers,” 1998 NPRA Annual includes over 10 years in petroleum refining.
Meeting, March 15–17, 1998. Prior to transferring to GED in 1998, he worked
8 Martin, G. R., “Refinery experience provides guidelines for centrifugal pump selection,”
for the Texaco Honduras Refinery and the Tex-
Oil & Gas Journal, March 11, 1996, pp. 80–86. aco Panama Refinery in Central America in
9 Golden, S. W., “Minimize capital investment for refinery revamps,” Hydrocarbon Processing,

January 1997, pp. 103–112. process engineering and refinery operations.

10 Martin, G. R., and B. E. Cheatham, “Keeping down the cost of revamp investment,” He holds a BS degree in chemical engineering
Petroleum Technology Quarterly, Summer 1999, pp. 99–107.
11 Barletta, A. F., C. Macfarlane, and M. Smith, “Crude unit revamp increases diesel yield,”
from the Honduras National University and an MS degree in chem-
Petroleum Technology Quarterly, Spring 2000, pp. 45–51.
ical engineering from Georgia Institute of Technology.

F/5M/7-2000 Article copyright © 2000 by Gulf Publishing Company. All rights reserved. Printed in USA.
 Technical Articles
Index

British HSE - Inspection - Non-Destructive Testing (NDT)

Corrosion and fouling in alky's by STRATCO
Cost of Corrosion
Delayed Coking Fundamentals
Frac tower diagram
Fractionation Efficiency Increase
Jaeger air stripping of VOC's
Jaeger Air stripping problems
Jaeger Process info
Jaeger scrubbing
Jaeger steam stripping
Kister-trouble free design of refinery fractionators
Linas Technology - Distillation of 21 century
Linas Technology - Distillation of 21 century-1
Linas Technology - Distillation of 21 century-2
Linas Technology - Distillation of 21 century-3
New life for cracker margins
OSHA PRESSURE VESSEL GUIDELINES
OSHA Technical Manual Petroleum Refining Processes
Revamping HDS Units
Sibneft - History of Oil in Russia
Succeed at Plant Debottlenecking
Sulzer Chemtech atmo-vac distillation—Atmo
Sulzer Chemtech Atmo-vacuum distillation
Sulzer Chemtech atmo-vac—Vacuum
Sulzer Chemtech Fluidized Catalytic Cracking Unit
Sulzer Chemtech Vacuum Distillation Unit ★
Technical Program Paper Detail – AIChE
Trouble free design of frac's Kister
Troubleshooting overhead ejectors
What caused tower malfunctions last 40 years by Kister
 HSE - Inspection / Non-Destructive Testing (NDT)

Inspection/Non Destructive Testing

This document provides detailed and specific guidance on inspection and Non-Destructive Testing
(NDT) in support of the Level 2 Criteria 5.2.1.3f, 5.2.1.6i, 5.2.1.11(63)e, 5.2.2.2 (79), 5.2.4.1(88)e,
5.2.4.3 and 5.2.4.4.

Related Technical Measures Documents are Training and Maintenance Procedures.

This document assists in the assessment and inspection of NDT applied on plant and how that
supports the continued safe operation of the plant.

1. Introduction: a description of NDT, what it can and cannot do and how it fits with safety
management.
2. Regulatory Requirements: how NDT meets the requirements of current regulations and a
description of a Written Scheme of Examination.
3. COMAH Safety Report: what information on NDT would be expected in the safety report.
4. HSE Follow Up COMAH Inspection: what the HSE Inspector should look for in the COMAH
follow up inspection of the site.
5. NDT Process & Management: a description of how dutyholders should initiate, specify and
apply NDT and how the results should be utilised. Also detail on how dutyholders should manage
NDT on a site.
6. Techniques & Capabilities: an overview of the common NDT techniques and the advantages
and limitations. Includes a description of common trade names.
7. Checklist for the Inspection of NDT: Aide Mémoire to assist follow-up HSE site inspection
8. Terminology and Current Trends: Glossary of terms and what they actually mean.
9. Case Studies: 2 case studies on the adequacy of NDT programmes to detect defects in pressure
vessels.
10. Sources of further, more detailed, information are given at the relevant position in the text and
other significant standards are listed.

1. Introduction

Pressure vessels, storage tanks and other safety critical components (including pipework and valves)
are designed to contain liquids, gases and solids such that a loss of containment does not occur. Leaks

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 HSE - Inspection / Non-Destructive Testing (NDT)

or the mechanical or structural failure of these items of equipment may result in a major accident on-site.

The presence of flaws in critical components may result in the integrity of such systems being
compromised and increase the likelihood of failure.

Non-Destructive Testing (NDT) is the application of measurement techniques in order to identify

damage and irregularities in materials. NDT often provides the only method of obtaining information
about the current 'health' of process plant.

If done well, NDT can provide useful information to assist in the management of plant safety. If
inappropriate NDT is applied or NDT is not applied correctly, then the results are likely to give a false
impression of the integrity and safety of the plant.

NDT is a measurement of a The distinction between what would be considered changes in

physical property or effect from material properties and what would be considered a defect is not
which the presence of damage or distinct. This can lead to NDT missing defects and also producing
irregularity can be inferred. It is not false calls i.e. a defect is reported when in fact the signal is not
a measurement of an absolute produced by a defect. Also, NDT is applied to a greater or lesser
parameter such as temperature or extent by human operators who introduce human error and
pressure. subjectivity into the process.

NDT is rarely 100% effective at detecting defects of concern. Like

all measurements, defect positioning and sizing measurements
with NDT techniques are subject to errors. As these techniques
are often a combination of separate measurements, these errors
can be significant.

NDT techniques fall into two ● techniques which only detect and size defects/damage
categories: present on the surface of a component;

● techniques which can detect and size defects/damage

embodied within a component.

A brief description of the common techniques applied to process

plant is given in Techniques. The basic NDT techniques have
changed little over the years but with improvements in technology
and the demand to maximise plant productivity new techniques
and variations on old ones have been developed, along with
various approaches to NDT. These are clarified in Terminology
and Current Trends below.

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 HSE - Inspection / Non-Destructive Testing (NDT)

The quality of the NDT applied to Extra steps are required in the development and application of the
a component cannot be easily test to provide confidence in its ability to identify the damage or
assessed by subsequently irregularities of concern. The Inspection Process and its proper
observing the component or the Management are discussed in more detail below.
results obtained.

NDT is a primary recovery Correct selection and application of an NDT technique can
mechanism for errors in design, provide confidence that a component or piece of plant does not
construction and operational contain defects of the type which the technique was capable of
activities. detecting.

When applied in a manufacturing environment it is used to

provide confidence that there are no defects of concern over a
certain size which may have been introduced by the
manufacturing process. In this case NDT is just one of a number
of quality control activities aimed at producing a component or
piece of plant to a particular specification.

In service NDT provides If such deterioration is detected then NDT can quantify the
confidence that the operation of damage and provide input to the justification for maintenance or
the plant is not causing monitoring actions.
deterioration in its integrity beyond
its design parameters.

Ad hoc NDT can be used to All techniques have strengths and weaknesses regarding the
check that unexpected damage types and parameters of the damage mechanism they can detect.
mechanisms are not occurring.
Either the ad hoc NDT needs to be targeted at a hypothetical
damage mechanism or the damage mechanism that can be
reported as not detected is defined by the capabilities of the
technique.

The types of defect / flaw and degradation that can be detected using NDT are summarised as:

● Planar defects - these include flaws such as fatigue cracks, lack of side-wall fusion in welds,
environmental assisted cracking such as hydrogen cracking and stress corrosion cracks; cold
shuts in castings etc;
● Laminations - these include flaws such as rolling and forging laminations, laminar inclusions and
de-laminations in composites;
● Voids and inclusions - these include flaws such as voids, slag and porosity in welds and voids
in castings and forgings;
● Wall thinning - through life wall loss due to corrosion and erosion;
● Corrosion pits - these are localised and deep areas of corrosion;
● Structural deformities such as dents, bulges and ovality.

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 HSE - Inspection / Non-Destructive Testing (NDT)

The application of NDT to support the manufacturing requirements or the continued operation of plant is
subject to certain Regulations.

There are some common misconceptions regarding NDT which are still prevalent in industry:

No defects found and reported All NDT techniques have strengths and weaknesses and no one
means that there are no defects technique is capable of detecting all types of defect nor is there
in the component. one which is 100% reliable.

A defect measured at 5 mm All measurements are subject to errors particularly if obtained

means that the defect actually is under site conditions.
5 mm.

If NDT reports defect growth or Identification of growth / no growth requires comparison of two
non-growth then this is actually measurements each with their own errors. The growth
the case. measurement is then subject to the combined errors.

100% Inspection Coverage. This does not necessarily mean that 100% of the component has
been inspected.

Hard copy results can't lie. Hard copy results are only as good as the techniques and data
used to produce them.

NDT to a national standard is National standards are only relevant to specific circumstances
always appropriate. and only include knowledge available at the time of development.
The NDT requirements need to be checked against the standard
to see if it is relevant to the particular situation.

2. Regulatory Requirements

The requirements of the various general regulations can be summarised as:

Pressure Systems Safety Regulations 2000 cover the operation of pressure systems.

Regulation 4 (2) states that plant should be "properly designed and properly constructed from suitable
materials so as to prevent danger."

Regulation 8 (1) requires the "owner or user must have a WSE for the periodic examination by a
competent person where:" according to Regulation 8 (1) (b) "a defect may give rise to danger."

WSE is a Written Scheme of Examination which specifies for each part of the pressure system the

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damage mechanism that may be expected, the examination interval and the method of examination.
Any NDT required will be specified on the WSE. Further information on Written Schemes is given on the
HSE Website.

The Pressure Equipment Regulations 1999 cover the manufacture of new pressure vessels in line
with the European Pressure Equipment Directive and apply to pressure equipment and assemblies with
a maximum allowable pressure PS greater than 0.5 bar. These regulations require:

● Pressure equipment must be designed and constructed so that all necessary examinations to
ensure safety can be carried out;
● Preparation of the component parts (e.g. forming and chamfering) must not give rise to defects or
cracks
● Permanent joints and adjacent zones must be free of any surface or internal defects detrimental
to the safety of the equipment.
● For pressure equipment, suitably qualified personnel must carry out non-destructive tests of
permanent joints.
● For pressure equipment of categories III and IV, the personnel must be approved by a third-party
organisation recognised by a Member State pursuant to Article 13.
● Pressure equipment must undergo a final inspection to assess - visually and by examination of
the accompanying documents - compliance with the requirements of the Directive. Tests carried
out during manufacture may be taken into account.

The old British Standard (BS 5500) for the manufacture of pressure vessels has been superseded by a
new European Standard BS EN 13445 but the design requirements that were in BS 5500 have been
kept as PD5500. This latter document sets defect Acceptance Criteria for the NDT applied at
manufacture. The acceptance criteria take into account the capabilities and limitations of the NDT
techniques so for Radiography it states "No cracks allowed" whilst for Ultrasonics it states conditions on
planar indications based on the height, length and amplitude.

The Control of Major Accident Hazards Regulations 1999 put a general duty on every Operator to
take all measures necessary to prevent major accidents and limit their consequences to persons and
the environment. The references to NDT that may occur in the safety report are described in COMAH
Safety Report.

NDT although not specifically mentioned in the above regulations has a role to play as part of the
Operator's demonstration in respect of mechanical integrity, that all necessary measures have been
taken. It provides confidence that plant is constructed to the required standard and is in good repair.
NDT can provide information to confirm or otherwise that unexpected damage is not occurring.

The Carriage of Dangerous Goods Regulations - Rail & Road. These regulations cover the transport
of dangerous goods by rail and road, including the examination and testing of the tanks used. The tanks
require to be examined and tested by the competent authority or approved person in accordance with
requirements approved and published in the Approved Tank Requirements. A certificate is required to
be produced which confirms the examination and test and also that the tank conforms to an approved
design and is suitable for the purpose for which it is intended.

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 HSE - Inspection / Non-Destructive Testing (NDT)

Management of Health & Safety at Work Regulations 1999. These require employers to make
suitable and sufficient assessment of the risks to employees and the public arising from the business
activities

The Provision and Use of Work Equipment Regulations 1998. These require the employer to ensure
that work equipment is so constructed or adapted as to be suitable for the purpose for which it is used
or provided and that work equipment is maintained in an efficient state, in efficient working order and
in good repair.

3. COMAH Safety Report

The information given on NDT in the safety report is likely to be in response to Criterion 5.2.4.3: the
report should show that systems are in place to ensure, for safety critical plant, that a competent person
examines systems at appropriate intervals. The information is likely to be general and details of the
NDT techniques would not be expected.

The safety report should provide evidence that Written Schemes of Examination are in place for
pressurised plant.

A safety report should declare whether examinations are performed by an in-house organisation or
bought in from a 3rd party. If the competent person is in-house then the report should show they are
independent from operations and have direct reporting to senior management.

Most of the information pertaining to the NDT applied and the use of the results in supporting the
integrity of the plant will need to be gathered during the HSE follow up COMAH inspection on-site.

4. HSE Follow Up COMAH Inspection

The HSE follow up COMAH inspection on site should be used to gather further information on the
following NDT activities:

4.1 Justification of Examination Intervals

● The interval between examinations may be based on guidelines offered by such organisations as
SAFed, CEOC and IoP.
● However, HSE interpretation of the legislation does not always agree with the advice given in
these guidelines. Also, care is required when interpreting these guidelines for specific situations.

Investigate the justification of the If standard guidelines have been used, look at the reasons for
examination intervals. selecting the chosen periodicity and if these are compatible with
the operating conditions of the plant.

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A risk based inspection (RBI) In this case damage mechanisms for each plant item should have
approach may have been used as been established and the item categorised according to risk.
an alternative to fixed interval
examinations.

The NDT should then be targeted This process can reduce the amount of NDT required from that
at the high/medium risk items with based on a fixed interval approach. However, it is then even more
a view to reducing the probability important that the correct NDT technique is used:
of failure and hence the risk.
● to look for the required damage mechanism,

● that the NDT is correctly applied,

● that the capability is understood and

● the results are fed back into the RBI process.

HSE's Best Practice for risk-based inspection can be used for

assessing the quality of the RBI process.

4.2 Management of NDT Process

NDT needs to be managed correctly to ensure that the theoretical capability of the technique is not
unduly impaired by incorrect or poor application. Companies should not only have procedures which
cover the management and application of the NDT but also evidence that they are being implemented.

Assess the plant owners' attitude If it is not given sufficient consideration then it is unlikely to be
to NDT. planned properly, good access is unlikely to be provided and
contractors are likely to be under undue pressure which will
prevent them performing the NDT properly. This is discussed
further under Management below.

Check whose quality system the Check that this actually happens.
NDT operators are supposed to be
applying the NDT under: the plant
owners' or the NDT companies'.

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The results of a Non Destructive Check that the requirements for the NDT are specified and that
Test are dependent on the type of the records provide sufficient information on what NDT technique
damage being sought and the was used and how it was applied. Is the technique likely to find
particular NDT technique applied. the defects of concern and is the capability of the technique
known? Further information on the various techniques is given
below.

The HSE's Best Practice Where NDT plays a key role in assuring the safety of the
documents (ultrasonics, surface component then additional steps should be taken to ensure that
techniques) give guidance on all the defects of concern are detected and that the NDT
assessing the role of the NDT and technique is applied correctly.
the effectiveness required of it in
reducing the risk of component
failure.

Finally the results of the NDT The presentation of the NDT results will often influence the
should be assessed and acted subsequent actions. Some reports will just state no defects found
upon by the plant owner. or that the NDT was acceptable i.e. there were no indications
observed above a certain acceptance criteria. Other computerised
techniques allow apparently detailed, colour plots to be produced
which create the impression of quality. Both these extremes can
disguise the fact that the NDT may have had limitations in defect
type detectable or that an insufficient sample area/volume may
have been examined.

The plant owner should be able to show how the assessment of

the NDT results has taken into account the limitations and errors
inherent in the technique applied.

These issues are covered by the questions in the Checklist designed to help in the HSE on-site
COMAH inspection.

Small companies are likely to buy in the competent person expertise and place reliance on the 3rd
party's expert judgement. The Operator should have the statutory records of inspection available, but
may not have immediate access to additional information about the examinations or the competence of
the 'competent person' organisation.

5. NDT Process & Management

5.1 NDT Process

The start of any NDT process is the identification of plant items which require NDT.

● For pressure systems this will be detailed in the Written Scheme of Examination.

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● For non-pressurised but hazardous plant this should be output from the systems which 'ensure
that safety critical plant and systems are examined at appropriate intervals by a competent
person'.

These two documents should also identify the damage mechanisms which could be expected to occur
in the plant item and hence should be detected, if present, by the NDT technique. As with any other
purchase or development (in line with ISO 9001) NDT should start with a specification of requirements.
For NDT this is a defect specification or description, which includes:

● A description of the damage mechanism - location, type, morphology, orientation;

● Whether the volume or surface requires NDT;
● The size of defect which needs to be detected and the sizing errors that can be tolerated.

NDT can be applied without The defect description is then defined by the capabilities of the
stating a particular defect to look technique applied and the plant item can only be passed clean of
for. defects detected by this technique.

Once the specification has been The NDT method should be specified in the Written Scheme of
produced then the appropriate Examination, if relevant, or documented elsewhere.
method and technique can be
selected.

All NDT should be applied under This is likely to be undertaken by the NDT company on behalf of
the control of a procedure which is the plant owner. The procedure, which may be supplemented by
produced and approved by a plant specific technique sheet, should be sufficiently detailed to
competent personnel (see define the technique to be applied.
Management).
The NDT technique can then be applied by a competent person
and the results reported. The report should highlight any
restrictions in the application of the technique and should list any
changes to the technique which were required by the particular
application.

Radiography is a popular NDT The results from large area NDT techniques such as corrosion
technique because the mapping, floor scanners are often presented as colourful
radiographic films provide a hard computerised plots. Although these visual outputs look
copy of the results. impressive, they do not show the limitations in the technique and
are not proof in themselves that the NDT was performed correctly.

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NDT is only able to lead to a If the result is no defects found, there may still be the need for
reduction in the probability of action taking into account the capability of the NDT technique and
failure if appropriate action is the nature of defects which may not have been found.
taken in response to the results
obtained. Standards governing Manufacturing NDT often specify
acceptance criteria in terms of the NDT measurement i.e. no
indication longer than … or no signal with amplitude greater
than…. This simplifies the assessment procedure and puts the
responsibility of deciding whether a defect indication is acceptable
or unacceptable on the NDT operator.

For in-service inspection, acceptance criteria are not as easy to

define. If manufacturing acceptance criteria are used, it should be
justified that these capture relevant degradation mechanisms
which may be present in operational conditions. They should also
be compatible with the both the plant item and the NDT technique
used.

The results of NDT can be fed ECAs involve the solution of mathematical formulae and, as a
directly in to an Engineering consequence, answers are often quoted to a number of decimal
Critical Assessment (ECA) so that places. However, the errors on the input information obtained
the fitness for purpose of the plant from the NDT results are likely to be in the order of millimetres. It
can be assessed. is important that the sizing errors in the NDT measurements are
estimated and taken into account in the ECA.

A number of codes can be Two of the more important codes are:

followed to assess flaws and
degradation. Many codes that ● BS7910: 2000 (Guide on methods for assessing the
have been prepared take into acceptability of flaws in metallic structures) which
account the accuracy of the NDT superseded earlier standards, PD 6493: 1980 (Guidance
test methods, however, some do on some methods for the derivation of acceptance levels
not and care should be taken for defects in fusion welded joints) and PD 6493:1991
when interpreting the results. (Guidance on methods for assessing the acceptability of
flaws in fusion welded structures).
● API 579 (Recommended practice for fitness for service).

All inputs into the ECA should be ● Have transients or worse case operating conditions been
justified: considered?
● Are the values for material properties correct?
● What assumptions have been made?

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The output from the ECA will If an RBI process is used then the results should be fed back into
determine the course of action the the risk assessment and appropriate changes made to the
plant owner should take. required action.

5.2 NDT Management

In order to have confidence in This requires proper management and control. Plant owners who
NDT results, it is important that the have a certified quality management system will have procedures
NDT, as a special process, is to control the instigation and purchase of NDT activities. They
applied correctly and the capability may also have procedures to cover the application of the NDT
of that process is known and although these will often be left to the NDT vendor.
understood.

In addition to a certified quality Shows that NDT companies have the systems in place to
system, UKAS accreditation adequately control the application of NDT.
required to either BS EN ISO/IEC
17025 Testing or EN 45004
Inspection

However, site practice can be Responsibility for the specification and control of the NDT is not
different from the documentation always clearly defined between the plant owner and the NDT
and the nature of NDT activities vendor. An ISO 9000 plant owner can hire in bodies from a UKAS
mean that they are not always accredited NDT vendor with the result that the operators work
subjected to the same control as is under neither quality system.
applied to other products and
services. Errors are common in unplanned NDT activities: operators
performing a planned job may be asked to 'inspect this item whilst
you are here'. In such a case the NDT performed is dependent on
the operator's experience; its appropriateness and capabilities are
not stated and records to allow future assessment or repetition
may not be produced. The control of NDT is discussed in more
detail in Best Practice documents for Ultrasonics and Surface
Techniques.

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NDT personnel are trained and The requirements for centrally administered schemes are laid
certified under either a central down in BS EN 473. Their employer should define their training in
certification scheme (e.g. PCN) or a written practice. Three levels of competency are defined:
an employer based scheme (e.g.
ASNT). ● Level 1 - qualified to carry out NDT operations according to
written instructions under the supervision of Level 2 or
Level 3 personnel.
● Level 2 - have demonstrated competence to perform and
supervise non-destructive testing according to established
or recognised procedures. This includes the ability to define
the limitations of application of the testing method and to
translate NDT standards and specifications into NDT
instructions.
● Level 3 - qualified to direct any NDT operation for which
they are certificated This is a supervisory qualification.

Full details are given for the PCN scheme in the Best Practice
documents for Ultrasonics and Surface Techniques.

Most NDT techniques require the This is one aspect which can be overlooked with inevitable
operators to have good eyesight consequences for the quality of the NDT performed.
and to have it checked annually.

Site NDT should be under the Level 2 qualifications are specific to a NDT method and, in the
supervision and support of a Level case of ultrasonics, to a particular geometry. Generic
3 operator and NDT procedures qualifications such as PCN may need supplementation by job
should be approved by Level 3 specific training for particular NDT technique applications. This
personnel, or equivalent. was highlighted by HSE's PANI project which investigated the
effectiveness of manual ultrasonics as applied on industrial plant.

Where a central certification The NDT vendor should provide the plant owner with evidence to
scheme does not exist for a show that the personnel have sufficient experience and training in
technique (which is the case for the application of that technique.
many but not all of the techniques
listed under Other Techniques)

All NDT should be controlled by a Some techniques such as magnetic particle inspection or dye
procedure approved by a Level 3, penetrant inspection are simple to apply in principle and there is a
or equivalent. temptation to just apply them without a procedure. Conversely
operators who have a wide experience of the technique may
apply advanced techniques and equipment and they may rely on
that experience to adjust the many variables instead of recording
them in a complete procedure.

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NDT is often applied in It is not sufficient to state that a component was inspected in
compliance with a national or accordance with a standard. Most standards have options on
international standard. various technique parameters and a procedure or technique sheet
should be produced to state what values are to be used. Approval
of the procedure by a Level 3 implies that the standard has been
assessed in the light of the plant item to be inspected and found
to be appropriate.

Some standards have been tried The harmonisation of standards across Europe has produced
and tested over many years but many standards which do not yet have a track record to support
the data and expertise on which them.
they were based and which
defines their capabilities and An exception is the Magnetic Particle Standard: "Method for
limitations is often not available. Magnetic particle flaw detection", British Standard BS 6072: 1981
where the supporting information is available as "Magnetic
particle flaw detection. A guide to the principles and practice of
applying magnetic particle flaw detection in accordance with
BS6072.", British Standard PD 6513: 1985.

NDT can be applied without a But only if all the parameters are recorded so that what has been
written procedure. applied can be subsequently assessed and if necessary repeated.

The HSE's Best Practice Where NDT plays a key role in guaranteeing the safety of the
documents (ultrasonics, surface component, additional steps would be expected to be taken to
techniques) give guidance on improve the reliability of the NDT, to ensure that all the defects of
assessing the role of the NDT and concern are detected and that the NDT technique is applied
the effectiveness required of it in correctly.
reducing the risk of component
failure. Such steps include:

● auditing the NDT with independent operators performing

repeat NDT on a sample of the volume inspected;
● repeating all of the NDT with different personnel or with
different NDT techniques;
● witnessing the inspection by independent third party;
● establishing capability through qualification.

The Best Practice documents also list other important measures

that should be considered when looking to ensure a high reliability
of inspection.

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The capability of an NDT This process is known as Inspection Qualification, Validation or

technique to detect and size Performance Demonstration. The amount of evidence gathered
specified defects can be assessed and assessed can be tailored to the importance of the NDT and
by the gathering of evidence so need not be prohibitive. Further information is given in the Best
based on physical reasoning, Practice documents for Ultrasonics and Surface Techniques.
theoretical modelling,
experimental work and previously
published work.

6. Techniques and Capabilities

● 6.1 Visual Inspection

● 6.2 Thickness Measurement
● 6.3 Defect Detection
● 6.4 Other Techniques
● 6.5 Common NDT Technique Trade Names

Detailed information on NDT techniques can be found elsewhere:

● HOIS,
● Best Practice RBI document
● Or from The British Institute of Non-Destructive Testing

A brief summary is given below. Terminology other than that relating to specific techniques is given in
the Terminology Section. In each section the information is presented in alphabetical order.

6.1 Visual Inspection

The simplest and easiest It is able to detect surface damage and distortion. However,
technique to apply and often access to the surface is required and the capability relies on the
called by the generic term illumination and the eyesight of the inspector.
'inspection' on process plant.
Many aids are available for visual inspection ranging from a
magnifying glass through endoscopes and boroscopes which
allow viewing of surfaces inaccessible to the eye alone, to fully
remote computerised video systems. In the latter case as 'seeing
is believing' care needs to be taken to ensure that the signal
processing of the image does not hide any defects.

6.2 Thickness Measurement

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The commonest damage found on Ultrasonics (high frequency sound) provides an accurate point
process plant is corrosion and so measurement of wall thickness.
techniques which allow remaining
wall thickness to be measured are The surface on which the transducer is placed needs to be clean
widely applied. and, as it provides a point measurement, the measurement
positions need to be selected with consideration of the type of
corrosion damage so that the minimum wall thickness can be
detected. When using a grid to survey a large surface area, the
pitch of the grid needs to be selected so that it will detect the
damage of concern.

Care needs to be taken when taking measurements on plant

which is painted or coated to ensure that the measurement is just
that of the remaining wall. Newer instruments have facilities to
assist the operator in this task but older equipment require more
care on the part of the operator.

Other thickness techniques Flash Radiography, Magnetic Flux Leakage, Pulsed Eddy
include: Currents and these are discussed below.

These techniques are more limited in their application by material

type, accuracy of measurement, wall thickness or geometry than
ultrasonics but offer other advantages such as speed of
application or the ability to inspect under insulation.

6.3 Defect Detection

Defect detection techniques fall into two categories:

● those that can only detect defects on or near to the surface of a component (Surface
Techniques);
● those which can detect both surface and embedded defects (Volumetric Techniques).

Surface Techniques

● Dye Penetrant Inspection (PT)

● Eddy Currents
● Magnetic Particle Inspection (MPI or MT)

Dye Penetrant Inspection (PT)

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Dye is drawn into any surface This NDT method can only detect defects which are open to the
breaking defects which are then inspection surface.
highlighted by the application of a
developer which draws the dye
back out of the defect.

Dye penetrant is the preferred Dye penetrant is better suited to the detection of volumetric
surface technique for non- defects like pits but is more susceptible to the surface condition
magnetic materials. than magnetic particle inspection. Detection of tight cracks will
require the dye to be left on the surface for a long time.

The component surface needs to Mechanical cleaning methods can lead to crack openings being
be cleaned prior to the application closed, subsequently preventing detection. Care needs to be
of dye penetrant inspection. taken with any technique which requires the application of
chemicals to plant to ensure that the chemicals are compatible
with the plant material. It is particularly important that only
chemicals with low halogen content are applied to stainless steel
to avoid the initiation of stress corrosion cracking.

Fluorescent dyes are used to increase the contrast of indications making them more visible to the
operator and hence increasing the sensitivity of the technique.

The HSE's Best Practice document on the procurement of Surface Techniques gives more details
regarding dye penetrant inspection.

Eddy Currents

When an alternating current is Any defect in the component which restricts the eddy current flow
passed through a coil close to a alters the balance between the applied and back EMFs and can
component surface, eddy currents be detected.
are induced and produce a back
EMF on the current in the coil.

The skin depth, which is a In ferro-magnetic material the skin depth is very small and the
function of the permeability of the technique will only detect surface breaking defects. In non-
material and the frequency, magnetic material it provides some sub-surface capability and can
determines the depth of give some indication of the depth of a defect.
penetration of the eddy currents.

Eddy current techniques are widely applied in the NDT of heat exchanger tubing.

Magnetic Particle Inspection (MPI or MT)

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Defects on the inspection surface Magnetic particles sprayed onto the surface are attracted to these
interrupt the lines of magnetic flux. defects identifying their position.

This NDT method only detects abrupt changes in the magnetic

field and therefore only supplies capability for defects that break
the inspection surface. Care needs to be taken to avoid false calls
which may arise due to changes in geometry or the presence of
residual magnetic fields.

Fluorescent magnetic inks are used to increase the contrast of indications making them more visible to
the operator and hence increasing the sensitivity of the technique.

Magnetic particle inspection is generally the preferred NDT method for the detection of surface cracks
in ferritic material. The HSE's Best Practice document on the procurement of Surface Techniques
gives more details regarding magnetic particle inspection.

Volumetric Techniques

● Radiography
● Ultrasonics

Radiography

Radiography is the detection of As it is sensitive to material loss, radiography is better suited to

material loss by the variation in the detection of volumetric defects such as slag or porosity.
applied radiation, g or x-ray, Detection of planar defects or cracks will depend on the gape or
passing through a component and opening of these defects and the misorientation of the radiation
impinging on a film. beam from the axis of the defect. In many cases, cracks will not
be detected.

Radiography is liked because it produces a hard copy of the

results - the film. It is unable to provide depth information
regarding defects without additional specialist techniques (eg
profile radiography may give depth information on large volume
defects).

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Defects are identified by abrupt The gradient of the curve of density against exposure determines
changes in the density of the how visible are small changes in exposure. Such changes can
developed film: the film density is arise from the presence of defects and so the ability to detect
related to the exposure it has them through changes in film density is of prime importance.
received from the radiation.
This characteristic of the film is its contrast. Contrast tends to
increase with film density and so high densities are beneficial in
the detection of defects. However, viewing high density films
requires good lighting conditions such as high light intensity, low
background light and film masking and there are practical limits
on the level to which density can be increased because of the
reduction in transmitted light intensity. Density in the range 2.0 -
3.0 is usually regarded as representing the best compromise
between contrast and viewing requirements.

Image quality indicators (IQI) are The IQI is placed on the object under test and imaged when the
commonly of the wire type, radiograph is taken. The smallest wire diameter, hole diameter or
comprising straight wires of step that is visible on the radiograph then gives a guide to the
differing diameters sealed in a sensitivity achieved.
plastic envelope, or ones which
use holes or steps in a block of The IQI type and its position are specified in the appropriate
metal. radiographic standard. It should be recognised that the sensitivity
established by an IQI relates only to the ability to detect changes
in section, wire size etc. This sensitivity is only indirectly related to
defect detectability.

The HSE's main concerns are that Incidents occur because of poor job planning (most notably with
a significant number of NDT site radiography); failure to use adequate local source shielding
contractors fail to adopt routine (collimation); or inadequate systems of work.
working practices capable of
keeping radiation exposures of
employees as low as reasonably
practicable.

The quality and sensitivity of a radiograph are measured by the density of the film and the use of an IQI.

Industrial Radiography is covered by the Ionising Radiations Regulations 1999 (IRR99) which mostly
came into force on 1 January 2000. Information regarding the requirements of the regulations is
available from the HSE website.

Ultrasonics

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Ultrasonics is the use of high In thickness checking the reflections from the wall surfaces are
frequency sound waves in a measured. In defect detection reflections from cracks, voids and
similar manner to sonar or radar: inclusions are detected and assessed.
sound pulses are reflected from
interfaces or discontinuities. The transfer of sound from the ultrasonic probe to the component
requires a coupling medium, which is usually water or gel. The
condition of the interface determines how much sound is
transferred into the component, how much is scattered and how
much noise is produced.

Ultrasonics requires a relatively Manual application over a large area is relatively slow and the
good surface finish. technique needs to be tailored to the defects requiring detection.
However, ultrasonics is able to provide both length and through
wall size information.

Some materials such as corrosion-resistant alloys (eg high nickel

alloys and austenitic steels) cause additional problems for
ultrasonics and require special techniques and appropriately
trained personnel.

Ultrasonics can be automated and hard copy results produced.

6.4 Other Techniques (in alphabetical order)

● AC-FM - Alternating Current Field Measurement

● Acoustic Emission
● Creep waves
● Digital Filmless Radiography
● Flash Radiography
● Leak Testing
● Long Range Ultrasonics
● Magnetic Flux Leakage (MFL)
● Phased Array Inspection
● Pressure Testing
● Pulsed Eddy Currents
● Radioscopy
● Remote Field Eddy Currents
● Replication
● Shearography
● Time of Flight Diffraction. TOFD
● Thermography

AC-FM - Alternating Current Field Measurement

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This is a non-contacting A uniform electric current is induced into the material to be

electromagnetic technique which inspected which produces a magnetic field which in turn will be
is used as a surface defect disturbed and flow around the edges of a defect if present. The
detection alternative to magnetic probes are constructed in order to detect these magnetic field
particle and dye penetrant disturbances. Software algorithms allow an estimate of crack
inspections in conducting depth and crack length to be obtained.
materials.

The technique is capable of It can cope with poor surfaces and can test through coatings.
detecting sub-surface defects on However, it requires skilled operators to apply it correctly.
non-magnetic materials.

Acoustic Emission

A passive technique in which an Signals originating in the plant item, which are above a specified
array of acoustic sensors are amplitude threshold, are recorded. Signals from crack
attached around the plant item propagation, corrosion products and leaks may be identified
under test. and located by triangulation.

A common application is in This is not a quantitative technique but gives a qualitative

monitoring above ground assessment of the condition of the tank.
storage tanks with the sound
being generated by the spalling of
corrosion products.

When acoustic emission is used to Operational noise may not be present when conducting a
detect crack growth it faces the hydraulic test but the stresses seen by the plant item may be
challenge of detecting the signal quite different to those seen in service.
generated by the growth in the
presence of operating noise.

Creep waves

This technique is another type of As it propagates it converts to a mode which travels into the
ultrasonic wave which travels component at an angle to the surface. This latter wave will
along the surface of a component. convert back to a surface wave if it hits a surface parallel to the
surface on which it originated.

The technique is often used for the detection of near surface defects as a complement to the time of
flight technique.

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Digital Filmless Radiography

Industrial radiography using Applications include process corrosion detection and

computer based or "filmless" measurement, particularly under insulation and coatings on
radiography systems can collect process pipework.
and analyse radiographic data,
completely replacing conventional This technology complements non-projection systems like SCAR
film. to provide a safe, rapid inspection. The system uses flexible, re-
usable phosphor plates to capture images. The exposed plate is
processed through a laser scanner, delivering the image to a high
resolution mono-monitor. After scanning the plate, the digital
image is interpreted, reported and digitally stored for future
retrieval and analysis.

The flexibility of this approach means that extra control is required of the process to ensure radiographs
are traceable and not distorted, deleted or over-written.

Flash Radiography

Originally developed to image It is normally applied to pipes up to 12" OD but can be applied to
rapidly moving dynamic events, items with diameters up to one meter given sufficient source to
flash radiography has found film distance and radiation output. The technique uses x-ray
application in the detection of equipment with a low radiation exposure time, fast x-ray films and
corrosion on pipe outside intensifying screens, or digital detection media. It saves costs
diameters under insulation. normally attributed to the removal and re-instatement of insulation
and associated scaffolding.

The beam is arranged tangentially It can also identify where lagging has become waterlogged.
to the pipe wall and corrosion of Contrast and resolution of the image are not as good as that for
the external wall shows up as a conventional radiography because of the limited radiation
variation in the profile of the pipe. available, the large grain film and the relatively large focal spot of
the sources.

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Recent developments have These involve hand held radiographic systems using a source
complemented flash radiography. such Gadolinium-153 in combination with solid-state scintillator
which converts the X-rays into electrons. The quality and output of
the source determines the maximum length of the beam path in
the lagging when looking for under lagging corrosion. Special
Gadolinium-153 equipment can allow measurement of pipe wall
thickness when shot through the centre of the pipe. The
limitations with regard to pipe and lagging diameter will depend on
the particular instrument used, notably the length of the fixed arm
holding the source opposite the detector and should be known by
the NDT vendor.

Leak Testing

This covers a variety of techniques They include:

which are used to identify leakage
paths through containment. ● Direct Bubble test - like mending a bike tyre.

● Vacuum Box - a local vacuum is drawn over a small area

in the containment. Any leakage path will prevent a full

vacuum.
● Tracer Gas Detection - relies on the detection of a tracer

gas such as helium or a halogen gas. These techniques

are semi-quantitative methods that detect the flow of the
tracer gas across a boundary.
● Pressure Change Test - detection of a leak by the

monitoring of absolute pressure, pressure hold, pressure

loss, pressure rise, pressure decay or vacuum retention.

Long Range Ultrasonics

This technique has found its main A particular type of sound wave, Lamb waves, are generated in
application in pipe NDT. the pipe wall which acts as a cylindrical wave guide allowing
propagation ranges of up to 50 m to be obtained. The waves are
reflected back from features including wall loss defects. The
frequency is less than that used in conventional ultrasonics at kHz
rather than MHz. Interpretation of the signals is complicated
because of the different modes of Lamb wave which propagate.
The technique is generally used as a screening tool to identify
areas worth more detailed NDT with alternative techniques.

Magnetic Flux Leakage (MFL)

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This technique relies on the In order to achieve this, the component wall needs to be close to
detection of the magnetic flux, magnetic saturation. The amplitude of the signal obtained from
which is 'squeezed' out of the any wall loss is proportional to the volume that is missing from the
metal wall under test by any region interrogated. This means that the amplitude does not
decrease in the wall thickness. necessarily correspond to the decrease in thickness of the wall.
The technique is not able to discriminate between material loss on
the near surface and material loss on the far surface.

Surface roughness, surface corrosion, distortion, build up of

debris on the magnets and any physical disturbance of the
scanning system as it moves across the component will adversely
affect the results.

MFL is a qualitative technique and It has found wide use in the NDT of tank floors because it is
is unable to give an accurate quick to apply and can detect material loss on both surfaces of
assessment of the remaining wall. the floor. The requirement for the sensor to be placed between
the poles of a magnet mean that the technique is unable to give
100% coverage of a floor up to vertical obstructions and side
walls. The wall thickness that can be inspected by magnetic flux
leakage is limited by the requirement to achieve magnetic
saturation.

The high level of set up effort makes the technique susceptible to

human error. Procedures need to be clear and sufficiently detailed
and operators need to be qualified and experienced in the
application of the technique.

Phased Array Inspection

Technology advances in materials A phased array transducer enables the ultrasonic beam to be
and computers have made it electronically focussed or swept in angle along the length of the
possible for ultrasonic phased array. One phased array transducer can therefore take the place
array transducers to be of a number of conventional transducers or reduce the scanning
manufactured in a similar sized requirement for the transducer.
case to conventional transducers.
This is new and advanced technique and operators need training
and experience of the technique additional to the conventional
ultrasonic qualifications.

Pressure Testing

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Pressure testing is normally a It involves the over pressurisation of a plant item (typically 10 to
requirement of design codes and 50 % over the design operating pressure) with a fluid to see if it is
is performed at the start of life and able to withstand the applied stress. A pneumatic test carries
subsequently. It is not always a more danger than a hydraulic test, releasing 200 times more
non-destructive test. energy should anything go wrong.

Arguments for and against The test may be complemented by the application of acoustic
pressure testing are complex and emission with the objective of trying to detect any crack growth,
beyond the scope of this which may be generated during the test.
document.

HSE have a Guidance Note GS4 on Safety in Pressure Testing, which is supported by Contract
Research Report CRR168: "Pressure Test Safety", 1998.

Pulsed Eddy Currents

This is a technique for detecting Unlike ultrasonic thickness measurement it measures average
corrosion and erosion and wall loss over an area (footprint).
measuring average remaining
wall.

A transmitter coil produces a The eddy currents in turn produce a second magnetic pulse which
magnetic pulse which induces is detected by the receiving coil. The system monitors the rate of
eddy currents within the decay of the eddy current pulse within the steel wall. The average
component wall. thickness is derived from the comparison of the transient time of
certain signal features with signals from known calibration pieces.

It is important that the operator is This technique is quick to apply, can test through non-
given information regarding the conductive and non-magnetic material (passive fire protection,
component to allow the NDT concrete) up to 100 mm thick. It is only suitable for low alloy
equipment to be set up correctly steels and is unable to differentiate defects on the top and bottom
and the results to be accurately surfaces.
interpreted.

Radioscopy

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Radioscopy is a digital version of The image is produced on a radiation detector such as a

radiography. fluorescent screen, rather than film, and is then displayed on a
television or computer screen. Often such systems work in real
time and can provide continuous NDT of objects. The recent
advances in detectors and computer technology mean that these
systems can offer advantages over the conventional film NDT
technique.

Remote Field Eddy Currents

This technique provides an The technique monitors the magnetic field produced by induced
alternative to eddy current NDT for eddy currents at some distance from the exciting coil. The system
ferro-magnetic tube inspection. gives poorer resolution and has a lower test speed than a high
frequency eddy current test. The technique is highly sensitive to
gradual wall thinning but detection of localised thinning requires
special probes and electronic control.

Replication

This involves the application of a The film is then removed and examined under a microscope.
temporarily softened plastic film Details such as cracks, surface inclusions and microstructure
onto the prepared surface of the can then be observed remotely from the plant item. A hard copy
item under test so that the surface of the results is also obtained.
profile is imprinted into the film.

Shearography

Shearography is used for Comparison of two sets of laser images produced before and
detection and characterisation of after the application of a load (thermal, tensile, pressure,
delaminations, debonds, and other vibratory) that causes the item under test to deform allows
defects in fibre reinforced calculation of relative deformation at each point on the object and
composites, rubber, and rubber/ highlights local variations in surface deflection. Local variations
metal parts. are characteristic of the defects such as delaminations and
debonds.

Time of Flight Diffraction. TOFD

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This is an ultrasonic technique Sizing can be accurate as the time difference between the signals
which uses the diffracted wave obtained from the top and bottom edges is used to predict the
produced by the edge of a planar size. TOFD requires two ultrasonic probes acting as transmitter
defect to detect and size such and receiver to be scanned as a pair either side of a weld.
defects.
It is relatively quick to apply compared to the conventional manual
pulse echo techniques and a hard copy image can be produced.
As a consequence TOFD is replacing radiography as a preferred
weld NDT technique.

However, TOFD has a number of ● The diffracted tip wave is relatively small in amplitude so
drawbacks which need to be the sensitivity of the NDT needs to be high which can then
considered: lead to false calls;
● Other techniques need to be applied to cover the near
surface region;
● As the weld thickness increases so does the number of
probe separations which are required to cover the
inspection volume;
● The technique requires optimisation for the defects of
concern;
● Skilled operators are required to operate the equipment
and interpret the images.

Thermography

An infrared camera or monitor is Variations in heat transfer through the wall may be attributable to
used to observe the actual wall thinning or the build up of scale. It may indicate the
temperature, or the variation over presence of wet insulation and the potential conditions for
an area, of the surface of a plant corrosion under insulation (CUI).
item.

Alternatively, a heat source can be Unexpected changes in the heat flow can be used to identify
used to heat the surface and the defects.
dispersion of the heat observed.

For containers containing hot or

cold liquid it is possible to observe
the level of the liquid in the item
non-invasively.

The size of defect which can be detected will depend upon the optical parameters of the system and the
resolution of the camera. In assessing the results the emissivity of any paints or coatings on the

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component need to be considered. Reflections of sunlight can also distort readings.

The technique is non-contacting and only line of sight to the surface under examination is required. It is
quick and easy to apply but can only detect defects and or faults which cause a change in heat flow or
the surface temperature of the item.

6.5 Common Trade Names

Fleximat This is a thin flexible strip containing an array of ultrasonic

transducers which can be permanently bonded to a component to
provide continuous corrosion monitoring of fixed locations.

Internal Rotary Inspection System An ultrasonic technique for the NDT of boiler and heat exchanger
- IRIS tubes consisting of a high frequency ultrasonic immersion probe
inside a rotating test head. The system provides coverage of the
full circumference and full wall thickness as the probe is scanned
axially along the tube. The head can be modified for defect
detection if required.

LORUS This is an ultrasonic technique which relies on bulk waves and

was designed specifically for interrogating the plate under the
shell on the annular ring of an above ground storage tank.
The probe does not need to be scanned backwards and forwards
and so is suitable for use on the restricted surface available on
the annular ring.

The sound floods the plate as it travels and is reflected from

corrosion defects on the top or bottom surface. The working range
is about 1 m but as the plate is flooded with sound it is unable to
discriminate between top and bottom defects.

Note: Although the acronym, LORUS, is derived from Long Range

Ultrasonic System, when compared to more recent techniques
referred to as long-range ultrasonics, the LORUS technique can
only be considered medium range (typically up to 1m).

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Saturation Low Frequency Eddy The (SLOFECÔ) technique is very similar to the magnetic flux
Current - SLOFEC leakage technique. However, instead of detecting the flux leakage
with a passive coil or a hall effect sensor, the SLOFEC technique
has an eddy current sensor.

The fact that the eddy currents are used to sense the distortion of
the magnetic field in a layer close to the surface of the component
means that this NDT system is able to inspect a greater wall
thickness and also able to cope with thicker non-magnetic
coatings than the magnetic flux leakage NDT system.

When the equipment is used on non-magnetic stainless steels the

detection technique becomes solely an eddy current NDT
technique.

Small Controlled Area This is a proprietary radiographic system which operates in a

Radiography - SCAR more controlled manner and hence a much smaller area than
traditional radiography. Proper application of the system will
reduce the controlled area to typically within 3 metres of the
emission point. This has the advantages of minimal disruption to
adjacent work areas and of reduced dose rates to classified
workers

7. Checklist for the HSE Inspection of NDT

This checklist covers the whole NDT process from planning through to assessment of results. It is
unlikely to be necessary to apply the checklist from start to finish. It is more likely that specific areas of
concern or criticality will need to be selected and addressed. Bold comments help to direct the
questions and interpret the answers.

7.1 NDT Planning

Identification of plant items which require NDT.

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Scope Is the plant governed by Pressure Systems Regulations?

● If so is there a Written Scheme of Examination for each

plant item?

Does the plant contain non-pressure but hazardous fluid?

● If so is it examined at appropriate intervals by a competent

person?

Is speculative NDT performed to identify unexpected damage

mechanisms?

The defect description will be defined by the capabilities of

the technique applied.

The plant item can only be passed clean of defects meeting

this capability.

Periodicity of NDT What are the examination intervals?

What is the justification for the examination intervals?

If standard guidelines:

● What are the reasons for selecting the chosen periodicity?

● Are these are compatible with the operating conditions?

If RBI:

● What are the damage mechanisms for each plant item?

● What is the risk category?
● If High/Medium, is NDT used to reduce risk?
● Are the results fed back into the RBI process?
● If so have appropriate changes been made to the required
action?

7.2 Management of NDT Process

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Plant owners' attitude to NDT. Is there sufficient independence between the NDT activity and
production/operations functions?

Does the NDT play a key role in assuring the safety of the
component?

Are additional steps taken:

● To improve the reliability?

● To ensure that all the defects of concern are detected?
● To ensure NDT technique is applied correctly?

Is the plant owner aware of the limitations and capability of NDT?

Is the quality of the results checked in any way?

Does the plant owner act on the results?

Is the plant owner an informed customer?

Q.A. Is there a certified quality management system?

Are there procedures to control the instigation and purchase of

NDT activities?

Are there procedures which cover the management and

application of the NDT?

Is the responsibility for the specification and control of the NDT

clearly defined between the plant owner and the NDT vendor?

Is there a system for maintaining inspections records?

Whose quality system are the NDT operators applying the NDT
under (the plant owners' or the NDT companies')?

Are the NDT companies UKAS accredited to either BS EN ISO/

IEC 17025 Testing or BS EN 45004 Inspection?

7.3 NDT Inspection Management

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Specification Is there a specification of requirements or defect description?

This should include:

● Location, type, morphology, orientation;

● Volume or surface that requires NDT;
● The size of defect which needs to be detected;
● The sizing errors that can be tolerated.

Is the NDT in compliance with a national or international standard?

Approval of the procedure by a Level 3 operator implies that

the relevance of the standard has been assessed for the
plant item to be inspected and found to be appropriate.

For pressure systems the NDT method should be specified in

the Written Scheme of Examination.

NDT Procedures Is there evidence of procedures to cover the application of the

NDT?

Whose are they [Plant Owners'? NDT vendor?]

Does this match up with the division of responsibilities?

Does a Level 3 approve the procedures?

Is the procedure, which may be supplemented by a plant specific

technique sheet, sufficiently detailed to define the technique to be
applied?

Is there evidence that the QA & NDT procedures are being

implemented?

Practice can be different from the documentation.

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Implementation What additional steps have been taken:

● To improve the reliability?

e.g. different NDT techniques, repeat independent

inspections or repeating all of the NDT with different
personnel or with different NDT techniques.

● To ensure that all the defects of concern are detected?

e.g. capability established through qualification or auditing

with independent operators repeating sample of volume
inspected.

● To ensure NDT technique is applied correctly?

e.g. witnessing the inspection by independent third party,

audits or measures listed in the Best Practice document.

Are NDT personnel trained and certified?

(e.g. either a central certification scheme such as PCN or

employer based such as ASNT)

Is the site NDT under supervision and support of a Level 3

operator?

Are PCN qualifications supplemented by job specific training for

particular NDT technique applications?

Where a central certification scheme does not exist for the

technique, can the NDT vendor or the plant owner show evidence
that the personnel have sufficient experience and training in the
application of the technique?

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Results Do reports highlight any restrictions in the application of the

technique?

Do they list any changes to the techniques which were required

by the particular application?

Are sufficient parameters recorded so that what has been applied

can be subsequently assessed and if necessary repeated?

Are the sizing errors in the NDT measurements estimated?

Is appropriate action taken in response to the results obtained?

If the result is no defects found, there may still be the need

for action taking into account the capability of the NDT
technique and the nature of defects which may not have been
found.

Assessment of Results How are the NDT results assessed?

Acceptance criteria If manufacturing acceptance criteria are used is there justification

for using them?

Are they compatible with the both the plant item and the NDT
technique used?

Engineering Critical Assessment Has the assessment of the NDT results taken into account the
(ECA) limitations and errors inherent in the technique applied?

Has a code been followed to assess flaws and degradation?

(BS7910: 2000 / PD 6493: 1980 / PD 6493:1991 / API 579)

Does the code take into account the accuracy of the NDT test
methods?

If not what care is taken when interpreting the results?

All inputs into the ECA should be justified.

Have transients or worse case operating conditions been

considered?

Are the values for material properties correct?

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What assumptions been made?

8. Terminology & Current Trends

Testing is the generic term given to a measurement of a property or the performance of an item to
assess whether it is fit for purpose.

Inspection is also a generic term but on process plant it is used in relation to the visual assessment of
plant condition.

Non-Invasive Inspections

Inspecting vessels for possible This can incur very high costs associated with releasing and
internal degradation has emptying the vessel, isolating it and preparing for it for entry. The
traditionally been performed from mechanical disturbances involved in preparing the tank for
the internal surface, e.g. by visual internal NDT and reinstating it may on occasions adversely affect
inspection. future performance of the tank. Also, the environment within the
empty tank may be hazardous for man access requiring additional
precautions to be taken for working in the confined space.

NDT performed from the outside If they are applied in lieu of internal NDT then evidence should be
of the vessel, i.e. non-invasively, provided to show that they are capable of achieving the same
without breaking the containment detection and sizing requirements. This may be in the form of
have the potential to reduce results from both previous invasive and non-invasive inspections
operating costs significantly. showing good correlation or a report on the capability of the non-
invasive inspection which can be compared with previous invasive
results.

Alternatively, non-invasive NDT can be applied in addition to the

internal NDT prior to an outage and during short shutdowns to
assist in the planning of internal NDT or to provide immediate
information on an identified potential problem with the minimum of
interference with other operations.

Non-invasive NDT techniques are more complex than the internal

NDT techniques and so require better planning, QA and project
management procedures. It is important to state the objectives of
the non-invasive inspection as this is likely to have an impact on
the approach to the NDT.

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HOIS2000 and Mitsui Babcock The HOIS project produced a decision tree to establish if non-
have carried out two research invasive inspection was acceptable (HSE staff can access the tool
projects into non-invasive on the HOIS website - HID CD5 will supply the password), and
inspection. the Mitsui Babcock project detailed the requirement to ensure
satisfactory inspection. Results from both of these projects are
being reviewed by HSE prior to being recognised as 'Good
Practice' documents.

Risk Based Inspection

Risk based inspection is the When implementing a risk based approach, safety concerns need
definition of the NDT requirements to take precedence over other influences such as business
based on the risk posed by a interruption and loss of earnings. The RBI approach identifies the
particular plant item. potential damage mechanisms and the required interval of
inspection: high-risk items requiring frequent NDT; low risk items
requiring infrequent or no NDT. This contrasts with the statutory
approach of standard fixed inspection intervals irrespective of risk
of failure.

To use this approach the plant Operator needs to demonstrate

that the risk assessment and NDT planning processes are being
implemented in an effective and appropriate manner.

The risk-based approach requires Information on integrity of plant can be generated from the design,
that the quality and veracity of the operational experience and NDT records, and from sound
information is tested and validated. knowledge of the deterioration mechanisms and the rate at which
deterioration will proceed. The approach is unreliable when there
is lack of, or uncertainty in, the key information required to assess
integrity.

NDT can then be planned at appropriate intervals using NDT

methods that are able to detect the type and level of deterioration
anticipated in order to allow an assessment of the current and
future fitness-for-service to be made.

Sample Inspections

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Rather than applying NDT to the Often a figure of 10% is used. This doesn't necessarily have any
total length of welds or number of scientific basis but is seen as being a reasonable amount without
components, NDT costs can be incurring undue cost. Such an approach is only viable if the
reduced by inspecting a reduced results from the 10% inspected can be legitimately extrapolated to
percentage or sample of the items. the 90% which wasn't inspected. i.e. if the damage mechanism is
equally likely to occur in all of the 100% and if it can be justifiably
assumed that if no defects are found in the 10% examined then
there will be no defects in the remaining 90%.

This approach is not applicable if the damage can occur

preferentially in one area over another or if random defects can
occur.

9. Case Studies

NDT Case Study 1

A process plant contained two stainless steel vessels which had been operating for 21 years. The
contents of the vessels were flammable, mildly toxic and contained 500 ppm of chlorides. The vessels
were operated from full vacuum up to 15 psi for 20 cycles per day. They contained an agitator which
was used in part of the process. Both vessels had been hydraulically tested to 70 psi when new but had
not been subjected to a test since.

The company philosophy was 'Leak before break' but they didn't think that stainless steel would break.
No leak detection equipment had been installed and reliance was placed on plant operators noticing the
smell or observing drips.

The plant owners hired a Competent Person from a large insurance company who produced the Written
Scheme of Examination (WSE) for the vessels. There was no evidence of shared decision making
between the plant owner and the insurance company. A generic WSE was put into use. This followed
SAFED guidelines on periodicity of inspection which was specified as:

External visual examination supplemented by a hammer test every 2 years.

Was this suitable?

The combination of stainless steel and chlorides immediately raises concerns regarding the
possibility of stress corrosion cracking. Whilst the cracks were likely to initiate on the inner
surface an external examination could detect the presence of through wall cracks. However,
stress corrosion cracks can be very tight and difficult to see with the naked eye. The hammer
test offers no benefit - who knows what a good vessel should sound like!

During a thorough examination of one of the vessels the Competent Person called for a small welded
repair to an external weld and for this to be followed by a hydraulic test. The vessel developed leaks at

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40 psi. Further investigation of the vessel found thousands of through wall cracks. The vessel had not
leaked in service because the contents were too viscous to pass through the tight stress corrosion
cracks.

The competent person modified the WSE for the 2nd vessel:

● Yearly examination instead of 2 yearly.

● Addition of internal examination from the access way.
● Addition of internal dye penetrant examination using red dye on 10% of welds.

Was this suitable?

The Internal inspection would be carried out from the small access way with agitator still in
place.

The failed vessel had shown most throughwall cracks in base. This region could not be inspected on the
second vessel from the access way.

Inspection of 10% of welds.

The failed vessel showed through cracks on parent plate and most welds. There was no justification for
limiting the inspection to welds only and for just inspecting 10% of them.

Dye Penetrant Inspection using red dye.

With the cracking on the internal surface there was a chance that the cracks may have been filled with
product and if this had been the case dye penetrant inspection would not have been effective.

Stress corrosion cracking can be tight and if so the dye penetrant indications would not reveal the
defects. Fluorescent dyes give a higher sensitivity and would give much better results in the confined,
dark space of the vessel.

Regulation 9 of The Pressure Systems Safety Regulations 2000 requires that a competent person
examines those parts of the pressure system included in the scheme of examination within the intervals
specified in the scheme. The actions above raise the question of how competent was the competent
person? Did they understand the damage mechanisms and the detection requirements?

The competent person, who was independent from the plant owner, did not involve their company NDT
expert in amending the WSE. When the expert was finally consulted they estimated that the probability
of detection, using the method stated, was less than 30%. In limiting the inspection to just 10% of the
welds then the overall probability of detecting a crack in a weld was just 3%. This is unacceptably low. A
probability of detection of only 50% may be acceptable for a regularly applied, non-critical inspection
whilst for a highly critical inspection the probability of detection would need to be up near 95%.

However, the examination of the second vessel did find two incidences of stress corrosion cracking
(SCC): one around the access way nozzle and a star crack in the plate. The nozzle was repaired by

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welding and the vessel was hydraulically tested to 60 psi. The star crack was to be monitored at the
next inspection in a year's time. No further review of WSE was performed and the vessel was put back
into service. The Competent person who put the vessel back into service was not the regular surveyor
for the site and it raises the question of whether they fully understood the process.

When the poor inspection and quick return to service was questioned the following excuses were
offered:

● The client needed the vessel back as quickly as possible.

● We worked day and night for 2 days.
● We have never seen this problem before.
● We follow SAFed guidelines.
● There is no better way to inspect this type ofvessel.

Conclusions

● The client placed a high dependency on the competent person to satisfy the 'sufficient' aspect of
the WSE.
● The Competent Person may not have understood the process.
● When part of a large company, the Competent Person system relies on the surveyor feeding
back information to Head Office which they will not be able to do if they lack understanding.
● The client imposed time pressures on the Competent Person.
● The Competent Person had access to experts in various disciplines but these were not used.
● The initial WSE was poorly thought out.
● The final WSE was even more poorly thought out.
● No attempt was made to estimate critical crack sizes or growth rates and the NDT selected did
not have a capability for measuring defect through wall size.

Finally, just because a leading competent person certifies the WSE it does not mean that it is sufficient:
the WSE should be scrutinised and the contents challenged wherever there is doubt on their
suitability.

NDT CASE STUDY 2

A number of large LPG storage vessels were due for their first thorough examination after 10 years of
use. The vessels had been designed and constructed to BS 5500 Class 1, with radiography used for the
detection of volumetric defects. The size of these vessels required that they were site constructed. Site
manufacture has the disadvantage that welding and inspection is open to the weather, and in similar
vessels Fabrication Hydrogen Cracking (FHC) has occurred, which is very difficult to detect with
radiography.

The Operator and Competent Person wished to change the inspection strategy to non-invasive
inspection to prevent disruption to operation. The period between inspections was not to be altered.

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It has been standard practice to inspect such vessels from the inside. Applying NDT from the outside
surface i.e. non-invasively can offer cost savings. However, it is important to ensure that the external
inspection achieves the desired detection capability. A relevant joint industry research project (phase 1)
had concluded that non-invasive inspection is best used where there was a history of invasive
inspections to enable a comparison of the results of the two methods to be made.

In this case there was no previous history, but the Operator had similar vessels on other sites, so
considered that they had previous experience in damage mechanisms that could occur.

Was this suitable?

In addition to checking for in-service degradation, the first in-service inspection is also used to show that
there are no defects in the vessel which were missed by the manufacturing inspections which could give
rise to integrity problems. The competent person and Operator offered no evidence to show that the
non-invasive inspection was fit for purpose: there were no results from previous invasive inspections to
use as a bench mark; there was no evidence to show that the non-invasive inspection techniques to be
used would give the same detection capability as an invasive inspection; no evidence was provided to
show that any Fabrication Hydrogen Cracking not detected during manufacture would be detected in
the in-service inspection.

The Operator carried out a number of studies to address these concerns. A detailed study of the LPG
supply chain was carried out to evaluate if any hydrogen sulphide or other trace components could have
been present, which could give additional damage mechanisms.

A test piece was manufactured, to simulate the main weld of the vessels with a number of defects
representing in-service defects and FHC. A number of NDT techniques were applied to the test piece,
and manual ultrasonic inspection gave the best results. An operator was then qualified on the test piece.

To evaluate the critical crack size, Engineering Critical Assessments were carried out on the vessels.
The ECAs assumed the fracture toughness of the material. As the material had a specified Charpy
value at -50oC, this value was converted to a fracture toughness and used for the low temperature
analysis. A LPG vessel is required to operate though a specific temperature range specified in the
LPGA Code of Practice, so assessments had to be made of the tolerable defect size at different
temperatures and pressures. The fracture toughness at other temperatures was taken from very limited
data available from The Welding Institute. The size of the tolerable defect was quite small, but the NDT
trials had demonstrated that defects half the tolerable size could be detected.

A trial inspection was carried out on the smallest vessel. To obtain access to the vessel welds a mobile
platform was used, with a target set for inspection of 10% of the weld, in the hope that a greater
coverage could be obtained in the period that the mobile platform was available. In addition magnetic
particle inspection was to be carried out in the regions of the support legs.

The inspection was carried out in high winds, and the platform was not available for use on two of the
days. Only 10 % of the weld length was inspected. The results for the inspection identified a number of
planar defects which exceeded the manufacturing ultrasonic acceptance criteria, but were smaller than
the maximum allowable size.

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 HSE - Inspection / Non-Destructive Testing (NDT)

Was this suitable?

The NDT of the vessel was performed from a mobile platform in windy conditions.

While the wind speed allowed the inspection to be carried out on 3 days, NDT requires a stable
work platform to ensure reliable results.

The ECA was based on material assumptions which had limited validity. The fracture toughness
was based on measurements on the parent plate and was then used for the low temperature
assessment. For the assessments carried out at different temperatures, parent plate data was
taken from a limited database, but, to be conservative, a lower bound value should have been
used. If FHC had occurred it would have been in the heat affected zone of the weld, which the
work did not address.

Inspection on the other vessels used mechanised ultrasonic inspection, which was not affected by wind,
and provided inspection data to a computer, which was analysed later. Only 10% of the weld length was
inspected on each vessel. In one case the defect was sized at 4 mm high, compared with the maximum
tolerable defect height of 6 mm.

Was this suitable?

The use of mechanised ultrasonic testing was a considerable improvement over the manual
technique. However, no allowance was made for the sizing errors of the inspection. The sizing
accuracy of the mechanised ultrasonic inspection would have been +/- 2 mm. This means that
the 4 mm defects detected could have actually been at the maximum tolerable size of 6 mm. No
justification was given for limiting the inspection to 10 % coverage of the weld and this was not
extended even when defects near the tolerable size had been detected.

Conclusions:

● The Competent Person changed the inspection strategy to non-invasive inspection without the
benefit of information from prior invasive inspection or other evidence to justify the decision.
● By use of a test specimen it was demonstrated that the intended non-invasive NDT technique
was capable of detecting and sizing in-service and manufacturing defects.
● With no prior inspection data available, inspecting only 10% of the weld does not give a strong
demonstration of the vessel integrity. Having detected defects that were on the limit of the
tolerable size, increased coverage should have been carried out.
● Allowance should have been made for the sizing error of the NDT technique, and the acceptance
criteria set accordingly.
● The ECAs did not use lower bound material properties that may have been present in the welds.

10. Guidance and Codes of Practice Relating to Inspection/NDT

Please note that references quoted are current at June 2003: the originating organisation should be
contacted to establish the status and current version.

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 HSE - Inspection / Non-Destructive Testing (NDT)

There are many guides and codes of practice relating to inspection and NDT. These include:

Acceptance Standards and ECA

● BS EN 25817: 1992, ISO 5817:1992, Arc Welded joints in steel. Guidance on quality levels for
imperfections.
● BS 7910: 1999, Guide on methods for assessing the acceptability of flaws in metallic structures.

Accreditation

● BS EN 45004:1995, General criteria for the operation of various types of bodies performing
inspection.

Personnel Qualification

● BS EN 473:2000, Non-destructive testing. Qualification and certification of NDT personnel.

General principles
● ISO 9712: 1999 Non-destructive testing - Qualification and certification of personnel.
● PCN/GEN/2000, General requirements for qualification and certification of personnel engaged in
Non-destructive testing, BINDT.

Inspection Techniques

● BS EN 1714: 1998, Non-destructive examination of welded joints - Ultrasonic examination of

welded joints.
● BS EN 1435:1997, Non-destructive testing of welds - Radiographic testing of welded joints
● BS EN 571-1 1997 Non-destructive testing. Penetrant testing. General principles
● BS EN 1290 1998 Non destructive examination of welds: Magnetic particle examination of welds:
Method

Information regarding these British Standards can be obtained from the BSI web Site.

American Petroleum Institute Guides.

● API 510 Pressure vessel inspection code: Maintenance inspection, rating, repair, and alteration
● API RP 572 Inspection of pressure vessels,
● API RP 576 Inspection of pressure-relieving devices
● API 579 Recommended Practice for Fitness for Service

Institute of Petroleum Guides

● Institute of Petroleum, Model Code of Safe Practice for the Petroleum Industry: Part 12: Pressure
Vessel Systems Examination. 2nd Edition. 1993. ISBN 0471 939366.
● Institute of Petroleum, Model Code of Safe Practice for the Petroleum Industry: Part 13: Pressure

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 HSE - Inspection / Non-Destructive Testing (NDT)

Piping Systems Examination. 2nd Edition. 1993. ISBN 0471 939374.

Safety Assessment Federation - SAFed

● Guidelines for the Production of Written Schemes of Examination and the Examination of
Pressure Vessels Incorporating Openings to Facilitate Ready Internal Access Ref: PSG4 April
2003.
● Pressure Systems: Guidelines on Periodicity of Examinations Ref: PSG1 (ISBN 1 901212 10 6).
Date of Publication: May 1997
● Shell Boilers: Guidelines for the Examination of Shell-to-Endplate and Furnace-to Endplate
Welded Joints Ref: SBG1 (ISBN 1 901212 05)
Date of Publication: April 1997
● Shell Boilers: Guidelines for the Examination of Longitudinal Seams of Shell Boilers Ref: SBG2
(ISBN 1 901212 30 0). Date of Publication: May 1998

Engineering Equipment Manufacturers and Users Association - EEMUA

● 159 Users' Guide to the Inspection, Maintenance and Repair of Above-ground Vertical Cylindrical
Steel Storage Tanks (3rd edition 2003) ISBN 0 85931 1317

Page last updated: 16/02/2004

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 CORROSION AND FOULING IN
SULFURIC ACID ALKYLATION UNITS

Presented By

Randy Peterson
Manager of Process Engineering

STRATCO, Inc.
11350 Tomahawk Creek Parkway
Suite 200
Leawood, KS 66211

November, 2001

Copyright 2001 STRATCO, Inc.

 Table of Contents

I. Introduction.........................................................................................................................Page 1

II. Reaction Section ............................................................................................................Pages 1-6

III. Refrigeration Section and Depropanizer.....................................................................Pages 6-10

A. Refrigeration System .............................................................................................Page 6

B. Depropanizer Systems ....................................................................................Pages 7-10

C. Refrigerant Recycle Pumps .................................................................................Page 10

IV. Net Effluent Treating and Fractionation ...................................................................Pages 11-20

A. Net Effluent Treating Systems .....................................................................Pages 11-18

B. Deisobutanizer Tower...................................................................................Pages 19-20

V. Alkylation Unit Blowdown Section..........................................................................Pages 21-23

VI. Conclusion ........................................................................................................................Page 23

-i-
 CORROSION AND FOULING
IN SULFURIC ACID ALKYLATION UNITS

I. INTRODUCTION

Causes and controls for many of the potential corrosion and fouling problems in a sulfuric
acid alkylation unit are reviewed in this paper. Corrosion and fouling should not be serious
problems in a plant which is well designed, operated, and maintained. Process conditions
which provide the best yield and product quality also tend to minimize corrosion and
fouling. However, problems are not uncommon in plants operated at rates significantly
above the design capacity, or where maintenance and operating problems go unresolved.

The Six "Commandments" for Corrosion and Fouling Control

Most corrosion and fouling in a sulfuric acid alkylation unit can be minimized by following
the six common sense "commandments" listed below:

1. Maintain good reaction conditions.

2. Avoid high velocity or temperature in sulfuric acid service.
3. Don't let acid sit stagnant.
4. Minimize the amount of water entering the process.
5. Keep acidic streams dry and "wet" streams neutral. Use corrosion-resistant
materials where they mix.
6. Maintain good treating operations.

Throughout the rest of this paper, these rules will be related to specific examples in each
section of the alkylation unit.

II. REACTION SECTION

Commandment 1: Maintain good reaction conditions.

The reaction conditions which produce the best quality alkylate at the lowest acid
consumption also tend to minimize corrosion and fouling throughout the alkylation unit.
Low reaction temperature, low space velocity, high isobutane-to-olefin ratio, and high acid
strength all favor the alkylation reaction and tend to suppress undesirable side reactions. In
particular, the rate of corrosion and erosion in an alkylation Contactor reactor is minimized
at low temperature and high acid strength (especially low water content). Contactor reactors
in low acid strength service tend to have the highest rates of corrosion.

-1-
II. REACTION SECTION (Cont'd)

Reactor operation also affects corrosion and fouling outside the reaction section. Good
reactor operation can minimize the formation of SO2, which can overload the depropanizer
feed treatment system, and alkyl sulfates, which can overload the net effluent treating
system. The formation of SO2 increases as reaction temperature increases.

Commandment 2: Avoid high velocity or temperature in sulfuric acid service.

In general, the acid strength in the reaction section of a sulfuric acid alkylation plant should
not be allowed to fall below about 89 wt%. The overall corrosion rate of carbon steel is very
low at this acid strength, so general corrosion is usually not a problem. However, some
precautions are necessary to prevent localized corrosion.

Iron sulfate, a corrosion product of carbon steel and sulfuric acid, forms a protective film on
the surface of the steel. This coating has low solubility in strong acid at low temperature, but
it is soft and can be stripped away if subjected to high fluid velocity. For this reason,
STRATCO sizes sulfuric acid lines that are to be in continuous service to keep the velocity
below 1 foot per second (0.3 meters per second) in new alkylation units. Corrosion and
erosion accelerates dramatically if the velocity exceeds 3-5 feet per second (0.9-1.5 meters
per second). This problem is not uncommon in older alkylation units which are being
operated well beyond their design capacity. Even at low velocity, localized turbulence can
occur in and around valves, meter stations and near the suction and discharge of pumps in
acid service. STRATCO therefore specifies Alloy 20 valves and piping in these areas of the
unit to resist erosion and corrosion.

The stability of the protective sulfate coating also depends on temperature. This is probably
the reason Contactor reactors operated at high temperatures tend to have a higher corrosion
rate than reactors operated in the preferred temperature range of 45-50°F (7-10°C). We have
found that there tends to be a substantial decrease in the life of a Contactor reactor tube
bundle as the reaction temperature exceeds about 55°F (13°C). With ideal reaction
conditions, a Contactor reactor tube bundle may last up to 15 years. However, some refiners
who operate at temperatures exceeding 60°F (16°C) have reported Contactor reactor tube
bundle life of < 5 years in low acid strength service, and some have reported failures of new
tube bundles after only one year of operation at 65°F (18°C).

Care must also be taken with the installation of heat tracing to avoid hot spots on acid lines.
For this reason STRATCO recommends electric tracing instead of steam tracing in acid
service, and the tracing must not actually touch the surface of the metal.

-2-
II. REACTION SECTION (Cont'd)

Commandment 3: Don't let acid sit stagnant.

Spent acid contains polymers and other acid soluble hydrocarbon compounds which are
reactive. These compounds continue to react in stagnant acid, producing solids which can
plug lines and instrument connections. Plugging in level instrument connections is
prevented by periodically flushing them with net effluent from the net effluent pump. Net
effluent flushing connections should be piped to all the sight glasses and level instruments in
acid service. Flushing of these instruments every other week (at a minimum) will usually
prevent plugging. In addition, any lines in intermittent spent acid service (such as startup
lines) should be drained after use to prevent plugging.

Commandment 4: Minimize the amount of water entering the process.

The amount of water in the Contactor reactor feed impacts the rate of corrosion as well as
acid consumption. The corrosivity of spent acid depends more on its water content than the
actual wt% H2SO4. For example, 90 wt% acid that contains 5% water and 5% hydrocarbon
is more corrosive than 90% acid that contains 2% water and 8% hydrocarbon. Higher water
content in the spent acid can also reduce alkylate product quality. It is therefore important to
monitor the performance of the feed/effluent exchangers and the feed coalescer. The feed
coalescer operating temperature affects the solubility of water in the feed. New plants are
typically designed for a feed/effluent exchanger outlet temperature of 55-60°F (13-16°C). A
theoretical water balance can be done on the feed coalescer based on the solubility of water
in the feed at the temperatures upstream and downstream of the feed/effluent exchangers. An
evaluation of the coalescer should be completed if the amount of water actually removed
drops below 80% of the theoretical amount.

Operation of a sulfuric acid alkylation unit with MTBE or TAME raffinate feed can also
affect the amount of water entering the process. Operation with MTBE raffinate should
normally provide higher octane alkylate product with acid consumption equal to or less than
operation with FCC olefins. However, some refiners have reported accelerated corrosion of
the Contactor reactor impeller, wear ring, feed nozzles, and tube bundles while processing an
MTBE raffinate. This problem is related to the presence of oxygenates such as methanol and
dimethyl ether (DME) in the alkylation unit feed. Ideally, MTBE raffinate should contain
less than 500 ppm DME and less than 40 ppm methanol.

-3-
II. REACTION SECTION (Cont'd)

Commandment 4: Minimize the amount of water entering the process.

Methanol and DME in the olefin feed react with the sulfuric acid alkylation catalyst,
consuming acid and producing water as a byproduct, which dilutes the acid. Typical
reactions are:

CH3-OH + H2SO4 → CH3HSO4 + H2O

CH3-O-CH3 + 2 H2SO4 → 2 CH3HSO4 + H2O

Methanol and DME are also harmful because they increase the solubility of water in the
hydrocarbon feed, rendering the alkylation unit feed coalescer less effective. The additional
water in the feed and produced by the above reactions causes accelerated corrosion in the
Contactor reactor mixing zone.

With good operation of the MTBE plant water wash and methanol recovery column, very
little methanol should be present in the alkylation plant feed. In contrast, the water wash
does not remove DME. Its concentration in the raffinate stream depends on MTBE reaction
zone operation. STRATCO recommends that refiners with MTBE or TAME units upstream
of their alkylation unit analyze their olefin feed regularly for oxygenates. Bear in mind that
the refinery laboratory gas chromatograph must be specially configured to fully detect all the
DME present in MTBE raffinate. We have found that refinery labs tend to underreport the
actual amount of DME present.

Steps should be taken to prevent conditions or upsets which increase the oxygenates in the
alkylation unit feed. The use of molecular sieve dryers to remove oxygenates and water
from the MTBE raffinate may be very attractive depending upon the concentration of
oxygenates in that stream.

If the reactors are experiencing a high corrosion rate due to oxygenates in the feed, one
possible remedy is to increase the acid spending strength up to about 92 wt%. This action
will increase the series acid rate through the unit and reduce the concentration of water in the
spent acid.

-4-
II. REACTION SECTION (Cont'd)

Commandment 5: Keep acidic streams dry, and "wet" streams neutral. Use corrosion-
resistant material where they mix.

Localized low acid concentration can occur in zones where wet hydrocarbon and acidic
streams are mixed, even though the final mixture may not be in the corrosive range. In this
context, a hydrocarbon stream is considered wet if it contains dissolved water. Free water
need not be present. By this definition, the olefin feed and recycle isobutane in most plants
are considered wet streams. In regions where wet and acidic streams are mixed, corrosion-
resistant material is substituted for carbon steel. Some examples are shown below in Figure
1. To resist corrosion in the Contactor reactor mixing zone, STRATCO typically uses Ni-
Resist 1 for the impeller, and Alloy 20 for the hydrocarbon feed nozzle. The piping wherein
the olefin feed and refrigerant recycle are combined is a point at which a wet stream and an
acidic stream are mixed is. (In this case the refrigerant recycle is the acidic stream, because
it contains a small amount of SO2.) STRATCO specifies Alloy 20 piping from two feet (610
mm) upstream of the feed/refrigerant recycle mix point to the Contactor reactor feed nozzle.
Figure 1
Zones where Wet Streams and Acidic Streams are Mixed

Olefin Feed
from
Coalescer
***=Alloy 20
***

2’
Emulsion 2’
to Settler *** Refrigerant
Coolant Recycle
Acid
Out Recycle
***

Ni-Resist 1

Coolant
In

-5-
II. REACTION SECTION (Cont'd)

Contactor Reactor Tube Bundle Fouling

The tube side of the Contactor reactor tube bundle sometimes becomes fouled, especially in
plants which are operated at high acid spending strength and/or above design capacity.
Contactor reactor tube bundle fouling is not directly related to one of the six commandments.
The cause is usually acid carryover from the acid settler, which becomes viscous and coats
the inside of the Contactor reactor tubes due to the lower temperature inside the tube bundle.
Tubeside fouling is most common in the Contactor reactor/settler system with the highest
strength acid, because higher strength acid is more viscous. The normal solution is to
temporarily stop feed to the affected Contactor reactor and allow it to warm up. Warming
the acid film inside the tubes reduces its viscosity, and allows it to be washed away when
feed is reintroduced.

III. REFRIGERATION SECTION AND DEPROPANIZER

The refrigeration section and depropanizer provide excellent examples for application of
Commandments 5 and 6:

Commandment 5: Keep acidic streams dry, and "wet" streams neutral. Use corrosion-
resistant material where they mix.

A. Refrigeration System

The refrigerant stream in an effluent or auto-refrigerated alkylation unit is acidic

because of the presence of sulfur dioxide (SO2) generated in the reaction section. SO2
will not cause corrosion of carbon steel unless there is water present. Corrosion in
the refrigeration system should therefore not be a problem, because under normal
circumstances no water should be present.

To insure against corrosion in the refrigeration section, potential sources of water

must be eliminated. A potential source of water is inadequate drying of refrigeration
section equipment prior to startup. Another potential source of water is a tube leak in
a water-cooled refrigerant condenser. It is important for the process pressure in the
refrigerant condensers to be higher than the cooling water pressure, so water will not
enter the process in the event of a leak.

-6-
III. REFRIGERATION SECTION AND DEPROPANIZER (Cont'd)

B. Depropanizer Systems

SO2 must be removed from the alkylation unit propane product in order to meet LPG
specifications, but there are differing schools of thought as to where it should be
removed. In a dry depropanizer system, SO2 is removed by caustic treating the
propane product. In a wet depropanizer system, it is removed by caustic treating the
depropanizer feed. Since the propane product is a considerably smaller stream than
the depropanizer feed, a dry depropanizer system usually has a smaller, less costly
caustic treatment system than a wet depropanizer system. Caustic consumption for
both systems should be about the same. Both the wet and dry depropanizer systems
have proponents, and both are used successfully.

1. Dry Depropanizer System

In a dry depropanizer operation, the column is fed directly from the

refrigerant accumulator. As long as no water is present, the SO2 which
concentrates in the depropanizer overhead will not cause corrosion.
However, if water is present it will also concentrate in the depropanizer
overhead system, and rapid corrosion can occur. When problems occur, it is
often difficult to track down the source of water.

Inadequate dryout of a dry depropanizer system prior to startup can lead to

corrosion. Another possible source of water is the propane product treatment
system downstream of the depropanizer. In alkylation units with
predominantly butylene feed, the amount of propane in the unit feed can be
so small that it is difficult to maintain a continuous flow of propane product.
If propane product flow is stopped, it may be possible for water from
downstream treatment to back up into the depropanizer overhead system. In
some refineries which send only butylene feed to the alkylation unit, a small
amount of propane product is recycled to the reactor section to insure the
flow of depropanizer overhead product is not interrupted.

Some refiners with dry depropanizer systems have found it necessary to use
stainless steel for condenser tube bundles and the top trays of the
depropanizer. Others use salt dryers on the depropanizer reflux to be certain
the overhead system stays dry. One of the challenges of operating a dry
depropanizer system is that there is no easy way to make certain the system is
staying dry. Equipment inspection (or failure between inspections) is often
the first indication something has gone wrong.

-7-
III. REFRIGERATION SECTION AND DEPROPANIZER (Cont'd)

B. Depropanizer Systems (Cont'd)

Commandment 6: Maintain good treating operations.

2. Wet Depropanizer System

With a wet depropanizer system, the depropanizer feed from the refrigerant
accumulator is treated to remove SO2 before it is charged to the column. The
treatment system, as shown in Figure 2, typically consists of a caustic wash
followed by a coalescer to prevent any caustic carryover to the depropanizer
tower.

The caustic wash removes SO2 via the following reactions:

SO2 + NaOH → NaHSO3

SO2 + 2 NaOH → Na2SO3 + H2O

The most important parameters in operation of the caustic wash are good
mixing, the ratio of caustic to hydrocarbon feed, and the caustic strength.

Figure 2
Wet Depropanizer Feed Treatment

-8-
III. REFRIGERATION SECTION AND DEPROPANIZER (Cont'd)

B. Depropanizer Systems (Cont'd)

2. Wet Depropanizer System (Cont’d)

A variety of mixing devices including eductors, mixing valves, and static

mixers have been used to improve contacting between the depropanizer feed
and the caustic. STRATCO specifies an Alloy 20 Koch SMV five stage
static mixer for this service. The static mixer is to be designed to produce a
300-400 micron calculated average droplet size at the outlet of the mixer.
Note that this droplet size is based on Koch’s correlations and might not be
applicable to mixers provided by other vendors. A typical ∆P for a mixer of
this type is 8-15 psi (0.6-1.1 kg/cm2). The static mixer pressure drop should
be monitored with a ∆P indicator. An unexplained reduction in pressure drop
is likely to indicate damage to the static mixer. In order to avoid corrosion
in the mixing zone, the piping from four (4) feet (1.2 meters) upstream of the
caustic/hydrocarbon mix point to the static mixer, and the piping from the
static mixer to the caustic wash drum is made of Alloy 20 (Refer to
Commandment 5). The caustic circulation rate is maintained in the range of
20-30% by volume of the depropanizer feed stream.

Fresh caustic strengths above 15 wt% should be avoided. Plugging of the

depropanizer feed/caustic static mixer and associated piping due to
precipitation of caustic salts has been known to occur when concentration of
the fresh caustic is too high. Fresh caustic strength is normally in the 10-12
wt% range. In batch systems the caustic should be replaced at a minimum
strength of 3 wt%, as measured by titration. The caustic solution will lose
water over time to the depropanizer feed stream. Concentrated caustic
solutions tend to emulsify and carry over, so makeup water should be added
to replace the water that is lost. The simplest way to do this is to add process
water as necessary to hold a constant interface level in the caustic wash
drum. It is important to use water, NOT fresh caustic, to maintain the
interface level. Another approach to water makeup is to add water to
maintain the specific gravity of the caustic in the range of 1.10 to 1.15.

Properly maintained and operated, a wet depropanizer system will produce

on-specification LPG product with little or no corrosion in the depropanizer
overhead system. However, poor depropanizer feed treatment will allow SO2
to break through while supplying a source of water, and accelerated corrosion
will result.

-9-
III. REFRIGERATION SECTION AND DEPROPANIZER (Cont'd)

B. Depropanizer Systems (Cont'd)

2. Wet Depropanizer System (Cont’d)

One of the advantages of a wet depropanizer system is that it is fairly easy to

make certain the caustic wash system is effective. The primary indicator of
good caustic wash operation is the pH of the water collected in the
depropanizer overhead accumulator boot. STRATCO recommends sampling
this stream at least once a day. With adequate treatment its pH should not fall
below 6-6.5. It is also wise to monitor the iron content of the accumulator
water, and the volume of water drained per day from the boot.

C. Refrigerant Recycle Pumps

Some refiners have reported fouling of refrigerant recycle pump suction strainers
with ice crystals. This problem usually occurs during startup. It is caused by the
presence of water in the flash drum, which normally operates at about 16°F (-9°C).
The source of this water is usually traced to inadequate drying of the refrigeration
system prior to startup, or to poor operation of a wet depropanizer which allows
water into the flash drum via the depropanizer bottoms stream.

There are two basic responses to this problem. One, the bottoms temperature of the
depropanizer can be increased to flash more water out of the bottom of the tower.
Two, the flash drum pressure can be increased such that the temperature in the
process is not below freezing 32°F (0°C). The former is the preferable of the two
since it does not result in a hotter Contactor reactor temperature.

- 10 -
IV. NET EFFLUENT TREATING AND FRACTIONATION

The net effluent stream from the suction trap/flash drum contains traces of free acid, as well
as alkyl sulfates formed by the reaction of sulfuric acid with olefins. If these impurities are
not removed they can cause corrosion and fouling of heat exchangers and fractionation
equipment.

Commandment 6: Maintain good treating operations.

A. Net Effluent Treating Systems

The most common net effluent treatment systems consist of either a caustic wash
followed by a water wash, or an acid wash followed by an alkaline water wash.

1. Caustic Wash/Water Wash Systems

The preferred caustic wash/water wash system configuration is shown in

Figure 3 on the following page. Free acid can be neutralized and mono-alkyl
sulfates (acid esters) decomposed at low temperature in the caustic wash.
Decomposition of di-alkyl sulfates (neutral esters) generally requires a
somewhat elevated temperature, and it is therefore important for at least one
of the treatment steps to operate at a temperature of about 120°F (49°C).
Caustic embrittlement is usually a concern if the caustic wash is heated. It is
less of a concern if the water wash is heated, because of its much lower
caustic concentration.

The following saponification, hydrolysis, and neutralization reactions are

examples of those that occur in the caustic wash drum:

Saponification (C4H9)HSO4 + NaOH → NaSO4(C4H9) + H2O

Hydrolysis (C4H9)HSO4 + H2O → C4H9OH + H2SO4

Neutralization H2SO4 + 2 NaOH → Na2SO4 + 2 H2O

- 11 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

A. Net Effluent Treating Systems (Cont'd)

1. Caustic Wash/Water Wash Systems (Cont’d)

Figure 3
Caustic Wash/Water Wash System

In batch operation, the minimum spent caustic strength recommended is 3

wt%. Makeup process water should be added to hold constant interface level,
or to keep the specific gravity of the caustic solution in the range of 1.10 to
1.15. Caustic circulation should be 20-30% by volume of the net effluent
rate.

- 12 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

A. Net Effluent Treating Systems (Cont'd)

1. Caustic Wash/Water Wash Systems (Cont’d)

A variety of mixing devices have been used to improve contacting of the net
effluent and caustic solution. STRATCO currently recommends an Alloy 20
Koch SMV static mixer with 5 mixing elements. The static mixer should be
designed to produce a 300-400 micron calculated average droplet size at the
outlet of the mixer. Note that this droplet size is based on Koch’s correlations
and might not be applicable to mixers provided by other vendors. A typical
∆P for a mixer of this type is 8-15 psi (0.6-1.1 kg/cm2). Turndown is an
important consideration in sizing a static mixer for this service. This pressure
drop should be monitored with a ∆P indicator to verify the integrity of the
mixer. The overall mixing pressure drop is sometimes adjusted through the
use of a globe valve, either in series or in parallel with the static mixer. Alloy
20 piping is recommended from 4 feet (1.2 meters) upstream of the
caustic/net effluent mix point to the static mixer, and from the static mixer to
the settling drum (see Commandment 5).

An unheated water wash will dilute any caustic carried over from the net
effluent caustic wash. However, it is preferable for the water wash to be
heated to aid in decomposition of di-alkyl sulfates in the net effluent. Most
water wash systems are not heated, so this is a potential revamp project for
better treating and possibly improved heat integration.

The following types of reactions should occur in a heated water wash:

(C4H9)2SO4 + NaOH → NaSO4(C4H9) + C4H9OH

(C4H9)2SO4 + 2 H2O → 2 C4H9OH + H2SO4

H2SO4 + 2 NaOH → Na2SO4 + 2 H2O

- 13 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

A. Net Effluent Treating Systems (Cont'd)

1. Caustic Wash/Water Wash Systems (Cont’d)

The recommended net effluent outlet temperature from the water wash drum
is 120°F (49°C). The preferred method of heating the water wash is to heat
the recirculating water stream in a heat exchanger against either low pressure
steam and/or hot alkylate product. Hot makeup water provides a heat source
in some plants. Net effluent preheaters upstream or downstream of the water
wash were incorporated in some earlier designs, but there were problems
with heat exchanger corrosion and fouling in this service.

The type of static mixer recommended for the water wash is the same as that
recommended for the caustic wash. Static mixer pressure drop should be
monitored.

The pH of an unheated water wash drum should naturally tend to be alkaline

because of caustic carryover from the caustic wash. However, if the water
wash is heated it is important to be certain enough caustic is present to
neutralize the liberated acid. Caustic should be added to maintain a pH of 11
+ 1. If pH control is employed, this unit operation is the same as the alkaline
water wash used after the acid wash in current designs. Makeup water
should be added as needed to maintain the water conductivity in the water
wash at 5,000-8,000 µmhos/cm. Conductivity is a function of total dissolved
solids (TDS) concentration. Excessive solids (>10,000 µmhos/cm) can cause
a tight emulsion which can carry over from the water wash into the
fractionation section. A high TDS level has also been associated with
corrosion and fouling of heat exchangers in the circulating water system.

- 14 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

A. Net Effluent Treating Systems (Cont'd)

2. Acid Wash/Alkaline Water Wash Systems

The net effluent treating system employed in most of the recent STRATCO
designs is an acid wash followed by an alkaline water wash. This type of
system is illustrated below in Figure 4.

Mono-alkyl sulfates (acid esters) have a much higher affinity for sulfuric acid
than for hydrocarbon. For this reason it is possible to extract most of the
mono-alkyl sulfates and some of the di-alkyl sulfates from the net effluent
stream with fresh sulfuric acid. The acid wash extracts mono-alkyl sulfates
from the net effluent stream and recycles them to the reaction section with
the incoming sulfuric acid catalyst. The main advantage of a net effluent acid
wash over a caustic wash is reduced caustic consumption and waste
generation. In theory, the acid wash also has a yield advantage. In practice,
however, this yield advantage is too small to measure.

Figure 4
Acid Wash/Alkaline Water Wash System

STATIC MIXER NET EFFLUENT

STATIC MIXER
NET EFFLUENT *** *** TO DIB
*** * ***
* *
4′ 4′ *
* 4′ * * 4′
* *
*
* *
E.P.

ALKALINE
FC LC ACID WASH pH LC
WATER WASH
FRESH ACID
FE
CORIOLIS ALKYLATE
PRODUCT
FRESH ACID FC DIB/DEB
TO CONTACTOR
BOTTOMS
REACTORS FE
CORIOLIS
FC SPENT
ALKALINE
ACID CIRC. FC WATER
MAKEUP PUMP
WATER
FC
CAUSTIC FROM DEPROP.
CAUSTIC WASH DRUM

ALKALINE WATER
CIRCULATION PUMP

- 15 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

A. Net Effluent Treating Systems (Cont'd)

2. Acid Wash/Alkaline Water Wash Systems (Cont’d)

The acid wash is operated with an acid recirculation rate of about 3-5% by
volume of the net effluent rate. A variety of mixing devices have been used
for contacting the net effluent and acid phases. STRATCO recommends an
Alloy 20 Koch SMV five stage static mixer for this service. The static mixer
is to be designed to produce a 300-400 micron calculated average droplet size
at the outlet of the mixer. Note that this droplet size is based on Koch’s
correlations and are not applicable to mixers provided by other vendors. A
typical pressure drop for a mixer of this type is 8-15 psi (0.6-1.1 kg/cm2).
This pressure drop should be monitored with a ∆P indicator to verify the
integrity of the mixer.

To resist corrosion in the mixing zone, Alloy 20 is the recommended piping

material from 4 feet (1.2 meters) upstream of the net effluent/acid mix point
to the acid wash drum (see Commandment 2). The flowrate of fresh acid
from storage to the acid wash drum is reset by the acid wash drum level
controller. The acid wash drum is equipped with an electrostatic precipitator
to minimize acid carryover (in most of STRATCO’s latest designs, a less
costly acid wash coalescer is substituted for an acid wash drum with EP).

The functions of the alkaline water wash are 1) to neutralize any acid carried
over from the acid wash, 2) to remove mono-alkyl sulfates not extracted in
the acid wash, and 3) to decompose di-alkyl sulfates. The chemical reactions
are the same as those previously described for the caustic wash/water wash
system.

The recommended net effluent outlet temperature from the alkaline water
wash is 120°F (49°C). In most of the recent STRATCO designs the
recirculating alkaline water stream is heated by heat exchange with hot
alkylate product. The alkaline water recirculation rate should be 20-30% by
volume of the net effluent rate. To resist corrosion in the mixing zone, Alloy
20 is the recommended piping material from 4 feet (1.2 meters) upstream of
the net effluent/alkaline water mix point to the drum. An Alloy 20 Koch
SMV five stage static mixer is recommended. Static mixer pressure drop
should be monitored, and typically should be within the range of 8-15 psi
(0.6-1.1 kg/cm2). Caustic is added to the alkaline water wash on pH control.
The pH should be maintained at 11 + 1.

- 16 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

A. Net Effluent Treating Systems (Cont'd)

2. Acid Wash/Alkaline Water Wash Systems (Cont’d)

Makeup process water should be added as needed to maintain the water

conductivity in the alkaline water wash at 5,000-8,000 µmhos/cm. The
alkaline water blowdown rate is reset by the interface level controller on the
drum.

In STRATCO’s latest designs, the net effluent from the alkaline water wash
drum flows to a water wash coalescer in which traces of caustic and TDS
from the alkaline water wash are removed. Fresh process water is fed to the
net effluent stream upstream of the coalescer to remove any alkaline water
droplets that may carry over from the alkaline water wash drum. The water
added is then removed in the coalescer.

3. Bauxite Treatment

Bauxite treating is an effective way to remove free acid and alkyl sulfates
from the net effluent. The bauxite bed adsorbs the sulfur compounds without
adding water to the net effluent, as occurs in a caustic or water wash. Since
the deisobutanizer feed and the recycle isobutane remain dry, both corrosion
and acid consumption should be reduced, and alkylate octane should be
higher. An upstream acid wash coalescer can reduce the load on the bauxite
beds by extracting most of the alkyl sulfates without adding water to the net
effluent stream. Figure 5 shows a sketch of an acid wash/bauxite treatment
system.

- 17 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

A. Net Effluent Treating Systems (Cont'd)

3. Bauxite Treatment (Cont’d)

Figure 5
Acid Wash/Bauxite Treatment System

FRESH ACID

NET EFFLUENT
FROM FEED/
EFFLUENT OPEN CLOSED
EXCHANGER ACID
COALESCER
STATIC
MIXER

FRESH ACID
TO CONTACTOR
REACTORS
BAUXITE BAUXITE
TREATER TREATER
(OPERATING) (ON STANDBY)

NET EFFLUENT
CLOSED TO DIB PREHEAT
EXCHANGER
OPEN

The bauxite must be regenerated periodically (approximately every 15 days),

and replaced after 35-50 regenerations. The bauxite is regenerated by
washing with hot water and drying with superheated steam followed by dry
hydrocarbon vapor. The used wash water is a waste stream which must be
neutralized with caustic and then disposed of properly. Only two bauxite
treatment systems are known to be in commercial operation in alkylation
service. One of these has an acid wash upstream, which reduces loading on
the bauxite beds and therefore reduces the required frequency of
regeneration.

- 18 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

B. Deisobutanizer Tower

The main purpose of the various net effluent treating systems is to prevent corrosion
and fouling in the fractionation section. In most sulfuric acid alkylation units the first
tower in the fractionation section is the deisobutanizer (DIB). If alkyl sulfates are
allowed to stay in the DIB feed, the high temperature in the bottom of the tower and
the reboiler can cause them to react, forming solid "tars" and releasing SO2. The
solids can foul the DIB reboilers and lower trays. The SO2 goes to the overhead
system of the tower, where it combines with water to cause corrosion. DIB reboiler
fouling and corrosion in the DIB overhead system are some of the most common
problems in sulfuric acid alkylation units.

The primary indicator of the effectiveness of the net effluent treatment system
(except for bauxite treating systems) is the pH of the water collected in the DIB
accumulator boot. With good net effluent treatment the pH of this water should not
drop below 6-6.5. STRATCO recommends checking the pH of the DIB overhead
accumulator at least once per day. It is also wise to monitor the iron content of the
accumulator water, and the volume of water drained per day from the boot.

Some refiners have taken steps to make up for inadequate net effluent treatment.
Ammonia or some type of amine is sometimes added to the DIB overhead system to
control pH and reduce corrosion. Filming agents are also used for corrosion control.
The solids which foul the reboilers and lower trays are water soluble, and some
refiners have developed on-line water washing procedures for the column. These
measures, while useful for extending run times and equipment life, are weak
substitutes for good reactor and net effluent treatment operations.

- 19 -
IV. NET EFFLUENT TREATING AND FRACTIONATION (Cont'd)

B. Deisobutanizer Tower (Cont’d)

Another common source of DIB fouling is aqueous carryover from the net effluent
treatment system. The water vaporizes inside the tower leaving behind salt deposits,
usually in the region of the feed tray. This problem can occur 1) if the TDS level in
the water wash or alkaline water wash is not properly controlled, 2) if the water wash
drum is undersized for the current throughput, or 3) if the water wash mixing device
produces too many very fine droplets. In the short term, this problem can be
addressed by water washing the tower. In the long run it is better to correct the
problem in the net effluent treatment system. If the water wash or alkaline water
wash drum is undersized, the usual remedy is to add a coalescer or electrostatic
precipitator to eliminate the carryover. In addition, coalescing media can be added
into the water wash drum to increase water removal from the hydrocarbon phase. In
STRATCO’s latest designs, a water wash coalescer is installed between the alkaline
water wash drum and the DIB tower.

Contamination of the alkylation unit feed with aromatics can also cause water to
carry over from the water wash. The aromatics are sulfonated in the reactors, and
can then react with caustic to produce "soaps," which cause tight emulsion in the
water wash. The only sure remedy for this problem is to keep aromatics out of the
alkylation unit feed stream.

Most treatment problems are related to inadequate mixing of the organic and water
(or organic and acid) phases. However, it is important to note that problems can
also be caused by mixing that is too intense, forming an emulsion which will not
separate in the drum. This can lead to caustic or water carryover which can foul the
trays in the top of the DIB, or to acid carryover from the acid wash which can cause
severe corrosion.

In one alkylation unit the mixing in the acid wash static mixer was too intense,
causing acid carryover, while the mixing in the alkaline water wash was inadequate.
Interestingly, pH control in the alkaline water wash appeared to function normally.
The acid was apparently so finely dispersed in the net effluent that the alkaline water
wash could not remove it. The problem was uncovered when the DIB feed line
failed after less than two weeks of operation. Subsequent inspections also uncovered
substantial corrosion of the alkaline water wash drum in the region of the interface.
The problem was eliminated by correcting the intensity of mixing in the two treaters,
and by adding an electrostatic precipitator to the acid wash.

- 20 -
V. ALKYLATION UNIT BLOWDOWN SECTION

The blowdown section of the alkylation unit provides good examples for application of
Commandments 3 and 5:

Commandment 3: Don't let acid sit stagnant.

Stagnant acid is a potential source of fouling problems in the acid blowdown system. Level
taps must be flushed periodically to prevent plugging and false level readings. It is also
advisable to run the spare acid blowdown pump regularly to prevent solids from plugging the
pump suction and discharge lines.

Commandment 5: Keep acidic streams dry, and "wet" streams neutral. Use corrosion-
resistant material where they mix.

A sulfuric acid alkylation unit requires two relief systems. Wet, non-acidic vents and reliefs
can go directly to the refinery flare, but a separate system is required for acidic reliefs from
the reaction section, the refrigeration section, and the acid wash drum. A constant sweep of
dry fuel gas or nitrogen on the acid relief system is advisable to prevent stagnation.

Figure 6
Acid Blowdown Drum and Blowdown Vapor Scrubber

- 21 -
V. ALKYLATION UNIT BLOWDOWN SECTION (Cont'd)

In current designs, the acid blowdown drum is the common collection point for the acid
relief system. As shown in Figure 6, the vent gas from the acid blowdown drum is sent to
the blowdown vapor scrubber where it is scrubbed with caustic to neutralize acidic
components (i.e. SO2). Vapor from the blowdown vapor scrubber goes to the refinery flare.

The blowdown vapor scrubber is a carbon steel tower typically equipped with six trays. To
prevent localized corrosion at the vapor feed nozzle, a section of the vapor feed line is
constructed of Alloy 20 and equipped with a spray nozzle to begin neutralization of the
vapor stream before it enters the scrubber. Caustic is recirculated from the bottom of the
scrubber to the top tray and the inlet spray nozzle by the scrubber circulating caustic pump.

The acid blowdown drum receives spent acid continuously when the unit is in operation, and
is also the destination for spent acid from the Contactor reactors and acid settlers during
shutdowns. Since the acid strength in the acid blowdown drum should not be below about 89
wt% for most units, corrosion of the carbon steel vessel should not be a serious problem.
However, it is important to ensure that the spent acid remains dry. Potential sources of water
must be eliminated through good design and operating practices.

The acid blowdown system should be dried thoroughly before it is put into service. Water
could potentially enter the acid blowdown drum via the blowdown system from the reaction
section, because water and caustic connections are included for neutralization and flushing of
the Contactor reactors and acid settlers during a shutdown. It is important for the water and
caustic connections to be blinded when the acid blowdown drum is in operation, and for the
spent acid line to the blowdown drum to be blinded when water and caustic are in use.

We must also be on the lookout for unusual sources of water. In most plants the acid
knockout pot on the suction trap is connected to the acid blowdown drum. This should be
fine in normal operation, but in one unit while alkylate was being recycled during startup
some water got into the suction trap because of a distillation upset. This water was sent to
the acid blowdown drum, which became hot to the touch because of the heat of dilution. In
other plants sight glass drains from the reaction section are routed to the acid blowdown
drum. Water could potentially enter the acid blowdown system via these drains during
maintenance.

In another unit, water got into the acid sump drum after purging acid lines with water during
unit startup. After acid was drained from the startup acid lines through the strong acid sewer,
the acid mixed with water in the sump drum. The result was catastrophic failure of the drum.

- 22 -
V. ALKYLATION UNIT BLOWDOWN SECTION (Cont'd)

Another potential source of water to the acid blowdown drum is reverse flow of caustic and
water from the blowdown vapor scrubber. It is important for the vapor line from the acid
blowdown drum to the blowdown vapor scrubber to be routed and sloped so that reverse
flow of liquid caustic cannot occur. STRATCO recommends to design a loop seal in this
vapor line, elevated above the hydrocarbon relief header, so caustic would overflow
preferentially to the flare knockout drum instead of to the acid blowdown drum in the event
the blowdown vapor scrubber is flooded with caustic. The dry nitrogen or fuel gas sweep on
the acid relief system should help prevent reverse flow of wet vapor from the scrubber,
which could occur when the acid blowdown drum is being pumped down.

VI. CONCLUSION

Most corrosion and fouling problems in sulfuric acid alkylation units can be controlled
by adhering to the six common sense rules presented at the beginning of this paper.
Many problems can be avoided by careful monitoring of operating parameters such as
feed coalescer performance, treating system performance, and the pH of water collected
in distillation column overhead accumulators. Particular caution is warranted in plants
operated significantly above design capacity. Corrosion and fouling should be
considered as part of any plant expansion study.

- 23 -
 APPENDIX U
PETROLEUM REFINING

Ui
Uii
 APPENDIX U

PETROLEUM REFINING

GREGORY R. RUSCHAU, PH.D.1 AND MOHAMMED A. AL-ANEZI2

SUMMARY AND ANALYSIS OF RESULTS

Corrosion Control and Prevention

Petroleum refining is an industry that is undergoing intense amounts of scrutiny in the United States from
regulatory agencies and environmental groups. As a result, releases of pollutants caused by corrosion leaks are
becoming a high-consequence event. The Clean Air Act of 1990 has forced refineries to implement a number of
costly measures to reduce their impact on the environment, both in the types of products they produce and the
manner in which they operate.

The total cost of corrosion control in refineries is estimated at $3.692 billion. Of this total,
maintenance-related expenses are estimated at $1.767 billion annually, vessel turnaround expenses account for
$1.425 billion annually, and fouling costs are approximately $0.500 billion annually. The costs associated with
corrosion control in refineries include both processing and water handling. Corrosion-related issues regarding
processing include the handling of organic acids (broadly referred to as napthenic acid corrosion) and sulfur species,
particularly at elevated temperatures, as well as water carried over in processing vessels and pipelines. Water
handling includes concerns with corrosives such as H2S, CO2, chlorides, and high levels of dissolved solids.

Opportunities for Improvement and Barriers to Progress

As with oil production, the lifeblood of a refinery is the production system. Failure in any processing vessel,
particularly the major feedstock lines, costs significantly more in lost production than the cost of prevention and
maintenance. Unlike oil and gas production, refining margins are dictated on both ends by commodity prices since
the input feedstock crude oil is purchased at the market price and the output product is sold at each individual
commodity price.

Because the economics of refining are wholly dependent on world market prices, the amount spent on
corrosion control is dictated by current economic conditions in the industry. Since 1981, the number of operating
refineries in the United States has dropped from 324 to 163. The industry has seen a trend toward refining more
highly acidic oils (which can be refined at a higher margin) since the early 1990s, which increases potential
corrosion problems, but may extend the economic life of some existing refineries.

Recommendations and Implementation Strategy

The majority of pipelines and vessels in refineries are constructed of carbon steel. Opportunities for significant
savings exist through the use of low-alloy steels and alloy-clad vessels, particularly as increasingly higher fractions
of acidic crude are refined.

1
CC Technologies Laboratories, Inc., Dublin, Ohio.
2
Saudi Arabian Oil Company (Saudi ARAMCO), Dhahran, Saudi Arabia.

Uiii
Appendix U – Petroleum Refining

Increasing regulation and pressure from environmental groups have essentially forced the refiners to
implement defensive strategies. This is compounded by overseas market forces such as the Organization of
Petroleum Exporting Countries (OPEC), which can control the price of feedstock crude oil, making long-term
planning difficult. In a commodity price-driven industry that is struggling to compete in the world market,
investment in more effective corrosion control strategies often takes a backseat to across-the-board cost-cutting
measures.

Summary of Issues

Federal regulations such as the Clean Air Act of 1990 have increased
operating costs due to stricter controls on releases. In addition, more
Increase consciousness of corrosion acidic crude oil is being refined because of the higher net margins
costs and potential savings. possible; a stronger approach to corrosion control will enable these more
aggressive crudes to be safely refined in the United States, otherwise, the
refining industry will continue to move overseas.
A longer-term vision must be incorporated into facility design and
Change perception that nothing can maintenance to enable U.S. refiners to remain competitive. This includes
be done about corrosion. the use of some exotic materials, such as ceramics, which can provide a
longer service life in high-temperature operations.
More efficient processing vessel design would reduce the carryover of
Advance design practices for better corrosives from one process to the next. Improved water separation, CO2
corrosion management. stripping, etc. would help isolate the problem areas and would allow
corrosion control efforts to be focused farther upstream.
Fitness-for-service principles will need to be applied to vessel inspections
Change technical practices to
rather than following existing protocol, which may be inadequate. Risk-
realize corrosion cost-savings.
based models would enable the maintenance staff to prioritize inspections.
Management may have to shift its focus from ensuring compliance with
Change policies and management
existing regulations to a more active strategy to prevent releases. Zero-
practices to realize corrosion
leak policies and programs would be implemented in plants to emphasize
cost-savings.
commitment to this strategy.
Flexible life prediction models are needed that can show how a change in
Advance life prediction and the feedstock crude affects all vessels downstream. Also needed are
performance assessment methods. improved inspection and monitoring techniques for in-plant piping
systems, both for aboveground and buried lines.
Processes in refineries are largely computer-controlled, but corrosion
Advance technology (research, control methods lag behind in technology. Computer-aided mitigation
development, and implementation). systems, perhaps integrated with existing process control modules, could
be used to track the changing corrosivity of existing processes.
Requiring contract services such as nondestructive inspection companies,
maintenance painters, and corrosion control specialists to provide NACE-
Improve education and training for
certified personnel or at least personnel who meet some minimum
corrosion control.
training/education requirements before they are allowed to work on-site
would improve the level of knowledge in the industry.

Uiv
 Appendix U – Petroleum Refining

TABLE OF CONTENTS
SECTOR DESCRIPTION.........................................................................................................................................U1

REFINING CAPACITY OF THE UNITED STATES..........................................................................................U1

Refined Products ..........................................................................................................................................U2
Types of Crude Oil .......................................................................................................................................U3
Elements of the Refining Operation .............................................................................................................U4

AREAS OF MAJOR CORROSION IMPACT.......................................................................................................U6

Water-Related Side Corrosion......................................................................................................................U6
Processing-Related Corrosion ......................................................................................................................U6
Naphthenic Acid Corrosion..........................................................................................................................U7
Sulfur............................................................................................................................................................U7

CORROSION CONTROL METHODS ...................................................................................................................U8

Materials in Refinery Construction ..............................................................................................................U8
Carbon Steel...................................................................................................................................U9
Austenitic Stainless Steel ...............................................................................................................U9
Ferritic and Martensitic Steels .....................................................................................................U10
Other Alloys.................................................................................................................................U10

CORROSION MANAGEMENT ............................................................................................................................U10

Economics of Refining...............................................................................................................................U10
Capital Expenditures ..................................................................................................................................U11
Operational Expenditures ...........................................................................................................................U12
Fouling .......................................................................................................................................................U13
Acidic Crude Oils.......................................................................................................................................U14
Failure Costs...............................................................................................................................................U15

CASE STUDY.........................................................................................................................................................U15
Corrosion-Related Failure in Refinery .......................................................................................................U15

REFERENCES .........................................................................................................................................................U16

LIST OF FIGURES
Figure 1. Past and predicted future refining capacity in the United States....................................................U1

Figure 2. 1996 Outputs from refineries by end-product usage ......................................................................U3

Figure 3. Flowchart diagram of a typical refining process ............................................................................U5

Figure 4. Margins of U.S. refiners since 1977.............................................................................................U11

Figure 5. Incremental costs for corrosion control of carbon steel distillation column ................................U14

Figure 6. Stress corrosion cracking near a weld ..........................................................................................U15

Uv
Appendix U – Petroleum Refining

LIST OF TABLES
Table 1. U.S. daily average supply and disposition of crude oil and petroleum products, January 1997 ....U2

Table 2. Typical approximate characteristics and properties and gasoline potential of various crudes .......U4

Table 3. Comparison of the relative costs of various alloys.........................................................................U9

Table 4. Environmental costs at a refinery.................................................................................................U12

Uvi
 Appendix U – Petroleum Refining

SECTOR DESCRIPTION

Petroleum is the single largest source of energy for the United States. When measured in British thermal units,
the nation uses twice as much petroleum than either coal or natural gas, and four times more petroleum than nuclear
power, hydroelectricity, and other renewable energy sources. On average, every citizen in the United States
consumes 9.1 kg (20 lb) of petroleum per day. This primary dependence on petroleum for energy has been a reality
for decades, with petroleum's share of the domestic energy mix peaking at 49 percent in 1977.

REFINING CAPACITY OF THE UNITED STATES

U.S. refineries represent approximately 23 percent of world production. The United States has the largest
refining capacity in the world, with 163 operating refineries, having declined from a high of 324 refineries in 1981
and 205 refineries in 1990.(1)

Most refineries in the United States are concentrated on the west and gulf coasts, primarily due to access to
major sea transportation and shipping routes. The majority of the oil distillation capacity is currently centered in
large, integrated companies with multiple refining facilities. About 25 percent of all facilities are small operations
producing fewer than 50,000 barrels per day, representing 5 percent of the total output of petroleum products
annually.

In 1970, U.S. refineries supplied just under 15 million barrels of refined product per day. In 1996, U.S.
refiners supplied more than 18 million barrels per day of refined petroleum products. Total daily crude oil refining
capacity by the end of 1999 was 16,511,871 barrels per day. U.S. refiners rely on both domestic and foreign
producers for crude oil. Historical trends over the last 10 years indicate that imports of crude oil have been rising
steadily.

Future refining capacity in the United States is predicted to increase slightly and level off in the next 20 years,
as shown in figure 1. The curve illustrates how the United States experienced a steep decline in refining capacity in
the years following 1981. Between 1981 and 1989, the number of U.S. refineries fell from 324 to 204, representing
a loss of 3 million barrels per day (MMBD) in operable capacity, and a concomitant increase in refining capacity
utilization from 69 to 86 percent.
Millions of barrels per day

Figure 1. Past and predicted future refining capacity in the United States.(2)

U1
Appendix U – Petroleum Refining

Refined Products

Table 1 shows the average daily throughput of U.S. refineries in 1997.(3) On an annual basis, this translates
into a total of 5.7 billion barrels3 of refined product. Approximately 90 percent of all crude oil entering a petroleum
refinery is converted to fuel products, with the remaining 10 percent divided into non-fuel products such as asphalt,
lubricants, and waxes and petrochemicals such as polymer feedstocks and industrial solvents. Gasoline production
alone accounts for more than 46 percent of all production, as shown in figure 2.

Table 1. U.S. daily average supply and disposition of crude oil and petroleum products, January 1997.(3)

FIELD REFINERY UNACCOUNTED-FOR

COMMODITY PRODUCTION PRODUCTION IMPORTS
(thousand barrels per day) (thousand barrels per day) (thousand barrels per day)
Crude Oil 6,402 7,492

NGLs and LRGs* 1,782 528 246

Pentanes Plus 302 53
LPGs** 1,480 528 193
Ethane/Ethylene 634 26
Propane/Propylene 520 519
N Butane/Butylene 165 -28
161 11
OTHER LIQUIDS 267 740
Other Hydrocarbons/Oxy 247 77
Ounfinished Oils 421
Mogas Blend. Comp.*** 242
Avgas Blend. Comp.**** 20

FINISHED PETRO PROD. 19 15,075 1,285

Finished Mogas 19 7,288 320
Reformulated 2,217 136
Oxygenated 134 0
Other 4,937 184

Finished Avgas 16 0
Jet Fuel 1,491 100
Naptha-Type 0
Kerosene-Type 1,491 100
Kerosene 118 3
Distilate Fuel Oil 3,119 293
≤0.05 Sulfur 1,751 94
>0.05 Sulfur 1,368 198
Residual Fuel Oil 801 211
Naptha Petro Feed 180 106
Oth Oils Petro Feed 240 206
Special Napthas 47 10
Lubricants 168 7
Waxes 21 1

3
1 barrel = 158 L.

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 Appendix U – Petroleum Refining

Table 1. U.S. daily average supply and disposition of crude oil and petroleum products, January 1997 (continued).(3)

FIELD REFINERY UNACCOUNTED-FOR

COMMODITY PRODUCTION PRODUCTION IMPORTS
(thousand barrels per day) (thousand barrels per day) (thousand barrels per day)
Petroleum Cake 638 2
Asphalt & Road Oil 322 26
Still Gas 585
Misc. Products 41
TOTAL 8,470 15,603 9,763
*Natural Gas Liquids and Lead Replacement Gasolines
**Liquefied Petroleum Gas
***Motor Vehicle Fuel
****Aviation Fuel

Other Fuels
13.7%

Other Products
14.8%
Gasoline
46.5%

Fuel Oils
25.0%

Figure 2. 1996 Outputs from refineries by end-product usage.

Types of Crude Oil

Crude oils are complex mixtures containing many different hydrocarbon compounds that vary in appearance
and composition from one oil field to another. Crude oils range in consistency from water to tar-like solids, and in
color from clear to black. An average crude oil contains about 84 percent carbon, 14 percent hydrogen, 1 to
3 percent sulfur, and less than 1 percent each of nitrogen, oxygen, metals, and salts. Crude oils are generally
classified as paraffinic, naphthenic, or aromatic based on the predominant proportion of similar hydrocarbon
molecules. Mixed-base crudes have varying amounts of each type of hydrocarbon. Refinery crude base stocks
usually consist of mixtures of two or more different crude oils. Table 2 lists some typical properties for crude oil
sources from around the world.

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Appendix U – Petroleum Refining

Crude oils are also defined in terms of API (American Petroleum Institute) gravity number. The higher the
API gravity number, the lighter the crude. For example, light crude oils have high API gravities and low specific
gravities. Crude oils with low carbon, high hydrogen, and high API gravity are usually rich in paraffins and tend to
yield greater proportions of gasoline and light petroleum products. Crude oils with high carbon, low hydrogen, and
low API gravities are usually rich in aromatics. Crude oils that contain appreciable quantities of hydrogen sulfide or
other reactive sulfur compounds are called sour. Those with less sulfur are called sweet. Some exceptions to this
rule are the West Texas crudes, which are always considered sour regardless of their H2S content, and the Arabian
high-sulfur crudes, which are not considered sour because their sulfur compounds are not highly reactive.

For refining operations, the acidity of the crude oil is an important consideration for economic reasons. A
number of organic acids may be present in crude oil feedstocks. The extra costs associated with handling high-acid
crudes can be offset by a lower feedstock cost. Acidity is defined in terms of the total acid number (TAN), which is
a measure of the number of milligrams of potassium hydroxide (KOH) needed to neutralize 1 g of sample. A TAN
exceeding 1.5 to 1.8 mg KOH/g is considered corrosive; however, corrosion problems can occur in crudes with
TAN numbers as low as 0.3 for several reasons, including velocity and the nature of the acidic species present.

Table 2. Typical approximate characteristics and properties and gasoline potential of various crudes.(4)

API OCTANE
CRUDE PARRAFINS AROMATICS NAPTHENES SULFUR NAPH. YIELD
GRAVITY NUMBER
SOURCE (%VOL) (%VOL) (%VOL) (%WT) (% VOL)
(APPROX.) (TYPICAL)
Nigerian
37 9 54 0.2 36 28 60
(light)
Saudi (light) 63 19 18 2 34 22 40
Saudi (heavy) 60 15 25 2.1 28 23 35
Venezuela
35 12 53 2.3 30 2 60
(heavy)
Venezuela
52 14 34 1.5 24 18 50
(light)
USA
Midcont. - - - 0.4 40 - -
Sweet
USA
(W.Texas 46 22 32 1.9 32 33 55
Sour)
North Sea
50 16 34 0.4 37 31 50
(Brent)

Elements of the Refining Operation

Petroleum refining begins with the desalting (dehydration) of feedstock followed by distillation, or
fractionation, of crude oils into separate hydrocarbon groups. The resultant products are directly related to the
characteristics of the crude oil processed. Most distillation products are further converted into more usable products
by changing the size and structure of the hydrocarbon molecules through cracking, reforming, and other conversion
processes as discussed in this sector. These converted products are then subjected to various treatment and
separation processes, such as extraction, hydrotreating, and sweetening to remove undesirable constituents and
improve product quality. Integrated refineries incorporate fractionation, conversion, treatment, and blending
operations, and may also include petrochemical processing. An outline of the refining process is shown in figure 3.

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 Appendix U – Petroleum Refining

Figure 3. Flowchart diagram of a typical refining process.(5)

Crude oil often contains water, inorganic salts, suspended solids, and water-soluble trace metals. As a first step
in the refining process, to reduce corrosion, plugging, and fouling of equipment and to prevent poisoning the
catalysts in processing units, these contaminants must be removed by desalting (dehydration). The two most typical
methods of crude oil desalting – chemical and electrostatic separation – use hot water as the extraction agent. In
chemical desalting, water and chemical surfactants (demulsifiers) are added to the crude and heated so that salts and
other impurities dissolve into the water or attach to the water, and are then held in a tank where they settle out.
Electrical desalting is the application of high-voltage electrostatic charges to concentrate-suspended water globules
in the bottom of the settling tank. Surfactants are added only when the crude has a large amount of suspended
solids. Both methods of desalting are continuous. A third and less common process involves filtering heated crude
using diatomaceous earth.

After desalting, crude oil is continuously drawn from the top of the settling tanks and sent to the crude
distillation (fractionating) tower. Fractionation (distillation) is the separation of crude oil in atmospheric and
vacuum distillation towers into groups of hydrocarbon compounds of differing boiling-point ranges called fractions
or cuts.

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Appendix U – Petroleum Refining

Conversion processes change the size and/or structure of hydrocarbon molecules. These processes include
decomposition (dividing) by thermal and catalytic cracking, unification (combining) through alkylation and
polymerization, and alteration (rearranging) with isomerization and catalytic reforming.

Treatment processes are intended to prepare hydrocarbon streams for additional processing and to prepare
finished products. Treatment may include the removal or separation of aromatics and naphthenes, as well as
impurities and undesirable contaminants. Treatment may involve chemical or physical separation such as
dissolving, absorption, or precipitation using a variety and combination of processes, including desalting, drying,
hydrodesulfurizing, solvent refining, sweetening, solvent extraction, and solvent dewaxing.

Formulating and blending is the process of mixing and combining hydrocarbon fractions, additives, and other
components to produce finished products with specific performance properties.

Other refinery operations include light-end recovery, sour-water stripping, solid waste and wastewater
treatment, process-water treatment and cooling, storage and handling, product movement, hydrogen production, acid
and tail-gas treatment, and sulfur recovery. Auxiliary operations and facilities include steam and power generation;
process and fire water systems; flares and relief systems; furnaces and heaters; pumps and valves; supply of steam,
air, nitrogen, and other plant gases; alarms and sensors; noise and pollution controls; sampling, testing, and
inspecting; and laboratory, control room, maintenance, and administrative facilities.

AREAS OF MAJOR CORROSION IMPACT

A refinery operation may have in excess of 3,000 processing vessels of varying size, shape, form, and function.
In addition, a typical refinery has about 3,200 km (2,000 mi) of pipeline, much of which is inaccessible. Some of
these pipelines are horizontal; some are vertical; some are up to 61 m (200 ft) high; and some are buried under
cement, soil, mud, and water. The diameters range from 10 cm (4 in) up to 76 cm (30 in).

Water-Related Corrosion

Crude oil desalting and distillation generates considerable wastewater. Typical wastewater flow from a
desalter is approximately 8 L (2.1 gal) of water per barrel of oil processed. This water contains accelerative
corrosive components such as H2S, CO2, chlorides, and high levels of dissolved solids. The wastewater also
contains a fraction of crude oil, which may be recovered during the water treatment process.

In addition to generated wastewater, cooling water (either fresh water or saltwater) is used extensively in
refining operations. The corrosivity of the cooling water varies greatly depending on the process, so it is difficult to
describe typical cooling water problems; however, corrosivity is highly dependent upon the level and type of
dissolved solids and gases in the cooling water, including chlorides, oxygen, dissolved gases, and microbes.
Cooling water temperature can also affect corrosivity.

Processing-Related Corrosion

The top section of a crude unit can be subjected to a multitude of corrosive species. Hydrochloric acid, formed
from the hydrolysis of calcium and magnesium chlorides, is the principal strong acid responsible for corrosion in the
crude unit top section. Carbon dioxide is released from crudes typically produced in CO2-flooded fields and crudes
that contain a high content of naphthenic acid.

Low molecular fatty acids such as formic, acetic, propionic, and butanoic acids are released from crudes with a
high content of naphthenic acid. Hydrogen sulfide, released from sour crudes, significantly increases corrosion of

U6
 Appendix U – Petroleum Refining

the crude unit top section. Sulfuric and sulfurous acids, formed by either oxidation of H2S or direct condensation of
SO2 and SO3, also increase corrosion.

Mitigation of this type of corrosion is performed by process changes, material upgrading, design changes, and
injection of chemicals such as neutralizers and corrosion inhibitors. Process changes include any action to remove
or at least reduce the amount of acid gas present and to prevent accumulation of water on the tower trays. Material
upgrading includes lining of distillation tower tops with alloys resistant to hydrochloric acid. Design changes are
used to prevent the accumulation of water. They include coalescers and water draws. The application of chemicals
includes the injection of a neutralizer to increase the pH and a corrosion inhibitor. The presence of many weak
acids, such as fatty acids and CO2, can buffer the environment and require greater use of neutralizers. Excess
neutralizers may cause plugging of trays and corrosion under the salt deposits.

A dew-point probe is typically placed in a location at least 38 °C (100 °F) above the calculated dew-point
temperature. The probe elements are then cooled internally by cold-air injection and the temperature at which the
first liquid drop forms is determined for the actual conditions in the tower. The injection point and the amount of
chemicals used depend on the knowledge of the temperature in the tower where condensation starts. With the
number of corrosive species present, the calculated dew point may be much lower than the actual dew point.

Naphthenic Acid Corrosion

High-temperature crude corrosivity of distillation units is a major concern of the refining industry. The
presence of naphthenic acid and sulfur compounds considerably increases corrosion in the high temperature parts of
the distillation units and, therefore, equipment failures have become a critical safety and reliability issue.
Naphthenic acid is the generic name used for all of the organic acids present in crude oils. Most of these acids are
believed to have the chemical formula R(CH2)nCOOH, where R is a cyclopentane ring and n is typically greater
than 12. In addition to R(CH2)nCOOH, a multitude of other acidic organic compounds are also present; however,
not all of them have been analyzed to date.

Isolated deep pits in partially passivated areas and/or impingement attack in essentially passivation-free areas
are typical of naphthenic acid corrosion (NAC). Damage is in the form of unexpected high corrosion rates on alloys
that would normally be expected to resist sulfidic corrosion. In many cases, even very highly alloyed materials (i.e.,
12 Cr, AISI types 316 and 317) have been found to exhibit sensitivity to corrosion under these conditions. NAC is
differentiated from sulfidic corrosion by the nature of the corrosion (pitting and impingement) and by its severe
attack at high velocities in crude distillation units. Crude feedstock heaters, furnaces, transfer lines, feed and reflux
sections of columns, atmospheric and vacuum columns, heat exchangers, and condensers are among the types of
equipment subject to this type of corrosion.

Sulfur

Other than carbon and hydrogen, sulfur is the most abundant element in petroleum. It may be present as
elemental sulfur, hydrogen sulfide, mercaptans, sulfides, and polysulfides. Sulfur at a level of 0.2 percent and
greater is known to be corrosive to carbon and low-alloy steels at temperatures from 230 °C (450 °F) to 455 °C
(850 °F).

At high temperatures, especially in furnaces and transfer lines, the presence of naphthenic acids may increase
the severity of sulfidic corrosion. The presence of these organic acids may disrupt the sulfide film, thereby
promoting sulfidic corrosion on alloys that would normally be expected to resist this form of attack (i.e., 12 Cr and
higher alloys). In some cases, such as in side-cut piping, the sulfide film produced by H2S is believed to offer some
degree of protection from naphthenic acid corrosion.

In general, the corrosion rate of all alloys in the distillation units increases with an increase in temperature.
Naphthenic acid corrosion occurs primarily in high-velocity areas of crude distillation units in the 220 °C to 400 °C

U7
Appendix U – Petroleum Refining

(430 °F to 750 °F) temperature range. No corrosion damage is usually found at temperatures greater than 400 °C
(750 °F), probably due to the decomposition of naphthenic acids or protection from the coke formed at the metal
surface.

Velocity and, more importantly, wall shear stress are the main parameters affecting NAC. Fluid flow velocity
lacks predictive capabilities. Data related to fluid flow parameters, such as wall shear stress and the Reynold’s
Number, are more accurate because the density and viscosity of liquid and vapor in the pipe, the degree of
vaporization in the pipe, and the pipe diameter are also taken into account. Corrosion rates are directly proportional
to shear stress. Typically, the higher the acid content, the greater the sensitivity to velocity. When combined with
high temperature and high velocity, even very low levels of naphthenic acid may result in very high corrosion rates.

CORROSION CONTROL METHODS

High-temperature crude corrosion is a complex problem. There are at least three corrosion mechanisms:

1. furnace tubes and transfer lines where corrosion is dependent on velocity and vaporization,
and is accelerated by naphthenic acid,
2. vacuum column where corrosion occurs at the condensing temperature, is independent of
velocity, and increases with naphthenic acid concentration, and
3. side-cut piping where corrosion is dependent on naphthenic acid content and is inhibited
somewhat by sulfur compounds.

Mitigation of process corrosion includes blending, inhibition, materials upgrading, and process control.

Blending may be used to reduce the naphthenic acid content of the feed, thereby reducing corrosion to an
acceptable level. Blending of heavy and light crudes can change shear stress parameters and might also help reduce
corrosion. Blending is also used to increase the level of sulfur content in the feed and inhibit, to some degree,
naphthenic acid corrosion.

Injection of corrosion inhibitors may provide protection for specific fractions that are known to be particularly
severe. Monitoring needs to be adequate in this case to check on the effectiveness of the treatment. Process control
changes may provide adequate corrosion control if there is the possibility of reducing charge rate and temperature.

For long-term reliability, upgrading the construction materials is the best solution. Above 288 °C (550 °F),
with very low naphthenic acid content, cladding with chromium (Cr) steels (5 to 12 percent Cr) is recommended for
crudes of greater than 1 percent sulfur when no operating experience is available. When hydrogen sulfide is
evolved, an alloy containing a minimum of 9 percent chromium is preferred. In contrast to high-temperature sulfidic
corrosion, low-alloy steels containing up to 12 percent Cr do not seem to provide benefits over carbon steel in
naphthenic acid service. Type 316 stainless steel [greater than 2.5 percent molybdenum (Mo)] or Type 317 stainless
steel (greater than 3.5 percent Mo) is often recommended for cladding of vacuum and atmospheric columns.

Materials in Refinery Construction

The selection of materials for refinery construction depends on the type of refinery, the type of crude oil
handled, and the expected service life for each vessel.(6) As with all materials selection, the life-cycle cost must be
considered in addition to purchase price. Table 3 lists some common alloys and their material costs relative to
carbon steel.

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 Appendix U – Petroleum Refining

Table 3. Comparison of the relative costs of various alloys.(7)

CONSTITUENTS COST
ALLOY CLASS EXAMPLE
RATIO(9)
Ni Cr Mo Fe Co Ti Cu Cb Al V
Carbon Steel C10 > 94 0.2
Low-Alloy Steel 1-1/4Cr 1/2Mo 1.25 0.5 balance 0.25
Type 316L 13.0 17.0 2.3 balance 1.0
Fe-Ni-Cr + Mo Alloy 800H 32.5 21.0 4.6 -
20Cb-3 35.0 20.0 2.5 balance 3.5 3.8
Alloy C2 54.0 15.5 16.0
Alloy C276 57.0 16.0 16.0 5.5 6.0
Ni-Cr-Mo
Alloy C4 54.0 16.0 15.5 3.0
Alloy 625 60.0 21.5 9.0 3.7 6.3
Alloy G 45.0 22.2 6.5 19.5 2.0 6.4
Ni-Cr-Fe
Alloy 600 76.0 15.0 8.0 -
Ni-Mo Alloy B2 balance 1.0 28.0 2.0 1.0 11.6
Ni-Cu Alloy 400 65.1 32.0 -
Nickel Alloy 200 99.9 -
Co-Base ULTIMET (R) 9.0 26.0 5.0 3.0 54.0 27.2
Ti-Base Ti-6Al-4V 90 6.0 4.0 -

Carbon Steel

Carbon steel is by far the most common structural material in refineries due primarily to a combination of
strength, availability, relatively low cost, and a resistance to fire. The low-alloy steels are specified for applications
that require higher properties than can be obtained with carbon steels. The workhorse refinery alloys for elevated
temperature service greater than 260 oC (500 oF) contain 0.5 to 9.0 percent chromium plus molybdenum. Normally,
at least 5 percent chromium is required to resist oxidation at temperatures in excess of 430 oC (800 oF). Currently,
most refineries use 9Cr-1 Mo tubes in coker heaters. For carbon steel and low-alloy steel, creep becomes a design
consideration at about 430 oC (800 oF) and 480 oC (900 oF), respectively. These alloys are used for pressure vessels,
piping, exchangers, and heater tubes.

Austenitic Stainless Steel

The austenitic structure provides a combination of excellent corrosion, oxidation, and sulfidation resistance
with high creep resistance, toughness, and strength at temperatures greater than 565 oC (1050 oF). They are,
therefore, often used in refineries for heater tubes and heater tube supports, and in amine, fluid catalytic cracking
(FCC), catalytic hydro-desulfurization (CHD) sulfur, and hydrogen plants.

They are susceptible, however, to grain boundary chromium carbide precipitation “sensitization” when heated
in the range of 540 oC (1000 oF) to 820 oC (1500 oF). Where “sensitization” is to be avoided, refineries prefer to use
the stabilized grades of Type 347 (with Cb) or Type 321 (with Ti).

The susceptibility of the austenitic stainless steels to stress corrosion cracking limits their use and requires
special precautions during operation and at downtime. At downtime, the precautions taken to prevent stress
corrosion cracking are either alkaline washing with a dilute soda ash and low-chloride water solutions and/or
nitrogen blanketing. The austenitic stainless steels are used for corrosion resistance or resistance to

U9
Appendix U – Petroleum Refining

high-temperature hydrogen or sulfide damage. Solid stainless steel vessels are rarely constructed. Strip-lined,
stainless-clad, or lined vessels are found in hydrocracking and hydrotreating services. Austenitic stainless steels also
find service as tubing in heat exchangers exposed to corrosive conditions.

Ferritic and Martensitic Steels

Other chromium-iron stainless steels with little or no nickel form crystallographic structures different from
austenitic. This stainless steel alloy contains less than 0.10 percent C, 11 to 13 percent Cr, balance Fe, and a ferritic
structure. When the ferritic stainless alloys are modified, they may be hardened and become what is called
"martensitic" by heat treatment. The ferritic and martensitic stainless steels are classified by the American Iron and
Steel Institute (AISI) as the 400 series. The most common alloys from this series found in refineries are types 410,
410S, 405, and 430 stainless steels. A common stainless steel for trays and lining in crude service is Type 410
stainless steel.

Other Alloys

The principal non-ferrous alloys in refinery processing equipment are the copper-based and copper-nickel
alloys; however, the use of copper-based alloys in NH3 or NH4 environments should be avoided.

Although admiralty brass was the original saltwater condenser tube material, it was found to be susceptible to
erosion-corrosion, particularly at tube ends. Aluminum brass, containing 2 percent aluminum, was found to be
somewhat more resistant to erosion in saltwater. Inhibition with arsenic is necessary to prevent de-zincification, as
in the case of admiralty brass. The stronger naval brass is often selected as the tube sheet material when admiralty
brass tubes are used in condensers. Generally, a bronze is a tin alloy of copper, although the term has been widely
used for other alloys, including some brasses. Cast brass or bronze alloys for valves and fittings are usually copper-
tin-zinc compositions, plus lead for machinability. Aluminum bronzes are often used as tube sheet and channel
material for exchangers with admiralty brass or titanium tubes exposed to cooling water.

The 70/30 copper-nickel alloy is used for exchanger tubes when better saltwater corrosion resistance than in
aluminum brass is needed, or when high metal temperatures in water-cooled exchangers may cause de-zincification
in brass. Monel is a nickel-copper alloy with 67 percent nickel and 30 percent copper. Monel has very good
resistance to saltwater and, under non-oxidizing conditions, to acids such as hydrochloric and hydrofluoric acids.
Monel has a better high-temperature resistance to cooling water than does 70/30 copper-nickel. Monel cladding and
Monel trays are commonly specified at the top of crude towers to resist HCl vapor and where the temperature is
below 205 oC (400 oF). Over 205 oC (400 oF), nickel-based alloys are attacked by H2S. For high temperature
strength and/or corrosion resistance, several nickel-based alloys are used for special applications such as expansion
bellows in FCC process units (Alloy 625), stems in flue gas butterfly valves (Alloy X 750), and in springs exposed
to high-temperature corrosives (Alloy X).

Titanium has excellent resistance to seawater, and it is also used for tubing in crude tower overhead
condensers. Overall, the use of titanium is extremely limited due to the high cost and the availability of suitable,
more economic alternatives.

CORROSION MANAGEMENT

Economics of Refining

Although the individual components are quite complicated, the large-scale economics of refining operations
can be defined in simple terms. Gross margin is the difference between the output of a refinery (refined products)
and the cost of the feedstock (crude oil and other chemicals). The net margin is the gross margin minus the
operating costs. Figure 4 illustrates the last 20 years of margins in the refinery industry.

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 Appendix U – Petroleum Refining

Figure 4. Margins of U.S. refiners since 1977.(8)

Capital Expenditures

The capital intensity of a process refers to the amount of capital needed to produce a unit of product. For U.S.
refining operations, capital intensity is measured by the ratio of net property, plant, and equipment (i.e., the balance
sheet value of productive long-term assets adjusted for depreciation) to refinery capacity (barrels per calendar day of
crude distillation capacity). Adjusted for general inflation (via the implicit gross domestic product deflator), the
refiners’ capital expenditures for U.S. refining doubled from 1989 to 1992.

A surge in capital expenditures occurred in the late 1970s through the early 1980s. During this period, the
major U.S. companies upgraded their refineries to process heavier, more sulfurous crude oil inputs into relatively
greater proportions of lighter products, particularly gasoline. These investments were largely premised on wide
price spreads between higher and lower quality crude oils and lighter and heavier refined products.

The decline in the price spread between differing qualities of crude oils in the 1990s contributed to the overall
deterioration in the gross margin evidenced in figure 4. The price decline between crude oils of differing qualities
was especially adverse for refiners who invested heavily in refinery upgrades to yield higher proportions of light
products. The refiners directed much of the surge in their refining investments in the late 1970s to the early 1980s
toward increasing their capability to process heavier, more sulfurous crude oils. For example, the capacity for
increased processing of heavy sour crude inputs, relative to basic crude distillation capacity, rose from 22 percent in
1974 to 30 percent in 1980 to 47 percent in 1993.

Unlike the earlier surge in refinery investments, the upswing in capital expenditures in the 1990s appeared to
be largely driven by increased expenditures for pollution abatement.(8) In particular, the Clean Air Act Amendments
of 1990 required production of oxygenated gasolines by late 1992, lower sulfur diesel fuels by late 1993, and
reformulated gasoline by January 1, 1995. The share of total U.S. refining capital expenditures for pollution
abatement increased from slightly more than 10 percent shortly before the Clean Air Act Amendments of 1990 to
more than 40 percent in recent years.

Although pollution abatement requirements clearly reduced the rate of return to refining/marketing assets,
these requirements appear to account for only a small part of the steep decline in the rate of return to U.S.

U11
Appendix U – Petroleum Refining

refining/marketing operations in the 1990s. The increase in pollution abatement operating costs over this period was
$0.07 per barrel of refined products sold, or 5 percent of the $1.52 per barrel decline in the net margin.

The cost of extra capital expenditures for corrosion control can be included in the operational expenditures for
refinery operations. If an operator chooses a corrosion-resistant alloy vessel for a refinery operation, then the extra
annual cost of this vessel amortized over the life of the vessel is included in the operational expenditures. If an
operator chooses carbon steel for the vessel, then the cost of corrosion control measures, such as anodes, chemical
treatment, and monitoring, are the only measurable capital expenditures, but annual costs of upkeep will greatly
increase operational expenditures. Economic justifications for such expenditures based on life-cycle costs continue
to be part of corrosion control decisions for refinery operations.

Operational Expenditures

The operating costs of refineries have steadily decreased in recent years due to technological advances and
improvements in efficiency. The 1996 operating costs were an average of $5.51 per barrel (bbl).(9)

It should be noted that direct costs for corrosion prevention and mitigation are extremely difficult to obtain, as
these are kept very “close to the vest” by the refining industry. While the reasons for this are unclear, it can be
assumed that the intense scrutiny that the entire petrochemical industry undergoes by environmental regulators and
community watchdogs has created a situation in which refiners prefer not to divulge the magnitude of their corrosion
problems. Thus, information for this sector has been gathered from a combination of some published surveys and
government sources.

One particular study(10) focused on operating costs at a single small refinery (53,000 barrels/day), concentrating
on the costs related to environmental protection. This project quantified air emissions, water discharges, and other
wastes generated at the facility. Moreover, it identified a range of options to reduce or prevent those releases, some
of which appeared more cost-effective than those required by existing rules.

At most refineries, operating costs are dominated by crude oil. Even small fluctuations in the price of crude oil
can overshadow other operating costs of the refinery. As a result, it is customary at the refinery level to track
"non-crude operating costs," excluding the cost of feedstock. The non-crude operating costs of this refinery are
shown in table 4.

Table 4. Environmental costs at a refinery.

PERCENTAGE OF 1992
ENVIRONMENTAL COST
NON-CRUDE OPERATING
CATEGORY
COSTS
Waste Treatment 4.9
Maintenance 3.3
Product Requirements 2.7
Depreciation 2.5
Administration, Compliance 2.4
Sulphur Recovery 1.1
Waste Disposal 0.7
Non-Recurring Costs 4.0
TOTAL 21.6%

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 Appendix U – Petroleum Refining

The analysis estimates that total environmentally related costs are 21.6 percent of total non-crude operating
costs. This total focuses primarily on capital, operating, and maintenance costs, and excludes contingent liability
costs. If these costs were added, the total could be higher. Remediation expenses are recorded as "non-recurring
costs."

At the outset of the project, prior to conducting the analysis, environmental personnel informally estimated
environmentally related costs at only 3 percent of the total non-crude operating costs. The magnitude of this
difference, as well as the magnitude of the costs, indicates the value of identifying and tracking environmental costs.

Maintenance costs (40 percent of which can be attributed to corrosion control) were estimated in the study(10)
to be 3.3 percent (rounded to 3 percent) of the non-crude operating costs (table 4). When scaling up to all processes,
this figure becomes:

0.03 fraction maintenance costs / 0.216 fraction environmental operating costs = 13.9% of
the total operating costs due to maintenance

$5.51/bbl operating costs(9) x 13.9% = $0.76/bbl maintenance costs

$0.76/bbl maintenance costs x 40% due to corrosion control = $0.31/bbl operating costs for
corrosion control

When multiplied by the annual refinery output in the United States (5.7 billion barrels in 1997), the total cost
of corrosion is ($0.31 x 5.7 billion = ) $1.767 billion per year.

Vessel turnarounds, during which a processing vessel is emptied, inspected, repaired (if necessary), and put
back into service, are mandatory in most cases due to U.S. Department of Transportation (DOT) regulations,
primarily due to suspected corrosion damage inside the vessels. The costs for these operations are capitalized rather
than included in the maintenance budget.

One refiner estimated the total cost of the turnarounds at one of their refineries (a 260,000-bbl per day plant).(4)
For the 3,000 processing vessels in this refinery, the total cost of turnarounds (5-year intervals for each individual
vessel) was $118 million. Therefore, turnaround costs per barrel are:

($118,000,000/yr x 1 turnaround/5 years) / (260,000 bbl/day x 365 days/yr) = $0.25/bbl

$0.25/bbl x 5.7 billion bbl/yr = $1.425 billion/yr for turnaround costs

It should be noted that the trend in this activity is to move toward risk-based inspections and longer intervals
(10 to 20 years) between turnarounds, which would significantly reduce the cost of corrosion maintenance, but
increase the risk factor dramatically. The validity of this strategy is yet to be determined since the number of
incidents with vessels outside the standard 5-year window will, in the future, help to define the proper risk
assessment.

Fouling
In addition to mitigation and maintenance costs, the component of lost production due to corrosion and related
problems must be considered. Fouling is the leading cause of diminished efficiency and productivity in refineries.
Fouling is a deposit buildup in refinery processes that impedes heat transfer and/or reduces throughput. The energy
lost due to this inefficiency must be supplied by burning additional fuel or reducing feed.

It is estimated that the cost penalty of fouling is in excess of $2 billion annually.(11) While most fouling is
caused by the deposition of heavier hydrocarbon species coming directly from the crude oil, a small undetermined
percentage is related to corrosion and scale deposits, either actively participating as loose corrosion products or by
scale acting as a substrate for hydrocarbon deposition.

U13
Appendix U – Petroleum Refining

It is not known exactly how much fouling is related to corrosion versus that related to deposits, which affect
only production rates. In the Drinking Water and Sewer Systems sector of this report (Appendix K), 50 percent of
the costs of fouling were corrosion-related. Applying this factor to the water handling half of the refining process,
the fouling-related corrosion costs in the refining sector are estimated to be $2 billion total costs x (1/2 fluid
volume on water handling portion of refining process x 50% corrosion-related fouling costs = $0.5 billion.

The estimate of the total annual cost of corrosion in refining applications is therefore:

$1.767 billion operational costs for corrosion

$1.425 billion turnaround costs
$0.500 billion fouling costs
$3.692 billion total cost of corrosion

Acidic Crude Oils

As was discussed earlier, the refiners’ willingness to accept the more corrosive, acidic crude oils has heavily
influenced U.S. refinery operations due to the lower cost of the feedstock.

It can be anticipated that the growth in expenditures for corrosion can be expected to increase at the rate of the
acidity in the crude oil refined. Therefore, this cost is part of the incremental maintenance cost, but in the near
future, this will become a significant expenditure.

For a typical carbon steel distillation column running acidic crude oil, there are additional costs associated with
corrosion coupons and probes for monitoring, nondestructive testing and analysis, and chemical treatment. It should
be noted that these costs, shown below in figure 5, have a wide variance associated with them.

12

10

8
Cost/bbl, cents

LO W
6
HIGH

0
C oupons P robes Inhibitors NDE TO TA L

Figure 5. Incremental costs for corrosion control of carbon steel distillation column.(12)

The total cost for chemical treatment and all associated costs in the column range from $0.01 per bbl to a high
of $0.11 per bbl. The figure is dominated by the chemical cost of the inhibitors. One study estimated that the total
inhibitor cost associated with refinery operations in the United States was $246 million in 1998.(13)

U14
 Appendix U – Petroleum Refining

Alternatively, a metallurgical upgrade in susceptible areas for a $120,000-bbl per day refinery is estimated to
be $12 million to $20 million, which could be economically feasible if the refinery has a long-term commitment to
processing acidic crudes. Based on a 20-year design life (typical for alloys), the incremental costs become $0.18 to
$0.30 per bbl (higher than the costs for treatment with corrosion inhibitor, but comparable). The increased chance
for success with the use of alloys relative to corrosion inhibitor treatments makes these options worth further study.

Failure Costs

The costs associated with catastrophic failures are very difficult to quantify since they include the costs of
equipment replacement, production loss, and sometimes lost lives and litigation. In addition to the direct costs,
indirect costs in publicity and increased scrutiny cannot be quantified.

Analyzing processing industry data for August 2000,(14) 9 incidents (fire, explosion, leak, or emergency
shutdown) were reported at refineries in the United States out of a total of 52 total incidents during that month. The
cause of each is still being investigated, but all of these incidents resulted in some loss of production and a
significant economic impact.

CASE STUDY

Corrosion-Related Failure in Refinery

This example clearly illustrates the hazards associated with amine absorber pressure vessels used in refineries.
On July 23, 1984, a refinery at Romeoville, Illinois, owned and operated by the Union Oil Company of California,
experienced a disastrous explosion and fire.(10,15) An amine absorber pressure vessel ruptured and released large
quantities of flammable gases and vapors. Seventeen lives were lost, 17 individuals were hospitalized, and more
than $100 million in damages resulted.

The National Bureau of Standards (NBS) conducted a detailed investigation, which included chemical
analyses, fracture mechanics analyses, stress corrosion cracking (SCC) susceptibility tests, and hydrogen cracking
susceptibility tests. Preliminary NBS test results indicated that the subject plate material (ASTM A516, Grade 70
carbon steel) of the amine absorber was susceptible to hydrogen-induced cracking. Furthermore, repair welds that
were done in the field and that had not been stress relieved, were especially sensitive to amine-induced corrosion
and cracking. Figure 6 is an example of SCC both parallel and perpendicular to the weld, but not in the weld. The
propagation of the crack clearly distinguishes SCC and reflects the different stresses along the weld area.

Figure 6. Stress corrosion cracking near a weld.

U15
Appendix U – Petroleum Refining

REFERENCES
1. Energy and Environmental Profile of the U.S. Petroleum Refining Industry, report by Energetics Inc. for
the U.S. Department of Energy, 1998.

2. Petroleum Industry of the Future, www.oit.doe.gov/petroleum, October 2000.

3. Petroleum Supply Annual 1999, Vol. 1, Energy Information Administration.

4. Petroleum Refining Corrosion, www.hghouston.com/refining.html#top, October 2000.

5. R.A. White, Materials Selection for Petroleum Refineries and Gathering Facilities, NACE, 1998.

6. Refinery Materials of Construction, www.corrmet.ndirect.co.uk/steel.htm.

7. C. Shawber and P. Manning, Haynes International, Personal Communication, August and September 2000.

8. The Impact of Environmental Compliance Costs on U.S. Refining Profitability, Energy Information
Administration, October 1997.

9. D. Ditz, J. Ranganathan, and R. Banks, Green Ledgers: Case Studies in Corporate Environmental
Accounting, World Resources Institute, May 1995.

10. Reduced Corrosion in Amine Gas Absorption Columns,

www.hydrocarbonprocessing.com/archive/archive_99-10/99-10_reduce-mogul.htm, October 2000.

11. Petroleum Project Fact Sheet – Fouling Minimization, U.S. Department of Energy, Office of Industrial
Technologies, January 1999.

12. J. Skippins, D. Johnson, and R. Davies, Corrosion Mitigation Program Improves Economics for Processing
Naphthenic Crudes”, Oil & Gas Journal 98, 2000.

13. Corrosion Inhibitors Market Analysis, Publications Resource Group Inc., 1999.

14. Process Incidents, August 2000, www.saunalahti.fi/ility/PI0008.htm, October 2000.

15. V. Novokshchenov, Proceedings of Fifth Middle East Corrosion Conference, Oct 28-30, 1991, Manama,
Bahrain, pp. 209-223

U16
 Tutorial: Delayed Coking Fundamentals

Paul J. Ellis
Christopher A. Paul

Great Lakes Carbon Corporation

Port Arthur, TX

Prepared for presentation at the AIChE 1998 Spring National Meeting

New Orleans, LA
March 8-12, 1998

Topical Conference on Refinery Processing

Tutorial Session: Delayed Coking
Paper 29a

Copyright 1998 Great Lakes Carbon Corporation

UNPUBLISHED

March 9, 1998

1
 ABSTRACT

Great Lakes Carbon Corporation has worked closely with refineries producing delayed coke in all forms,
fuel grade (shot or sponge), anode grade (sponge), and electrode grade (needle) since start-up of the
company's first calcining operation in 1937. With on-going research in the area of delayed coking since 1942,
Great Lakes Carbon (GLC) has operated delayed coking pilot units including an excellent large-scale pilot
unit with a coke drum 0.3 meter (1 ft) diameter by 2.1 meters (7 ft) long and has developed physical models
which explain coke formation in coke drums.

Knowledge of commercial delayed coking units as well as that of the GLC Pilot Delayed Coker is used in
this tutorial paper to describe the formation and uses of the three types of structures of delayed petroleum
coke: needle, sponge, and shot. Troubleshooting tips are included on many aspects of the delayed coking
drum cycle including: steam stripping, water quenching, coke cutting, drum warm-up, and drum switching
technique. Suggestions and descriptions of delayed coking unit hardware are included.

The objective of this tutorial paper is to acquaint the non-refinery technologist and further the knowledge of
refinery personnel with the delayed coking operation, delayed coking unit hardware, types of coke that can
be produced, coke formation models, and the uses of petroleum coke. Also, by discussing some of the
delayed coking process problems encountered in the industry, we hope to encourage new advances, solutions,
and improvements for the future.

2
WHAT IS DELAYED COKING?

Delayed coking is a thermal cracking process used in petroleum refineries to upgrade and convert petroleum
residuum (bottoms from atmospheric and vacuum distillation of crude oil) into liquid and gas product streams
leaving behind a solid concentrated carbon material, petroleum coke. A fired heater with horizontal tubes
is used in the process to reach thermal cracking temperatures of 485 to 505oC (905 to 941oF). With short
residence time in the furnace tubes, coking of the feed material is thereby “delayed” until it reaches large
coking drums downstream of the heater. Three physical structures of petroleum coke: shot, sponge, or
needle coke can be produced by delayed coking. These physical structures and chemical properties of the
petroleum coke determine the end use of the material which can be burned as fuel, calcined for use in the
aluminum, chemical, or steel industries, or gasified to produce steam, electricity, or gas feedstocks for the
petrochemicals industry.

CRUDE OIL REFINING: WHERE DOES DELAYED COKING FIT IN?

To understand the delayed coking process, one must understand how the delayed coker is integrated with
the rest of the refinery. Delayed coker feed originates from the crude oil supplied to the refinery. Therefore,
brief descriptions of each of the processing steps preceding the delayed coking unit are provided below. A
basic refinery flow diagram is shown on the following page in Figure 1.

Crude Oil Desalting

Crude oil contains around 0.2% water in which is mixed soluble salts such as sodium chloride and other
metals which are on the edge of the sphere of water. In desalting, crude oil is washed with around 5% water
to remove the salts and dirt from the crude oil . The water, being heavier than the oil, drops out of the
bottom, and the cleaned oil flows overhead with around 0.1% water.

Atmospheric Distillation

The desalted crude oil is heated in a tube furnace to over 385°C (725°F), just below the temperature that
cracking of the oil can occur, then flashed into a distillation column. The primary products are straight run
gasoline, kerosene, jet fuel, diesel, atmospheric gas oil (AGO) and atmospheric reduced crude.

Vacuum Distillation

The atmospheric reduced crude (ARC) is then heated to around 395°C (743°F) and flashed into a vacuum
distillation column that is operated at low pressures, 10 mm Hg absolute desired but more common
25 to 100 mm Hg absolute. The desired aim is to lift the maximum amount of oil boiling below 565°C into
heavy vacuum gas oil (HVGO) reducing the production of vacuum reduced crude (VRC), the main
feedstock to the delayed coker. The HVGO and the AGO are the principal feedstocks to a fluid catalytic
cracking unit (FCCU) for the production of gasoline and diesel. Improving vacuum distillation is one of the
best methods for increasing gas oil yield in a refinery while at the same time reducing the amount of vacuum
reduced crude (coker feed). This enables higher refinery throughput rates to be achieved.

3
 Figure 1.
Basic Refinery
Crude Oil Gasoline
Storage Diesel
Fluid
Cat. Delayed
Cracker Coke
Desalter Drums

Atm.
Distillation

Tube HGO Switch Valve

Furnace

Coker
Fractionator
Tube
Furnace Vacuum
Distillation Tube
Furnace
Great Lakes Carbon Corporation

Vacuum Reduced Crude Processing Options or End Uses

Delayed Coking
Visbreaking - Primary function is to reduce viscosity of the oil with some production of heavy gas oil.
Resid FCC - Residuum Fluid Catalytic Cracking, metals deactivate catalyst, must use passivating chemicals
to reduce unwanted reactions
Resid Hydrocracking - Feed is contacted with a catalyst and hydrogen at high temperature and pressure
to remove sulfur, nitrogen, and some aromatic compounds with some conversion to lighter liquid products.
ROSE - Residual Oil Supercritical Extraction for production of metal free gas oil, asphaltenes and resins
Propane Deasphalting / Bright Stock - Solvent extraction of heavy lubrication oils
Road Asphalt
Roofing Asphalt - May require air blowing to increase hardness
Fuel Oil - Burner and slow RPM marine diesel

HISTORY OF THE DELAYED COKING PROCESS

“Petroleum coke was first made by the pioneer oil refineries in Northwestern Pennsylvania in the 1860's.
These primitive refineries boiled oil in small, iron stills to recover kerosene, a valuable and much needed
luminescent. The stills were heated by wood or coal fires built underneath which over-heated and coked the
oil near the bottom. After the distillation was completed, the still was allowed to cool so the workmen could
dig out the coke and tar before the next run [1].” The use of single horizontal shell stills for distillation of
the crude was used until the 1880's, with the process sometimes stopped before bottoms coked to produce
a heavy lubricating oil. Multiple stills were used to process more fractions by running the stills in series with
the first still producing the coke. In the 1920's the tube furnace with distillation columns (bubble cap
distillation trays patented by Koch ushered in the modern distillation column) were being built with the
bottoms from the distillation column going to wrought iron stills in which the total outside of the horizontal
still was in direct contact with the flue gases. This produced the maximum amount of heavy gas oil. Some

4
of these units were still in operation after World War II. Operators assigned as decokers used picks, shovels,
and wheelbarrows and had rags wrapped around their heads to protect against the heat.

The coke that was produced in the horizontal stills had a high density, low volatile matter (VM) content of
around 8 wt%, and less than 1 wt% moisture. One problem was that ash content was high, around 1 wt%
compared to under 0.2 wt% in most modern delayed cokers. Conners [1] thought that this was due to the
lack of desalting and washing of the crude oils processed at that time.

The origin of the vertical coke drum was probably from thermal cracking of gas oil for the production of
gasoline and diesel fuel. From 1912 to 1935 the Burton process developed by Standard Oil at Whiting,
Indiana converted gas oil to gasoline with the production of petroleum coke. Dubbs and other thermal
cracking processes also produced petroleum coke. Lack of an adequate supply of crude oil and the lack of
a heavy oil market caused land-locked middle American refineries to process the heavy fuel oil (atmospheric
distillation bottoms and vacuum distillation bottoms) in a delayed coker to produce more gasoline and diesel
fuel. Decoking the drums was difficult. “Manual decoking was a hot and dirty job. ...various mechanical
devices were tried. One of the common systems employed was to wind several thousand feet of steel cable
on holding devices in the drum. The cable was pulled by a winch, to loosen the coke. Coke was also
removed by drilling a small hole, then a large hole, after which beater balls on a rotating stem knocked out
the remaining coke [1].”

The first delayed coker was built by Standard Oil of Indiana at Whiting, Indiana in 1929 [2,3]. The
development of hydraulic decoking came in the late 1930's. Shell Oil at Wood River, Illinois presented a
paper on hydraulic decoking 4.0 m (13 ft) diameter Dubbs units and stated that they had patents along with
Worthington Pump Company on hydraulic decoking bits and nozzles [4]. Standard Oil of Indiana had patents
on the original cutting nozzles used by Pacific Pump [5]. A very similar nozzle is currently used in the new
compact combination coke cutting unit. A pilot hole is drilled down through the coke in the drum using high
pressure water, and then the coke is cut out with a drilling bit with horizontal water nozzles. Roy Diwoky
while at Standard Oil Whiting was one of the key people in developing the hydraulic decoking in the 1930's.
Diwoky in May 1952, while Executive Vice President of Pan Am Southern Corp. (Owned by Standard Oil
of Indiana), worked with Great Lakes Carbon Corporation to produce the first needle coke in a delayed
coker. Bernard Gamson, the Director of Research and Development for Great Lakes Carbon at the time,
stated in a report that Diwoky was “the father of delayed coking [3].”

Delayed coking combined a number of the features and improvements from the development of the thermal
cracking process. The use of pressure as well as heat for cracking and separating the heater from the coker
and the use of two drums enabled the delayed coker to operate on a continuous basis. The number of cokers
built before 1955 was small, with a surge in delayed coker construction between 1955 to 1975 at 6% per year
and an 11% growth rate during the 1965 to 1970 period [1]. The growth of delayed cokers was in step with
the growth of fluid catalytic cracking and rapid decline in thermal cracking. A fluid coker, similar to a fluid
catalytic cracker except that fluid coke is circulated instead of catalyst, was first built in 1954 at Billings,
Montana. Five more fluid cokers were built in the late fifties, and one in 1970. In 1958, the head of
petroleum refining engineering at Colorado School of Mines, J.O. Ball, stated that there would not be any
more delayed cokers built. Ball thought all new cokers would be fluid cokers, and that a delayed coker was
just a garbage can in the refinery. Today there are 49 operating delayed cokers in the U.S. and only six fluid
cokers / flexicokers.

5
MODERN DELAYED COKING PROCESS

The delayed coker is the only main process in a modern petroleum refinery that is a batch-continuous process.
The flow through the tube furnace is continuous. The feed stream is switched between two drums. One
drum is on-line filling with coke while the other drum is being steam-stripped, cooled, decoked, pressure
checked, and warmed up. The overhead vapors from the coke drums flow to a fractionator, usually called
a combination tower. This fractionator tower has a reservoir in the bottom where the fresh feed is combined
with condensed product vapors (recycle) to make up the feed to the coker heater.

Delayed Coking Drum Cycle

Since the feed stream is regularly switched between drums, a cycle of events will occur on a regular interval
depending on the delayed coking unit feed rate, drum size, and throughput capacity. Most typical delayed
cokers currently run drum cycle times of about 16 hours with one drum filling on-line while its counterpart
is off-line for stripping, cooling, and decoking. Drum cycle event approximate time requirements for such
a cycle are shown below in Table 1. Shortening the cycle time is one method of increasing throughput on
delayed coking units. One refinery regularly runs 12 hour drum cycles and has attempted 10 and 11 hour
cycles, but cycles this short are extremely difficult due to minimum time requirements for each of the steps
of the drum cycle. Some of the more important drum cycle steps are described in detail in the following
sections.

Table 1. Typical Short Cycle Coking Operations

Drum Cycle Hours

Steam to Fractionator 0.5
Steam to Blow Down 0.5
Depressure, Water Quench and Fill 4.5
Drain 2.0
Unhead Top and Bottom 0.5
Cutting Coke 3.0
Rehead / Steam Test / Purge 1.0
Drum Warm-Up (Vapor Heat) 4.0
-----------------------------------------------------------------
Total Time 16.0

Drum Warm-Up (Vapor Heat). To prepare the cold empty coke drum to be put back on-line to
receive the hot feed, hot vapors from the on-line drum are circulated into the cold empty drum. The hot
415°C (780°F) vapors condense in the cold drum, heating the drum to a target temperature of around 340°C
(650°F). While the drum is heating, the condensed vapors are continuously drained out of the drum.

On-line Filling. After the cold drum has been vapor heated for a few hours, hot oil from the tube
furnace at about 485°C (905°F) is switched into the drum. Most of the hot vapors condense on the colder
walls of the drum, and a large amount of liquid runs down the sides of the drum into a boiling turbulent pool
at the bottom of the drum. The drum walls are heated up by the condensing vapors, so less and less vapors
are condensing and the liquid at the bottom of the drum starts to heat up to coking temperatures. A main
channel is formed similar to the trunk of a tree. As time goes on the liquid pool above the coke decreases

6
and the liquid turns to a more viscous type tar. This tar keeps trying to run back down the main channel
which can coke at the top causing the channel to branch. So the limbs of the “tree in the drum” appear [6].
This progresses up though the coke drum. Sponge coke, which includes needle coke, is formed from this
liquid which remains in a quiescent zone between the main branches or channels up through the coker. The
liquid pools in the quiescent zones slowly turn to solid coke. Shot coke has a different type of coke structure
indicating that it is produced while suspended in the vapor phase in the drum. This will be discussed in detail
later in the paper.

On top of the liquid layer is foam or froth. Paraffinic type feedstocks with some sodium present foam
readily compared to aromatic feedstocks which tend to have smaller foam heights. Higher temperatures
greatly decrease the height of the foam. At high temperature, needle coke has very small or no foam present.
After the coke drum is filled, the hot oil is switched to the new drum.

Steam-Stripping / “Hot Spots.” Steam must be flowing before the switch and immediately after
the switch; otherwise, the yet unconverted liquid feed on top of the coke bed will run down the channels
which will coke or solidify and plug the channels. The plugging of the channels causes problems in cooling
the coke since sections of the coke bed will be isolated from the steam and cooling water by the plugged
channels. This is the cause for “hot spots” and “steam eruptions” when cutting the coke. Cold water from
the cutting nozzle hits the exposed hot coke which results in a steam explosion. This is particularly hazardous
when the pilot hole is being cut, since the drum is filled with a large quantity of hot water. A steam explosion
during pilot hole cutting can cause the hot water to erupt out of the top of the drum and has caused fatalities
in the past.

Steam stripping also serves to transfer heat from the hot bottom section of the coke bed to the
unconverted liquid present at the top of the coke drum. Adequate steam stripping increases the amount of
recovered gas oil yield while at the same time reduces the amount of volatile matter and pitch left in the top
section of the coke drum. After the steam has been flowing up through the coke bed for about thirty minutes
with the vapors going to the fractionator, the vapor line is vented to blowdown system. Steam is increased
for a short time or in some cases water is immediately introduced at the bottom of the drum which instantly
flashes to steam. The steam is backed out and the flow of cooling water is gradually increased. The top
vapor temperature in the drum may increase slightly at first before cooling due to the increased flow of steam
up through the coker.

Water Cooling / “Drum Bulging.”. The rate of cooling water injection is critical. Increasing the
flow of water too rapidly can “case harden” the main channels up through the coker without cooling all of
the coke radially across the coke bed. The coke has low porosity (the porosity comes from the thermal
cracking) which then allows the water to flow away from the main channels in the coke drum. Porosity of
delayed coke has been measured experimentally in the past by measuring water flow through cores about the
size of hockey pucks cut from large chunks of needle coke from different areas of a commercial coke drum.
Most of the coke cores were found to have no porosity except the coke right at the wall which had some
porosity . This explains problems that have been found to occur with drums bulging during cool down. If
the rate of water is too high, the high pressure causes the water to flow up the outside of the coke bed
cooling the wall of the coke drum. Coke has a higher coefficient of thermal expansion than does steel (154
for coke versus 120 for steel, cm/cm/°C x 10-7). This was measured in the transverse direction from a chunk
of needle coke. The coefficient of thermal expansion for raw sponge coke is probably even greater than that
of the needle coke tested.

7
DELAYED COKING UNIT HARDWARE

A basic coker operation flow diagram is shown below in Figure 2 to illustrate some of the delayed coking
unit hardware.

Combination
Figure 2.
Fractionation Basic Coker Operation
Column Cooler
Gas Compressor
Reflux Drill Rigs
Cut water
Kerosene
Naphtha
Diesel
Waste Heat Boiler
Pump-Around
Hvy Coker Gas Oil HCGO

Feed Vapor Overhead Coke

Drums
VRC

Feed Coker
Pump Tube
Furnace

Switch Valve
Great Lakes Carbon Corporation

Feed Preheat

In some refineries, delayed coker feed which is usually vacuum reduced crude (VRC) arrives at the coker hot,
straight from the vacuum distillation unit, but in most cases, delayed coker feed is relatively cold coming from
tankage. The feed is preheated by heat exchangers with gas oil products or in some rare cases by a fired
coker preheater (tube furnace). In some refineries, the convection section of the main coker furnace is used
to preheat the cold feed. The hot coker feed, ranging from 360 to 400oC (680 to 750oF), then enters the
bottom of the fractionator / combination tower where the fresh feed is combined with some condensed
product vapors (recycle) to make up the feed to the coker heater. The fractionator bottom provides some
surge storage capacity for the incoming fresh feed, and in some units, heat is transferred to the fresh feed by
flowing a split of the fresh feed above the drum overhead vapor entrance to the fractionator. This practice
usually results in increased amounts of heavy coker gas oil recycle in the furnace charge.

Coker Charge Pumps

The coker charge pumps located between the fractionator bottom and the coker heater are normally driven
by an electric motor with a steam-driven turbine pump as a backup. The pressure is in excess of 35 bars
(500 psig) with a mechanical seal operating up to 382°C (720°F).

8
Coker Tube Furnace

The coker tube furnace is the heart of the delayed coking process. The heater furnishes all of the heat in the
process. The outlet temperature of a coker furnace is typically around 500°C (930°F) with a pressure of 4
bars (60 psig).

Coker Furnace Design.

Delayed coker furnace design objectives according to Elliott [7] are:
• High in-tube velocities resulting in maximum inside heat transfer coefficient
• Minimum residence time in the furnace, especially above the cracking temperature threshold
• A constantly rising temperature gradient
• Optimum flux rate with minimum practicable maldistribution based on peripheral tube surface
• Symmetrical piping and coil arrangement within the furnace enclosure
• Multiple steam injection points for each heater pass

Normally the modern-day furnace has two to four passes per furnace. The tubes are mounted
horizontally on the side and held in place with alloy hangers. The furnace tubes are around 100 mm ID with
6 to 12 mm wall thickness and are at least a 9% chrome alloy. Higher alloy tubes are being used with the
more rapid steam spalling and steam-air decoking methods. Aluminized tubes have been tried, but offer no
advantage. Multiple burners are along the bottom of the radiant wall opposite from the tubes and are fired
vertically upward. The burners for each firebox are controlled by the temperatures of tubes in that firebox
only. The control thermocouple for the firebox should be three or more tubes back from the outlet to prevent
coke forming on the thermocouple. The outlet thermocouple is initially read and an off-set from the control
thermocouple is then used to control the furnace. Tall furnaces are advantageous since the roof tubes are less
likely to have flame impingement and overheating by both radiation and convection. Normally just the radiant
section of the heater is used to heat the oil for a delayed coker. The upper convection section of the coker
heater is used in some refineries to preheat the oil going to the fractionator or for other uses such as steam
generation.

The typical gas burners in a delayed coker furnace are 3 MM BTU size. Adams [8] stated that the
burners will produce flame height of around 0.33 meter per 1 MM BTU. Elliott [7] and others state that the
average radiant flux rate should be below 9000 BTU/HR/FT2 with cold oil velocity of 2 meters/sec (6 ft/sec)
or mass velocity of 1800 kg/sec/meter2 (400 lb/sec/ft2) or greater. Velocity steam is added at around 1 wt%
of the feed. This helps increase the velocity in the tube furnace, and reduces the partial pressure in the drum
so that more gas oil product is carried out of the drum. The specific heat of the steam is less than the oil, so
steam is not a good source of heat in the drum. The main use for the steam is that it keeps the velocity
flowing in the tube furnace if the oil flow is momentarily is lost or decreased which reduces the chance of
coking up the furnace tubes.

Heater Tube Decoking. When coke forms in the heater tubes, it insulates the inside of the tube
which results in elevated temperatures on the outside of the tube. With good operational practices, coker
furnace run lengths of 18 months are possible before decoking of the tubes is needed. When temperatures
approach 677°C (1250°F) on the exterior skin thermocouple, the furnace must be steam spalled and/or
steam-air decoked or cooled down and cleaned by hydraulic pigging.

9
 Steam Spalling. Steam spalling was probably first practiced by Exxon but was perfected by
Lloyd Langseth while operating the cokers at Arco in Houston, Texas in the 1970's. He was able to operate
a coker furnace over four years without shutting down by practicing on-line steam spalling. The only reason
he had to shut down was that Texas had a law that required steam boilers to be inspected every five years.
On-line steam spalling requires replacing the oil with steam in the pass and then heating and cooling the tubes
to snap or spall off the coke inside the tube. The steam and coke go into the drum. The main problem is in
controlling the velocity and speed of spalling off the coke. Too rapid spalling can plug the tube outlet, and
too high steam velocity can erode the metal in the elbows. In one refinery, return bends failed after the
second steam spalling. Steam spalling requires that the delayed coker be supplied with four passes or more.
Attempts to steam spall a two-pass furnace has been tried, but the large amount of steam being handled
caused problems in the fractionator.

Steam-Air Decoking and Pigging. The usual method of decoking the tubes in a coker furnace is
to take the furnace off-line, steam spall, then burn the coke out of the tubes by steam-air decoking. After
steam-air decoking, the tubes need to be water washed since the salts still remain in the tubes and will cause
rapid coking of the tubes. A new method of decoking the tubes is to steam spall, and then use water
pressure to push Styrofoam pigs with studs and grit on the exterior through the tubes and around u-bends
(even u-bends with clean-out plugs). The pigs scrape out the coke without scratching the tube walls. Early
methods of pigging coker heaters left scratches on the tube walls, but with the grit-coated pigs, pigging just
polishes the inside of the tube wall. Pigging is faster than steam-air decoking, and refiners generally have
longer campaigns on the heater compared to steam-air decoking.

Heater Tube Deposits. Iron sulfide is probably not totally removed in steam-air decoking. Coke
deposits have very high content of iron, silica and sodium. Deposits recovered from return bend clean-out
plugs are sometimes long cylindrical shapes and in another case looked like a thick scallop shell. These
deposits were mostly sodium and calcium.

Transfer Line and Switch Valve

Transfer Line. The line from the furnace to the switch valve and on to the drum is referred to as the
transfer line. The transfer line must be very well insulated to prevent coking and plugging. The shorter the
line the better. Long transfer lines with many crosses and tee’s used for clean outs will rapidly coke and
increase the pressure on the furnace which usually results in increased fouling of the tubes in the furnace.
Flanges near the drums are difficult to insulate without causing the joints to leak. Some transfer lines have
a pressure relief valve in the line, but most furnaces and transfer lines are designed to withstand the maximum
pressure the charge pump can produce in case of an accidental switch into a blinded valve.

Switch Valve. The switch valve is a three-way valve with ports to the two drums and a port
(recirculation line) back to the fractionator which is used in startup and shutdown. Older cokers used a
manually operated Wilson-Snyder valve which was a tapered plug valve that required unseating before
rotation. The newer units and retrofits are using ball valves which are usually motorized. One problem with
the ball valves is that many separate steam purge lines are required to keep coke from forming on the seal
bellows. If the steam purges are not monitored they can decrease the temperature of the oil going to the coke
drum resulting in high volatile matter coke being produced.

10
Coke Drums

The coke drum diameters range from 4 to 9 meters (13 to 30 ft) with the straight side being around 25 meters
(82 ft) with a 1.5 meter diameter top blind flange closure and a two meter diameter bottom blind flange in
which the 15 to 30 cm diameter inlet nozzle is attached. Both the top blind flange and the bottom must be
removed when decoking the drum. Usually the drum is constructed from 25 mm of carbon steel and is clad
internally with 2.8 mm of stainless steel for protection against sulfur corrosion. The pressure ranges from
1 to 5.9 bars, typically around 2 to 3 bars. The vapor outlet nozzles, 30 to 60 cm diameter, are located at
the top of the drum. Pressure relief valves are also located on the top of the drum on modern cokers.. The
outside of the drum is insulated with around 10 cm (4 in.) of fiberglass insulation with an aluminum or
stainless steel covering. The coke level in the drum is usually determined with three nuclear backscatter
devices mounted on the outside of the drum.

Overhead Vapor Lines

The vapor overhead line runs from the top of the coke drum to the fractionator. The temperature in the line
is around 443°C (830°F) . The temperature is decreased by about 28°C (50°F) by injecting hot heavy coker
gas oil into the line as quench oil. This prevents coking in the line. The heavy coker gas oil is a wash oil
coating the inside of the pipe. If the liquid layer dries out, coke starts to form. Some refineries leave the
insulation off the overhead lines to help drop the temperature and keep the inside wetted. Prevention of coke
in the line is important since this will increase the pressure in the coke drum thus increasing reflux of gas oil
in the drum. Decreasing coke drum pressure increases liquid yield (decreases coke yield). Also, high
pressure drops in overhead lines can cause foaming in the coke drum during the drum switch. Vapor line
sizes are very large in order to obtain the minimum amount of pressure drop. One refinery used two 760 mm
(30 inch) vapor lines in parallel.

Antifoam Injection System

Injection of silicon antifoam should always be furthest away from the vapor overhead line outlet at the top
of the drum to prevent silicon from being carried overhead into the vapor lines to the fractionator. The
heaviest possible antifoam that can be handled in the refinery should be used. Lower viscosity antifoams
appear to break down at lower temperatures and are not as effective. Usually a carrier stream is used to carry
the antifoam into the drum, heavier carrier material would not be as easily flashed off in the drum. Several
refineries are using less antifoam and having less problems with foam since starting continuous injection of
antifoam. A Dow Chemical Company representative stated in 1981 that it is easier to prevent a foam than
it is to kill a foam. Also, when a foam is broken down, it still leaves a mist which can cause coking in the
bottom of the fractionator. A rule of thumb is that antifoam should cost around $0.10 per ton of coke
produced. Costs different than this may indicate that too much or too little antifoam is being used.

Coker Fractionator

The fractionator or combination distillation tower separates the coker overheads into gases, gasoline, diesel,
heavy coker gas oil (HCGO), and recycle. An oversized fractionator can be used to maximize the amount
of diesel product and minimize the heavy coker gas oil to the FCCU. Hot overhead vapors can cause coking
in the lower section of the fractionator if trays are not kept washed (wet). The major amount of heat is
removed in the heavy coker gas oil section by trapping out the oil and then extracting the heat with heat

11
exchangers or steam boilers. This pump-around HCGO is then pumped back into the tray above the trap-out
tray. Some of the HCGO is sprayed below the trap-out tray to wash and cool the hot vapors. Trap-out trays
can be used to catch some of this oil and reduce the amount of recycle oil going back to the furnace. Packing
can be used in fractionators to reduce the pressure drop, but it is critical to keep the packing wet to prevent
coking in the packing. The pressure in the fractionator and also the coke drums is controlled by the gas
compressor at the top of the fractionator.

The fresh feed from the vacuum distillation (VRC) should go directly to the bottom of the tower since the
effective temperature of distillation is higher than in the fractionator. Originally when some cokers were
designed to coke atmospheric reduced crude, the feed was sprayed into the fractionator above the vapor inlet
to fractionate out more light ends in the feed. If VRC is injected above the vapor it condenses out part of
the HCGO into the bottom of the fractionator increasing the recycle to the coker furnace. The bottom of
the fractionator should be operated at as high a temperature as possible without causing coking in the bottom
in order to keep the tube furnace duty low. Normally the temperature in the bottom ranges from 343°C
(650°F) to 382°C (720°F) without coke formation in the bottom of the fractionator. A slotted stand pipe
in the bottom of the fractionator feeds the furnace charge pump.

Hydraulic Coke Cutting System

Cut Water Pump. High pressure water is used to cut the coke out of the drum. Water pressures
range from 86 bars (1250 psig) to 275 bars (4000 psig) and flow rates range from 2.8 cubic meters per
minute (750 GPM) to 4.7 cubic meters per minute (1250 GPM). Cut water pumps are multistage barrel type
or split case multistage pumps which were originally developed for feed water pumps for steam boilers. The
pumps are usually powered with an electric motor, but some older units use steam-driven turbines.

Cutting Equipment. Derricks are built on top of the drum so that the drill stem (5 to 6 inch extra
heavy pipe) can be moved with a winch and cable. The high pressure water flows through an API 10,000 psi
drilling hose to the top of the drill stem. The drill stem is rotated with an air motor at the top through a
rotary joint. The cutting nozzles are the pilot bit with down facing nozzles and the cutting bit with nozzles
facing outward. New units have both nozzles incorporated into a single drilling head.

Coke Cutting Technique. A pilot hole approximately one meter in diameter is drilled from the top
of the drum to the bottom. The pilot hole must be cut down through the coke with minimum weight on the
bit, since if pushed, the bit can follow the main channel in the coke drum, bend, and stick the drill stem in the
coke. After completing the pilot hole, the pilot bit is changed to the cutting bit, and the bottom of the hole
is belled out and opened up to around two meters in diameter to prevent plugging. The bit is then pulled
to the top of the drum and cutting begins by spiraling downward at four to six RPM with vertical movement
of one-half meter per revolution of the drill stem. Usually a vertical four meter section will be cut by moving
the drill stem up and down until the coke is all cut out of the section. Normally around 15 to 20 minutes are
required to drill out the pilot hole and three to four hours to cut the coke. The coke can be cut directly into
rail cars, cut into a crusher car and the coke pumped hydraulically, or cut into a pit or pad with cranes or end
loaders moving the coke.

12
COKE FORMATION, PROPERTIES, AND STRUCTURE

Crude Oil Origin

In order to understand the components of petroleum coke, we must review the origin of the vacuum reduced
crude (coker feed) and crude oil from which it originates. The formation of crude oil is thought to be derived
from ancient remains of animals and plants. The organic matter was squeezed out of the strata probably by
the connate water or water that was originally in formation. Ancient stream beds, reefs, and sand beaches
had the porosity to allow migration of the oil water and provided a conduit for this oil and water material to
flow. The final requirement is a trap for the oil. Oil is not in an open pool but is trapped between layers of
sand or in cracks of limestone. The basic trap is the anticline where the formations were pushed up into a
dome shaped area where the oil and gas accumulated. Also, faulting of geologic formations pushed oil up
against some impervious formations forming fault traps. The salt domes in the Gulf Coast of the United
States have been big producers of oil. Spindletop near Beaumont, Texas was a prime example. Salt laid
down in the bottom of a drying sea, and these formations were covered by over five miles of sediment. The
resultant pressure liquefied the salt which then started to migrate toward the surface. The geologic
formations are pushed up as the salt plug punches through the formations. The salt, being impervious to the
oil, forms an excellent trap.

The vanadium and nickel are in the crude oil as porphyrins or metal chelates. Originally the metals were
probably magnesium (chlorophyll) and iron (hemoglobin). The ratio of the metals to each other is due to
when and how they were buried. Some of the vanadium and nickel can be loosely held between the
asphaltene molecules (intercalation). The other metals are complexed onto the water droplets and probably
were due to the structure that the oil migrated through. Crude oils such as the Paraffinic Pennsylvanian crude
contain very small amounts of asphaltenes. It is possible that the asphaltenes dropped out of the oil phase
but are still down in the formation.

Parts of Crude Oil

Crude oil contains three different fractions. The “Oil” is the hydrocarbon: paraffinic, naphthenic, and
aromatic which also contain sulfur and nitrogen. The second part of the crude oil, the resins, coat the
asphaltene fraction so that it can be peptized into the crude oil. The resins are a brown, sticky hydrocarbon
which contain nitrogen, oxygen, and sulfur, are soluble in n-pentane but insoluble in propane, and have
molecular weights greater than 3000. The asphaltenes contain the chelated metals, vanadium, nickel, and
possibly some calcium along with sulfur, oxygen, and nitrogen. During crude oil distillation, the asphaltenes
are not volatilized and remain in the vacuum reduced crude along with most of the resin fraction. Jakob [9]
thought that all the resins and asphaltenes dropped out in the coker and the remaining coke was made from
the oil fraction. With higher temperatures and lower pressures, the hydrocarbon part of the coke could be
reduced but not the resin and asphaltene fraction. The amount of coke produced in a delayed coker is always
more than the Conradson or Ramsbottom carbon residue percentage by a factor of about 1.6.

Desalter’s Influence on Coke Properties

The crude oil desalter is one of the most critical pieces of equipment in the refinery for producing good
quality anode grade coke (coke low in metals suitable for calcination and use in the aluminum industry) and
keeping a coker furnace on line. Crude oil contains around 0.2% water in which is mixed the soluble salts

13
such as sodium chloride and other metals which are on the edge of the sphere of water. In desalting, the
crude oil is washed with 5% water to remove the salts and dirt from the crude oil . The water, being heavier
than oil, drops out of the bottom, and the cleaned oil flows overhead with around 0.1% water.

Without good desalting, caustic (sodium hydroxide) or filming amines must be added to eliminate the chloride
corrosion in the overheads of the distillation and vacuum distillation lines. The chlorides are usually in the
form of salt (sodium, calcium, and magnesium chloride) The salt content can vary from 50 to 300 lbs per
1000 barrels of crude. Since the number of droplets are high, around 9 x 1011 , the amount of dirt and other
metals on the outside of these water droplets is appreciable. The magnesium chloride causes most of the
corrosion since it breaks down at low temperatures in the distillation column liberating chlorine which forms
hydrochloric acid that attacks the overhead lines in both the vacuum and atmospheric distillation units.

Sodium is a catalyst for burning of carbon (air and carboxyl reactivity in baked anodes) and also causes rapid
tube fouling in the coker tube furnace. The mechanism for rapid fouling of tubes due to sodium is not fully
understood, but it is known that if the tubes are not water washed after steam-air decoking to remove the
crystals of salt, the unit will rapidly foul. Iron in fine particles, probably iron sulfide, is very difficult to
remove in a desalter, but some chemical companies can do a better job than others. Metals that the desalter
does not remove will end up in the coker feed and ultimately in the delayed coke.

Coke Physical Structure

Coefficient of Thermal Expansion. To determine a quantitative value describing coke structure,

the coke is calcined, ground to a flour, mixed with coal tar pitch, extruded to orientate particles into 13 mm
rods, baked to 850°C, and graphitized to 2900°C, and then the difference in expansion at 0°C and 50°C is
measured for Coefficient of Thermal Expansion (CTE) determination. Typical values of CTE corresponding
to coke structure are: needle coke (acicular), 0 to 4; sponge coke, 8 to 18; and shot coke (isotropic),
> 20 (cm/cm/°C x 10-7).

Shot Coke. The production of shot coke in a delayed coker requires high concentrations of
asphaltenes in the feedstock, dynamics (velocity and/or turbulence) in the coke drum, and high coke drum
temperatures. A coker feedstock high in oxygen content can also produce shot coke. When asphaltene
content compared to the Conradson carbon residue content of the coker feed is high, the production of shot
coke is very likely. The present trend in refineries is to run heavier crudes with higher asphaltene contents
and to improve operation of the vacuum distillation unit to produce a heavier VRC with a higher asphaltene
content. This trend toward increased production of shot coke has been observed in refineries which originally
ran atmospheric reduced crude in the delayed coker, never making shot coke, that started producing shot
coke after a vacuum distillation unit was installed.

Shot coke is produced as the oil flows into the coke drum. With the light ends flashing off, small
globules of heavy tar are suspended in the flow. These tar balls rapidly coke due to the exothermic heat
produced by asphaltene polymerization. (Cokers going from sponge coke production to shot coke
production have seen the drum overhead temperature increase by as much as 3°C.) The balls then fall back
into the drum as discrete little spheres two to five millimeters in size. In the main channel up through the
drum, some of the spheres will roll around and stick together forming large balls as large as 25 centimeters.
When these large balls are broken, they are found to be composed of many of the two to five millimeter size
balls. Normally, small shot coke balls from different delayed cokers will be nearly the same size; however,

14
Mexican Mayan VRC has been found to produce larger shot coke balls upon delayed coking than does
Venezuelan VRC. It is thought that smaller balls are made when very high feed rates are used in the coker.
Aromatic feeds, such as decant oil from the FCCU, can help eliminate shot coke formation. All other
methods of eliminating shot coke such as decreasing temperature, increasing drum pressure, and increasing
recycle ratio, will all increase coke yield (decrease more valuable liquid yields) which is not desired.

It is very difficult to produce shot coke spheres in a pilot delayed coker. Spherical shot coke can only
be produced in pilot delayed cokers if the velocity in the drum and the temperature in the drum are both very
high. In a batch (pot) coker, the typical spherical form of shot coke cannot be produced at all; but the shot
coke micro-structure in the batch-produced coke can be seen with a microscope, and the batch-produced
coke does have a high CTE value similar to the spherical form.

Shot coke is unique in that the small spheres two to five millimeters in diameter each have a slick
shiny exterior coating of needle or acicular type carbon. The inside of each sphere contains isotropic or
amorphous type coke as originally described by Marsh and Bacha [10]. Shot coke cannot be used in making
aluminum anodes because the outer needle coke layer of the shot sphere has a very low coefficient of thermal
expansion while the inside of the sphere, being isotropic, has a very high coefficient of thermal expansion.
When rapidly heated in a calcining kiln, the outer layer is cracked and pulled away from the center; thus
when used in an anode with a coal tar binder, the binder adheres to the outer layer (egg shell). This results
in many cracks between the ball and the skin causing the anode to crack and dust in an aluminum smelter cell
[11].

Sponge Coke. Sponge coke is named for its sponge-like appearance and is produced from VRC with
a low to moderate asphaltene concentration. If sponge coke meets strict property specifications, it is
considered anode grade sponge coke suitable for calcination for use in making carbon anodes for the
aluminum industry. Otherwise, if sponge coke meets the more lenient fuel grade specifications, it can be used
in its raw form for fuel.

The biggest problem for refineries producing anode grade sponge coke is obtaining the low volatile
matter (VM) required. The metals and sulfur are strictly controlled by the crudes being processed, but the
VM is in the control of the delayed coker operators. Temperature in the drum is the most critical item, along
with cycle time and drum pressure. Longer residence time at temperature helps to decrease the VM.
Increased recycle can increase the temperature in the drum. Insulation of the transfer line and coke drum,
especially the upper sections of the coke drum, are critical for obtaining low VM coke. Poor insulation and
other bad practices on the delayed coker require higher temperatures in the tube furnace, which results in
shorter campaigns and more downtime for decoking of the furnace. Monitoring the seal steam to prevent
decreasing the temperature in the transfer line, elimination of seal oil on the pressure relief devices on the
transfer line, and minimizing the amount of carrier oil for the antifoam all help in increasing the temperature
in the drum in order to decrease VM of the resultant coke.

Raw or “green” sponge coke must be calcined before it can be used in making anodes. The density
of the calcined coke is critical for producing good carbon anodes. The higher the density, the more carbon
can be incorporated into the anode, and the longer the anode will last. Vibrated bulk density (VBD) of the
calcined coke must be greater than 86 (grams/100 cc). The best single property that correlates from the raw
coke is the Hardgrove Grindability Index (HGI). Raw coke with lower than 70 HGI usually can be calcined
to produce an 86 VBD. Volatile matter is another good property used to correlate how well the raw coke

15
will calcine. Structure is a strong factor in calcinability also, since cokes with low CTE must have volatile
matter much lower than a more isotropic type coke to produce the same density. Porosity of the calcined
coke should be low and is also a function of the raw coke volatile matter. The ash in the calcined coke is
normally around 0.2 % with vanadium and nickel combination under 500 ppm. Sodium and calcium are very
strong catalysts for air burn of an anode. Vanadium, nickel and iron and other metals causes increased
carboxyl reaction in the bottom of the anode. The sulfur in the anode must be below 3.5% to prevent the
sulfur from increasing the electrical resistance of the cast iron connection between the anode and the power
rod. Normally, sulfur is more of an environmental and scrubbing problem. Sulfur can cause the real density
of calcined coke to decrease due to an increase of the porosity and micro cracking of the calcined coke.
Sulfur does help reduce reactivity (air and carboxyl) by reacting with the caustics which are strong catalysts.

Aluminum production requires around one-half kilogram of carbon per kilogram of aluminum
produced. Anode grade coke must be low in metals concentration since the exhaust from the aluminum cell
is being scrubbed with the alumina used as feed to the aluminum cells. Therefore, any metals in the coke
would get into the alumina and into the aluminum metal produced. The carbon is used in the aluminum
smelter as a means of carrying electrical power into the cell. It takes around 15 KW of power per kg of
aluminum produced. A carbon with some porosity must be used since gases coming off the cell would block
the power going into the cell if the anode was not porous. The high temperature along with the very
corrosive fluoride salts used in the aluminum cell and the problem with the evolution of the gases makes the
discovery of a non-consumable anode difficult.

Needle Coke. Needle coke, named for its needle-like structure, is produced from feedstocks without
asphaltenes present, normally FCCU decant oils. Needle coke is the premier coke, used in graphite electrode
manufacturing (used in steel arc furnaces) and commands a high price (calcined ultra-premium non-puffing,
$500 per ton); but needle coke requires special feedstocks, special coking, and special calcination to obtain
the optimum properties that it requires. The Shea patent on needle coke [12] gives an accurate description
of the formation of needle coke, still relevant today. Most needle coke is produced from hydrodesulfurized
decant oil (due to the low sulfur requirement for non-puffing coke, that can be nearly flash graphitized in the
new direct current (DC) length-wise graphitization method, without splitting the electrode). The principle
requirement for needle coke is that the CTE must be 2.0 or below (low CTE is required to prevent spalling
due to the thermal stresses on the tip of the electrode which can be as high as 2000°C/cm). Needle coke must
have low sulfur (<0.6 wt%) and nitrogen contents in order to be non-puffing during graphitization to 2900°C
(measured by a special dynamic puffing test that is proprietary). Needle coke must also have a maximum
amount of coarse sizing (>6 mm), a minimum amount of fines (<1 mm), good density (>78 grams/100 cc;
4/6 mesh test), low ash content (<0.3%; any ash leaves a void when graphitized), and a high real density
(2.13 grams/cc).

Even with all the property specifications, an electrode manufacturer will not accept a calcined needle
coke for production until they have actually run a trial lot through the plant and trials on the electric arc
furnace. Most graphite plants want a needle coke with low variability so that they can set up the optimum
pitch level, extruding and baking to produce a good electrode. The most popular electrode is the 24 inch (60
cm), with a demand for larger than 30 inch (76 cm) for DC single electrode furnaces. Obtaining good needle
coke is still a “black art” for excellent graphite electrodes. The principal property that the electric arc steel
mill wants in a graphite electrode is a low amount of graphite per ton of steel melted. In single electrode
DC furnaces, the amount of graphite per ton is below 2 kg/ton . With better practices and foamy slag, AC
furnaces (using three electrodes due to three-phase electric power) have approached this level.

16
 Pilot Delayed Coker and Coke Formation Model.

In the 1970's, Great Lakes Carbon obtained a design for a pilot delayed coker and modified and
improved the design to produce needle coke identical to that which is commercially produced. Non-puffing
needle coke was in short supply at that time with prices for raw coke running over $600 per short ton. The
pilot unit had a 305 mm (12 in) diameter drum 2.13 meters (84 in) tall, a gas fired tube furnace (both
convection and radiant sections), self-generated recycle (up to 400%), an operating pressure of up to 6.8 bars
(100 psig), adiabatic drum conditions, load cells on feed tanks, overhead receivers, gas meters, and a Ranarex
(for determining molecular weight of overhead gases). A very good material balance could be obtained.
The unit produced around 70 kg (>150 lbs) of coke which was calcined in a gas fired pilot calciner. The
coke drum was decoked by cutting the head off the drum with a metal band saw and then trepanning the coke
out with a large core drill.

Due to the extreme hardness of needle coke, the specified auger type cutter could not be used for
decoking and a core bit was designed and built. By using the core bit, the coke came out in long sections
which could then be sliced with a band saw. From studying these slices of coke, GLC discovered how coke
forms in the drum. This discovery was confirmed by sampling many railcars of commercially produced needle
coke and observing a puzzling strange-looking coke found on the very large piece cut from the lower section
of the drum near the knuckle. This coke was typical except the center contained a dull looking, very friable
coke. Also, looking up at the bottom of the coke bed after the bottom blind flange was removed on a delayed
coker, a hole was observed going up in the bottom of the coke drum. From these observations of
commercial delayed coke and from studying the slices from the pilot unit, the “tree in the drum” channel
branching theory was formulated. This coke formation model is illustrated in Figure 3 shown below [6,13].

Figure 3. Great Lakes Carbon Coke Formation Model: How Coke Forms in the Drum

Structure Orientation In Drum. Samples of coke cut both vertically and horizontally showed that
the coke was oriented in the drum. The coefficient of thermal expansion is much lower in the vertical
direction compared to horizontal, 132% lower in the raw coke, 505% lower in the calcined, and 2850%
lower in the graphitized sample (see Table 2). Gas bubbles, formed from cracking, migrate upward during
coke formation in the liquid, orientating the mesophase chain growth.

Table 2. Coke CTE by Orientation

CTE (cm/cm/°C x 10-7) Vertical Horizontal

Raw Coke 117 154
Calcined 850°C 1.8 9.1
Graphitized 2900°C 0.2 5.7 [13]

Chemical Property Distributions. The ash is mainly in the lower sections of the coke drum with
a high percentage in and near the channels up through the coker. The ash being a particulate drops out of
the oil, or the wall of the channel traps out the particle, otherwise known as the “fly paper effect.” A
cross-section of the coke in the drum was cut into small cubes which were analyzed for ash content. The ash

17
level was five to ten fold higher near the channels compared to the rest of the section. In an experiment
where a paint pigment sized chromium oxide was pumped through the tube furnace into the coke drum in
an attempt to get good dispersion in the coke (chromium oxide is a puffing inhibitor in needle coke), all the
chromium oxide dropped out in one spot in the lower section of the coke drum up about six inches from the
inlet. This is normally the first spot where the main channel starts to branch. Several runs were made with
identical results. Injecting the pigment through the top of the drum distributed the material uniformly in the
coke, indicating that there is some back mixing in the top of the drum either in the froth layer or in the liquid.

Iron, silicon, and ash are in the coke as particulates. These metals concentrate in the lower section of the coke
drum as shown in Table 3 below. Vanadium and nickel are in crude oil as metal chelates or porphyrins in the
asphaltene fraction. It was puzzling that vanadium and nickel are not uniformly distributed in the drum until
it was understood that some of the metals are intercalated in the structure and are not chemically bonded, so
they drop out early in the coke drum similar to the ash and particulates. Volatile matter (VM) in the coke
drum is normally high in the top of the drum due to the short residence time of the material. The sulfur is
uniformly distributed in the drum unless the feedstock to the drum is changing as the drum is filled [13].

Table 3. Property Distributions in the Coke Drum (wppm)

Top Middle Bottom
Iron 40 440 1300
Silicon 60 150 380
Ash 1000 2100 4500
Vanadium 267 310 380
Nickel 130 160 230
VM (wt%) 11.8 9.5 9.4 [13]

18
USES OF PETROLEUM COKE

Raw Petroleum Coke

Fuel Coke. Fuel grade coke (shot or sponge) is used in the production of cement and with fluidized
bed boilers (using limestone for sulfur removal) for generation of steam and electricity. The important
properties for pulverized fuel coke is the cost per BTU, high HGI, and sulfur content. Vanadium in
petroleum coke does not cause corrosion on boiler tubes as does vanadium in heavy fuel oil.

Metallurgy Uses. Some raw petroleum coke, if the sulfur is low enough, can be blended into feed
for slot ovens which produce blast furnace coke. Petroleum coke increases the physical strength and density
of the coke when blended with coal.

Gasification. Partial oxidation of petroleum coke in a gasification process enables raw petroleum
coke to be used to produce steam, electricity, or gas feedstocks for the petrochemicals industry.

Calcined Petroleum Coke - Other Uses

Some calcined petroleum coke is used in production of titanium dioxide (in the chloride process), production
of carbon monoxide for production of plastics, as a feedstock for continuous particle thermal desulfurization
for special low sulfur carbon raiser (steel ladle additive), or as carbon raiser in cast iron and steel making.

19
LITERATURE CITED

1. Conners, J.W., “Changes in Petroleum Coke Quality and Future Prospects,” Union Oil Company of
California, February.12, 1981.

2. Gamson, Bernard W., “Reflections,” Great Lakes Carbon Cororation Research and Needle Coke
Production, Private Letter, September 15, 1983.

3. Peters, C.F., “Additional Reflections,” Great Lakes Carbon Corporation, Research and Needle Coke
Production, Private Letter, October 14, 1983.

4. Court, W.F., “Hydraulic Decoking of Coke Chambers,” Refiner & Natural Gasoline Manufacturer,
Vol.17, No.11, November 1938, A Gulf Publishing Company Publication, pp. 568-573.

5. Purton, Robert M. , Ingersoll-Dresser Pump Company, Telephone Conversation, January 1998.

6. Ellis, Paul J., Hardin, Edward E., “How Petroleum Delayed Coke Forms In A Drum,” Light Metals 1993,
pp. 509-515.

7. Elliot, John D. , “Delayed Coker Design and Operation: Recent Trends and Innovations,” Foster Wheeler
USA Corporation, 1996.

8. Adams, Harry A., “Basic Principles of Delayed Coking,” Adams Consulting Enterprises, Inc.,
January 14, 1994, pp. 1-32.

9. Jakob, Russ R., “Coke Quality and How to Make It,” Hydrocarbon Processing, September 1971,
pp. 132-136.

10. Marsh, H., Clavert, C. , Bacha, J., “Structure and Formation of Shot Coke - A Microscopy Study,”
Journal of Materials Science, March 13, 1984, Chapman and Hall Ltd., pp. 289-302.

11. Ellis, Paul J., Bacha, John D. , “Shot Coke,” Light Metals 1996, pp. 477-484.

12. Shea, F.L. “Production of Coke From Petroleum Hydrocarbons,” U.S. Patent 2,775,549 to Great Lakes
Carbon Corporation, December 25, 1956.

13. Hardin, Edward E., Ellis, Paul J., “Pilot Delayed Coker,” Light Metals 1992, pp.609-615.

20
 FRACTIONAL DISTILLATION TOWER
Less than 104 F
Methane, Ethane,
Propane, Butane
for heating +
Liquid Petroleum Gas cooking + plastics
Reforming Chemicals
104 - 212 F
Unit
Naphtha

104 - 401 F Gasoline & Motor Fuel

Gasoline

350 - 617 F
Jet and Heating Full
Kerosene

Diesel Fuel & Heating Oils

482 - 662 F
Gas Oils
Lubricating Oil Plant

572 - 700 F
Lubricating Motor Oil Cracking
Oil Grease Unit
Polishes
Heated
Crude
Oil Crude Oil
Asphalt
Heavy Oil or
700 - 1112 F Fuel Oils Industrial Fuel
Ship Fuel
Coker
Bitumen
Increase Fractionation Efficiency
and Capacity by Reviewing Liquid
Flow Characteristics:
Propylene Fractionation Example

Charles Nolidin
Karl Kolmetz
Increase Fractionation Efficiency and Capacity
by Reviewing Liquid Flow Characteristics:
Propylene Fractionation Example

1. Introduction
2. Distillation Liquid Flow History
3. Micro and Macro Liquid Flow
4. Propylene Case
5. Summary
 Introduction
1. In Early Distillation Liquid Flow was
only marginally considered.

2. There can be two types of liquid flow

that may either assist or impede
heat and mass transfer; a macro and
micro flow phenomena.
 Introduction

3. This presentation will focus on

the liquid flow in distillation and
present an example in Propylene
Fractionation.
 Distillation Liquid Flow History

1. In Early Distillation Liquid Flow was

only marginally considered.
 Introduction
John French “ Art of Distillation” 1651
 Distillation Liquid Flow History

2. A Kerosene Hydrotreater built in

1952 had one bubble cap per tray
on the Stripper Tower which was
one meter in diameter.
Distillation Liquid Flow History

1 Bubble cap per Tray

1m
 Distillation Liquid Flow History

3. The liquid and vapor traveled

through the center of the column
and mixed on the outer ring
of each bubble cap.
 Distillation Liquid Flow History

5. Early trays were dual flow trays

which then progressed to dual flow
bubble cap trays and down comers
were added.
Distillation Liquid Flow History
 Distillation Liquid Flow History

6. Later when tray efficiency became

more important, flow on the tray
began to be studied and several
types of trays with changed the flow
patterns were developed.
Distillation Liquid Flow History
Two Pass Sieve Deck Tray
Distillation Liquid Flow History
Distillation Liquid Flow History
UOP Multiple Down-Comer Tray 1968
Distillation Liquid Flow History
Distillation Liquid Flow History
Distillation Liquid Flow History
Shell Hi Fi Tray - 1972
 Distillation Liquid Flow History
Trends in Destillatie, The Chemical Engineering Journal, 63 (1996),
pp 167-180.
 Micro and Macro Liquid Flow
1. Micro Liquid flow is when you
consider the small flow interactions
on a tray.

A. Parallel flow across a Circular

Column

B. Poor distributors in a packed

column
Micro and Macro Liquid Flow
 Micro and Macro Liquid Flow
2. Macro Liquid flow is when you
consider the large flow interactions
across a column.

A. Lack of re-mixing between

packed beds

B. Lack of mixing on tray due to

vertical flow
 Propylene Case
1. Two-column in series C3 Splitter was
constructed to produce polymer
grade (99.50 %) propylene.

2. The towers were equipped with 258

dual flow trays. The trays are
corrugated into a sinusoidal wave,
with alternate trays installed with
the waves at right angle.
Propylene Case
Figure 1
Dual Flow Tray Loading Schematic

Rain Space

{
Tray Spacing

{ Average
Froth Height

Corrugated Tray Deck

 Propylene Case
3. The propylene service was
commissioned in late 1999. It achieved
both the nameplate capacity and
propylene product quality.

4. Unfortunately, the propylene loss in the

propane recycle stream was observed to
be significantly higher than the original
design heat and material balance.
 Propylene Case
5. During a high load test carried out in
July 2000, data was collected to
pinpoint the high propylene loss was
attributed to lower tray efficiency.

6. The efficiency was determined to be

in the range of 45% verses the
design of 70%.
 Propylene Case
7. The average propylene in the
propane recycle was averaging 45%,
much higher that the designed 8%.

8. A gamma scan was carried out in

early 2001 to eliminate potential tray
damage. The scan showed all the
trays were still intact.
 Propylene Case
9. However, the liquid density profile
showed mal-distribution occurring
after the first 30 trays of each
column.

10. The decision was made to inspect

the column on the results of the
gamma scan.
 Propylene Case
11. The tower was opened for inspection
during the February 2000
Turnaround.

12. The trays are intact and level but

large 152.4 mm (6”) I-beams and U-
Channels were found laid
perpendicular across the centerline
of each tray.
Propylene Case
Propylene Case
 Propylene Case
13. The I-beams and U-Channels
effectively divided each tray into
four quadrants with no chance for
re-mixing across the tray.

14. Without re-mixing, the bulk liquid

travels down the column as four
different flow paths.
 Propylene Case
15. Below a certain number of trays, the
heavy and light key composition in
the four quadrants becomes
significantly different.

16. If the liquid flow was inconsistent,

the gas flow will follow the path of
least resistance.
 Propylene Case
17. The top of a column will move in a
typical meteorological disturbance.

18. If any hydraulic flow instability were

developed it would remain down the
column. This hypothesis is consistent
with the results from the gamma
scan.
 Propylene Case
19. A decision was made to install six
vapor and liquid re-distributors every
thirty trays.

20. The U-Channel was constructed in

three parts and the middle part was
removed on each of the 14th and
15th trays between the
re-distributors.
Propylene Case
Typical Re-distributor
Propylene Case
Typical Re-distributor
 Propylene Case
Vent

C
W Reflux
Drum
T-30
Fee T-30
d
T-60
T-60

T-90
Side T-90
Dra
w
T-
128
T-
Q QO 130
W

Propylene
Liquid Product
Recycle

Propylene
Stripper Propylene
Rectifier
 Propylene Case
21. With the addition of the vapor and
liquid re-distributors, the tray
efficiency of the column was
increased 10%.

22. The propylene in the propane

recycle was reduced from 45% to
below 10%.
 Propylene Case

23. The tower maximum capacity

achievable before was 112%, and
has presently run as high as 115%.
 Summary
1. Reviewed Early Distillation Liquid
Flow and tray development.

2. Reviewed the two types of liquid

flow that may either assist or
impede heat and mass transfer; a
macro and micro flow phenomena.
 Summary

3. Reviewed the liquid flow in a

Propylene Fractionation example.
The example may conform to the
lack of mixing on a tray case which
could result in lower than expected
tray efficiency.
Air Stripping of VOCs
from Water
VOC removals in excess of 99.99 % are possible with air stripping when
a packed column is used. The packing provides the necessary surface
and turbulence to allow the air to strip the organics effectively down to
very low levels.

In a conventional air stripper, the contaminated water is introduced at

the top of a packed bed. The packing is usually a randomly dumped
plastic media intended to provide a very large effective surface area.
Water flows down the packed bed in opposite direction to ascending
air. The air strips the VOC and exits out the top of the stripper. Clean
water leaves the bottom of the packed section into a sump.

THE JAEGER ADVANTAGE

Jaeger Products has extensive experience in the successful design of

air stripping systems for organic removal and recovery. No other mass
transfer equipment manufacturer has the number of successful installa-
tions. Jaeger=s mass transfer products have become the standard in
air stripping installations.

Our engineering staff can provide you with a complete process design,
and with the necessary engineering, to specify the stripping column in
detail and supply you with all process specification for the peripheral
equipment. We have a complete line of packings, trays, and tower
internals that can satisfy any air stripping needs.

Superior performance by design

JAEGER PRODUCTS, iNC.

1611 Peachleaf, Houston, Texas 77039
Phone:(281) 449-9500 Fax: (281) 449-9400
(800) 678-0345
http://www.jaeger.com
Air Stripping of VOCs from Water

Jaeger Products, Inc.

Houston, Texas

A very prevalent water pollution problem pertains to contamination by organic compounds that offer
potentially very important health risks. These organic contaminants are called volatile organic compounds
(VOCs) when they have a certain propensity to evaporate away from the water into air. Many of these VOCs
are only partially miscible with water but in general they all present a certain solubility. Table 1 lists some
of the more common VOCs found in waters as well as some possible sources. The problem of reducing
VOCs in water applies to ground waters, surface waters, and waste waters alike. The origin of the water
has some important design implications that will be discussed later. This write-up will discuss the
application of air stripping (also known as packed column aeration) for the removal of VOCs from water.

TABLE 1
SOME COMMON VOCs FOUND IN WATER

VOC FOUND IN SOURCE

benzene ground waters gasoline leaks
waste waters process drains and effluents
toluene/xylene ground waters gasoline leaks
trichloroethylene solvent leaks into water table
tetrachloroethylene
trichloroethane
dichloroethanes
trihalomethanes source waters chlorination/ozonation of
waste waters treated waters
vinyl chloride waste waters plastics manufacture
carbon tetrachloride ground waters solvent spills
naphthalene ground waters diesel spills
acetone waste waters solvent spills
ground waters
methyliso-butyl ketone ground waters gasoline leaks
chlorobenzenes waste waters process spills
source waters solvent spills

AIR STRIPPING . . .WHAT IS IT AND HOW DOES IT WORK?

The contaminants shown in Table 1 are considered VOCs because of their "volatility. Volatile compounds
have a natural tendency to leave the water phase and evaporate into air or other gases that are free of the
contaminant. This high volatility makes these organic compounds easy to remove from water by putting
the contaminated water in contact with air that is free of the pollutant. This process is called air stripping
since the VOC is "stripped" out of the water with air. In essence, the VOC is removed from the water and
transferred to the air.

JPI\1996ARST.DOC
Obviously air stripping by itself is not a solution since the pollution is only transferred from one phase to
another. Treatment of the VOC laden air that leaves an air stripper is necessary and several alternatives will
be discussed later.

VOC removals in excess of 99.99 % are possible with air stripping when a packed column is used. Packed
columns offer very high removal efficiencies at very low pressure drops. The packing provides the
necessary surface and turbulence to allow the air to strip the organics effectively down to very low levels.
Other high pressure drop devices have been used in air stripping service such as perforated trays, bubble
aerators and membranes, but undoubtedly packed columns are the devices of choice in air stripping since
they can achieve extremely low effluent concentrations and they are the most economical to operate. This
is especially true in applications exceeding feed flows of 50 gpm where the cost of operating the air
compression systems required for tray strippers become very large.

PACKED AIR STRIPPERS OR TRAY STRIPPERS?

The main advantages packed strippers have over tray strippers are:

a) Packed strippers are more economical and effective for larger flows (above 50 gpm).
b) Packed strippers are more economical when large fractional removals are required (in excess
of 98%), such as in the case of waste-waters and highly contaminated source waters.
c) Packed strippers operate at a fraction of the pressure drop of trayed ones. This reduces the
horsepower requirements for air movement in the stripper.
d) Packed strippers are more economical when low volatility VOCs are to be removed. These VOCs
require higher air/water ratios that are better handled with packed columns.
e) Packed strippers offer better turndown and operability range than tray strippers.

WHAT DOES A PACKED AIR STRIPPER LOOK LIKE?

In a conventional air stripper, the contaminated water is introduced to the top of a packed bed. The packing
is usually a randomly dumped plastic media that is intended to provide a very large effective surface area
per unit volume for mass transfer. Water flows down the packed bed in opposite direction to ascending air.
The air strips the VOC out and exits the top of the stripper. Clean water leaves the bottom of the packed
section into a sump.

The liquid is carefully distributed at the top of the tower by means of a specially designed liquid distributor.
The packing is supported at the bottom of the bed by a support plate that is carefully designed not to
constrain the gas flow. In cases where the required packed bed depth exceeds 25 feet, separately
supported beds are installed with a collection-redistribution stage between them.

Figure 1 depicts an air stripping system complete with off-gas treatment facilities. Materials of construction
for the packing are generally polypropylene or polyethylene. The tower vessel is generally manufactured in
FRP, aluminum, or stainless steel. Tower internals such as distributors and supports can be manufactured
out of polymeric materials, aluminum, or stainless steel. In drinking water applications, all materials should
be compatible with food service and meet FDA requirements.

Filtered air is introduced into the stripper by means of a blower. The blower can be installed in forced or
induced draft configurations depending on the site. Forced draft arrangements, such as the one shown in
Figure 1, are more common and better suited for off-gas treatment.

JPI\1996ARST.DOC
HOW IS AN AIR STRIPPER DESIGNED?

The purpose of this section is to detail some of the general criteria used in designing air strippers for VOC
removal for ground and waste water applications. These guidelines are general and can be applied without
modification to the more conventional VOCs such as trichloroethylene (TCE), perchloroethylene (PCE),
chloroform, benzene-toluene-xylene (BTX), dichloroethylene (DCE), trichloroethane, and carbon tetrachloride.
These criteria apply only when air is the stripping medium and for dilute concentrations of VOCs in water
(below 10,000 ppb).

JPI\1996ARST.DOC
HENRY'S LAW CONSTANT

The value of the Henry's law constant plays an important part in determining the required air to water ratio
in the stripper. It can also be important in the determination of the number of transfer units and the height
of a transfer unit through its effect on the stripping factor.

Reliable data on Henry's constants (H) are not easy to find, especially since these constants are dramatically
affected by temperature and other solutes present in the water. In general H increases with temperature
and concentration of inorganic salts in the water. The effect of temperature should always be considered
whereas the effect of inorganic salts is usually neglected since this represents a conservative assumption.

Henry's law is usually expressed as:

yP = H'x

where y is the mole fraction in the gas phase, x is the mole fraction in the liquid phase, P is the total
pressure and H' is Henry's constant. For the purpose of this document we will consider the Henry's constant
to be in pressure (atmospheres) units and that the total pressure is always very close to one atmosphere:

y= Hx

The numerical values of H and H' are the same since P=1.

The following table shows some estimates of values of H for some common VOCs:

Compound H Temperature deg. C

trichloroethylene 500 20
perchloroethylene 800 20
1,1,1 trichloroethane 700 20
chloroform 180 20
methylene chloride 125 20
o-dichlorobenzene 71 20
BTX 240 20 (other data indicate a value of
400 for H at 20 deg. C)
carbon tetrachloride 1183 20
methyl-ethyl ketone 1.7 20
methyl-isobutyl ketone 7.1 20
ethylbenzene 389 20
1,1,2,2 tetrachloroethane 20 20

These numbers appear to be conservative and could be used for design. Always consult Jaeger Products
for designs at different temperatures. Jaeger will assist you in the estimation of the correct value of H
based on a very extensive, field-proven data base.

JPI\1996ARST.DOC
AIR TO WATER RATIO

The ratio of the volumetric air flow to the volumetric water flow (CFM air/CFM water) is commonly referred
to as the air to water ratio. The recommended value of the air to water ratio varies for different VOCs in
accordance with variations of the Henry's constant and the hydraulic stability of the column. Considerations
of whether off-gas treatment will be required also affect the selection of the design air/water ratio since
larger air volumes will be significantly more expensive to handle in an off-gas treatment unit.

The air to water ratio (A/W) is related to the stripping factor S by the following equation:

S = 0.00075 H (A/W)

Normally, stripping factors of between 10 and 15 work well in cases where no off-gas treatment is required
so that as an initial guess one can use

A/W = 16, 000/H

to estimate the air to water ratio needed for the column. This guideline works well for components like
trichloroethylene and BTX. It should not be used strictly for components that have Henry's constants below
100 since it would yield unreasonably high values of A/W. In reality, values of A/W above 500 are rare in
VOC applications.

Lower stripping factors and air to water ratios are mandated when off-gas treatment is considered. Designs
using stripping factors between 3 and 7 are not uncommon. In these cases the design procedure for the air
stripper becomes more sensitive and delicate. Please consult with Jaeger when your application requires
stripping factors below 10.

MASS TRANSFER EFFICIENCY

There are several ways to predict the mass transfer performance of a given packing under the required
operating conditions. In every case, the required height of packing will be the product of HTU times NTU.
(Packed height = number of transfer units * height of a transfer unit). The values of HTU and NTU to be
used will depend on the stripping factor and the inherent efficiency of the packing. NTU is a variable that
relates exclusively to the stripping factor and the degree of removal. HTU relates to the stripping factor,
liquid load, and the packing efficiency.

The equation to determine the value of NTUol (number of transfer units) for a VOC stripper is:

NTUol = (S/S-1) ln [ (1-1/S)xin/xout + 1/S]

S = stripping factor, xin = inlet concentration, xout = outlet concentration

At values of S = 12 or above, the previous equation can be approximated by

NTUol = ln (xin/xout)

To determine HTU, one can use a correlation or experimental data adapted to the conditions of the design
and applicable to the packing being considered. Methods based on correlations can be fairly reliable if
applied carefully to systems within the proven limits of the correlations. The best method currently available
to the public is a modification of the Onda method that was developed by the Separations Research Program
JPI\1996ARST.DOC
at The University of Texas at Austin. The use of the Onda method without modification is adequate in many
cases and is also used extensively.

Experimental data is hard to come by and hard to validate but it represents the best basis for design. It is
always advisable to compare a design based on experimental data with a correlation method since this will
provide a good sense for the importance of the different variables on the design. Typical design values of
HTU derived from performance data on Jaeger Tri-Packs7 packings are shown in Table 2. The reader should
understand that the selection of the proper value of HTU for design should be done by the packing supplier.
Jaeger Products will provide guaranteed designs for VOC strippers so be sure and consult with us when
finalizing a design.

TABLE 2
HTU VALUES FOR JAEGER TRI-PACKS7 IN VOC STRIPPING

All data for trichloroethylene at 60 deg. F

PACKING A/W ratio Liquid load HTU

2" Jaeger Tri-Packs7 60 25 gpm/ft 2.9 ft

2" Jaeger Tri-Packs7 90 35 3.2

3 " Jaeger Tri-Packs7 60 25 3.4

3 " Jaeger Tri-Packs7 90 35 3.7

EXAMPLE DESIGN

The accompanying chart gives quick estimation of the packed tower requirements for some of the easier
VOC removal applications. In order to use the chart:

a) Determine the necessary air/water ratio for the selected contaminant

A/W = 16 000/H (with H is the Henry's Law constant in atmospheres).
b) With the total water flow to be treated in a single stripper and the desired removal, read the
necessary packed height.
c) The packing size (Jaeger Tri-Packs7 1, 2, or 3.5 inches) is specified next to the column diameter.

NOTE: CONSULT JAEGER FOR DESIGNS REQUIRING HIGHER REMOVALS, LOWER AIR/WATER RATIOS,
LOWER TEMPERATURE, OR DIFFERENT LIQUID LOADING RATES THAN THOSE SPECIFIED IN THE CHARTS

LIQUID AND GAS DISTRIBUTION

The performance of an air stripper is in many cases wrongly related only to the packing itself. In reality, the
packing performs only as well as the initial liquid and gas distribution allows it to. Badly designed liquid
distributors and inlet air nozzles are the most common problems found in non-performing air strippers. Care
should be taken to design and install proper distribution devices in the stripper. For this reason, Jaeger
Products also offers a complete line of liquid and gas distributors to be used in air strippers.

The design and selection of a proper distributor is not simple and cannot be detailed here. For Jaeger to
supply a performance guarantee on a VOC stripper, we would have to supply the internals as well as the
packing or at least have access to review the drawings prior to installation. Contact Jaeger for complete

JPI\1996ARST.DOC
JAEGER PRODUCTS, iNC.
HIGH PERFORMANCE TOWER PACKINS
AND COLUMN INTERNALS SELECTION CHART FOR AIR STRIPPERS FOR VOC REMOVAL
(281) 449-9500 (for estimation purposes only)
Fax (281) 449-9400
(800) 6778-0345

Number in table indicate required packed bed height in ft for a given removal, and air/water ratio. * Recommended packings
size in parenthesis. Numbers to be used only when stripping factor is above 10 and temperature above 55 deg. F.
Consult Jaeger Products for values outside the range of the table.
CONTACT JAEGER PRODUCTS FOR A GUARANTEED DESIGN

Air to water ratio ----------------------------------------------------------- 20 40 60 80 100 150 200

Water flow (GPM) Tower ID (in.) Removal (%)
10 14 90 8' BED 8 8 8 7
*(1" TTRIPACKS) 95 10 10' BED 10 10 9
99 16 16 16' BED 16 15
50 16 90 8 8 8 8 7
*(1" TTRIPACKS) 95 10 10 10 10 9
99 16 16 16 16 15
100 *(2" TTRIPACKS) 24 90 9 9 9 9 9 9 9
95 11 11 11 11 11 11 11
99 17 17 17 17 17 17 17
200 *(2" TTRIPACKS) 36 90 9 9 9 9 9 9 9
95 11 11 11 11 11 11 11
99 17 17 17 17 17 17 17
300 *(3 1/2" TTRIPACKS) 48 90 10 10 10 10 9 9 9
95 12 12 12 12 11 11 11
99 18 18 18 18 18 18 18
500 *(3 1/2" TTRIPACKS) 60 90 10 10 10 10 9 9 9
95 12 12 12 12 11 11 11
99 18 18 18 18 18 18 18
750 *(3 1/2" TTRIPACKS) 72 90 10 10 10 10 9 9 9
95 12 12 12 12 11 11 11
99 18 18 18 18 18 18 18
1000 *(3 1/2" TTRIPACKS) 94 90 10 10 10 10 9 9 9
95 12 12 12 12 11 11 11
99 18 18 18 18 18 18 18
HOW TO USE THIS TABLE: 1) Calculate the required air/water ratio for the component of choice: A/W = 16,000/H where H is the Henry's constant
in atmospheres. 2) At the design flow rate and rewuried percent removal, read the required packed bed height under the specified A/W RATIO.
designs on the optimum internals for your VOC stripper.

AIR POST-TREATMENT

The air emitted from an air stripper frequently needs to be treated to remove the VOC before the air is
exhausted to the atmosphere. This is commonly accomplished by gas phase carbon adsorption or by direct
combustion. The size and cost of the post-treatment equipment is highly dependent on the air rate so that
there is a great incentive to design air stripping systems with as little air flow as possible. This makes the
design of the stripper very critical since all safety factors tend to be removed. Careful evaluation of the
proposed air stripping system at low air flows is required and Jaeger can be of assistance here.

HOW CAN JAEGER HELP YOU IN AIR STRIPPING APPLICATIONS?

Jaeger Products has extensive experience in the successful design of air stripping systems for organic
removal and recovery. No other mass transfer equipment manufacturer has the number of successful
installations. Jaeger s mass transfer products have become the standard in air stripping installations.

Our engineering staff can provide you with a complete process design, with the necessary engineering to
specify the stripping column in detail, and supply you with all process specifications for the peripheral
equipment as illustrated in Figure 1. Our data base is very extensive and chances are there are very few
organics we have not tackled. We can simulate and optimize a complete air stripping and off-gas treatment
unit using the most advanced and comprehensive models. Our calculations will account for unusual
vapor/liquid equilibria and will incorporate the best mass transfer efficiency rating methods available.

We have a complete line of packings, trays, and tower internals that can satisfy any air stripping needs. The
performance of the system depends heavily on the correct internals selection as well as on a good process
design; Jaeger can assist you with both so that total responsibility is easily identified. Although we normally
do not provide turn-key systems, we can direct and/or assist you in such a project. We can also put you
in contact with a systems manufacturer that would provide a turn-key project with Jaeger engineering and
hardware.

JPI\1996ARST.DOC
SOME PITFALLS IN THE DESIGN OF AIR STRIPPERS

Reliability of Henry's constant data

The design of an air stripper depends heavily on the value of the Henry's constant for the target VOC. The
literature abounds with experimental values of Henry's constants but unfortunately they do not always agree
with other published values or with values apparent from field trials and installations.

The Henry's constant is a thermodynamic variable that depends only on temperature and composition. Many
misguided efforts have tried to link the value of Henry's constant to mass transfer performance by regressing
values of H from actual stripping data. This is wrong and dangerous since a fundamental thermodynamic
value is held dependent on totally unrelated things such as liquid distribution, packing shape and size,
column levelness, gas distribution, instrument accuracy, etc.

Values of H derived in such manner should never be used for design since they will prove unreliable in
scale-up and will undoubtedly supply wrongful answers.

The correct procedure is to determine values of H from good experimental data of volatility and solubility
and to determine column efficiency separately using the proper value of H. Consult with Jaeger for proper
and rigorous experimental data interpretation. No other mass transfer company has the level of experience
Jaeger has in using field data to come up with reliable design. There is no magic involved, only the use of
good engineering concepts and thermodynamic principles.

Designs at very low air/water ratios

The necessity of off-gas treatment in air stripping installations has created a very strong incentive for
designing air strippers at very low air flows. Very often the cost of the treatment of the exhaust air
dominates the total treatment cost and is roughly proportional to the air flow.

When the air/water ratio is very low (i.e., stripping factors below 3), the design of the air stripper becomes
very sensitive to the value of the Henry's constant and the packing efficiency. Inaccuracies in the estimation
of these variables of less than 20% can result in calculated packed bed depths that are inadequate by 80
to 120%. It is in these cases where Jaeger Products, Inc., can help you determine the exact tower
dimensions and requirements for reliable results. Crude Henry's constant extrapolations and mistaken mass
transfer data interpretation can lead to grossly under-designed strippers. Consult with Jaeger Products, Inc.
when there is a need to minimize air flows. We have more and better experience than anyone in the
business when it comes to tough designs.

Designs at very low concentrations in the effluent

A pervasive problem in the application of air strippers is the fact that removal of VOCs to very low levels,
say below 0.5 ppb, becomes very difficult if attention is not paid to the detailed design and layout of the air
stripper. Very slight defects in the gas and liquid distribution can lead to strippers not performing properly.
Many designers add extra safety factors to the packed bed depth to compensate for perceived problems
at very low concentrations. The fact is that the best way to ensure proper performance at such low levels
is to ensure that the gas is distributed properly at the bottom of the tower. Guidelines on the layout of the
bottom section of the stripper that set the minimum distances between the gas inlet and the bottom of the
packed bed are very important as is the use of proper feed pipes and spargers. The air entering the tower
from outside needs to be totally free of VOCs for these applications to be successful and measures must
be taken to ensure that the air entering the tower is never contaminated with any of the exhaust air.

JPI\1996ARST.DOC
Cooling of the effluent water

Contact between the air and the water in an air stripper leads to cooling of the water by evaporation. The
amount of cooling depends largely on the water temperature, the air temperature and humidity, and the
air/water ratio. Very efficient packings for VOC removal will also be very efficient cooling devices leading
to potential problems, especially in cold environments. These problems can range from lack of the desired
removal caused by low water effluent temperatures to actual freezing of the water in the stripper. It is
always advisable to check the heat balance of the system for the cold conditions to ensure that freezing has
been properly addressed.

Misuse of safety factors in design

Many performance specifications for air strippers include health safety factors in the inlet and outlet
concentration requirements. Typically the effluent concentration is set at the detection limit of the VOC in
question and the inlet concentration is an absolute maximum that will very rarely (if ever) present itself.

Unfortunately, some mass transfer device suppliers take advantage of this fact when presenting a design
in a competitive situation. Their designs will be based on removals that are below the specified ones with
the "hope" that the specified levels will never present themselves and the performance of the stripper
challenged. Under-designs, such as these, give the false and dangerous impression that a packing is far
better in performance than others. In reality, these suppliers are cutting corners at the expense of
performance reliability of the stripper and are providing designs that do not meet the specified removal but
only meet the effluent characteristic.

The user must be aware of this practice and protect himself or herself against it by strongly requiring and
verifying that all calculations and designs be based on meeting the specified outlet concentration, and the
specified inlet concentration. Jaeger Products, Inc. will always offer you a sound, safe, and economical
design without second-guessing the meaning of the specification. This is why we will always guarantee
our air stripper designs.

Fouling and plugging of packings

Paradoxically, the high mass transfer efficiency provided by the packing in a stripper promotes the
deposition of insoluble metal oxides and bacterial growth. Packings with high surfaces will be more efficient
but will promote fouling as well. There is no magic cure for fouling. The composition of the water, the
irrigation and aeration rates, and the operating temperature have much more to do with how rapidly a tower
will foul than the type of packing used.

If the contaminated water contains free iron or other minerals, the action of the stripping air could cause
some of these compounds to precipitate out and foul the packing media. Organic contaminants promote
biological growth that accentuates the fouling problem. The degree of fouling is usually a function of several
factors beyond the actual inlet conditions. Among them:

1) If the packings or any of the other internals are exposed to ultra violet light, then algae growth
will be accelerated. Algae formation creates an excellent base for mineral deposits, such as
iron, manganese, and calcium.

2) A packing that stays completely and continually wet, thereby constantly washing itself of the
participate, seems to resist fouling and plugging. Fouling and plugging can also be

JPI\1996ARST.DOC
 accelerated by poor initial liquid distribution.

The reality is that all air strippers will eventually lose some of their efficiency and capacity due to fouling
if the water is not pre-treated before entering the tower. The degree of fouling as well as the amount of time
for the fouling to affect the performance of a stripper is a function of all of the above factors plus other
unique characteristics of a particular site. It must also be noted that in many cases the fouling process is
so slow that a contaminated site is essentially cleaned before fouling is a problem.

The best answer to the problem is a combination of good design and pre-treatment. Jaeger Products, Inc.
can assist you in both activities to provide an effective solution to the problem of fouled packing. Our
engineers will properly select the right internals for your tower to assure good liquid distribution and will also
detail the recommended pre-treatment and maintenance options. Pre-treatment involves the continuous
addition of a chemical or compound to the water to keep the minerals from precipitating and to prevent
algae build-up during the stripping process.

All packings foul and, we assure you, the solution to your fouling problem can be found in good maintenance
practices, good monitoring of your process conditions, and good overall process design. A fouling problem
will not be resolved by trying a different packing unless important compromises are made in mass transfer
efficiency. Severely fouled packed beds are inefficient and cause high pressure drop. They can also be very
dangerous since support plates are generally not designed to handle the weight of packing heavily laden
with inorganic salts. In some extremes, the weight of the packed bed can increase by a factor of 10 or more
as the packing fouls.

What can you do to minimize your risk of fouling?

Operate at higher water loads to eliminate dead spots in the packing

Manipulate your water chemistry by pH adjustment
Optimize liquid distribution in the bed
Pre-treat your feed
Clean and maintain packing frequently (at least every year)

There are some fairly effective "in-situ" cleaning techniques that can be very helpful as long as the bed is
not fouled too severely and it has remained wet. One method is to clean while continuing to operate the
tower by pre-treating the feed water with a slight overdosing of our chelating agent. This excess chemical
will break the structure of the oxidized minerals into colloidal suspension and wash them away with the
tower discharge. The rate of clean-up is a function of water chemistry, amount of fouling and percent of
clean-up dosage. Rates of 110% typically clean moderately fouled towers in 90-120 days.
Do not let anyone tell you that there is such a thing as a "perfectly anti-plugging" packing. Such a thing is
called something else, maybe a tray or a spray column, or maybe just wishful thinking. Jaeger Products,
Inc. can assist you in analyzing an existing or potential fouling problem and can provide you with a viable
solution. Jaeger is the only mass transfer equipment supplier that has the in-house capability and expertise
to tackle a tough fouling problem.

JPI\1996ARST.DOC
GUIDELINES FOR EXPERIMENTAL PROCEDURES AND HARDWARE SET-UP FOR AIR STRIPPING
TESTS OF VOCs FROM WATER

The set-up of a test column to perform air stripping experiments to evaluate the removal of VOCs from water
under different conditions has to be looked at carefully if repeatable and reliable results are to be obtained.
When the objective is to measure commercial packing performance and to make comparisons in efficiency,
the proper column set-up becomes critical.

Of equal importance are the experimental and analysis procedures that need to be followed in order to
obtain reliable data. Care has to be taken in how the feed and effluent samples are handled between
collection and actual analysis. Plastic packings are very susceptible to the effects of aging and
contamination by inorganic depositions; the researcher must be careful to assure that all the packings
evaluated have undergone the same conditioning.

This report will address some of the most important hardware and experimental technique considerations
that are relevant when evaluating plastic packings in VOC stripping applications.

Hardware Considerations

Stripping Column - Material. Stainless steel, aluminum, or plastic (preferably fiberglass reinforced polyester)
should be used in test columns to prevent the effects of water-induced- corrosion. If a plastic column is
selected, then care should be taken so that the packed bed section of the column is not exposed to sunlight
since this will harm plastic packings and will promote biological activity. All other internals should be either
stainless steel, aluminum, or plastic.

Stripping Column - Size. The column diameter should be at least nine times the nominal diameter of the
packing under study. A design with ratio of less than 9 to 1 could be subject to severe maldistribution,
channeling and wall flow effects that would translate into poor performance. Furthermore, below this ratio,
different packings would exhibit different degrees of performance deterioration so that comparisons made
under these circumstances cannot be extrapolated to larger columns. Some packings with nominal
diameters of 3 or 3.5 inches exhibit performances similar to 2 inch packings. In these cases, it is probably
adequate (but not preferable) to test such packings in columns as small as 18" in diameter.

The packed height selected for the experiments should be the same for all the packings tested, since
different packings have different abilities to maintain liquid distribution so that their local mass transfer
efficiency is depth dependent. For commercial size packings, a packed bed should be at least 5' deep and
should not exceed 25'. A 10' bed is considered standard. Care has to be taken to ensure that the column
is not so deep that the outlet concentrations go below the minimum reliable detection and quantification
point.

Liquid Distribution. The quality of liquid distribution in an experimental set-up is of critical importance to the
reliability of the data for scale-up or for performance comparison. Different packings react differently to initial
maldistribution and these effects are almost impossible to quantify. This means that the use of an inferior
distributor will have different effects on different packings. Usually, the more modern, high efficiency and
low pressure drop random and structured packings, are affected more severely by initial maldistribution than
the older random packings.

Drip pan or ladder type (perforated pipe) distributors are probably the best type for experimental work since
they can offer substantial turn-down, if enough head is provided. A third choice would be the spray cone,
as long as extreme care is taken to ensure that the nozzles deliver solid and homogeneous cones (which

JPI\1996ARST.DOC
by the way, is rarely the case).

The geometric coverage provided by the distributor should always be in excess of 80%. This means that
at least 80% of the cross section of the column receives the average liquid load. Manufacturers of liquid
distributors can quantify this and should be consulted when procuring a distributor. The pour point density
should be at least of 5 points/square foot and ideally around 10 points/square foot. It seems that there is
little effect of the pour point density beyond 10 points/square foot.

The liquid distributor should be installed so that mass transfer above the packed bed is minimized so that
one sees the performance of the bed without the end effect caused by the distributor. This is a clear
disadvantage of the spray nozzle configuration since the angle of the spray cone has to be minimized and
still provide homogeneous coverage. The distributor should not be located more than three feet from the bed
and ideally, this distance should be on the order of 12" or less.

Gas Distribution. This particular issue is often disregarded in the design of experimental strippers and can
be of extreme importance. Unfortunately, few guidelines exist for the design of effective gas distributors.
A good gas distributor will have a pressure drop on the order of 1/2 to 2 inches of water but can detract
from the capacity of the packing. The best way to ensure adequate gas distribution is to measure the gas
velocity profiles in the empty column, equipped only with the support plate, and if the profiles are
symmetrical, proper gas distribution can be assumed.

The location of the gas inlet nozzle should be perpendicular to the column axis and not tangential. It should
be located at least one column diameter or 20" below the support plate (whichever is less).

Bed Sampling. Intermediate bed samplers, when obtained properly, can be very valuable since they can
provide insight into liquid distribution issues and into how the performance of the packings change with bed
depth. Unfortunately, the insertion of a bed sampler into the column has to, by necessity, disturb the liquid
distribution characteristics of the bed. The design of the intermediate bed samplers has to be done very
carefully, particularly in small (36" and less) diameter columns where high efficiency packing is to be used.
Intermediate samplers should not be used at distances of less than 5' and the diameter of the sampler
inserted into the bed should never exceed 1/2". The length of the sampler should be the same as the
diameter of the column so that a representative sample can be obtained. Any other arrangement will only
introduce a maldistribution.

The sample to represent the bottom of the bed should always be taken at the bottom of the bed itself and
not in the sump below the air inlet if there is the opportunity for mass transfer to occur below the bed. If
proper care has been taken in the placement and design of the distributor, the top sample can be taken from
the distributor itself or from the feed line to it.

Wall Flow. The only important considerations to prevent the effects of wall flow are the proper selection
of the column to packing diameter ratio and the vertical alignment of the column. Trivial as it may sound,
checking the vertical alignment of the column is usually disregarded. Misalignment can have an important
effect on liquid distribution. The use of wall wipers is not necessary and can even be damaging to the
performance of high efficiency packings since they can introduce maldistributions the same way badly
designed samplers can.

Experimental Procedure

Packing Installation. Always follow the recommendations of the packing manufacturer with respect to
installation and storage. Plastic and ceramic packings are delicate and should be handled with care.

JPI\1996ARST.DOC
Packing Aging. It has been amply demonstrated by various researchers that the performance of plastic
packings changes with time due to changing surface characteristics and the slow loss of molding
compounds used in the fabrication of the packing. The surface of a plastic packing suffers some changes
when first put in service. It has been noted that the mass transfer performance of a plastic packing improves
as much as 20 to 30% over a period of about 150 hours of operation. Beyond that, there is little
improvement in mass transfer efficiency and some loss of capacity and increased pressure drops can be
found due to biological growth and inorganic salt deposition. At a minimum, plastic packings should be
tested after about one week of being "aged" in water since it appears that the changes in the surface
characteristics are due to the effects of water swelling and roughing the polymer.

When trying to make performance comparisons among different packings, it is very important that all
packings considered are "aged", or for that matter, "not aged" similarly.

Variables to be Measured and Peripheral Instrumentation. Air and water flowrates, inlet and outlet
compositions, water and air temperatures, and pressure drops are the most important process variables
to be measured. The importance of properly measuring inlet and outlet water temperatures cannot be
underscored enough since the effects on the equilibrium constants can be very great.

The water chemistry can also be of great importance since it affects the equilibrium and can have an effect
on packing performance through the deposition of salts or the existence of dissolved gases. The best way
to evaluate packing performance, particularly for comparison purposes, is by using synthetic, controlled feed
streams. As always, the calibration and accuracy of all instrumentation should be checked frequently since
many instruments, especially flow, composition and pressure drop meters, can drift appreciably.

Sample Taking. All samples should be taken in such a way that the exposure to air is minimized. Sealed
sample bombs should be used and filled to the limit to reduce the potential for exposure to air. The samples
should be chilled immediately before transferring to the analysis vials. Once again, care has to be taken so
the contact with air is minimized in this step. The materials of construction of all sample handling equipment
should be stainless steel or glass. The amount of polymeric materials used in things such as gaskets, lids,
septa, syringes, etc. should be kept to an absolute minimum since VOCs absorb easily into them. Samples
should always be treated to prevent bio-degradation and the formation of other products after they have
been taken.

The useful life of VOC samples has been found to be 48 hours at best. After that time, variations are
observed that can be attributed to biological activity, absorption into the components of sample vials, and
volatilization. It is not recommendable to keep samples longer than 48 hours before analysis, and ideally,
one would perform the analysis immediately after the sample is taken. Samples should never be frozen.

Analytical Techniques. Gas chromatography is the best method for VOC analysis in water. The sample
concentration techniques that can be used vary and are usually the greatest source of error. EPA considers
trap and purge and extraction techniques as the ones of choice but they involve a great deal of sample
handling and offer numerous opportunities for introducing errors. We believe that a technique that reduces
sample handling by humans, such as head space analysis, can be more reliable and repeatable.

JPI\1996ARST.DOC
Calibration curves for the analytical instrument set-up should be run as frequently as possible, and actual
experimental sample analysis should always be mixed with analysis of known standard samples to ensure
accuracy. One should always be aware of the detection and reliable quantification limits of the technique.
This limit is not the same as the detection limit in many cases.

The following pages are provided to help expediate any technical or sales information which you may
require. Please copy, fill out and fax any pertinent information and we will be glad to assist you with a
design.

JPI\1996ARST.DOC
Problems with Air Stripping?

The Problem
One of the most common problems with air stripping towers is that
over time they become fouled with solids, resulting in the loss of effi-
ciency, capacity, and increased pressure drop. The added weight of the
entrapped solids can also be a problem for the packing and other inter-
nals.

The Variables
There are a number of variables that cause plastic packings to foul. The
total surface area of the packing per unit volume is one important vari-
able. However, the most important variables that cause fouling are the
chemistry of the system and the conditions of the process. The
shape of the packing elements, although important to the gas/liquid
contacting, has been proven to have little effect on plugging or fouling
problems.

The Truth
Recently, claims have been made that a particular shape of packing ele-
ment is more resistant to plugging than others. These claims are based
on “tests” in the field, where the variables are anything but controlled.
The difference in fouling rates is surely due to different operating chem-
istry and the conditions of the process, and not the shape of the pack-
ing. Unfortunately, there is no single “truly non-plugging” packing
shape.

The Solution
Over the years, Jaeger Tri-Packs have become the standard by which
G

plastic random packings are measured. In the laboratory, as well as in

the field, Jaeger Products, Inc. has accumulated a wealth of knowledge
on how to deal with packing fouling problems while optimizing your
stripping and absorption efficiencies. The next time you have a strip-
ping or absorbing application, let Jaeger Products put their experience
to work for you.

Superior performance by design

JAEGER PRODUCTS, iNC.

1611 Peachleaf, Houston, Texas 77039
Phone:(281) 449-9500 Fax: (281) 449-9400
(800) 678-0345
http://www.jaeger.com
 Figure 1. Plastic Pall ring fouled with iron precipitate from a groundwater stripper.

Figure 2. Fouling on trough distributor of an air stripper. Pour points can be severely plugged.
How Can You Deal Effectively With Packing Fouling?

Jaeger Products, Inc.

Houston, Texas

Fouling and plugging of packings

Random and structured packings are used in gas/liquid mass transfer operations such as distillation,
absorption, and stripping in order to provide available surface for mass transfer. The mechanisms for the
generation of active surface are varied, but can be summarized into two: formation of films and rivulets, and
formation of drops and drips. In both cases, the geometry of the system is such that the ratio of liquid
volume to surface area is very small.

This small ratio maximizes mass transfer efficiency, but also promotes precipitation of insoluble compounds.
A very common example is the precipitation of iron oxides onto plastic packing surfaces in air stripping
units. These strippers are generally used to remove organic contaminants from source waters to acceptable
limits. Oxygen from the air is simultaneously transferred into the water and this promotes the conversion
of iron to oxidation states that are insoluble in water. These insoluble iron oxides precipitate out of the water
and the crystals attach themselves to any available surface. As soon as a crystal attaches itself, it becomes
a "seeding" site for other crystals to adhere and grow. A complicating factor is that the heavily aerated
water is also an excellent medium for bacterial growth. Bacteria colonies in the water attach themselves
to the packing and provide numerous sites for inorganic deposition and vice-versa. Some forms of bacteria
will use the iron oxides as a nutrient.

Paradoxically, the high mass transfer efficiency provided by the packing promotes the deposition of the
oxides and bacterial growth. Packings with high surfaces will be more efficient but would tend to promote
fouling as well.

There is no magic cure for fouling. The composition of the water, the irrigation and aeration rates, and the
operating temperature have much more to do with how rapidly a tower will foul than does the type of
packing used. Figure 1 shows a conventional Pall ring that has fouled severely in an air stripping application.
Pall rings are not considered among the "high efficiency" packings, but they plug nevertheless. It is
interesting to note in Figure 2 how the trough distributor at the top of the same stripper shows severe iron
fouling as well. In other words, even the low-surface trough plugged significantly. Presumably this happened
because entrained liquid droplets adhered to the surface of the trough, evaporated, and deposited the iron.

Figure 3 shows a similar application where a structured plastic packing, with very high surface area,
suffered severe plugging by iron oxide. It seems that the large surface availability was a hindrance more
than an advantage in this application.

Figures 4 and 5 show pieces of a supposedly "non-fouling" plastic random packing that are actually severely
fouled. These pieces came out of an air stripping unit in an area with high iron water.

All packings foul and one can be sure that the solution to a fouling problem can be found in good
maintenance practices, good monitoring of process conditions, and good overall process design. Severely
fouled packed beds are inefficient and cause high pressure drop. They can also be very dangerous since
support plates are generally not designed to handle the weight of packing heavily laden with inorganic salts.

JPI\1996FOUL.DOC
Figure 4. Modern plastic packing severely with iron deposits. This plastic packing is said to
be “non-fouling” by the supplier.

Figure 5. Modern plastic packing severely with iron deposits. This plastic packing is said to
be “non-fouling” by the supplier.
In some extremes, the weight of the packed bed can increase by a factor of 10 or more as the packing fouls.

What can be done to minimize the risk of fouling?

Operate at higher water loads to eliminate dead spots in the packing

Manipulate water chemistry by pH adjustment
Optimize liquid distribution in the bed
Pre-treat feed with sequestering agents and biocides
Pre-treat with ozone or other strong oxidants
Clean and maintain packing frequently

There is no such thing as a "perfectly anti-plugging" packing. Jaeger Products, Inc. can provide assistance
in analyzing an existing or potential fouling problem and provide a viable solution. Sequestering agents,
ozone/detergent and acid cleaning technology in combination with sequestering agent chemicals are
available from Jaeger that will work on any packing in water service, even the competition's, as long as the
fouling is not excessive.

Keeping Groundwater Air Stripping Units Clean and Unplugged

As mentioned before, if the contaminated ground water contains free iron or other minerals, such as calcium
and manganese, the action of the stripping air could cause some of these compounds to precipitate and foul
the packing media. The degree of fouling is usually a function of several factors beyond the actual inlet
conditions. Among them:

1) If the packings or any of the other internals are exposed to ultra violet light, then algae growth
will be accelerated. Algae formation creates an excellent base for mineral deposits, such as iron,
manganese, and calcium.

2) A packing that stays completely and continually wet, thereby constantly washing itself of the
precipitate, seems to resist fouling and plugging. Thus, fouling and plugging can be accelerated by poor
initial liquid distribution.

The reality is that all air strippers will eventually lose some of their efficiency and capacity due to fouling,
if the water is not pre-treated before entering the tower. The degree of fouling and the amount of time for
the fouling to affect the performance of a stripper are functions of all of the above factors, plus other unique
characteristics of a particular site.

The best answer to the problem is a combination of good design and pre-treatment. Jaeger Products, Inc.
can assist you in both activities to provide an effective solution to the problem of fouled packing. Our
engineers will properly select the right internals for your tower to assure good liquid distribution and will also
detail the recommended pre-treatment and maintenance options. Pre-treatment involves the continuous
addition of chemicals to the water, to keep the minerals from precipitating and to prevent algae build-up
during the stripping process. Jaeger Products has put together a complete treatment/ maintenance package
to address plastic packing fouling in water service. Two processes are available: a pH adjustment process,
using a mild acid solution, and a process that sequesters ions of insoluble salts and prevents them from
precipitating.

Furthermore, ozone or chlorine can be used to attack biological fouling in the contactors as well as in the
packing itself. Without biological growth, the possibilities of inorganic fouling are greatly diminished. Ozone
is an unstable compound in air and has a very short life reverting quickly back to O2. Ozone emissions to the

JPI\1996FOUL.DOC
atmosphere or any post-treatment facilities should not represent a problem. Chlorine on the other hand can
present complications of THM generation. Consult with Jaeger Products, Inc. when deciding which
treatment technique to apply.

An additional option for controlling inorganic salt and oxide deposition is to pre-treat the feed water with
a sequestering agent that will maintain the solids in solution. The selection of the type and dosage of the
sequestering agent can be done very precisely to ensure that the effluent water meets all drinking water
standards and that it can be directed harmlessly to its desired destination. The inorganic polyphosphate
agents that are used in this application have been approved by EPA, the US Department of Agriculture, and
several state health agencies for use in potable water systems at concentrations well above those needed
for treatment. This technique can also be used to clean fouled packings since the sequestering agent will
tend to solubilize the deposited salts. The effectiveness of this wash can be substantially enhanced by
combining it with the ozone treatment procedure outlined above.

Keeping Wastewater Stripping Units Clean and Unplugged

Organic stripping from waste waters presents additional fouling problems than the ones outlined above.
Wastewaters typically have higher concentrations of organics, and in the case of steam strippers, operate
at higher temperatures. These high temperatures often result in inorganic salt precipitation that can severely
foul packed beds.

The use of sequestering agents in wastewater strippers should be evaluated carefully since the consumption
of chemicals can be very high due to the high concentrations and flows. Selection of a preventive process
based on sequestering agents can still be relevant where wastewaters have moderate inorganic
concentrations, such as the ones from chemical plants and oil refineries. Furthermore, correct design of a
steam stripping system will direct the majority of the salt deposition to the feed preheater section of the
process. Heat exchangers that can be easily cleaned are then a necessity.

The most significant fouling problem found in wastewater strippers is caused by bacteria and algae growth
promoted by high organic loads. This problem is more prevalent in air strippers, but it does present itself
in steam strippers as well. Methods to control biological growth in wastewater strippers do not differ much
from those outlined above for groundwater strippers. Good design of liquid distribution systems, combined
with manipulation of the water chemistry, offer the best possibilities for control. Ozonation of the
wastewater is certainly a viable alternative.

JPI\1996FOUL.DOC
How Can a Fouled Stripping Unit be Cleaned?

In many cases, preventive maintenance, such as described above, is not performed. Severely fouled towers
need to be cleaned, preferably without having to remove the packing and internals. Implementation of a
cleaning procedure is not trivial since one has to consider many issues ranging from the chemistry of the
system to the mechanical design of the stripper.

There are two major issues to be addressed in terms of selecting a proper cleaning protocol: liquid
distribution and proper chemistry.

First and foremost, it is essential that whatever cleaning solution is used reaches the fouled areas of the
packing. In towers where the fouling has been excessive, so that areas of the packing are completely
plugged off, it would be impossible to reach the most critical portions of the packing by trickling the liquid
down the media.

If the mechanical design of the tower allows for liquid-full operation, then filling the tower with cleaning
solution and recycling it is the best alternative. On the other hand, this procedure consumes large amounts
of cleaning solutions. The next best choice is to trickle the liquid down the packing at the maximum possible
rate and to feed gas (air, nitrogen, etc.) into the bottom of the tower at a rate that propitiates flooding. The
volume of liquid required in this approach is significantly less but there is the requirement of gas flow.

A stripper that is cleaned before severe fouling occurs, as described above, will be more readily irrigated
properly by the use of the tower distributor. Nevertheless, it is recommendable that the tower be flooded
with the cleaning solution as the first step in the cleaning process.

There are some fairly effective "in-situ" cleaning techniques that can be very helpful, as long as the bed is
not fouled too severely and it has remained wet. Washing the packed bed with a mild acid solution is an
effective technique for removing some inorganic deposits. Ozone injection can clean plastic packings by
breaking down the bacterial colony structure and allowing the salts to fall off the surface of the packing.
Ozone injection, in combination with detergent rinses, provides a good maintenance solution. The user does
need to perform the cleaning with certain regularity to prevent excessive buildup. The use of phosphates
as metal sequestering agents can also be very effective in keeping metal salts and oxides in solution during
heavy aeration.

Most towers can be completely cleaned in 24 to 72 hours. Cleaning time is, of course, a function of the
severity of the fouling problem and the size of the tower. In cases where a tower cannot be out of service
for the entire cleaning period, the process can be alternated with normal tower operation. Additionally, the
final filtration and neutralization process can be performed in a separate holding tank, thus allowing the
tower to be put back in service at the earliest possible time.

Figures 6 and 7 illustrate the effectiveness of the clean-up process. In both cases, the packings were
severely fouled with iron deposits and biological growth. The pictures show how different degrees of
treatment can achieve remarkable results. All these were packings that fouled while in air stripping service
without water pre-treatment. It should be noted that structural damage of the polymeric packing pieces has
never been a problem because of the relative short exposure times and the availability of oxidizable material.

Figure 8 shows a complete cleaning and treatment process for a groundwater stripper with heavy biological
and inorganic fouling tendencies. The inorganic polyphosphate is used to sequester iron, calcium and
manganese ions to prevent their deposition.

JPI\1996FOUL.DOC
The correct selection of the chemicals to be used in the cleanup is also of extreme importance. The nature
of the fouling needs to be identified as extensively as possible so that the proper combination of chemicals,
and in the proper order, can be used. There are some general guidelines that can be established:

1) Biological fouling (bacteria, algae). The chemical of choice for removal of biological fouling is an
oxidant or a free radical generator. Ozone used as described above is very effective and less sensitive to
liquid maldistribution effects. Furthermore, ozone is both an oxidant and a free radical generator. Other
oxidants that are commonly used include potassium permanganate and hydrogen peroxide. Chlorine can
also be used as a radical producer.

2) Inorganic fouling by basic salts and oxides (for example calcium carbonate, iron oxide, calcium
hydroxide, etc). These can be removed by weak acid solutions. Mineral acids, such as phosphoric and nitric,
are frequently used. Organic acids can also be used effectively. A sequestering agent can work in these
applications but it would be significantly slower.

3) Inorganic fouling by neutral or acidic salts and oxides (for example calcium sulfate, iron sulfate,
calcium chloride, etc). Acids will not be effective in removing fouling caused by these compounds. The best
solution here is a combination of a sequestering agent with colloidal agents that can break crystal-crystal
bonds and disperse the pieces. These fouling compounds are the most difficult to remove.

It should be noted that acid cleaning can, in some cases, form new acidic salts that precipitate and
aggravate the fouling problem significantly.

Service from Jaeger Products

Jaeger Products, Inc. offers its customers the necessary proprietary hardware and chemicals to perform
preventive as well as corrective maintenance. The chemicals can be specially formulated for the particular
application and can be supplied to the user on a pre-set schedule. Jaeger also offers the customer several
different service options.

1) Call Out - We will do a site assessment and provide a quotation for the complete cleanup. Our
quotation will include complete mobilization, tower modification (if needed), and all equipment and
treatment process.

2) Contract Call Out Program - We will do a site assessment, and design and build all the
necessary hardware to clean a tower by the end user or a qualified third party contractor. We provide all
chemicals as well as supervisory resources on a time and rate basis.

3) Contract Maintenance Program - We will do a site assessment, and design and build all the
necessary hardware to clean a tower by the end user or a qualified third party contractor. The tower will
be put on a regular maintenance program in which we will make regular follow-up visits to clean the tower
or to supervise the cleaning of the tower by the owner or a contractor. We will also supply all the necessary
chemicals.

4) Continuous Maintenance Lease Program - We will do a site assessment, and design and build
all the necessary hardware to clean a tower by the end user. Lease includes all chemicals, training of operating
personnel, product information updates, and four (4) annual inspections of equipment.

All equipment is packaged on a skid or trailer mounted, pre-wired, and factory tested and ready for operation.
In many cases, one system can be designed to service more than one unit.

JPI\1996FOUL.DOC
Figure 6. Fouled plastic packing cleaned by the use of ozone. First piece is as found in tower before
clean-up. Second piece appears as it was halfway through cleaning process. Third piece is the
final product.
Inorganic Polyphosphates

Jaeger Products, Inc. offers water pre-treatment technology using inorganic polyphosphates. These chemicals
are recognized as non-hazardous, therefore, permitted for human consumption and use in potable water
distribution systems. The chemical is supplied in liquid form. Dosages are calculated for every case after a
detailed analysis of the feed water is performed. They can be adjusted as the water composition varies with
time.

Polyphosphates maintain iron, calcium, and manganese in solution by complexing with the metal ions and
forming large, soluble clusters that prevent crystallization and deposition of the metal salts. In some ways,
these polyphosphates act as "molecular detergents" and can even be used to dissolve or disperse crystals by
tuberculation.

The temperature of the system, phosphate concentration, pH, and the reversion tendency or time stability of
the polyphosphates all play a very important role when assessing pre-treatment possibilities. The conditions
in a stripper can be very severe with respect to phosphate chemistry, since the dilutions are large, the pH can
vary from 2 to 13, and temperatures can be high (i.e., in steam strippers where conditions can exceed 220 deg.
F).

These stringent requirements mandate the use of polyphosphate blends that can operate at high temperatures
and low concentrations without significant reversion. The better polyphosphate blends offer synergistic effects
that cannot be found in single polyphosphate. The technology exists today to produce high-performance
polyphosphate blends that are uniquely suited for use in air and steam strippers. Some packing vendors offer
this expertise to the users, as do some specialized water treatment chemicals suppliers.

Summary

If you are going to install a stripping unit, you should make provisions for keeping the packing clean. A few
inexpensive modifications now can save time and money in the future. As a starting point, you may want to
request our write-up on air stripping volatile organic compounds (VOCs) from groundwater. Request Product
Bulletin No. 600AS.

If you are operating a stripping unit in fouling service, you should begin now to pre-treat the feed water.
Prevention is still the least expensive option. We have pre-treatment options to suit your specific needs.

If you have a stripping unit that is fouled, we have effective technology for cleaning it without the expense and
problem of removing the media from the tower.

For us to recommend a pre-treatment procedure or a procedure for cleaning an existing stripping unit, please
fill out the attached form. This form can be copied and faxed directly to our office.

JPI\1996FOUL.DOC
 900 Process Information

Design Questionnaires
Gas Scrubbing
Steam Stripping
Air Stripping
Problems with Fouling
JP-7 Chelating Agents
Handling/Installation/Operation Instruction

Superior performance by design

JAEGER PRODUCTS, iNC.

1611 Peachleaf, Houston, Texas 77039
Phone:(281) 449-9500 Fax: (281) 449-9400
(800) 678-0345
http://www.jaeger.com
 TERMS AND CONDITIONS
1. Jaeger Products, Inc. is referred to herein as "Seller". The person or company shall remain with Seller until Buyer actually receives the goods or equipment. All
purchasing or to which a Quotation is made is referred to herein as "Buyer". sales are final upon delivery of the goods or equipment to the Buyer.
14. No agent, employee or representative of Seller has any authority to bind Seller to
2. Seller represents that the goods or equipment sold pursuant to this order will be any affirmation, representation or warranty concerning the goods or equipment
free from defects in workmanship or material. SELLER MAKES NO OTHER sold pursuant to this order; and unless an affirmation, representation or warranty
W ARRANTIES, EXPRESS OR IMPLIED, WHETHER OF WORKMANSHIP, PERFORMANCE, made by an agent, employee or representative is specifically included within this
QUALITY, DURABILITY, MERCHANTABILITY OR FITNESS FOR ANY PARTICULAR order, it has not formed a part of the basis of this bargain and shall not in any way
PURPOSE OR USE OR OTHERWISE WITH RESPECT TO MATERIALS SOLD, OR WITH be enforceable.
RESPECT TO ANY PART OR LABOR FURNISHED DURING THE SALE, DELIVERY OR
SERVICING OF THE MATERIAL. IN NO EVENT SHALL SELLER BE LIABLE TO BUYER FOR 15. Any specifications, drawings, plans, notes, instructions, engineering notices or
ANY SPECIAL, INDIRECT, INCIDENTAL OR CONSEQUENTIAL DAMAGES ARISING OUT technical data of Seller furnished to Buyer shall be deemed to be incorporated
OF, OR AS THE RESULT OF, THE SALE, DELIVERY, SERVICING, USE OR LOSS OF USE OF herein by reference the same as if fully set forth. The Seller shall at all times retain
THE MATERIAL OR ANY PART THEREOF, OR FOR ANY CHARGES OR EXPENSES OF title to all such documents and Buyer shall no disclose such to any party other than
ANY NATURE INCURRED WITHOUT SELLER'S WRITTEN CONSENT, EVEN THOUGH Seller or a party duly authorized by Seller. Upon Seller's request or upon completion
SELLER HAS BEEN NEGLIGENT. THE OBLIGATIONS OF SELLER ARE LIMITED TO REPAIR and delivery of the products or services, whichever first occurs, Buyer shall promptly
OR REPLACEMENT OF DEFECTIVE MATERIAL OR, AT ITS SOLE OPTION, TO THE return to Seller all such documents and copies thereof.
REFUND OF THE PURCHASE PRICE. Buyer assumes all risk and liability for results of
using the material covered by this order, whether used separately or in 16. Neither party may assign any of its rights hereunder without the prior written
combination with other products, and agrees to hold Seller harmless for all consent of the other except that Seller shall have the right to assign to any
damages, direct or consequential, resulting from the use thereof. corporation into which it shall be merged, with which it shall be consolidated, or by
which it, or all or substantially all of its assets, shall be acquired.
3. Any affirmation of fact or promise made by Seller to Buyer which relates to the
goods or equipment sold under this order shall not be regarded as part of the 17. If the materials are to be exported, this order is subject to Seller's ability to obtain
basis of the bargain and shall not be deemed to create an express warranty that export licenses and other necessary papers within a reasonable period. Buyer will
the goods or equipment shall conform to the affirmation or promise. Any furnish all Consular and Custom declarations and will accept and bear all
description of the goods or equipment sold under this order shall not be regarded responsibility for penalties resulting from errors or omissions thereon.
as part of the basis of the bargain and shall not be deemed to create an express
warranty that such goods or equipment shall conform to the description. The 18. Seller certifies that any materials described on the front hereof which are
exhibition of a sample or model shall not be regarded as part of the basis of the fabricated by Seller will be fabricated in compliance with all applicable
bargain and shall not create an express warranty that the whole of the goods or requirements of Section 12 of the Fair Labor Standards Act, as amended, of Section
equipment shall conform to the sample or model. 204(c), (d), 212(b), 301-305, 401-403 and 501 of the Fair Labor Standards
Amendments of 1966, and of regulations and orders of the United States
4. The prices stated herein do not include any sale, use or other taxes unless s o Department of Labor issued under Section 501 thereof, and of Section 5(a) of the
stated specifically. Such taxes will be added to invoice prices in those instances in Occupational Safety and Health Act of 1970, as applicable to the fabrication of
which Seller is required to collect them from Buyer; provided, however, that if such materials.
Seller does not collect any such taxes and is later asked by or required to pay
such to any taxing authority, Buyer will make such payment to Seller or, if 19. This clause applies only in the event that the materials ordered herein are to be
requested by Seller, directly to such taxing authority. At Seller's option, prices may used in whole or in part for the performance of government contracts and where
be adjusted to reflect any increase in the costs of Seller resulting from state, the dollar value of such materials exceeds, or may in any one year exceed,
federal or local legislation, or any change in the rate charge or classification of $10,000:
any carrier.
In connection with the performance of work under this contract, the
5. All sales are based upon price per unit per order placed for delivery at one time contractor (subcontractor) agrees not to discriminate against any employee
and upon continuous production of the quantity prescribed. The price stated is or applicant for employment because of race, color, religion, sex or natural
firm for thirty (30) days from date of written quotation. origin. The aforesaid provision shall include, but not limited to, the following:

6. All orders, if accepted by Seller, are not subject to cancellation, change, The provisions of the Equal Opportunity Clause, as promulgated by Executive
reduction in amount or suspension of deliveries, except with Seller's consent and Order 11246 dated September 24, 1965, as amended, are incorporated herein
upon terms which indemnify Seller against direct, indirect, incidental and by reference.
consequential loss and damage.
Where the dollar value of said materials exceeds, or may in any one year exceed
7. All orders are subject to the provision that Seller is not obligated to make delivery $50,000, Seller shall attach upon Buyer's request a copy of the complete text of the
by any specified date, but will use its best efforts to make delivery within the time Equal Opportunity Clause, as promulgated by Executive Order 11246 dated
shown. Seller assumes no responsibility or liability for any loss or damage occurring September 24, 1965.
by reason of delay or inability to make delivery caused, directly or indirectly, by
acts of God; war; force of arms; fire; the elements; riot; labor disputes; picketing or 20. This document contains the entire agreement between Seller and Buyer and
other labor controversies; sabotage; civil commotion; accidents; any constitutes the final, complete and exclusive expression of the terms of the
governmental action, prohibition or regulation; delay in transportation facilities; agreement, all prior or contemporaneous written or oral agreements or
shortage or breakdown of or inability to obtain or non-arrival of any labor, negotiations with respect to such terms as are included herein or are the subject
material or equipment used in the manufacture or fabrication of the material matter hereof being merged herein. By way of illustration and not limitation, Buyer's
covered hereby; the failure of any party to perform any contract with Seller order shall be deemed to incorporate, without exception, all the terms and
relative to the production of the material covered hereby; or from any causes conditions hereof notwithstanding any order form of Buyer containing additional or
whatsoever beyond Seller's control, whether or not such cause be similar or contrary terms or conditions, unless Buyer shall have expressly advised Seller to the
dissimilar to those enumerated, and if delays from such causes occur, delivery contrary in a writing apart from the printed provision of such order form, and no
time shall be correspondingly extended. In the event that Seller's inability shall be acknowledgment by Seller of, or reference by Seller to, or performance by Seller
partial and not complete, Seller may allocate its remaining supplies and deliveries under an order of Buyer shall be deemed to be acceptance by Seller of any such
in any manner that is fair and reasonable among its regular customers, its additional or contrary printed terms or conditions, then such modifications may
contract customers and its own requirements. only be made in these terms and conditions by a written instrument signed by one
of Seller's executive officers.
8. Delivery of ten percent (10%) more or less than the quantity specified shall
constitute fulfillment of the order and any excess, not exceeding ten percent 21. In the event that any word, phrase, clause, sentence or other provision hereof shall
(10%), shall be taken and paid for by Buyer. violate any applicable statute, ordinance, or rule of law in any jurisdiction in which
it is used, such provision shall be ineffective to the extent of such violation without
9. Claims for shortage or rejection for defects must be made in writing within ten (10) invalidating any other provision hereof.
days after date of shipment. Merchandise shall be returned only upon Seller's
written authorization. Seller accepts no responsibility for merchandise returned 22. This document and the sale of any products hereunder shall be governed by and
without such authorization. Seller shall not issue credit on any merchandise which construed in accordance with the laws of the State of Texas.
has been altered or defaced in any way or upon which any additional operations
have been performed. Any lawsuit or other action based upon breach of this 23. Any Quotation furnished by Seller is subject to, and shall not become binding upon
contract or upon any other claim arising out of this sale (other than any action by Seller until, actual receipt by Seller of Buyer's written order based on all the terms
Seller for the pur-chase price) must be commenced within one year from the and conditions stated in the Quotation, without qualification. Seller may, at its sole
date of the tender of delivery by Seller. option, accept a verbal order from Buyer, whether or not Seller has furnished a
Quotation. If Seller has furnished a Quotation and Buyer places a verbal order,
10. Seller reserves the right at any time to alter or suspend credit or refuse shipment or Seller may rely on such Quotation with respect to the quantity and specifications of
cancel unfilled orders when in its opinion the financial position of the Buyer, or the the material ordered. On all verbal orders, acceptance of delivery of material by
status of Buyer's account, warrant it, or when delivery is delayed by fault of Buyer, Buyer or its agent shall be conclusive evidence against Buyer that such material
or Buyer is delinquent in any payments. conforms to the order. Modifications in Buyer's written confirmation of a verbal
order shall not be binding on Seller unless actually received by Seller prior to the
11. Seller reserves the right to make changes in goods or equipment manufactured time Seller places a non-cancellable order for the material or ships the material,
without incurring any obligation to make corresponding changes upon goods or whichever occurs first. Buyer agrees to indemnify and hold Seller harmless from any
equipment previously manufactured or sold. and all cost and expense (in-cluding without limitation any restocking charges)
incurred by Seller as a result of any misunderstanding or errors arising out of a
12. The finish of goods or equipment purchased pursuant to this order includes only verbal order by Buyer.
such finish as directly obtained from the manufacturing process.
24. All design data and recommendations provided by Seller are based on specific in-
13. Unless otherwise specified by Seller, all material and prices are F.O.B. Point of formation provided by Buyer with respect to temperature, operating conditions,
Origin, and payment will be net/cash 10 days form date of invoice. No deferment and other critical matter. Recommendations are provided at no cost, are not
of ship-ment at Buyer's request will be made except on terms that will indemnify intended to be relied on by Buyer, and do not constitute any warranty, express or
Seller against all loss and additional expense, including, but not limited to implied by Seller. Buyer agrees to indemnify, defend and hold harmless Seller from
demurrage, handling, storage and insurance charges. Identification of the goods and against any and all liability, cost, damage or expense arising out of or in
or equipment shall occur when delivered to the shipper by Seller, or ten days after connection with the furnishing of design data and recommendations to Buyer by
order confirmation is sent to Buyer, whichever shall first occur. Risk of loss of the Seller and releases Seller from any liability, cost, damage or expense to Buyer, its
goods shall pass to the Buyer upon identification. Title to the goods or equipment officers, directors, employees or invitees arising out of or in connection with the
provision of such design data and recommendations by Seller.
 RETURN TO:
Jaeger Products, Inc.
1611 Peachleaf St., Houston, TX
281- 449-9400 (fax)
jpadmin@jaeger.com
page 2
DISTILLATION
Designing packed distillation columns can become quite involved. A few up-front
questions will help both you and Jaeger determine how best to proceed:

Do you have a running steady-state Aspen or Hysys simulation for the column or column
train?
no
yes
If yes, could you send the simulation files to Jaeger? yes no

If sending a complete simulation file presents difficulties could you supply a spreadsheet which
summarizes, stage by stage: (page by page simulation results are not as useful as a summary compiled in Excel by you)
vapor and liquid mole fractions
vapor and liquid mass and molar flows
vapor and liquid mass densities
vapor and liquid molecular weights
vapor and liquid viscosities
vapor and liquid mass diffusivities
temperature and pressure

Do you want Jaeger Products, Inc. to complete a simulation for you?* yes no
Please email or fax us a preliminary process flow diagram (PFD) for the separation you envision.
We will contact you for further information as we need it.
*Depending upon the manpower required to complete the simulation we might charge you at the
rate of $200.00/hr. Part of this charge will be recoverable in the purchase price should you decide
to buy from Jaeger.

Plastic Packings Jaeger Decides

 1” 1.25” 2” 3.5”
Jaeger Tri-Packs
Cascade Mini-Rings  1A 2A 3A
Plasric Jaeger Rings 5/8” 1” 1.5” 2” 3”
Pastic Jaeger Saddles 1” 2” 3”

Materials: PP PE PVC CPVC PFA ECTFE Other

Metal Packings Jaeger Decides

Interpack (mm) 10 15 20 30

VSP (mm) 25 40 50

Top-Pak (mm) 50

Jaeger Max-Pak® metal structured packing. Other structured packing:

Materials: 304 316 316L 410 Inconel Hastelloy Other

Scrubbing Pollutants
from Vent Streams
Scrubbing is a common name given the unit operation normally known
as gas absorption. In this process, mass is transferred from the gas
phase into a liquid for the purposes of removing material from the gas
stream. Wet packed scrubbers can achieve extremely high contami-
nant removals and can operate at a variety of loads. Scrubbing can
achieve simultaneous removal of various contaminants as well as pro-
vide a measure of gas cooling and particle emission control.

Gas absorption will play a very important role in controlling pollutants

to bring industry into compliance with the requirements of The Clean
Air Act. The number of absorption applications will grow profusely in
the next 5 years.

THE JAEGER SCRUBBER ADVANTAGE

Jaeger Products, Inc. has extensive experience in the successful

design of scrubbing systems for ammonia, acid gases, and organics
removal. Our engineering staff can provide you with a complete
process design. We have a complete line of packings, trays, and
tower internals that can satisfy any scrubbing or absorption need such
as:

SO2 removal from stack gases

H2S removal for odor control
Removal of HCL, chlorine oxides, and chlorine
Alcohol, ketone, ether, and aldehyde absorption
Sulfuric and nitric acid emission control
Organics removal
CO2 and H2S removal from gas streams.
Vents from microelectronics facilities
Vents from chemical plants and refineries
Vents from the pulp and paper industry
Vents from rendering plants
Vents from fermentation plants

Jaeger Products will assist you in the design of all your absorption and
scrubbing needs. Our high efficiency products and our commitment to
excellence in engineering will assure you of a successful application
every time. Jaeger has engineered more pollution control systems
than any of its competitors and we can put this experience to work for
you. Give us a call.
Superior performance by design

JAEGER PRODUCTS, iNC.

1611 Peachleaf, Houston, Texas 77039
Phone:(281) 449-9500 Fax: (281) 449-9400
(800) 678-0345
http://www.jaeger.com
Waste Gas Scrubbing.

Jaeger Products, Inc.

Houston, Texas

Gaseous emissions containing inorganic as well as organic pollutants are highly regulated and becoming
more so in light of the Clean Air Act. Emission control of gas streams in an industrial setting can include
very large applications associated with power generation as well as small vent and emergency release
scrubbers. Packed column technology is well proven in most gas scrubbing applications, and is the
technology of choice in the majority of applications due to low initial and operating cost, high efficiency, and
reliability.

WHY USE SCRUBBING?

Scrubbing is a common name given the unit operation normally known as gas absorption. In this process,
mass is transferred from the gas phase into a liquid for the purposes of removing material from the gas
stream. Obviously, the material to be transferred has to be soluble in the liquid for the process to be
effective and herein lies one of the most important aspects of scrubbing: selection of the proper scrubbing
liquid.

In most air pollution control applications, the materials to be removed are usually acidic or basic gases even
though in some cases, organic vapors are also present. Most of the inorganic contaminants are sulphur,
phosphorous, halogen, or nitrogen oxides or acids. In other cases, ammonia or chlorine are of concern.

Wet scrubbers can achieve extremely high contaminant removals and can operate at a variety of loads.
Operation and control are simple, very stable, and ideal for remote, unattended locations. Scrubbing can
achieve simultaneous removal of various contaminants as well as provide a measure of gas cooling and
particle emission control. With the correct configuration, a scrubber can be used to remove acids and
alkalies, or even soluble organic compounds. Scrubbing produces a liquid waste stream in many cases that
has to be dealt with, but in the majority of the cases, this liquid stream does not require post-treatment
beyond neutralization.

WHAT IS SCRUBBING?

A waste gas stream is put in contact with a scrubbing liquid in a contactor, most commonly a packed
column. The scrubbing liquid exhibits high solubility for the contaminants in the gas and these migrate from
the gas into the liquid. In many instances, the liquid contains a reactive solute that enhances the degree
of absorption by reacting with the contaminant once it dissolves and effectively removing it chemically.
This chemical absorption process allows for effective scrubbing of fairly insoluble gases, such as SO2,
chlorine, and H2S.

The scrubbing process takes place in a packed column where the packing provides the necessary surface
area and turbulence to achieve the desired removal. The scrubbing liquid is distributed at the top of the
packed bed and it "rains" down flowing through the bed where it comes in intimate contact with the gas.
This contact allows the scrubbing liquid to remove a contaminant from the dirty gas.

JPI\1996SCTX.DOC
Important variables in the design of a packed scrubber include the following:

--types and amounts of contaminants to be removed

--gas flow, temperature, molecular weight, and humidity
--type and composition of scrubbing liquid
--amount of dust present
--available or allowable pressure drop for system
--effluent limitations in terms of composition, temperature, entrained liquid
--means for disposal of purge scrubbing liquid

PACKED SCRUBBERS VS. OTHER TYPES

Many different types of scrubbers are used in industry. Spray, venturi and jet scrubbers are generally used
when large amounts of solids are present and when heat transfer is required in quenching applications
Packed scrubbers are indicated when high efficiency is required and when flexibility is desirable.

The main advantages packed scrubbers have over others are:

a) Packed scrubbers are more economical and effective for larger flows.

b) Packed scrubbers are more economical when large fractional removals are required (in excess
of 98%).

c) Packed scrubbers operate at a fraction of the pressure drop of trayed ones. This reduces the
horsepower requirements for air movement in the tower.

d) Packed scrubbers are more economical when high volatility compounds are to be removed
because they offer very large transfer areas and the ability to operate counter-currently.

e) Packed scrubbers offer better turndown and operability range than any other type of scrubber.

WHAT DOES A TYPICAL PACKED GAS SCRUBBER UNIT LOOK LIKE?

The configuration of a scrubbing unit can vary depending on the characteristics of the material to be
removed from the gas stream . Figure 1 depicts a countercurrent waste gas scrubber that can handle
ammonia and acid gases in the same piece of equipment. This configuration would be typical of power
generation applications and of cases where combustion gases need to be cleaned. Two separate cycles
are present; the bottom section of the scrubber utilizes acidic water to remove ammonia and some of the
more soluble acids. Ammonia is highly soluble in water at low pH values, so the acid loop absorption can
be very effective. On the other hand, the acid components of the stream, including SO2, are fairly insoluble
in pure or acidic water but very soluble in caustic solutions at high pH values, thus the use of a caustic
absorption loop.

The caustic solution is routinely a solution of sodium hydroxide (NaOH) in water and the by-products of
absorption are a variety of soluble sodium salts. More flow is normally needed in the tower itself to provide
good contact than the replacement flow required for make-up of spent caustic. The most economical
arrangement involves a recycle loop for the scrubbing liquid with small make-up and purge streams.

JPI\1996SCTX.DOC
The acidic water solution loop operates in a similar manner except that the low pH value is achieved by the
absorption of acids from the feed gas. The purge from this cycle will be a sour water concentrated in
ammonia. Suitable disposal or treatment needs to be available for this stream.

Other gas scrubbing applications involve only the acid gas absorption step. These applications can be found
routinely in the pulp and paper industry where removal of HCL, chlorine, and chlorine oxides is needed from
vent streams. Sodium hydroxide or sodium carbonate solutions are frequently employed as the scrubbing
liquid.

Odor control applications involve scrubbing H2S out of air. This is essentially an acid gas absorption
application that uses a caustic scrubbing liquid. Vent stream from diverse manufacturing operations, such
a anodizing, plating, or other metals treatment often have acid gases present that can be easily removed
in a scrubber.

A very important application of scrubbing is the treatment of vent gases from process vessels and tanks.
These vents can sometimes be continuous and part of routine operations or they can be emergency
releases from safety devices. In any case, these vent streams are likely to contain a large number of
contaminants to be removed, ranging from acid gases to organic vapors. It is not unusual to have separate
scrubbers for inorganics and organics depending largely on water disposal constraints.

Finally, in many cases it is necessary remove moisture from process gases. This can be accomplished by
absorbing or scrubbing water out of the gas stream using a hygroscopic solvent. Solvents can be liquids,
such as sulfuric acid, glycols, or other heavy, hygroscopic compounds.

CROSS-FLOW SCRUBBERS

In many cases, and especially when a chemical reaction in the liquid phase is involved in the scrubbing
process, cross-flow scrubbers are used. These devices operate in a similar manner to the traditional
countercurrent devices with the exception that the gas flows horizontally in the packed bed. The liquid is
still fed at the top and flows downward through the packing, but the contact between the liquid and the gas
takes place in a cross-flow arrangement. Cross-flow scrubbers can be very effective especially when using
a caustic solution to scrub an acid gas. They exhibit lower pressure drops and slightly higher capacities and
have a horizontal profile that can be desirable when space restrictions exist. In general though,
countercurrent scrubbers are capable of higher removal efficiencies and are able to minimize the amount
of scrubbing liquid needed.

Jaeger Products, Inc. can assist you in the selection of the most economical flow arrangement for your
scrubber given the constraints of your site and required performance.

Independently of the objective of the scrubber, the necessary hardware to be installed in the tower is similar
regardless of scrubbing application. Figure 1 shows the packing and internals normally associated with a
wet countercurrent scrubber. Various materials are available on all packings and internals to accommodate
hot, corrosive, erosive, and dusty services.

JPI\1996SCTX.DOC
DOES RECYCLE OF THE LIQUID HELP?

Many scrubbers operate with a recycle liquid stream from the bottom to the top. This is done when the
material to be scrubbed reacts quickly in the liquid phase so that the actual concentration of the pollutant
in the liquid is always very low. Recycle allows the scrubber to operate at reasonable wetting rates without
the need for excessive make-up or purge of liquid. This arrangement is typical of ammonia and acid gas
scrubbers.

Recycle, on the other hand, cannot be used when the pollutant does not undergo reaction in the liquid
phase, such as scrubbing organic compounds into water or a solvent. In this case, once-through operation
of the liquid is indicated to prevent accumulation of the pollutant in the liquid inlet.

CHEMISTRY, CHEMISTRY, CHEMISTRY!!

Scrubbers that remove reactive pollutants, such as ammonia and acid gases, depend heavily on the proper
chemistry in the liquid phase for their effectiveness. It is of crucial importance that the designers and
operators of scrubbers understand the chemistry of the system since more scrubbers fail because of faulty
chemistry than any other reason.

This is of particular importance in systems that include a multitude of pollutants, since interaction among
them can be serious. An excellent example is the typical mixed vent from a microelectronics manufacturing
facility that includes HCL and ammonia. These two compounds react in the gas phase to make ammonium
chloride, a white microscopic powder that is extremely difficult to remove and results in pervasive white
plumes.

Jaeger Products has more experience than any other mass transfer supplier in tackling tough scrubbing
problems from the chemistry to the equipment.

SOME PITFALLS IN SCRUBBING SYSTEM DESIGN.

Several aspects of the design of scrubbing systems are very crucial and not immediately obvious. First,
is the accuracy and reliability of equilibrium and chemical reaction data. The thermodynamic model of
choice for scrubbing systems is one based on activity coefficients that can predict solubilities as a function
of temperature and composition. Two models that fit this function very well are the Wilson and Van Laar
activity coefficient models. Interaction parameters for these models are readily available for most of the
common contaminants.

Waste gases can be very fouling, especially when heavily laden with ash and inorganic salts. In typical
scrubbing configurations, most of the fouling will occur in the bottom of the scrubber and design provisions
are needed to allow for frequent cleaning. In many cases, especially when significant gas cooling is needed,
it is recommended that a quench tower be used ahead of the scrubber to take the brunt of the fouling and
cool the gases down. Cooler gases can be handled in a scrubber using less expensive materials of
construction, if adequate control logic protection is provided to prevent damage to the scrubber if the
quench were to fail. The quench tower can be packed or equipped with trays.

Design at low scrubbing solution make-up rates is desirable since it reduces the downstream processing
requirements. Nevertheless, care must be taken to assure that the scrubbing power of the solution is not

JPI\1996SCTX.DOC
diminished. Common practice limits the amount of absorption to a maximum of 25% of saturation (be it by

chemical reaction in the case of reactive absorption, or solubility when no reaction takes place) in the
recycle loop. Some operations run at total recycle for a period of time and then replace the solution charge
when it reaches 25% saturation. Others have a continuous make-up and purge.

Mass transfer will also occur from the liquid to the gas if conditions are favorable. Care must be exercised
to account for any evaporative losses of the scrubbing liquid to the gas stream. This is especially important
when non-aqueous scrubbing liquids are utilized.

Most absorption processes are exothermic since the materials removed exhibit significant heats of solution
and/or reaction with the scrubbing liquid. A good example of this is the large heat release encountered in
hydrochloric acid absorption. The design of the system, the selection of the liquid rate, and the selection
of the materials of construction should take into account the temperature rise of the liquid caused by these
heat effects.

THE SCRUBBER AND OTHER COLUMNS IN THE SYSTEM.

The contacting devices in the scrubbing system are where the transfer of mass takes place. Commonly,
they are vertical countercurrent vessels filled with a mass transfer device. In general, these devices are
either sieve trays, random packings, or structured packings (the level of efficiency and capacity follows the
same order and so does their sensitivity to fouling). In some other cases, cross-flow arrangements can be
had when height limitations are present.

The scrubbers are also equipped with liquid distributors and support plates for the packing. In the case of
deep bed requirements, intermediate liquid collectors and redistributors are also installed to ensure good
performance. Figure 2 shows different combinations of internals that can be installed in a scrubber. In most
cases though only combinations of trays and packings (with the associated internals) are used. Jaeger
Products, Inc. offers all internal devices necessary for scrubbers and quench columns in a variety of designs
and materials to suit the application.

HOW CAN JAEGER HELP YOU IN SCRUBBING APPLICATIONS?

Jaeger Products, Inc. has extensive experience in the successful design of scrubbing systems for ammonia,
acid gases, and organics removal. Our engineering staff can provide you with a complete process design,
specify the contacting column in detail, and supply you with all process specification for the peripheral
equipment as illustrated in Figure 1. Our data base is very extensive and chances are there are very few
contaminants we have not tackled.

We have a complete line of packings, trays, and tower internals that can satisfy any scrubbing or absorption
need. Since the performance of the system depends heavily on the correct internals selection, as well as
on a good process design, Jaeger can assist you with both so that total responsibility is easily identified.
Although we normally do not provide turn-key systems, we can direct and/or assist you in such a project.
We can also put you in contact with a systems manufacturer that would provide a turn-key project with
Jaeger engineering and hardware.

JPI\1996SCTX.DOC
 THE JAEGER ADVANTAGE

Typical Scrubbing Applications

SO2 removal from stack gases

H2S removal for odor and sulphur emission control
SO3 removal from sulfuric acid plant vents
Removal of HCL, chlorine oxides, and chlorine
Alcohol, ketone, ether, and aldehyde removal from vent streams
Sulfuric and nitric acid emission control
Gas dehydration using hygroscopic solvents
Organics removal by use of heavy, large molecular weight solvents
CO2 removal from gas streams
Vents from microelectronics manufacture
Vents from pulp and paper operations
Vents from rendering plants
Vents from chemical plants and refineries

Scrubbing facts

Capable of achieving very high removals and low gas effluent concentrations
Most economical removal technique at all gas feed concentrations
Cost effective at very low feed concentrations (ppm level)
Minimizes air emissions
Reduces loads to incineration
Can be operated at vacuum or pressure depending on process needs

Typical hardware for scrubbers

Sieve trays for fouling service (SS, Monel)

Metal random packings for many applications (SS, Monel)
Ceramic and carbon random packings for high temperature service
Plastic random packings for acid service (PP, GFPP, Noryl, PVDF, Teflon)
Metal structured packings for high efficiency/capacity (SS, Monel, Aluminum)
Tower internals in appropriate materials including distributors, supports and demisters

JPI\1996SCTX.DOC
SAMPLE CALCULATION FOR AN ACID GAS ABSORBER

The first step in sizing a scrubber is to determine the column diameter. This is done based on recommended
gas velocities. The second step is to determine the necessary liquid flow based on a recommended liquid
loading and the column diameter. Once the liquid load and the column diameter are determined, the
required packed bed depth will be determined. The variables that specify the packed bed depth are: Liquid
and gas loads (mass velocities), Removal (concentrations in and out), Packing efficiency, and Type of
system.

The calculation of the necessary design parameters vary with the type of application, the type of scrubbing
liquid, the gas concentration levels, and the type of gas being scrubbed. The information provided herein
is for illustration purposes only and is intended only to provide guidance. Consult Jaeger Products,
Inc. for a final and guaranteed design.

DESIGN INFORMATION: (generally supplied by the user)

Gas flow = 47 273 lb/h or 11 250 ACFM
Contaminant = H2S
Concentration in = 200 ppm
Concentration out = 0.1 ppm
Temperature = 68 deg. F
Pressure = 1 atm.

RECOMMENDATIONS: (based on Jaeger's experience and data base)

Use NaOH solution as scrubbing liquid
Design gas velocity = 300 ft/min (typical for H2S scrubbers)
Design liquid load = 3 gpm/ft2 (typical)

CALCULATIONS:
Column diameter = [ (11 250 acfm/300 ft/min)/0.785]1/2
= 6.9 ft-----Use 7 ft. (calculated from gas flow and recommended gas velocity)

Liquid flow = 3 X 72 X 0.785 = 115 gpm = 57 470 lb/h

(calculated from column diameter and design liquid load)

Gas mass velocity = 47 273 lb/h / 72 / 0.785 = 1229 lb/h/ft2

(calculated from gas flow and column diameter)

Liquid mass velocity = 57 470 lb/h / 72 / 0.785 = 1494 lb/h/ft2

(calculated from liquid flow and column diameter)

**Number of transfer units = NTU = ln ( Xin/Xout) = ln (200 ppm/0.1 ppm) = 7.6

(mass transfer term that depends on absorption factor and removal. Definition used here
applies only to very dilute gases and absorption with very fast chemical reaction)

** Height of a transfer unit = HTU = 19.4" for equivalent loadings and 3 1/2 in. TRI-PACKS7 (from
mass transfer data table in Figure 3)
(mass transfer term that depends on loadings and the characteristics of the packing)

Packed bed height = NTU X HTU = 12.3 ft ---------Use 14 ft

(**Consult Jaeger for values for your application)

JPI\1996SCTX.DOC
PRODUCT BULLETIN 600

PLASTIC JAEGER TRI-PACKS®

Properties Table
SPECIFICATIONS
Size (in.) 1 1 ¼” 2 3 ½”

Materials Geometric
Twelve standard, injection moldable plastics are available: Surface Area 85 70 48 38
(ft2/ft3)
Polypropylene (PP) TopEx® (LCP)
Polyethylene (PE) Kynar® (PVDF) Packing
Polypropylene Halar® (ECTFE) Factor 28 25 16 12
(1/ft)
Glass-Filled (PPG) Teflon® (PFA)
Noryl® (PPO) Tefzel® (ETFE) Void
Polyvinylchloride (PVC) Tefzel® Glass- Space 90 92 93.5 95
CorzanTM (CPVC) Filled (ETFE-G) (%)

Bulk Density
(lb/ft3) 6.2 5.6 4.2 3.3
Other plastics are available on request. (PP)

Sizes. Plastic Jaeger TriPacks® packings are madein four

sizes:
1” Nominal
11/4“ Nominal
2” Nominal
31/2” Nominal

IMPORTANT NOTE:
Design data presented in this bulletin are for
preliminary calculations only. Contact Jaeger before
finalizing calculations.

JAEGER TRI-PACKS® is a Registered Trademark of

JAEGER PRODUCTS, INC.
Superior performance by design

JAEGER PRODUCTS, iNC.

Figure 3. typical mass transfer performance data for
scrubbing applications using Jaeger tri-Packs®
Steam Stripping

Stream stripping for water clean-up is essentially a distillation

process where the heavy product is water and the light product is a
mixture of volatile organics. These organics are present in the feed
water, in relatively small concentrations. Since the volatility of the
organics is a very strong function of temperature, the high stripping
temperature inherent in stream stripping allow for the removal of
heavier more soluble organics that are not strippable with air. No
off-gas treatment is needed ad the only wastestream generated is a
small amount of very concentrated organics.

The Jaeger Advantage

Jaeger Products, Inc. has extensive experience in the successful

design of steam stripping systems for organic removal and recovery.
Our engineering staff can provide you with a complete process
design, and with the necessary engineering, specify the contacting
column in detail. We have c complete line of packings, trays, and
tower internals that can satisfy any steam stripping need.

Typical Steam Stripping Applications

Benzene removal from wastewaters

Sour water (h2O and NH3) stripping
Acetone removal/recovery from wastewaters
Oxygenate (MTBE. MEK) removal/recovery
Removal of chloroform, bromoform and
other halogenated organics from water
Removal of organics from quench waters
Organics recovery from leachates
Alcohol (ethanol, propanol, IPA, butanol)
removal from water
Solvents recovery or removal
(tetrahydrofuran, hexane, heptane)

Superior performance by design

JAEGER PRODUCTS, iNC.

1611 Peachleaf, Houston, Texas 77039
Phone:(281) 449-9500 Fax: (281) 449-9400
(800) 678-0345
http://www.jaeger.com
Removal of Organics From Water Using Steam Stripping

Jaeger Products, Inc

Houston, Texas

Dilute mixtures of organic materials in water can be concentrated by a process known as steam stripping.
The end products of such operation are a clean water stream almost devoid of organic materials, and a
highly concentrated organic stream suitable for recycle to a process or for disposal. The use of heat in the
form of steam as a separating agent offers significant advantages over other methods, such as inert gas (air)
stripping.

WHY USE STEAM STRIPPING?

Steam stripping for water clean-up is essentially a distillation process where the heavy product is water and
the light product is a mixture of volatile organics. These organics are present in the feed water in relatively
small concentrations. The process of steam stripping takes place at high temperatures compared to air
stripping, usually very close to the boiling point of water. Since the volatility of the organics is a very strong
function of temperature, the high stripping temperatures inherent in steam stripping allow for the removal
of heavier, more soluble organics that are not strippable with air.

Another very important feature of steam stripping is the fact that no off-gas treatment is needed, and that
the only waste stream generated is a small amount of very concentrated organics. These are easily dealt
with by incineration, biological treatment, or recycled to process.

In summary, steam stripping is a good solution for wastewater streams that contain fairly soluble,
non-volatile organics and where no off-gas stream is desired. On the other hand, steam striping does
necessitate the presence of steam (or process heat) and would tend to be more capital intensive than air
stripping. Ideal settings for steam stripping are oil refineries, petrochemical, and chemical plants.

WHAT IS STEAM STRIPPING?

A wastewater stream is heated and put in intimate contact with steam in a packed or trayed tower. The
combined effects of the steam and heat, or temperature cause organic material to transfer from the liquid
to the vapor phase. This material is then carried out with the vapor. As contacting proceeds down the
tower, the wastewater becomes leaner in the organic material while the vapor phase becomes more
enriched as it travels up the tower.

Steam is injected at the bottom of the tower to provide heat and vapor flow. Clean water leaves the bottom
of the tower. The wastewater is fed at the top of the tower and the steam leaves the top heavily laden with
organic material. This steam/organic combination is condensed and processed further as detailed in the
next few pages. The net effect achieved in the steam stripper and condenser is that a contaminated
wastewater and steam are injected into the tower and a clean water stream is obtained. A low-volume,

JPI\1996STMT.DOC
but concentrated water/organic mixture, is also obtained as a by-product.

WHAT DOES A TYPICAL STEAM STRIPPING UNIT LOOK LIKE?

The configuration of a steam stripping unit can vary depending on the characteristics of the organic material
to be removed, and on what is to be done with it in terms of disposal and recycle. As a minimum, a steam
stripping unit will look like the unit depicted in Figure 1. It is important to note that heat recovery from the
bottom product is necessary for economical operation. Operations at reduced pressure do not need
recovery exchangers, but operate at lower temperatures and larger steam rates. The towers also tend to
be a bit larger in vacuum operations.

Steam requirements for stripping vary with the operating pressure, the type of organic, and the degree of
organic removal/recovery. Further, steam requirements for heat balance purposes need to be accounted
for. A very important consideration in the design of a steam stripper is the fact that the column needs to
be capable of handling enough steam flow to operate without the benefit of the recovery exchanger. This
feature will be needed during start-up and when the exchanger is out of service for cleaning.

Some organic materials are not totally miscible in water and separate into a distinct organic phase when
the concentration exceeds the solubility limit. Most aromatics and halogenated organics fall in this
category. Steam stripping applications for these types of compounds can be very effective, since a good
part of the concentration of the organic can be accomplished in a decanter as indicated in Figure 2. In this
case, the water layer is recycled to the stripping column for reprocessing. The design of the decanter poses
some interesting questions since the water flow is generally significantly larger than the organic flow.
Furthermore, in some cases (benzene, toluene, etc), the organic layer is the lighter of the two liquid phases.
In applications involving halogenated organics, the organic liquid is heavier than water. Needless to say,
good models to predict the phase behavior of the system in question are essential.

Figures 3A and 3B are refined versions of the flowsheet in Figure 2. These arrangements are needed when
better organic recoveries are needed from more dilute streams. The selection between Figure 3A and 3B
depends solely on the equipment sizing. Figure 3A is used when required steam flows are larger (less
volatile compounds).

Figure 4 is applicable when the organic material to be removed exhibits very high solubility in water. In this
case, a refluxed distillation column is needed to achieve high organic concentrations.

Other variations on the same flowsheets shown above include the use of reboilers instead of direct steam
injection and operation at reduced pressure to reduce operating temperature.

JPI\1996STMT.DOC
CHEMISTRY, CHEMISTRY, CHEMISTRY!!

It is of crucial importance that the designers and operators of steam strippers understand the chemistry of
the system, since lack of operability and maintenance problems occur frequently because of faulty
chemistry.

This is of particular importance in systems that include a multitude of pollutants, since interaction among
them can be large. An excellent example is the typical mixed wastewater from a chemicals manufacturing
facility that includes inorganic acids, organic pollutants, and dissolved gases. As the gases, such as CO2
and or NH3, are stripped, the pH of the water changes causing potential solids precipitation. This is
aggravated by the fact that steam stripping temperatures often exceed the precipitation temperature for
salts, such as calcium carbonate.

The volatility of the compounds to be stripped is often affected by the water chemistry present. Accurate
predictions of the volatility are of extreme importance for proper stripper design; the operators of stripping
systems should always be aware that changes in the chemistry of the incoming water can affect the
removal efficiencies observed in the stripper.

Jaeger Products, Inc. has more experience than any other mass transfer supplier in tackling tough stripping
problems from the chemistry to the equipment.

SOME PITFALLS IN STEAM STRIPPING SYSTEM DESIGN.

Several aspects of the design of steam stripping systems are very crucial and not immediately obvious.
First is the accuracy and reliability of equilibrium data. Steam stripping is a situation where the old reliable
Henry's law just isn't applicable due to the broad concentration ranges, high temperatures, extensive
interactions between components, and the existence of two liquid phases. The thermodynamic model of
choice for steam stripping systems is one based on activity coefficients that can predict immiscibility. No
model fits this function better than the NRTL activity coefficient model (non-random two liquid model
developed by Prausnitz and co-workers). Pilot and laboratory tests to establish the adjustable parameters
in the NRTL model for the mixture in question are advisable, but solubility and vapor pressure data can
suffice as a good approximation.

Wastewaters can be very fouling, especially when the temperature is raised and inorganic salts precipitate.
In typical steam stripping configurations, most of the fouling will occur in the recovery exchanger and
design provisions are needed to allow for frequent cleaning. In the absence of a recovery exchanger, the
brunt of the fouling will be taken by the stripper itself. In such cases, the use of trays can be a way to avoid
plugging even though packings would yield better performance characteristics. The use of sequestering
agents is also a good solution for reliable and lengthy operation.

Materials of construction need be some grade of stainless steel or a high performance plastic due to the
varied and changing nature of the water chemistry. Capital savings by use of lesser materials of
construction generally translate into severe problems and added expense later.

Start-up of any steam stripper requires heating of the feed water to the operating temperature in the
stripper. This added heat has to be supplied in the form of steam at the bottom of the stripper. Design
provisions need to be made to accommodate this larger, but temporary, steam flow in the stripper. This
capability is also desirable to allow for continued operation while cleaning of a fouled recovery exchanger
takes place.

JPI\1996STMT.DOC
Design at low stripping steam rates is desirable since it reduces the downstream processing requirements.
Figure 5 illustrates how sensitive the process is to steam flow. Optimum designs require stripping factors
between 1.5 and 4. These stripping factors mandate more stages for separation and taller packed heights.
Design under these conditions becomes very sensitive to the reliability of the equilibrium data and the mass
transfer models. This is also the case where excellent packings and internals are necessary and where
vendor experience in design of steam stripping systems is invaluable.

THE STEAM STRIPPER AND OTHER COLUMNS IN THE SYSTEM.

The contacting devices in the steam stripping system are where the mass transfer takes place. They are
vertical countercurrent vessels filled with a mass transfer device. In general, these devices are either sieve
trays, random packings, or structured packings (the level of efficiency and capacity follows the same order
and so does their sensitivity to fouling).

The columns are also equipped with liquid distributors and support plates for the packing. In the case of
deep bed requirements, intermediate liquid collectors and redistributors are also installed to ensure good
performance. Figure 6 shows different combinations of internals that can be installed in a steam stripper.
In most cases though, only combinations of trays and packings (with the associated internals) are used.
Jaeger Products, Inc. offers all internal devices necessary for steam strippers and distillation columns in
a variety of designs and materials to suit the application.

HOW CAN JAEGER HELP YOU IN STEAM STRIPPING APPLICATIONS?

Jaeger Products, Inc. has extensive experience in the successful design of steam stripping systems for
organic removal and recovery. Our engineering staff can provide you with a complete process design, and
with the necessary engineering, specify the contacting column in detail, and supply you with all process
specification for the peripheral equipment as illustrated in Figure 7. Our database is very extensive and
chances are there are very few organics we have not tackled. We can simulate and optimize a complete
steam stripping and solvent recovery unit using the most advanced and comprehensive models. Our
calculations will account for unusual vapor/liquid equilibria and will incorporate the best mass transfer
efficiency rating methods available.

We have a complete line of packings, trays, and tower internals that can satisfy any steam stripping needs.
The performance of the system depends heavily on the correct internals selection as well as on a good
process design; Jaeger can assist you with both so that total responsibility is easily identified. Although
we normally do not provide turn-key systems, we can direct and/or assist you in such a project. We can
also put you in contact with a systems manufacturer that would provide a turn-key project with Jaeger
engineering and hardware.

JPI\1996STMT.DOC
 THE JAEGER ADVANTAGE

Typical Steam Stripping Applications

Benzene removal from waste waters

Sour water (H2S and NH3) stripping
Phenol recovery
Acetone removal/recovery from waste waters
Oxygenate (MTBE, MEK) removal/recovery
Removal of chloroform, bromoform and other halogenated organics from water
Removal of various organics from quench waters
Concentration and organics recovery from leachates
Alcohol (ethanol, propanol, IPA, butanol) removal from water
Solvent recovery or removal (tetrahydrofuran, hexane, heptane)

Steam stripping facts

Capable of achieving very high removals and low effluent concentrations

Most economical removal technique at feed concentrations above 0.1% weight organics
Cost effective at feed concentrations as low as 200 ppm
Can produce a re-usable concentrated product
Minimizes air emissions
Reduces loads to incineration
Can be operated at vacuum or pressure depending on needs with little penalty
Can be made very energy efficient with heat recovery
Fouling is a continuous concern

Typical hardware for steam strippers

Sieve trays for fouling service (SS, Monel)

Metal random packings for most applications (SS, Monel)
Plastic random packings for acid service (GFPP, Noryl, PVDF, Teflon)
Metal structured packings for high efficiency/capacity (SS, Monel, Aluminum)
Column internals to include: distributors, redistributors, supports, and mist eliminators

JPI\1996STMT.DOC
HANDLING / INSTALLATION / OPERATING

INSTRUCTIONS

PACKINGS

COLUMN INTERNALS

Jaeger Products, Inc.

1611 Peach Leaf
Houston, TX 77039
 TABLE OF CONTENTS

RANDOM PACKING, SHIPMENT, STORAGE AND HANDLING I

RANDOM PACKING INSTALLATION INSTRUCTIONS II

MAX-PAKJ INSTALLATION INSTRUCTIONS III

INTERNALS SHIPPING AND HANDLING IV

INTERNALS INSTALLATION INSTRUCTIONS V

OPERATION AND MAINTENANCE VI

MANUFACTURER'S WARRANTY VII

JPI\INSTLMAN.DOC Rev.9 (03/97)

I. RANDOM PACKING SHIPMENT, STORAGE, AND HANDLING

A. Shipment Description - Jaeger Tri-Packs7 packings are shipped to location by either UPS, LTL
motor freight a contract dedicated truck. The standard shipping containers for packing are
corrugated cardboard boxes containing 10 cubic feet of packing each. UPS shipments may
arrive in shipments over several days depending on size of shipment and UPS truck
availability. LTL and contract trucks are loaded without pallets to facilitate the maximum cubic
footage possible per truck. Larger plastic shipping bags (approx 60 ft3) are available upon
special request and at an additional charge.

B. Storage and Handling - Plastic packings are shipped in cardboard boxes which are not designed
for outside storage. Packings can be shipped in bags (option at additional cost) for long term
outside storage. Care should be taken when handling boxes to prevent boxes from tearing open
or corners becoming crushed.

1. Inside Storage of Standard Boxes.

a. Stack boxes on wooden pallets or on level floor so that there is no danger of water or
condensation. Make sure the entire bottom of each box is fully supported.

b. Do not stack boxes more than three high and tape boxes together as a group. If the
packing is to be stored in climate with high humidity (90% or more) for extended
period of time, then reduce stacked height to two boxes.

c. Do not put any load on top of stacked boxes. Never allow any person to climb on or
store other equipment on top of any of the boxes.

2. Outside Storage of Standard Boxes.

a. Outside storage of standard boxes is not recommended, however if no other alternative

is available, follow recommended stacking as inside storage to maximum of two high.

b. Completely cover boxes with weatherproof canvas or plastic. It is important that there
is air circulation among the boxes so that condensation does not form on the inside of
the cover.

c. Boxes of packing stored outdoors will not support any load. Make sure rain water, ice,
or snow do not build up on top of the covered boxes.

3. Storage of Optional Bulk Bag

All bags should be kept closed to avoid direct sunlight on packing. Bags are designed for
storage of up to 6 months and require minimal care. Bags are supplied with corner straps
for lifting and can be requested with bottom chutes for installing from bottom of bag into
tower.

JPI\INSTLMAN.DOC Rev.9 (03/97)

Page 3
II. RANDOM PACKING INSTALLATION INSTRUCTIONS

A. General Information
There are many ways to install random packing. Installation procedures vary depending on the
column diameter, size of packing, packed bed height, and position of other internals.
Generally, if common sense is used problems can be avoided. For small diameter columns of
less than 30 inches, it is usually easier to load through lower level manways or use a chute to
lower the packing to the support level. For larger columns, it may be possible to place men
inside so that complete boxes or bags can be lowered and manually dumped. Extreme caution
should be used in this procedure to make sure that the packing support is designed to hold the
combined load of the workman and the packing. Another method is to attach a rope on the top
of a bucket and another on the bottom and lower the packing to the desired level. The rope
attached to the bottom would then be used to tip the bucket over. Finish all welding or other
work that produces hot sparks before installing any packing. Test the dimensional
compatibility of the individual packing pieces with the openings in the support grid. Make sure
that it is not possible for packing to squeeze through any openings.

B. Installation Procedures

1. Distribute packing pieces in a random manner and never let them free fall more than four
(4) feet. In cold weather (temperature below 45EF) particular care must be taken in
handling and installing plastic packing since the ductility of the plastic resin will be greatly
reduced.

2. Make sure the packing fills all the space in the packed bed section. Pay particular attention
to the manway and irregular spaces.

3. Check the position of sensors when they become buried in packing. Be careful not to drop
packing on sensitive process measurement devices.

4. If workmen are in the tower, use plywood to distribute their load over as large an area as
possible. Never exert a concentrated load onto a few of the individual packing pieces.
Check with the packing support manufacturer for loading limits before placing men inside
the tower.

5. Before completing the packing installation, consult the installation procedures of the bed
limiter, liquid distributor, and other internal components to see if partial installation of
these parts is required before packing installation is completed.

6. Be careful not to leave any foreign materials in the packed bed section. Make sure all
plywood, boxes, and bags are removed.

7. If procedures are not clearly understood by the personnel installing the packing, they should
call the packing manufacturer for additional information required.

JPI\INSTLMAN.DOC Rev.9 (03/97)

Page 4
 C. Special Considerations for Ceramic Packings

In addition to instructions and guidelines for installation of random dumped packing, the
following offers additional information to protect the fragile ceramic packings. Ceramic
packings may be dumped randomly or stacked depending on the type of packing. Most
ceramic packings up to 3" are dumped. Cross partition which can range to 6" diameter are
generally installed with the column filled or partially filled with water.. If possible, a
minimum of 4 feet of water should be maintained above the packing depth during installation.
This volume of water will offer buoyancy that will reduce the impact of each element and
offer maximum randomness of each element providing the ideal packed bed. It is to be noted
that not all columns or towers are designed to sustain a full hydrostatic load. It is imperative
that you consult with the vessel supplier before proceeding to place a full hydrostatic load on
the vessel.

III. MAX-PAKJ INSTALLATION INSTRUCTIONS

The installer should read these instructions in their entirety before beginning installation.

A. General Information

Max-PakJ Structured Packing is fabricated from sheets of corrugated sheet metal that are
welded together in a specific pattern using a proprietary electric welding procedure that makes
the packing significantly sturdier than other structured packings. These corrugated sheets are
stacked to form the packing elements. These corrugated sheets will end in a vertical position
in the tower so that the axis of the corrugations rungs at a 45E angle to the axis of the column.

The complete cross section of the column is to be filled with packing so that layers of
structured packing, having the same diameter of the column and depth of one element, are
formed. Each layer is rotated 90E with respect to the one immediately below. Each layer is
comprised of several packing elements of varying length and width, but always of the same
depth.

Each packing element, and hence each layer, is 12.25" deep in the vertical direction. The
width of the element will vary depending on the manhole size used to introduce the packing
in the vessel, and can be from approximately 6" to 15". The length of the elements depends
on the tower diameter, and will be a maximum of 6'. In many cases, a complete circular
element is supplied, to be inserted through a body flange in the column. These circular
elements can be anywhere from 6" to 48" in diameter.

The portion of each packing element that comes in contact with the vessel wall is fitted with
two or three wiper bands designed to pick up liquid flowing down the wall, and to minimize
gas flow along the wall. These wiper bands have tabs that are to be bent outward at a 45E
angle before inserting the packing element, or positioning against the wall of the vessel.

JPI\INSTLMAN.DOC Rev.9 (03/97)

Page 5
 The packing elements are designed to be inserted through manholes, and to be handled by not
more than two people at a time. THE CORRUGATED SHEET METAL THAT IS USED TO
FORM MAX-PAKJ IS EXTREMELY SHARP, AND CAN INFLICT SERIOUS CUTS AND
LACERATIONS. INSTALLERS SHOULD WEAR LEATHER GLOVES, LONG SLEEVES,
LONG PANTS, STEEL TOE SHOES, AND THICK SOLES AT ALL TIMES.

B. Installation

1. The packing layers are formed one at a time, from the bottom of the bed upward. The first
layer of packing will rest on a grating support (Jaeger PS02), and will be oriented 90E with
respect to the main pattern of the grating support.

2. The elements that go against the wall of the vessel will be installed first, and the
construction of the layer should progress from the outside toward the center of the tower.
Contact between the individual elements of one layer should be maximized, and no visible
gaps between elements should be present once the layer is completed. The final elements
to be installed in a layer will be those toward the center. When an element needs to be
positioned between two elements already in place, a slide horn should be used to slide the
element down. This slide horn will consist of a thin sheet of metal of the appropriate
dimension, and should slide out from between the elements once the layer is finished.
Always make sure the slide horns are removed from each layer before proceeding with the
next layer.

3. A check of levelness of the layer should be performed before proceeding the next layer.

4. Installers should not walk directly on the packing to avoid potential injury, and damage to
the edges of the corrugated sheets. Wooden planks that sit on top of the packing elements
should be used to spread the load.

INSTALLERS SHOULD ALWAYS FOLLOW PLANT, SHOP, AND OSHA RULES AND
PROCEDURES FOR VESSEL ENTRY, SCAFFOLDING, MONITORING, ETC. THESE
INSTALLATION INSTRUCTIONS ARE INTENDED ONLY TO PROVIDE A GUIDELINE
FOR HOW TO POSITION THE PACKING IN A VESSEL, AND ARE NOT TO BE
CONSTRUED IN ANYWAY AS REPLACEMENT FOR SAFETY PROCEDURES. JAEGER
WILL NOT BE RESPONSIBLE IN ANYWAY FOR THE SAFETY OF OTHERS PERFORMING
INSTALLATIONS OF MAX-PAKJ STRUCTURED PACKING.

IV. INTERNALS SHIPPING AND HANDLING

Shipment Description - Internals manufactured by Jaeger Products are generally shipped crated
or palletized. The internals are generally shipped via LTL motor freight or by a contract dedicated
truck. No special handling or storage requirements are required as long as standard freight
handling/reporting procedures are followed. All crates should be inspected for any damage upon
receipt and notations made on original freight bill. Fiberglass internals should be of special
inspection as they are more susceptible to hidden damages than those made of metal or

JPI\INSTLMAN.DOC Rev.9 (03/97)

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 thermoplastics. Damage to fiberglass internals typically appears as a lighter color or "white" spot
indicating possible delamination of the fiberglass structure.

V. INTERNALS INSTALLATION INSTRUCTIONS

A. General Information

The following information and instructions for installation of internals are guidelines
representing the most common applications. Most internals are custom made and may have
slightly different parts, particularly clips and sealing devices that may be required. Drawings
are submitted with most internals and should be referred to for specific details of your
equipment. Should there be any questions concerning installation of internals, please call your
equipment supplier or Jaeger Products at (800) 678-0345.

1. Support Plate, Gas Injection Type - Model PS1

The installer should confirm that proper support beams, if any, or support rings are
installed and secured before installation of the multi-beam support. The multi-beam support
is made in sections to allow installation through manway, or body flange openings.
Sections are installed from each end working toward the center. Hardware, if any, is to be
attached to each section as it is installed. Refer to drawings for clip orientation details.

2. Support Plate, Grating Style - Model PS2

The installer should confirm that proper support beams are installed and secured before
installation of the support gratings. The support gratings are made in sections to allow
installation through manway openings. Sections are installed from each end working
toward the center. Hardware, if any, is to be attached to each section as it is installed. Refer
to drawings for orientation details.

3. Mist Eliminators, Mesh Type - Model ME1

The installer should confirm that proper support beams or annular ring, if applicable, are
installed and secured before installation of the mist eliminator. The mesh pad is sectioned
for installation through manway openings. Grids, if applicable are attached as an integral
unit. Installation should begin at each end, working toward center. Each section is secured
to the annular ring and support beam(s) with tie wires supplied with the mist eliminator.
When complete, pad should be securely anchored to the annular support rings and support
beam(s). The mesh pad portion is oversized to gain snug fitting of mesh over the entire
surface area. Refer to drawings for section layout and location of beams.

4. Mist Eliminator, Vane type - Model ME2 (Vertical up-flow of gas)

The installer should confirm that proper support beams, if required, and support rings are
installed and secured before installation. The vane type mist eliminator is sectioned for
installation through manway, or body flange openings. Installation should begin at each
end, working toward center. Care should be taken to be sure the unit is installed in the

JPI\INSTLMAN.DOC Rev.9 (03/97)

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 proper flow direction. There are indicators on the sections to establish direction of flow.
Each section is nested onto another. In some plastic fabrication cases, inverted T-beams
are included and installed to support adjacent sections. The T-beam should rest on the
annular support ring in an inverted position so that the top of the T-beam is on the same
plane as the support ring. (In metal fabrication, the T-beam is normally not required.)
Anchoring the unit to the support ring is not normally required. In special cases, multi-
layers of vane sections are used, which are stacked one on top of the other, and the
orientation and direction are maintained as before. Refer to project drawings for details.

5. Mist Eliminator, Vane type - Model ME2 (Horizontal gas flow)

This unit is similar to the vertical gas flow ME02, except for it s orientation and
installation. The vanes are normally slid into position through an opening above or on the
end side of it. When made in sections, they are installed sequentially, per the required
layout. Make sure that the liquid drains, hold-in supports, etc., are adequate before the start
of operation and the installation openings are shut properly. Refer to project drawings for
details.

6. Mist Eliminator, Packed Bed Type - Model ME4

A packed bed mist elimination unit is installed in the same fashion as the components that
make up the item. Support grating is installed by sections starting with the end pieces and
working toward the middle. After securing the support grating, the packing is installed to
the desired level being sure to distribute the packing evenly. A bed limiter is installed on
top of the packing much like the support grating. Refer to project drawings for hardware
details.

7. Liquid Distributor - Orifice Riser with Feed Pipe - Model LD1

Before installation, the unit has to be checked thoroughly to make sure all the sections,
hardware, supports, feed connections, etc., are available. The lower flat sections called the
orifice plates with rectangular boxes, or circular risers, are installed starting with the end
pieces working towards the center. The gaskets have to be properly placed first along the
support ring and the joints. The support ring clamps, tray joint hardware, and the proper
seal-plates, if any, have to be properly installed. The middle section is installed last, which
is equipped with manway sections (for larger units). The manway section and manway
pieces, if nay, are installed last. These have welded nuts to the bottom sides and requires
seal plates in the corners. After aligning the orifice sections properly, to make sure the feed
pipe with the flanged nozzle connection will fit correctly, all the hardware is tightened and
the feed pipe is installed on the top, connected to the inlet flange of the vessel. The
chimney bottoms are normally fitted with anti-migration (a.m.) features so that an extra
Bed Limiter (BL1) is not required. Refer to project drawings for details.

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 8. Liquid Distributor - Pipe Branch Type - Model LD3

The installer should make sure proper support, and header flange connections with
hardware and gaskets (by others) are provided. The header is first installed by securing it
to the flange connection and end supports with the holes at the bottom. The branch pipes
are then screwed to the header pipe into the couplings provided, making sure the holes on
the branches are properly pointing downwards. Finally, check on all the supports and make
sure they are properly secured. Refer to project drawings for details.

9. Liquid Distributor - Trough Type - Model LD4

The installer should confirm that proper support beams, if any, and annular rings are
installed and secured before installation of the distribution plate. The removable feed pipe
should be installed last, after making sure the orientations are correct. The installation will
not allow for men to stand on sections. The distributor is sectioned for installation through
vessel manways. The lower box sections with the V-notches are called troughs and are
installed first, starting with the shorter end sections and working toward the center.
Leveling devices (optional) are located at the end of each trough. The trough should be
leveled as it is installed and secured. The larger channel is called a parting box. Depending
on design, there could be multiple units. The parting box is installed perpendicular to the
troughs and secured to the troughs. The feed points from the parting boxes should be
aligned properly with each trough. If a support beam is installed, it should be located
directly under the troughs, and at right angles to support the weight under the parting box.
The installer must make sure the troughs and parting box are properly oriented to the feed
pipe before it is installed at the end. Refer to project drawings for complete details.

10. Bed Limiter - Model BL1

The bed limiter is installed in sections. If an annular ring is provided it should be assembled
from the outside working toward the center while securing it at the ring. Bed limiters often
simply lay directly on the packing and thus will not require an annular ring for support.
Refer to project drawings for details.

11. Redistributor - Orifice Riser with Feed-type - Model CR1

This unit is assembled in a similar way as the LD1 above, after making sure all the sections,
hardware, supports, feed connections, gaskets, etc., are available. The lower flat sections
called the orifice plates with rectangular boxes, or pipe risers with caps, are installed
starting with the end pieces working towards the center. The gaskets have to be properly
placed first along the support ring and the joints. The support ring clamps, tray joint
hardware, and the proper seal-plates, if any, have to be properly installed. The middle
section is installed last, which is equipped with manway sections (for larger units). The
manway section and manway pieces, if any, are installed last. These have welded nuts to
the bottom sides and require seal plates in the corners. After aligning the orifice sections
properly, to make sure the relevant feed pipe (FP1) with the flanged nozzle connection will
fit, and be supported correctly, all the hardware is tightened and the feed pipe is installed
on the top, connected to the inlet flange of the vessel. (Anti-Migration features are usually

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 included so that no additional Bed Limiter (BL1) is required.) Refer to project drawings
for details.

12. Redistributor - Orifice Riser Type - Model CR1

The unit is assembled the same way as the CR1 above. The only difference being, there is
no feed pipe to worry about, and the number of manway pieces in the center manway
section may be two, or one depending on how the nuts on the section/manway joints are
located. This unit is normally fitted with anti-migration features requiring no bed limiter.

13. Collector/Redistributor - Orifice Drip Tube Type - Model CR2

This redistributor is made in sections with drip tubes and rim seal bands to sit on support
clips, also has a feed pipe above and Max-PakJ Packing below. Before its installation, the
Max-PakJ Structured Packing below and the support clips, with bolts attached, have to
be installed. Proper care has to be insured that the orientation is correct, that there is at
least a 1"~2" gap from the top surface of the support clips to the upper surface of the Max-
PakJ packing below all the clips are leveled properly and feed pip nozzle is 10" above
before the clips are welded to the wall of the vessel. Please refer to project drawings for
details. Before welding of the CR2 sections inside the tower, the rim seal banding has to
be bent outwards from the rim to at least 45°. The seal-welding of the sections are carried
out along the floor and the vertical overlap joints followed by installation with proper
hardware, including washers on the clips. The feed pipe is installed last.

15. Collector/Tray - Chimney Riser Type - Model CRS

This unit is assembled the same way as the LD1 or CR1 models, except it does not have
liquid distribution holes. It has much taller risers with caps and may be installed with
hardware or seal-welded in the tower. Sometimes it includes a sump for proper delivery
of the collected liquid to a nozzle on the vessel wall or another distributor below. Refer to
project drawing for details.

16. Gas Sparger - Model GD3

A gas sparger is a device used to enhance the initial distribution of the gas prior to contact
with the packed bed. Gas spargers are typically composed of a horizontal pipe with slots
or holes where the gas is discharged. The sparger is connected to the gas inlet nozzle
usually with a flanged connection, however, in some cases it may be seal-welded for
permanent installation. Depending on the length of the sparger, the unit may require
support at the opposite end of the inlet nozzle. This is typically accomplished by means of
a flange, cradle, or welded bracket. Due to design, and/or manway considerations, the unit
may utilize flanged sections. Be sure that all parts are properly oriented and appropriate
gaskets and hardware are used. After installation, the slots/holes should be pointing
downward. Refer to project drawings for details.

17. Feed Pipe (Liquid Distribution System) - Model FP1

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 The feed pipe is the initial stage of the distribution system and is vital to the performance
of other components of the distribution system. The feed pipe is attached to the inlet liquid
feed nozzle, usually with a flanged connection. The FP1 can consist of single use of or
combination of headers and branch pipes. Design considerations will dictate the number
of headers, branches, flanged sections, and use of nozzles. A properly installed feed pipe
allows full and even liquid flow to the distribution system without excessive turbulence or
spillage. When installed, check to be sure that all parts are secure and level. Should you
notice excessive turbulence or uneven performance after startup and you have verified
installation procedures, contact Jaeger Products for further assistance. Refer to project
drawings for details.

VI. OPERATION AND MAINTENANCE

A. Operation

Operation of the vessel should be done in accordance to the procedures outlined by the vessel
manufacturer and according to the design specifications.

B. Maintenance

Inspection of the packing should be done on a regularly scheduled basis. Maintenance of the
packing is determined by the amount of fouling (collection of deposits such as iron, carbonate,
and bacteria) that has accumulated on the packing. Past experience indicates a minimum
inspection period of six (6) months to check for excess fouling. The optimum inspection time
will vary from area to area depending on the rate of concentrations of deposits from your water
source. The packing can be removed and cleaned if fouling is not permitted to get excessive.
The packing should always be kept wet until cleaning can be done to prevent the deposits
solidifying and become much more difficult to remove.

C. Cleaning of Packing

Packing can be cleaned using different methods and is often a combination of common sense
and effort depending on labor and equipment available. A regularly scheduled maintenance
program can reduce or eliminate difficult cleaning efforts.

There are many chemicals and acids that can be used in the cleaning of packed towers. Good
results have been evident using 3-5% HCl or muriatic acid, as well as citric acid. The approval
and recommended use of any specific chemical or acid should be indicated by the manufacturer
of all components of the vessel or system, including the packing and internals.

If the packing is not severely fouled, the cleaning can be done without removing it from the
tower and basically involves circulation and recirculation of the cleaning agent through the
packed bed. The circulation rate should be at the maximum that can be achieved through the
process equipment. In some cases, it may be necessary to install a temporary distribution
system to achieve adequate distribution and effectiveness. The recirculated solution should be
filtered to remove solids as they become dislodged from the packing. The length of time
required to clean the packings will depend on the amount of fouling present, the effectiveness

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 of the cleaning agent chosen and the quality of the liquid distribution. If acids are used, the
solution may require neutralization before proper disposal.

If the packing is severely fouled where distribution of the cleaning agent would be difficult, the
packing is typically removed and cleaned using similar techniques, but in baskets or wash
basins. The packing can be soaked longer with this method and agitated to aid in the removal
of the fouled matter. However, this method will most likely result in more broken packing
elements and the need for replacement of some media. Consideration should be given to the
cost effectiveness of this method versus total packing replacement.

D. Preventative Maintenance

Preventative maintenance products are available for use in most applications and can be very
effective in eliminating or reducing the need of periodic cleaning. Towers with low liquid rates
in the range of 150 gpm or less are good candidates for a pretreatment system. Systems with
flow rates over 150 gpm should be evaluated to determine cost effectiveness based on
comparison of cost of pretreatment chemicals and the cost of cleaning. Cleaning costs should
include all labor and equipment cost as well as the cost of acids and neutralizing agents
required for disposal.

Jaeger Products specializes in chemical pretreatment systems in applications involving hard

water with high concentrations of free iron, calcium, manganese and other minerals. Our
product called JP-7 is injected into the water source on a continual basis using a low
maintenance chemical feed pump. The JP-7 as a pre-treatment will keep the minerals in
solution preventing oxidation within the packed tower thus preventing a fouling problem.

Bacterial fouling presents a different approach to the preventative maintenance procedure. JP-7
is not a biocide, but it will assist in breaking the gelatinous membranes and allowing
penetration of other agents. Bacterial fouling can often be reduced by common method of
adding chemicals such as those containing some form of chlorine. However, depending on the
type of microorganism growing, it can manifest itself as a very severe problem with no single
cost-effective solution. Bacterial fouling will not clean with acids and care should be taken in
their use as the use of acids could exasperate a bacterial fouling problem.

In the worst cases, it is best to attempt to determine the type of organism causing the heavy
growth and use a product that will kill the growth to allow for easy washing through the
packing. Some chemicals may seem to be effective for a period of time and then become
ineffective due to immunities built up by the microorganisms. Others may be successful in
killing the growth but leave remaining residue in a consistency or form that is extremely
difficult to remove.

E. Disposal of Packing

Should it become evident that the packing cannot be cleaned and are to be replaced, one must
dispose of the packings properly. Typical disposal of "spent" packings is usually done at

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 regional landfills. The packing itself does not pose an environmental problem to the landfill,
however the solids present with the fouling may be considered hazardous depending on your
application. Disposal of "spent" packings will usually require a test at the request of the landfill
to determine toxicity or presence of contaminants that may leach over time. Based on their test,
the landfill will then assign the proper classification of landfill suited for your disposal.
Landfills usually charge by volume. To reduce the amount of volume that is to be land filled,
it is recommended to chip the packing into smaller pieces. This can be done with common
chippers available through most equipment rental companies and reduces the volume by 50-
90% depending on the amount of fouling on the packing.

VII. MANUFACTURER S WARRANTY

Seller represents that the goods or equipment sold pursuant to this order will be free from
defects in workmanship or material for a period of one (1) year and, if applicable, will be
manufactured using the resin specified in this order and identified on Seller's invoice to Buyer.
Seller makes no other warranties, expressed or implied, whether of workmanship, performance,
quality, durability, merchantability or fitness for any particular purpose or use or otherwise with
respect to materials sold, or with respect to any part or labor furnished during the sale, delivery,
servicing of the material. In no event shall Seller be liable to buyer for any special, indirect,
incidental or consequential damages arising out of, or as the result of, the sale, delivery,
servicing, use or loss of use of the material or any part thereof, or for any charges or expenses
of any nature incurred without Seller's written consent, even though Seller has been negligent.
The obligations of Seller are limited to repair or replacement of defective material or, at its sole
option, to the refund of the purchase price. Buyer assumes all risk and liability for results of
using the material covered by this order, whether used separately or in combination with other
products, and agrees to hold Seller harmless for all damages, direct or consequential, resulting
from the use thereof.

SHOULD YOU HAVE ANY QUESTIONS CONCERNING THE INSTALLATION OF

PACKING AND/OR INTERNALS BY JAEGER PRODUCTS, INC., PLEASE CALL YOUR
SUPPLIER OR CONTACT JAEGER PRODUCTS, INC.

(800) 678-0345
(281) 449-9500

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 MASS TRANSFER: DISTILLATION

Trouble-free design of
refinery fractionators
A review of factors most frequently the cause of distillation towers falling short
of design objectives. Analysis of case histories provides guidelines for identifying
potential troublespots in the most important fractionators
Henry Z Kister
Fluor Corporation

A
two-phase survey was recently refinery fractionators,
completed by Fluor of all the case and these form the
histories related to malfunctions in basis for the current
refinery towers that have been document- analysis.
ed over the last 50 years. Altogether, 400 As with other Fluor
case histories were found in the literature. surveys, certain
The first phase identified the most com- ground rules were
mon root causes of problems in refinery applied to limit the
fractionators (towers), but did not exam- scope. Only specific
ine the troublespots in each specific ser- incidents were
vice. This phase yielded general included. For exam-
guidelines for trouble-free design, but did ple, a statement such
not address issues related to each specific as “leakage from
fractionator. chimney trays in
In the second phase, case histories of refinery vacuum tow-
tower malfunctions are analysed specifi- ers can be reduced by
cally for each of the major refinery frac- seal-welding” does
tionators. Each case history teaches a not constitute a case Figure 1 Uplifted packing in wash section of a vacuum tower
lesson. Together, these lessons are the best history. On the other
tool for understanding the potential trou- hand, a statement such as “one vacuum It clearly shows that the vacuum tower is
blespots in each service, and for drawing tower experienced severe chimney tray by far the most troublesome refinery ser-
guidelines for trouble-free design of each leakage at low-rate operation. Seal weld- vice, which is where the survey begins.
service. ing tray sections reduced leakage to
I have previously described the Fluor acceptable levels” does. Vacuum tower malfunctions
survey methodology in Distillation Opera- Also, incidents of corrosion and foul- The 86 case histories reported for the vac-
tion (McGraw-Hill, New York,1990). All ing were included only if a feature unique uum tower is almost double the number
the case histories used as a basis for the to the column design, operation, or con- reported for the atmospheric crude tower,
survey were extracted from the published trol contributed to their occurrence. For which is the next most troublesome refin-
literature. There were 900 total cases, of instance, an incident where the wrong ery tower. When a vacuum tower per-
which 400 were for refinery towers. In corrosion inhibitor or antifoulant was forms poorly, valuable distillate is lost to
about one quarter of these, the specific applied does not qualify as a case history the resid, and poor distillate quality poi-
service was not stated or the service was in this survey. A case where fouling was sons FCC catalyst. The wash section of
one that did not have enough cases caused by insufficient liquid flow, maldis- the fractionator is the most critical sec-
reported on to permit detailed analysis. tribution, or poor process control, does. tion and also one where most of the mal-
This left about 300 cases for the main Finally, optimisation case studies
(where capacity was raised or pressure Top causes of vacuum tower
Fractionator malfunctions drop lowered by replacing trays by pack- malfunctions
ings) are outside of the scope of the sur-
Number of vey. The objective of the current survey is No. Description Cases
cases to identify the issues that make towers fall
1 Damage 27
1. Vacuum towers 86 short of achieving these design capacities.
2 Coking 21
2. Atmospheric crude fractionators 45 There is some overlap in the tabulation
3 Intermediate draws 17
3. Debutanisers 37 of cases for each fractionator. For 4 Misleading measurements 10
4. FCC main fractionators 33 instance, a coked chimney tray case study 5 Plugging 9
5. Deethanisers 23 will be listed once under “coking” and – Installation mishaps 9
6. Depropanisers, C3/C4 splitters 22 another time under “intermediate draws”. – Abnormal operation (startup,
7. Alky main fractionators/isostrippers 17 This means that adding the individual shutdown, commissioning) 9
8. Coker main fractionators 15 malfunctions may yield a number greater 8 Maldistribution 6
9. Naphtha splitters 11
than the number of malfunctions report- – Weeping 6
10. Deisobutanisers 8
ed for the service. Table 1 lists the main 10 Condenser 4
11. Amine towers 8
fractionators surveyed and the concise
Table 1 number of cases reported for each service. Table 2

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 MASS TRANSFER: DISTILLATION

functions were reported. The wash sec-

tion of the vacuum tower is therefore the
most troublesome tower section in the
refinery (Figure 1, on previous page).
Table 2 shows the common malfunc-
tions reported in vacuum towers. With 27
case histories, damage tops the list. Most
of this damage can be readily prevented.
Table 3 shows the most common causes.
Foremost are water-induced pressure
surges, which account for one third of the
reported damage incidents. In three of the
nine cases reported, the source of water
was poor draining of stripping steam
lines. In another three, pockets of water
lying in the piping of spare pumps
entered the hot tower when these pumps
were connected to the hot tower.
A lesson in this case is that many, pos-
sibly most, vacuum tower damage inci-
dents can be prevented by design and
operating procedures that adequately
drain the tower steam lines in wet towers, Figure 2 Broken, non-standard flange in the spray header supplying wash oil to the
and that positively prevent water from wash bed of a vacuum tower
spare pump piping from entering the
tower. A joint designer/refiner “hazop” flanges and gaskets and properly allowing sory Committee, San Antonio, Texas, Nov
should focus on these troublespots. for thermal expansion in the header 2001].
The next source of damage in Table 3, design. Coking of the wash section (Figure 3) is
insufficient mechanical strength, is also Damage due to high base level con- a close second in Table 2 with 21 reported
readily preventable. It should be recog- tributed three out of the 27 damage case case studies. Table 4 gives a breakdown of
nised (as can be readily seen from Table 2), histories. This again is an issue that can be the causes. Excess stages and vaporisation
that damage is a major issue in a vacuum at least alleviated by good level monitor- occur in wash beds that are either too tall
tower, and that heavy duty internals ing, alarms, and well-designed trip sys- or contain packings that are too efficient.
design should be used. Although the tems. Another three damage-related In either case, the additional stages inten-
heavy duty design would not be able to case-histories were caused by packing sify the vaporisation of the wash oil, leav-
withstand a major pressure surge, it would fires. This type of damage is more difficult ing little liquid to reach and wet the lower
weather the smaller pressure surges. Some to prevent due to the difficulty of clean- sections of the bed. These lower sections
good heavy duty design practices have ing the packings, especially when coked. of the bed dry and coke. Poor modelling
been described by Shieveler [Shieveler G H, Nonetheless, much progress has been and simulation is another cause of coking.
Use heavy-duty trays for severe services; Chem reported in developing preventive mea- Golden et al stress that the heavy ends of
Eng Progr, Aug 1995]. sures, and is discussed in two excellent the crude must be correctly characterised
Special attention should be paid to grid papers by Bouck and Markeloff [Bouck D in the simulation and that the feed entry
installation and tightening. In two of the S, Vacuum Tower Packing Fires; API Operat- to the tower must be modelled by a series
five cases, poorly fastened grids disinte- ing Practices Symposium, 27 April 1999. of flash steps that correctly represent the
grated in service. Through-bolting has Markeloff R, Packing fires; FRI Technical Advi- physical sequence of steps between the
been far more effective than J-bolting for
keeping grid together, and should be rou-
tinely specified.
Spray distributors and their headers are
prone to damage (Table 3). Again, this
damage can be easily prevented by sound
design, good installation, and thorough
inspection and testing. Water testing
spray nozzles and headers can readily
detect damage (Figure 2). Header damage
can be prevented by using standard

Causes of damage in vacuum

fractionators

No. Description Cases

1 Water-induced pressure surges 9
2 Insufficient mechanical strength 5
3 Broken nozzles or headers of
spray distributors 4
4 High bottom liquid level 3
– Packing fires 3

Table 3 Figure 3 Coking of fouling resistant grid in the wash section of a vacuum tower

110
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 MASS TRANSFER: DISTILLATION

chimney tray that caused entrainment to be troublesome in three cases. With

Causes of coking in vacuum into the wash bed. Three cases were trapout trays being used only in tray tow-
fractionators reported where maldistribution of vapour ers, this issue is experienced mainly in
or liquid led to coking. vacuum lube towers that have trays and
No. Description Cases Intermediate draw malfunctions is a not packing. The remaining case histories
1 Excess stages and vaporisation 4 close third of the most common vacuum describe coking and excessive hydraulic
– Poor modelling and simulation 4 tower malfunctions, with 17 reported gradients.
3 Insufficient wash, reason not case histories (Table 5). Foremost is leak- With 10 case histories, misleading mea-
reported 3 age from total draw chimney trays. Leak- surements are in the 4th spot in Table 2.
– Misleading measurement 3 age of the HVGO chimney tray represents Three of these 10 are the troublesome
– Liquid, vapour maldistribution 3 good distillate degraded into resid with chimney tray level measurements previ-
Not reported 4 no beneficial effects whatsoever, as leak- ously mentioned. Other troublesome
ing liquid poorly distributes in the wash cases have been reported with short coil
Table 4 bed and does little washing. Leakage of outlet thermocouples (two cases), ambi-
LVGO into the HVGO section lowers the ent changes affecting vacuum measure-
heater outlet and flash zone [Golden S W, HVGO boiling point and can reduce heat ments with ordinary gauges (two cases),
Vacuum Tower Troubleshooting; AIChE Spring transfer, even limiting vacuum on the bottom level, heater fuel flow rate, and
Meeting, 1994]. tower. reflux to a packed bed distributor. The
When these principles are overlooked, This leakage needs to be avoided with lessons from these case histories to instru-
the simulation underestimates wash oil totally seal-welded chimney trays. Special ment specifications are self-explanatory.
vaporisation, leading to the drying up and techniques, as recommended by Lieber- With nine case histories, plugging (as
coking reported in four cases. man, are effective and need to be incor- distinct from coking) is in the 5th spot in
In three other reported coking inci- porated to avoid tray buckling due to Table 2. Of the nine cases, five were plug-
dents, it was stated that the wash flowrate thermal expansion [Lieberman N P, Process ging of spray headers. One case was
was insufficient but no specific reason Design for Reliable Operation, 2nd ed; Gulf Pub- reported of plugged packing, plugged
was given. It is likely that in those cases lishing, Houston, Texas, 1988]. quench pipe, plugged instrument line
too, either the number of stages was Level measurement on chimney trays and plugged ejector. In two cases, the
excessive, or the modelling/simulation has been troublesome in three reported plugging was by corrosion products. Mea-
were poor, or both. In three other cases, cases. This may lead to overflow or sures found effective for alleviating plug-
coking was produced by either a mislead- entrainment. While overflow is equiva- ging in the wash spray headers, which is
ingly low coil outlet temperature signal lent to leakage, entrainment from the one of the most common troublespots,
that caused excessive firing, or a faulty overflash chimney tray can induce cok- are to provide good wash oil filtration and
level measurement on the overflash ing. Leaking trapout trays were reported to specify an all-stainless-steel wash oil

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fluid dynamics (CFD) is essential. Improv-

Intermediate draw malfunctions ing the vapour horn design has eliminat- Top causes of FCC main
in vacuum towers ed operating problems and improved the fractionator malfunctions
performance of several vacuum towers.
No. Description Cases A surprisingly high number of case
histories place tray weeping in the equal No. Description Cases
1 Leaking total draw chimney trays 6 1 Plugging 9
2 Level measurement on total draw 8th spot in Table 2. This is not an issue
with most vacuum towers that are all – Abnormal operation 9
chimney trays 3
3 Liquid maldistribution to
– Leaking trapout tray 3 packed, but appears to be a major issue packing liquid 6
4 Chimney tray coking 2 with the trayed vacuum towers, mostly
4 Intermediate draws 5
– Excessive hydraulic gradients on in lube service. In two of these, leakage
5 Water-induced pressure surges 4
chimney trays 2 at a draw tray made it impossible to draw
6 Pressure control 3
6 Others 1 sufficient product. In two others, weep- – Vapour maldistribution 3
age at pumparound trays starved the
Table 5 pump and reduced heat transfer. In two
more, poor separation between side-cuts Table 7
line downstream of the wash oil filter. resulted. Blanking valves, using high
Sharing the 5th spot in Table 2 is installa- turndown valves, and in the case of the and air premature upon shutdown,
tion mishaps. Three case histories were side cuts, drawing from seal-welded another from a packing fire at the
associated with spray header installation, chimney trays, was the solution. turnaround. The cause of damage in the
and another three with grid or packing Condenser issues complete Table 2. other three cases was not reported.
assembly. Other cases describe problems Condenser problems raise pressure in the Seven abnormal operation (startup/
with tower out-of-roundness, using car- tower and thus reduce distillate recovery. shutdown/commissioning) incidents
bon steel bolts where stainless steel was Two of the four cases reported dealt with were reported. Four of these resulted in
specified and poor installation of strip- excess lights, one with ejector plugging, four of the damage incidents listed, one
ping trays. and one with flash equilibrium at the pre- led to an explosion, another to a fire and
Also sharing the 5th spot in Table 2 is condenser. one to a chemical release. Three of the
abnormal operation incidents. Five of pressure surges listed in Table 6 under
these describe incidents during startup Crude tower malfunctions Damage resulted from poor dehydration
where a pocket of water entered the hot The three most common atmospheric during startup or pump switchover. Poor
tower and created a pressure surge. Poor crude tower malfunctions (Table 6) are blinding and unblinding led to one
blinding/unblinding contributed two case plugging, intermediate draw malfunc- reported case of explosion and another of
histories, one case is related to pressur- tions and damage. chemicals release.
ing/depressuring and one to flushing. There were nine plugging incidents The next four entries in Table 6 are well
Maldistribution problems, other than reported in atmospheric crude towers: below the top four, and have three to
those attributed to coking, plugging or Four in the wash section or gasoil four reported malfunctions. Four installa-
damage, are in the 8th spot in Table 2 pumparound, three in the top section or tion mishaps were reported, all involving
with a surprisingly low number of case top pumparound, and two in the strip- trays or chimney trays. Leaks of a
histories (six). Of the reported six, four ping section. No cases of plugging were pumparound exchanger, a pump seal,
were vapour maldistribution, three of reported in the middle of the tower. In the and resid to atmosphere were the three
these originating in the flash zone and wash section, the most common cause of reported leak cases. Controlling liquid
one in the previously mentioned tray plugging was entrainment from the flash flow to the wash section has been a spe-
chimney. The other two cases reported zone or from the vapour overhead of a cial challenge, contributing three more
liquid maldistribution problems. Fluor’s preflash drum. In the top of the tower, the case histories.
experience has been that maldistribution, plugging was by scale and corrosion prod- Trouble-free designs properly dis-
especially of vapour from the flash zone, ucts, corrosion inhibitors, and salting out. entrain the vapour in the flash zone and
has been far more troublesome than sug- Five of the nine incidents resulted in preflash drum and properly desalt the
gested by the low spot of this item in plugged trays, two in plugged downcom- crude in order to minimise plugging in
Table 2. ers. In one case, a packed bed plugged, in the fractionator. Specifying fouling-resis-
Vapour horn design and good distribu- another, a liquid distributor to the pack- tant hardware in the wash zone and
tion of liquid to the wash bed are central ing plugged. upper trays is good practice. Downcomer
for achieving trouble-free performance of The large number of intermediate draw trapouts and chimney trays are the most
the wash bed. An expert hydraulic analy- incidents is well in line with Fluor’s expe- important internals for ensuring trouble-
sis, often with the aid of computational rience: chimney trays and downcomer free operation. They need to be designed
trapouts make or break fractionators. Of and inspected carefully, not just left to
Top causes of atmospheric crude the nine, seven took place with down-
tower malfunctions comer trapouts, two with chimney trays. Top causes of malfunctions in
Four of these involved choking or restric- debutanisers, incl stabilisers and
tion in the outlet liquid line, while in two depentanisers
No. Description Cases
others, leakage at the drawoff restricted
1 Plugging 9
the recovery of a side cut. No. Description Cases
– Intermediate draws 9
Four of the nine damage incidents 1 Control 10
– Damage 9
4 Abnormal operations 7 reported were due to water-induced pres- 2 Vapour cloud release 5
5 Installation mishaps 4 sure surges. Two of these were caused by – Installation Mishaps 5
6 Condenser Problems 3 undrained stripping steam lines, one by a 4 Feed arrangement, tray towers 4
– Poor control of wash 3 water pocket in a spare pump, and one by – Reboiler draw arrangements 4
– Leaks 3 plugged drainholes in the bottom seal
pan. One case of damage resulted from
Table 6 exposing column internals to cold water Table 8

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incidents.There were nine plugging inci-

Top causes of malfunctions in dents reported in FCC main fractiona- Top causes of malfunctions in
C3/C4 splitters and depropanisers tors. Of these, four were salting-out deethanisers (absorbers and
(excl those in alky units) incidents that plugged trays near the top strippers)
of the tower, and were overcome by
online water washes. Three were inci- No. Description Cases
No. Description Case dents in which grid in the slurry section 1 Reboiler draw and return
1 Reboiler draw and return coked up due to vapour or liquid-maldis- arrangements 6
arrangements 7 tribution. One incident was plugging due 2 Excessive tower base level 4
2 Tower flooding by excess to catalyst carryover into upper packed – Control 4
base level 6
sections, and another was plugging of a – Component accumulation 4
3 Vapour cloud release 5 line draining the main feed line, both – Side draw arrangements 4
Table 10 during startup.
There were also nine abnormal opera- Table 9
others, as these will make or break the tion (startup/shutdown/commissioning)
fractionator. incidents reported, two of which were FCC main fractionators than in atmo-
Prevention of water entry, by ensuring previously described. These two inci- spheric crude fractionators, so it comes as
adequate drainage on stripping steam lines dents, plus two others, occurred during little surprise to find liquid maldistribu-
and eliminating dead pockets inside the liquid circulation and dehydration. In tion to packings in a prominent spot in
tower, is central for damage prevention. three of these four, a pressure surge and Table 7. Of the six reported incidents, two
Proper inspection of equipment is a must major damage resulted, the other was the involved liquid maldistribution to the
to prevent the installation mishaps. Proper catalyst carryover. The remaining five slurry pumparound section, the others to
control of the liquid flow rate to the wash incidents include poor unblinding caus- various fractionation sections. Intermedi-
section is the prime control consideration ing a toxic release; switching over oxygen ate draws in FCC main fractionator have
in the tower. and nitrogen purge gas causing explo- been troublesome in five reported case
sions; trip failure on the reflux drum caus- histories, more in chimney trays than in
Main column malfunctions ing liquid carryover and major downcomer trapouts. Finally, four cases
There are some similarities with regard to compressor damage; a major leak due to of water-induced pressure surges were
FCC main fractionator malfunctions the thermal shock while opening or clos- reported, three of which led to major
(Table 7) and atmospheric crude tower ing the valve in the tower inlet; and a damage.
malfunctions, but there are also major startup pressure control problem resulting Two other malfunctions are also shown
differences. Two malfunctions top the from steam condensation. in Table 7: Vapour maldistribution, all
list: plugging and abnormal operation Packings are used more frequently in cases dealing with grid in the slurry sec-

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tion, and pressure control. Plugging and There is a distinct link between the or pan causing liquid to bypass a once-
coking can be alleviated by providing ade- feeds and reboiler draw arrangements. through thermosiphon reboiler (two
quate on-line wash facilities near the top Both constitute “points of transition”, ie, cases); a reboiler tube leak and slug flow at
of the fractionator, by using plugging- where a stream enters or leaves the tower. the reboiler outlet pipe.
resistant trays there, and by preferring These points of transition are some of the Slightly behind, with six case histories, is
shed decks or disk and donut trays to grid major troublespots in a tower. The lesson tower flooding by excess base level. Two of
in the wash section. Shed decks and disk for debutanisers is that all points of transi- these resulted from the type of reboiler
and donut trays are far less sensitive to tion need to be critically examined for problems previously discussed. False level
vapour or liquid maldistribution than potential bottlenecks, both at the design indica tions led to two others, and frothing
grid, and therefore far less prone to cok- (or debottleneck) and when trouble- or foaming at the tower base led to the
ing during upsets. shooting. remaining two. Clearly, the lessons learned
Startups and shutdowns are major are that troubleshooting and trouble-free
issues in FCC main fractionators, and a Deethaniser malfunctions debottlenecks of C3/C4 splitters should focus
good design of these needs to hazop what The strippers and absorbers are included in on the reboiler piping and the bottom
can go wrong and take preventive mea- the deethaniser malfunctions. Topping the sump.
sures. Intermediate draws and liquid dis- list (Table 9) with six case histories are Similar to debutanisers, depropanisers
tributors are the weakest links in the reboiler draw and return arrangements. and C3/C4 splitters have experienced a high
internals design, and need to be designed Three of the six cases report excessive pres- number of vapour cloud releases, mostly
and inspected carefully, not just left to sure drop in the process inlet or outlet due to line rupture (three cases), but also
others. Finally, pressure controls as well as pipes of a kettle reboiler. The high-pressure due to poor blinding or plugging/freeze ups
liquid flow control to the wash section drop either caused the tower base liquid of valves. Some major blasts resulted. The
are major considerations in these frac- level to rise above the reboiler return inlet, vapour cloud lessons described under
tionators. or back liquid up on the chimney tray feed- debutanisers extend to depropanisers and
ing the reboiler to the top of the chimneys. C3/C4 splitters.
Debutaniser malfunctions Insufficient heat during coke drum
Due to similar functions, stabilisers and switchover was reported in two cases, one Other fractionators
depentenisers have been lumped together of them due to weeping from the draw tray For other refinery fractionators, the num-
with debutanisers. Over 70% of the cases, to a once-through thermosiphon reboiler. ber of case studies reported was less than
however, were contributed by debutanis- Four case histories were reported of base 20, a sample too small for a detailed anal-
ers. level exceeding the reboiler return. Two of ysis. Nonetheless, some observations are
Table 8 shows that the most common these were due to high-pressure drop in significant and require more detail, includ-
malfunctions experienced in debutanisers the kettle piping (those previously men- ing coker main fractionators, alky unit
are widely different from those experi- tioned), the other two due to absence of or main fractionators/isostrippers, naphtha
enced in the vacuum, crude and FCC frac- to poor level indication. As with debu- splitters, deisobutanisers and amine
tionators. Topping the list with 10 case tanisers, control issues are also important absorbers/regenerators.
histories is controls, an item that showed in deethanisers, and account for four case A total of 15 malfunction case histories
low down (if at all) on the main fraction- histories. Also, with four case histories, of coker fractionators have been report-
ator malfunctions list. Of the 10 cases, five component accumulation in deethanisers ed. Of the 15, seven described fouling by
reported difficulties with pressure and con- is a problem. coking or carryover of coke, while five
denser controls. In all five, a total con- Either ethane or water or both accumu- others described damage due to water-
denser was used with partial flooding of late and can lead to cycling, capacity bot- induced pressure surges. There is no
the condenser. In two of the five, the prob- tlenecks, and in the case of water, also doubt that coking and water-induced
lem was induced by presence of non-con- corrosion. Finally, choking of side draws pressure surges are the major issues with
densables. Composition control or the with entrained gas bubbles has been a these fractionators.
assembly of a control system contributed problem in four case histories. A total of 17 malfunctions have been
the other control case histories. The lessons learned from this documen- reported for alky unit main
Vapour cloud release and installation tation are that the points of transition in fractionators/isostrippers. Of these, four
mishaps share the second spot in Table 8. deethanisers (the side draws as well as the described plugging, mostly by scale or cor-
Three of the five case histories of vapour region below the bottom tray, including rosion products; three described explo-
clouds ended in explosions, and one more the reboiler draw and return lines) require sions, either due to vapour cloud release or
in a fire. Some of these were accompanied thorough design, review, and inspection, due to HF carryover in the hydrocarbons
by injuries and heavy damage. Line frac- and must not just be left to others. Preven- and a violent reaction in a caustic bed
ture (two incidents), poor blinding (two tion of component accumulation and care- downstream; and three others described
incidents), and freeze-ups in leaking ful review of the control systems are also accumulation of either ethane or water in
valves (two incidents) were some of the prime considerations that make the differ- the overhead system.
contributing factors. Hazops of debutanis- ence between a troublesome and trouble- A total of 11 naphtha splitter malfunc-
ers should consider some of the lessons free deethaniser. tions have been reported. Four of these
learned from previous vapour cloud reported plugging, mainly by scale and cor-
releases to positively eliminate further Splitter malfunctions rosion products; three reported reboiler
accidents. Malfunctions in C3/C4 splitters also include issues; two were a result of poor installa-
With four case histories, poor feed depropanisers other than those in alky tion; and two reported control problems.
arrangements closely follow, leading to a units, which are uniquely different (Table Control problems with other refinery
capacity bottleneck or an efficiency loss in 10). Similar to deethanisers, reboiler draw fractionators have also been reported. For
the feed region. Also with four case histories and return arrangements lead the list with example, of the eight total deisobutaniser
are reboiler draw arrangements, including seven reported cases. Again, the main prob- malfunctions that have been reported, six
vapour entrainment choking the reboiler lems have been excess pressure drop in involved control problems. Five of these
draw lines and liquid leaking from a trapout inlet and outlet lines of a reboiler causing were temperature control issues that can
tray to a once-through thermosiphon base liquid level to exceed the reboiler be particularly troublesome with narrow-
reboiler, thus “starving” the reboiler. return inlet (two cases); leaking draw tray boiling mixtures.

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Eight case histories were specifically Plugging, abnormal operation issues,

reported for refinery amine absorbers. Of liquid maldistribution to packings and
these, six reported foaming, and three grid, intermediate draws and water-
reported scale and corrosion products induced pressure surges are the key trou-
catalysing the foam and causing plugging. ble spots in FCC main fractionators. The
Seven other case histories of amine main causes of plugging in these fraction-
absorbers and/or regenerators were report- ators are salting out near the top and cok-
ed without stating whether they came ing of grid beds in the slurry section.
from refineries or natural gas plants. Of These can be alleviated by online water
these, foaming was the issue in five. There washes and by avoiding grid in the slurry
is no doubt that foaming, and to a lesser section, respectively. Tower dehydration
degree fouling with scale and and corro- and liquid circulation are important start-
sion products, is the prime issue in amine up operations that can turn troublesome
towers. and lead to pressure surges or catalyst car-
ryover into the less fouling-resistant
Lessons learned regions.
The vacuum tower is by far the most trou- Control issues, especially pressure con-
blesome refinery fractionator. Damage, trols, are the primary source of problems
wash bed coking, and intermediate draws in debutanisers. Vapour cloud releases
are prime trouble spots. Water-induced have led to explosions and major damage
pressure surges are the leading cause of in debutanisers. Installation mishaps,
damage. Most of the damage incidents in tower feed entry arrangements, and
the vacuum towers are preventable. reboiler draw arrangements have also
Hazoping the possibility of water entry, been major trouble spots. Critical review
using heavy duty mechanical designs, of the control system, especially the pres-
inspecting, testing and correctly design- sure/ condenser controls, learning from
ing spray headers, paying attention to past vapour-cloud accidents, using
level measurement, and applying special hazops to minimise the possibility of
designs and procedures to prevent pack- vapour cloud releases, and sound design
ing fires can drastically reduce damage of feed entry piping and of tower base
incidents. arrangements are the key for trouble-free
Wash bed coking can be alleviated by debutanisers.
using short beds of relatively inefficient In deethanisers, the points of transi-
packing by correctly simulating the wash tion (the side draw arrangements and the
zone and by avoiding insufficient wash. region below the bottom tray, including
Intermediate draw malfunctions can be the reboiler draw and return lines) are
alleviated by seal-welding draw trays and prime trouble spots and are key to trou-
water-testing them at the turnaround and ble-free design and operation. Prevention
by paying attention to reliable level mea- of water and ethane accumulation in this
surement on these trays. Other measures tower is also important.
that promote trouble-free operation are Reboiler draw and return arrange-
good instrument specifications, good fil- ments, and tower base level are the key to
tration of the wash oil followed by stain- trouble-free depropanisers and C3/C4
less steel piping downstream of the filters, splitters. Depropanisers also are prone to
and good distribution of vapour and of vapour cloud releases, and lessons
the wash oil to the wash bed. learned from past vapour cloud incidents
Plugging, intermediate draws, damage should be incorporated in the design and
and abnormal operation incidents are the operation of depropanisers and C3/C4
most troublesome malfunctions in atmo- splitters.
spheric crude towers. Plugging is most Trouble-free operation of coker frac-
common near the tower top, at the wash tionators focuses on preventing coking
zone, or in the stripping section. Choking and water-induced pressure surges; of
and leakage are the prime intermediate alky main fractionators focuses on plug-
draw issues. Water-induced pressure ging, vapour cloud and component accu-
surges are the most common cause of mulation prevention; of deisobutanisers
damage. Plugging problems can be allevi- on composition control; and of amine
ated by eliminating the source of fouling absorbers and regenerators on foaming
and/or by using plugging-resistant inter- and plugging prevention.
nals. Downcomer trapouts and chimney
trays need to be designed and inspected
carefully, not just left to others.
Prevention of water entry by proper Henry Z Kister is a Fluor Corporation
draining of stripping steam lines and by fellow and director of fractionation
sound dehydration procedures at startup technology at Aliso Viejo, California, USA.
is critical. Proper inspection of equip- He has over 25 years’ experience in design,
ment and adequate control of liquid control and startup of frationation processes
flowrate to the wash section are also and equipment. He obtained his BE and ME
important for promoting trouble-free degrees from the University of New South
operation. Wales, Australia.

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 Russian
Linas Technology - Distillation of 21 century version
®

General Information About companies Linas-Techno and Linas Technology International

About us
Corporation
Introduction in Linas
technology Companies Linas-Techno and Linas Technology International Corporation carry up a responsibility for the
Industrial applications development, design and industrial applications of Linas distillation technology. Both companies have a special
of Linas Technology agreement for close co-operation around the world.
and pilot plant
projects
Patent protection of 1. Company Linas-Techno (Novosibirsk, Russia)
Linas technology
Publications and 1.1. The company Linas-Techno has been established in 1994 for carrying out of research works in the field of
presentations of Linas development of new distillation technologies. C 1994 on 1999 main principles of Linas technology at a laboratory
technology and small pilot level have been developed. Since 1999 the company Linas-Techno has started to use Linas
Commercial proposals technology in industry.
Contact us 1.2. Core activity of the company is based on commercial use unique Linas distillation technology. The company
has:
http://www.linas.ru
1. The full package of the know-how on Linas technology.
2. Several patents protecting a priority on Linas technology.
3. The official right on development of production schedules (that is new technological processes) and the
design documentation on installations of rectification units for chemical, petrochemical, oil refining and
other applications (the License of State Organization Gosgortekhnadzor 00ПP #013253).

1.3. For commercial applications of Linas technology in the industries the company Linas-Techno contains the
following departments:

1. Technological and research departments with several pilot plants for a development of concrete
industrial applications of Linas technology.
2. The design department.
3. The sales department for connections with clients and suppliers of equipments.
4. Department on installation and start-up in commercial operation Linas distillation units.

1.4. The company Linas-Techno is supervised by highly skilled managers and experts. The Linas-Techno is the
growing company. For last year several young engineers has been involved in a staff of the company. The skilled
staff of the company is capable to solve all possible problems in industrial use of Linas technology.

1.5. The Linas-Techno has the adjusted and confidential relations with suppliers of the standard equipment and with
manufacturers of the non-standard equipment. It allows carry out delivery and manufacturing of the equipment
quickly.

1.6. The company Linas-Techno has adjusted good communications with the design companies which are carrying
out projects on industrial platforms for distillation units. Due to these connections there is an opportunity to assist
clients in the decision of problems on designing industrial platforms.

1.7. Linas-Techno good and reliable business relations with a company Sinetik delivering for the Linas-Techno
element and program base for automatic operation system were established. Automatic operation system is based on
Siemens's elements.

2. Company Linas Technology International Corporation (New York, USA)

2.1. The company Linas Technology Int. Corp. has been established in 2002 in the USA for a representation of
Linas technology outside of Russia around the world. The company is legally registered in the state New York.

2.2. The company carries basic activities related to Linas technology outside of Russia. Mainly it is the following
actions:

A. Patent protection of Linas technology;

 B. Management of relations with potential clients and foreign partners;
C. Presentation of Linas technology and Linas rectification units;
D. Development and improvements of Linas technology;
E. Financial management;
F. Design of Linas rectification units.

up^
web-master © 1999-2004 Linas Technology
 Linas Technology - Distillation of 21 century
®

General Information Introduction in Linas distillation technology

About us
Introduction in Linas For understanding of the Linas distillation technology we have to make the analysis of modern industrial distillation.
technology It will allow to reveal the basic falls and contradictions of existing industrial distillation and to show solutions of
Industrial applications these problems in Linas technology.
of Linas Technology
and pilot plant 1. Importance of distillation
projects
Modern industry requires more and more very pure chemicals. Therefore separation technologies become more and
Patent protection of
more important, complex and expensive . The most common separation technology is distillation.
Linas technology
Modern industrial distillation was established some 40 to 50 years ago. Distillation consumes huge amounts of
Publications and energy and it can generate more than 50% of plant operating cost. Significant efforts by thousands of researchers
presentations of Linas and developers around the world during the second half of the twentieth century did not bring a sufficient
technology improvement in industrial distillation. Main efforts were concentrated on local improvements and the main
Commercial proposals principles of the modern distillation were not changed for a long time.
Contact us In general distillation is a very conservative process. Huge rectification towers at chemical and petrochemical plants
are a symbol of modern industry. Very often the height of rectification towers exceeds 50 m. Costs of design,
manufacturing, transport, operation and repairs become enormous. Rectification towers looks like dinosaurs and
http://www.linas.ru clash with the technological image of the 21st century.

2. Core fundamental problems of modern distillation

For more than 50 years of an existence of modern distillation the opinion was gradually created, that significant
improvement of industrial distillation is impossible. However the careful and detailed analysis of principles and
applications of modern industrial distillation shows several fundamental contradictions and disadvantages. These
contradictions are concentrated in an arrangement of heat and mass exchange processes and reflux.

2.1. Arrangement of the heat and mass exchange processes in conventional distillation

Indeed the distillation is based on heat and mass exchange processes between two or more compounds.
In the modern industrial distillation the basic and actually the main attention is given to the mass exchange
process. All efforts of technologists and engineers are directed on improvements of mass exchange processes
between substances. Complex and expensive packing and trays of the most complex forms are created to improve
the mass exchange processes. Thus to the heat exchange processes between substances it is not given due attention.

From our point of view the basic problem of modern distillation consists in absence of the control over the
heat exchange processes and a particular to the energy of condensation of substances. Moreover, the control
over the energy of evaporation and condensation is the primary process, and it can operate and rule the mass
exchange process.

Let's consider the elementary case of water and alcohol distillation (Fig.1). After evaporation of both substances,
there is their condensation on surface of packing or trays. Thus it generates the large amount of the heat of the
condensation.
 Fig.1.
Scheme of elementary step of a water-alcohol distillation

Modern distillation practically does not pay an attention to the heat of condensation and does not operate with him.
It results to that a part of the condensed water can again evaporate and to rise above on a column. In an ideal step
after the first act of the evaporation and condensation all amount of water should be removed at a bottom of the
column (See the figure 1), and all amount of alcohol should be evaporated again and to be removed at the top of the
column. It means, that the energy of the water condensation is necessary to remove from a distillation space or to
make a heat transfer to a process of alcohol evaporation. However in modern columns there are no devices or
technology which would supervise this process.

Thus, the energy of a condensation is not supervised and it results in an efficiency reduction of a separation,
increase in height of columns and expenses of energy.
It is necessary to organize a certain device or technology which would supervise the energy of condensation of
separated substances. Introduction of such mechanism would lead to sharp decrease in a height of a column and an
efficiency of distillation. The modern conventional distillation is based on tray and packing rectification columns.
From our point of view it is nearly impossible to introduce any energy operation device inside of modern column to
control and supervise distribution of energy inside of distillation space above mentioned columns.

2.2. Arrangement of reflux

The second weak point of the modern industrial distillation is the arrangement of reflux and its distribution along a
height of rectification towers.
As the reflux a pure distillate is usually used. The reflux is entered through a special device at a top of the column
(Fig.2).
Simple analysis of the mass exchange processes in a conventional column shows, that a huge surplus of pure
distillate on top of a column is not required. Moreover, the reflux at the top of a column is absolutely not necessary.
The amount and structure of a reflux should be various along the height of the column. The bottom part of the
column needs a plenty of reflux with a large amount of the higher boiling component. With increase in the height of
the column a content of the higher boiling component in reflux should decrease and an amount of the reflux should
decrease too (Fig.2. Desirable distribution of the reflux). This arrangement of the reflux reduces a total amount of the
reflux and makes the reflux more effective.
 Fig.2.
Scheme of reflux distributions for a conventional and Linas distillations (V is vapor).

Modern industrial distillation uses a different arrangement of reflux. There is no the arrangement of the reflux
composition and an amount of the reflux along the height of a column. It results in a sharp increase in expenses of
energy and increase in height of columns.

Thus, from our point of view the basic lacks of modern industrial rectification are the following:

1. Modern design of rectification towers does not pay attention to real microbalance of heat and mass
exchange processes.
2. Mass and heat exchange processes in every point of conventional columns are not correlated to each
others.
3. Modern arrangement of a composition and an amount of reflux does not correlated to real distribution
of low and high boiling compounds along rectification towers.

This results in columns being very tall and requires more energy for the distillation process. Finally this results in a
high cost of an operation.

2.3. Linas distillation technology

A small group of enthusiastic and highly professional researchers and engineers made an attempt to develop a new
breakthrough distillation technology. This attempt was successful. They have developed a totally new solution for
industrial distillation and have developed the new industrial distillation technology called Linas technology.

Linas technology is based on the modified very much film distillation.

Indeed, the conventional film distillation has several attractive advantages such as simple construction, the very low
flow resistance and good separation ability. The film distillation has the lowest height of the theoretical tray (equal 5
mm) among all distillation technologies, but only if vapor velocity is around 1 cm/s. Therefore applications of the
film distillation are very limited and not really applicable in a large scale industrial distillation. At high velocity of
vapor along vertical surfaces a film ceases to be uniform and heat and mass exchange processes become unstable
and all advantages of film rectification are not realized.

This main problem of the conventional film distillation was solved by Linas technology.

Linas technology is based on vertical tubes with a length from 0.5 m till 3 m and a diameter 6-25 mm.

Linas technology concentrates the main attention on the energy of a condensation of separated compounds.
Temperature of Linas tube's wall (Fig.3, temperature TW) is fixed on certain level along a height of tube between TA
and TB. Under distillation conditions the compound B with higher boiling point (TB) condenses always on the walls
of tubes and removed down in a liquid film form. The compound A with lower boiling point (TA) is evaporated
always from a surface of the film and leave tube's space as the final distillate. Indeed the energy and the energy
barrier are the moving force of Linas technology. Special device outside of Linas tubes is responsible on the
arrangement of the temperature TW on the certain level. In many cases TW is varied along the height of the tubes
from the temperature close to TA to the temperature close to TB. Access of energy of the B component's
condensation is removed outside of the rectification space by the additional condenser (Fig.5).

Fig.3.
Moving force of Linas technology (TA and TB are boiling temperatures of compounds A and B).
Linas technology is based on so called internal reflux. The temperature different of TW along the height of the tubes
manages the reflux. Indeed the A and B from moving up vapor phase are condensed on the moving down film. This
condensed mixture of A and B is the Linas reflux. The composition of the reflux is very much different. At low
level of the tubes the composition of the reflux is close to evaporated mixture of A and B and with large
concentration of B. Amount of the reflux is relative large. At high level of the tubes the composition of the reflux
contains mainly the component A. The amount of the reflux becomes small (Fig. 4). Principles of Linas reflux
arrangement is presented on Figure 4.

Industrial applications of Linas technology give:

1. Stable distillation film under a velocity of a vapor stream inside of Linas's column up to 1.5-2.0 m/s.
2. Adaptation of a heat and mass exchange processes inside the Linas column to the physical properties of
separated compounds.
3. New arrangement of the reflux process. All evaporated distillate is a final product according to the
scheme below. Linas reflux process takes place inside the Linas distillation towers.
4. Three to ten fold reduction of the height of rectification towers and 50 to 100 time reduction in the
amount of separated compounds inside the Linas column compared with conventional rectification
towers.

Fig.4.
Relative distribution of an amount of the reflux and ratio of components A and B in the reflux. (The total amount of
reflux is presented as the square inside of red line)

Technologic scheme of Linas rectification tower is little bit different from the conventional one (Fig.5). There is no
a reflux line back to the tower from the distillate drum. The Linas tower contains an additional device (the
condenser) compare to the conventional tower.
From technical point of view the external different between technologic scheme of Linas rectification tower and the
conventional tower is relative small.
 Technologic scheme of Linas and conventional rectification towers.

Linas scale up process

Linas technology includes the totally simple scale up process. For example the design of Linas rectification part
included in Linas rectification tower for crude oil refinery with a capacity from 80 MTY till 50.000 MTY is
presented below.

The basic element of above described rectification part is the single Linas distillation tube and the capacity of Linas
rectification towers can be increased by the number of the tubes. Indeed there is no difference in parameters of
distillation processes between one tube with a capacity 80 MTY and 621 tubes with a total capacity 50.000 MTY. It
means that the design of any industrial unit can be done very fast if there is data for one distillation tube.

A conventional scale up process is very complex and does not give any warranty for a final successful result.
Increasing of a capacity always gives larger diameters and heights of columns.

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 Linas Technology - Distillation of 21 century
®

General Information Industrial applications of Linas Technology and pilot plant projects
About us
Introduction in Linas Because of a limited financial support the only several industrial applications of Linas technology were developed
technology for last four years. These applications are presented below.
Industrial applications
of Linas Technology
and pilot plant
projects 1. Distillation of trifluoride-methyl sulfonic acid
Patent protection of
The first industrial application of Linas technology was accomplished four years ago for the distillation of
Linas technology trifluoride-methyl sulfonic acid at the Angarsk chemical plant (East Siberia). The height of the column is only 1.4
Publications and meters. After four years of a continuous operation the rectification column did not demonstrate any real distillation
presentations of Linas problems.
technology
Commercial proposals
Contact us
2. Linas technology in oil refinery applications
http://www.linas.ru Linas distillation technology was applied for the oil refinery two years ago. First oil refinery SMR-8(10) (8000-
10.000 MTY of crude oil) is in test operation in city Miass (Russia) for 1.5 years. At the end of 2003 SMR-8(10)
was moved to city Kemerovo (West Siberia). Certified gasoline, diesel oil and fuel oil were produced by the oil
refinery. The height of the rectification part of the tower is 1.5 m. In operation, the rectification tower contains 1.5
kg of compounds only. Industrial operation of SMR-8(10) confirms the advantages of the new technology.
A general view of SMR-8(10) is presented on a photo below. The rectification tower is very small and it is even
difficult to find the tower in the photo

Technical data of oil refinery SMR-8(10) (Linas-Tekhno) based on Linas technology and a conventional oil refinery
presented in Table 1.
Indeed the technical parameters of Linas rectification tower are superior to conventional rectification tower.

Table 1

Type of plant SMR-8(10) conventional

Capacity 8000-10000 MTY 7000 MTY
Type of feed Crude oil Crude oil
Yield of useful products >99 % wt. >99 % wt.
Pressure 0,01 MPa 0,17 MPa
Weight of unit 4,2 t (column + furnace + reboiler) 5,6 t (column only)
Height of column 1,5 m 10 m
Diameter of column 500 mm 520 mm
Consumption of furnace 13-20 kg/h 12,6-22 kg/h
Electric power 8kW 27kW
o o
Operation conditions -55 C - (+55 C) -30 C - (+55oC)
o

Type of operation continuous continuous

Next generation of Linas oil refinery SMR-50 (50.000 MTY of crude oil) will be in an industrial operation in July
2004. The rectification tower of SMR-50 is improved very much compared the tower of SMR-8(10). The diameter
of the rectification tower is 1.0 meter. The height of the rectification part is 1.5 meter and the height of the stripping
is 1 meter. The tower contains the original design diesel oil tray combined with an evaporator. It gives the
possibility to remove small amount of light and heavy fractions from the diesel oil and improve very much a quality
of diesel oil. The rectification part together with the diesel oil tray is combined in Linas rectification unit. Total
height of the column is around 5 meters. General scheme of the tower of the oil refinery SMR-50 is presented on
Fig.1.
In conventional distillation the towers for the oil refineries contain additional small towers to make the diesel oil or
kerosene with good quality. The Linas tower does not need the additional towers. The Linas rectification unit (Fig.1)
with the height 2.5 meter makes a good quality gasoline and diesel oil.
Linas rectification unit has a very lower pressure drop and therefore the temperature of a crude oil feed could be
established on level from 280 till 315oC. It is a lowest temperature of a crude oil feed among oil refineries.

Several SMR-50 will be built up and start up at the end of 2004 and the beginning of 2005 in Russia.
Linas rectification tower for the oil refinery with a capacity 50.000 MTY is much smaller than an average size of
conventional rectification tower with the same capacity 50.000 MTY (Fig.2).

Fig.1
General scheme of Linas rectification tower including in oil refinery SMR-50

Fig.2.
 Comparable size of Linas and conventional rectification towers for the oil refinery with the capacity 50.000 MTY

Companies Linas-Tekhno Inc., and Linas Technology International Corporation plan to develop and built up in a
nearest future the Linas oil refinery towers with a capacity 150 MTY and 500 MTY. Heights of rectification towers
for SMR-150 and SMR-500 will be around 6 meters.

Several Linas distillation processes on a pilot plant level were developed.

3. Distillation of waste lube oil

Very compact pilot distillation unit was built up for a regeneration of waste lube oil in 2002. Height of a
rectification part is 1.5 meter only. Total height of the rectification tower is 2.5 meters. Around 75-85 % wt. of
waste lube oil is regenerated and could be used again as lube oil for ship diesel engines. Based on the pilot plant
data the design of an industrial plant with a capacity 75.000 MTY was done.

4. Distillation of ethanol-water mixture

The pilot plant for alcohol-water distillation was designed and built up at beginning of 2003. The pilot plant was
tested in the Netherlands in company Zeton BV and presented in Achema (Germany) in May 2003.

The pilot plant contains the rectification part with a height 1.5 meter, stripping part with a height 0.5 meter.
Diameter of the rectification tower is 0.2 meter. Standard feeding is 15 kg of alcohol-water mixture per hour. Feed
contains 6-20 % wt. of alcohol. Concentration of alcohol in the distillate is around 90 % wt. In September 2003 the
technology was improved and a distillate with the alcohol concentration around 94 wt. % was obtained.

Photo of the pilot plant is presented below.

5. Pilot plant for oil refinery application

The pilot plant was designed and built up in August 2003. The feed capacity of the plant is around 7-10 kg of crude
oil per hour. The pilot plant is supplied by a full automatic system. The photo of the pilot plant is presented below.

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 Linas Technology - Distillation of 21 century
®

General Information Commercial proposals

About us
Introduction in Linas At moment companies Linas-Techno and Linas Technology International Corporation concentrate the main business
technology activity on oil refinery because of a limited financial support.
Industrial applications Oil refineries SMR-8(10) and SMR-50 with the annual capacity 10 and 50 MTY are the commercial products.
of Linas Technology
and pilot plant
projects
1. Oil refinery SMR-8(10)
Patent protection of
Linas technology First installation SMR-8 with the capacity 8000 or 10000 MTY has been started in trial operation in March, 2002 in
Publications and Miass (Russia). Installation has passed the long period of tests and after lines of improvements is prepared for
presentations of Linas constant commercial operation.
technology At the moment after change of the owner the installation transported and will begin a permanent commercial
Commercial proposals operation in city Kemerovo in January 2004. Basic feature SMR-8 is a using of Linas distillation technology. It
Contact us results in significant reduction of height of the rectification tower and to sharp increase of safety of operation of
installation and simplification of managerial process by installation.
Installation is executed by a modular principle and can be transported by three trucks.
http://www.linas.ru

The price of SMR-8 begins from 430.000 euros (without delivery to a place of permanent operation and charges for
start-up of installation).
Term of performance of contract is 12 months from the moment of the first payment.
The companies Linas-Techno and Linas Technology Int. Corp. carry up the following kinds of actions:

1. Design of installation (the non-standard equipment) and automatic system.

2. Manufacturing installation (the non-standard equipment) and a complete set of installation by the
standard equipment, materials and so forth accessories and carrying out of tests.
3. Delivery of installation to the prepared platform the customer in block performance with a high degree
of readiness and necessary accessories for installation, carrying out of tests and commissioning of
installation.
4. Preparation and training of the personnel for operation of installation.

2. Oil refinery SMR-50

First installation SMR-50 will be in commercial operation by summer of 2004. Now first contract on delivery SMR-
50 is eight months in work.
Several contacts is planned for manufacturing and delivery SMR-50 in the near future.
Structurally installation SMR-50 is similar to installation SMR-8 and different only the increased productivity.
Distinctive feature is very small height of rectification tower (about 5 м) and the increased safety of operation of the
installation, based on application of Linas technology.
Installation is executed by a modular principle and can be transported on 6-7 trucks.
At the moment the price of installation begins from 1.030.000 euros (without delivery to a place of permanent
operation and charges of the company for start-up of installation). The price of installation depends appreciably on a
complete set, performance and a market situation.

PFD of SMR-50 (33K)

3. Possible commercial proposals

The companies Linas-Techno and Linas Technology Int. Corp. has made several processes at a pilot plant level and
are ready to offer potential customers development and delivery of plants. At the moment there is an operational
experience on the following processes:
 1. Regeneration of waste lube oil.
2. Distillation of ethanol.

***

The companies is ready to develop also new distillation industrial processes on the basis of Linas technology at
presence of the commercial order.
The following kinds of rectification can be quickly feasible:

1. Division of the hydrocarbons received at oil refining.

2. Development rectification towers for primary division of oil by capacity of 100-250 thousand tons per
year.
3. Allocation of useful mineral oil from a various sort of waste products of petrol industry.

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Shell Global Solutions Impact 4, 2003

New life for cracker margins

Hydrocracker revamp produces benefits for German refinery

A hydrocracking unit revamp has boosted margins by almost $4 million per year
and achieved payback on the investment in less than a year for a refinery in
Germany.

The hydrocracker unit’s fractionation section was underperforming, and the

refinery needed to run it to give a higher throughput and increase distillates
production. Refinery management asked Shell Global Solutions to conduct a
feasibility study on upgrading the unit.

“The main constraint in the unit was the overhead condensing capacity of the
main fractionator that prevented the removal of all the potential kerosene from
the column feed,” says Dico Engel, process engineer, distillation group, Shell
Global Solutions International BV. “The result of this poor performance was that
the diesel side cut from the mild vacuum column still contained a significant
amount of kerosene, while the recycle oil contained a lot of valuable diesel.”

The study report recommended a design revamp for the hydrocracking unit that
involved increasing the effectiveness of the reboiler furnace to raise the column
bottom temperature. This would enable evaporation of all the kerosene from the
bottom liquid. Shell Global Solutions recommended a new
mid-circulating reflux positioned just below the heavy naphtha draw-off tray to
condense the additional vapours generated by the reboiler furnace. Furthermore,
high-liquid-capacity trays (Shell HiFi† trays) were proposed to accommodate the
increased vapour and liquid traffic in the tower.

The revamp has given the refinery the flexibility to switch production between
different products. “As well as achieving better fractionation, we can now adapt to
the changing demands of the market, whether the need is for heavy naphtha,
motor gasoline, jet fuel or diesel,” says the hydrocracking unit manager.

“This was a fast-track project because the revamp had to be realised during the
refinery shutdown period,” he continues. “This type of project is possible only if
you have a good technical partner who understands your needs and works with
you. We finished the project on time and without overrunning the budget. We are
now able to run an additional 15% of fresh feed, with the same run length: that’s
payback on the investment in less than a year.”

• Contact: Jan de Vries

Email: jan.devries@shell.com
†
Shell HiFi is a Shell trademark
 OSHA TECHNICAL MANUAL - SECTION IV: CHAPTER 3

U.S. Department of Labor

Occupational Safety & Health Administration

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SECTION IV: CHAPTER 3

PRESSURE VESSEL GUIDELINES

Contents:

I. Introduction
II. Recent Cracking Experience in Pressure Vessels
III. Nondestructive Examination Methods
IV.
Information for Safety Assessment
V.
Bibliography

Appendix IV:3-1. Recordkeeping Data for Steel Vessels and Low-Pressure Storage Tanks

1. INTRODUCTION.

1. Recent inspection programs for metallic pressure containment vessels and tanks have
revealed cracking and damage in a considerable number of the vessels inspected. Safety and
hazard evaluations of pressure vessels, as also presented in OSHA Instruction PUB 8-1.5,
need to consider the consequences of a leakage or a rupture failure of a vessel.

2. Two consequences result from a complete rupture:

■ Blast effects due to sudden expansion of the pressurized fluid.

■ Fragmentation damage and injury, if vessel rupture occurs.

3. For a leakage failure, the hazard consequences can range from no effect to very serious
effects:

■ Suffocation or poisoning, depending on the nature of the contained fluid, if the leakage
occurs into a closed space.
■ Fire and explosion (physical hazards for a flammable fluid).
■ Chemical and thermal burns from contact with process liquids.

4. Only pressure vessels and low pressure storage tanks widely used in process, pulp and paper,
petroleum refining, and petrochemical industries and for water treatment systems of boilers
and steam generation equipment are covered in this chapter. Excluded are vessels and tanks
used in many other applications. Also excluded are other parts of a pressure containment
system such as piping and valves.

5. The types and applications of pressure vessels included and excluded in this chapter are
summarized in Table IV:3-1. An illustration of a schematic pressure vessel is presented in
Figure IV:3-1.

NOTE: Though this review of pressure vessels excludes inspection or evaluation of safety
release valves, the compliance officer should be aware that NO valves or T-fittings should be
present between the vessel and the safety relief valve.

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OSHA TECHNICAL MANUAL - SECTION IV: CHAPTER 3

6. Most of the pressure or storage vessels in service in the United States will have been designed
and constructed in accordance with one of the following two pressure vessel design codes:

■ The ASME Code, or Section VIII of the ASME (American Society of Mechanical
Engineers) "Boiler and Pressure Vessel Code."
■ The API Standard 620 or the American Petroleum Institute Code which provides rules
for lower pressure vessels not covered by the ASME Code.

In addition, some vessels designed and constructed between 1934 to 1956 may have used the
rules in the "API-ASME Code for Unfired Pressure Vessels for Petroleum Liquids and Gases."
This code was discontinued in 1956.

7. Vessels certification can only be performed by trained inspectors qualified for each code.
Written tests and practical experience are required for certification. Usually, the compliance
office is not equipped for this task, but is able to obtain the necessary contract services.
TABLE IV:3-1. VESSEL TYPES

Vessels included: Vessel types specifically excluded:

Stationary and unfired Vessels used as fired boilers

Used for pressure containment of gases and Vessels used in high-temperature processes (above 315° C,
liquids 600° F) or at very low and cryogenic temperatures

Constructed of carbon steel or low alloy steel Vessels and containers used in transportable systems

Operated at temperatures between -75° and Storage tanks that operate at nominally atmospheric pressure
315° C (-100° and 600° F)

Piping and pipelines

Safety and pressure-relief valves

Special-purpose vessels, such as those for human occupancy

FIGURE IV:3-1. SOME MAJOR PARTS OF A PRESSURE VESSEL.

2. RECENT CRACKING EXPERIENCE IN PRESSURE VESSELS.

1. DEAERATOR SERVICE.

1. Deaeration refers to the removal of noncondensible gases, primarily oxygen, from the
water used in a steam generation system. Deaerators are widely used in many
industrial applications including power generation, pulp and paper, chemical, and
petroleum refining and in many public facilities such as hospitals and schools where
steam generation is required. In actual practice, the deaerator vessel can be separate
from the storage vessel or combined with a storage vessel into one unit.

2. Typical operational conditions for deaerator vessels range up to about 300 psi and up
to about 150° C (300° F). Nearly all of the vessels are designed to ASME Code,
resulting in vessel wall thicknesses up to but generally less than 25 mm (1 in). The
vessel material is almost universally one of the carbon steel grades.

3. Analysis of incident survey data and other investigations has determined the following
features about the deaerator vessel cracking.

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OSHA TECHNICAL MANUAL - SECTION IV: CHAPTER 3

■ Water hammer is the only design or operational factor that correlates with
cracking.
■ Cracking is generally limited to weld regions of vessels that had not been
postweld heat treated.
■ Corrosion fatigue appears to be the predominant mechanism of crack formation
and growth.

4. The failures and the survey results have prompted TAPPI (Technical Association of Pulp
and Paper Industry), the National Board of Boiler and Pressure Vessel Inspectors, and
NACE (National Association of Corrosion Engineers) to prepare inspection, operation
and repair recommendations. For inspection, all recommendations suggest:

■ Special attention to the internal surface of all welds and heat-affected zones
(HAZ).
■ Use of the wet fluorescent magnetic particle (WFMT) method for inspection.

5. The TAPPI and the NACE recommendations also contain additional items, such as:

■ Inspection by personnel certified to American Society for Nondestructive

Testing's SNT-TC-1A minimum Level I and interpretation of the results by
minimum Level II.
■ Reinspection within one year for repaired vessels, 1-2 years for vessels with
discontinuities but unrepaired, and 3-5 years for vessels found free of
discontinuities.

2. AMINE SERVICE.

1. The amine process is used to remove hydrogen sulfide (H2S) from petroleum gases
such as propane and butane. It is also used for carbon dioxide (CO2) removal in some
processes. Amine is a generic term and includes monoethanolamine (MEA),
diethanolamine (DEA) and others in the amine group. These units are used in
petroleum refinery, gas treatment and chemical plants.

2. The operating temperatures of the amine process are generally in the 38 to 93° C
(100° to 200° F) range and therefore the plant equipment is usually constructed from
one of the carbon steel grades. The wall thickness of the pressure vessels in amine
plants is typically about 25 mm (1 in.).

3. Although the possibility of cracking of carbon steels in an amine environment has been
known for some years, real concern about safety implications was highlighted by a
1984 failure of the amine process pressure vessel. Overall, the survey found about
40% cracking incidence in a total of 294 plants. Cracking had occurred in the absorber/
contactor, the regenerator and the heat exchanger vessels, and in the piping and other
auxiliary equipment. Several of the significant findings of the survey were:

■ All cracks were in or near welds.

■ Cracking occurred predominantly in stressed or unrelieved (not PWHT) welds.
■ Cracking occurred in all amine vessel processes but was most prevalent in MEA
units.
■ WFMT and UT (ultrasonic test) were the predominant detection methods for
cracks; internal examination by WFMT was the preferred method.

4. Information from laboratory studies indicate that pure amine does not cause cracking
of carbon steels but amine with carbon dioxide in the gas phase causes severe
cracking. The presence or absence of chlorides, cyanides, or hydrogen sulfide may also
be factors but their full role in the cracking mechanism is not completely known at
present.

3. WET HYDROGEN SULFIDE.

1. Wet Hydrogen Sulfide refers to any fluid containing water and hydrogen sulfide (H2S).
Hydrogen is generated when steel is exposed to this mixture and the hydrogen can
enter into the steel. Dissolved hydrogen can cause cracking, blistering, and
embrittlement.

2. The harmful effects of hydrogen generating environments on steel have been known
and recognized for a long time in the petroleum and petrochemical industries. In
particular, sensitivity to damage by hydrogen increases with the hardness and strength
of the steel; and damage and cracking are more apt to occur in high strength steels, as
follows:

■ Significant cracks can start from very small hard zones associated with welds;
these hard zones are not detected by conventional hardness tests.

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OSHA TECHNICAL MANUAL - SECTION IV: CHAPTER 3

■ Initially small cracks can grow by a stepwise form of hydrogen blistering to form
through thickness cracks.
■ NACE/API limits on weld hardness may not be completely effective in preventing
cracking.
■ Thermal stress relief (postweld heat treatment, PWHT) appears to reduce the
sensitivity to and the severity of cracking.

3. Wet hydrogen sulfide has also been found to cause service cracking in liquified
petroleum gas (LPG) storage vessels. The service cracking in the LPG vessels occurs
predominantly in the weld heat affected zone (HAZ). The vessels are usually spherical
with wall thickness in the 20 mm to 75 mm (0.8 in. to 3 in.) range.

4. Recommendations for new and existing wet hydrogen-sulfide vessels to minimize the
risk of a major failure include:

■ Use lower-strength steels for new vessels.

■ Schedule an early inspection for vessels more than five years in service.
■ Improve monitoring to minimize breakthrough of hydrogen sulfide.
■ Replace unsafe vessels or downgrade to less severe, usually lower-pressure,
service.

4. AMMONIA SERVICE.

1. Commercial refrigeration systems, certain chemical processes, and formulators of

agricultural chemicals will be sites of ammonia service tanks. Careful inspections of
vessels used for storage of ammonia (in either vapor or liquid form) in recent years
have produced evidence of serious stress corrosion cracking problems. The vessels for
this service are usually constructed as spheres from one of the carbon steel grades,
and they operate in the ambient temperature range.

2. The water and oxygen content in the ammonia has a strong influence on the
propensity of carbon steels to crack in this environment. Cracks have a tendency to be
found to be in or near the welds in as-welded vessels. Cracks occur both transverse
and parallel to the weld direction. Thermal stress relieving seems to be a mitigating
procedure for new vessels, but its efficacy for older vessels after a period of operation
is dubious partly because small, undetected cracks may be present.

5. PULP DIGESTER SERVICE.

1. The kraft pulping process is used in the pulp and paper industry to digest the pulp in
the papermaking process. The operation is done in a relatively weak (a few percent)
water solution of sodium hydroxide and sodium sulfide typically in the 110° to 140° C
(230° to 285° F) temperature range. Since the early 1950's, a continuous version of
this process has been widely used. Nearly all of the vessels are ASME Code vessels
made using one of the carbon steel grades with typical design conditions of 175° to
180° C (350° to 360° F) and 150 psig.

2. These vessels had a very good service record with only isolated reports of cracking
problems until the occurrence of a sudden rupture failure in 1980. The inspection
survey has revealed that about 65% of the properly inspected vessels had some
cracking. Some of the cracks were fabrication flaws revealed by the use of more
sensitive inspection techniques but most of the cracking was service-induced. The
inspection survey and analysis indicates the following features about the cracking.

■ All cracking was associated with welds.

■ Wet fluorescent magnetic particle (WFMT) testing with proper surface
preparation was the most effective method of detecting the cracking.
■ Fully stress-relieved vessels were less susceptible.
■ No clear correlation of cracking and noncracking could be found with vessel age
and manufacture or with process variables and practices.
■ Analysis and research indicate that the cracking is due to a caustic stress
corrosion cracking mechanism although its occurrence at the relatively low
caustic concentrations of the digester process was unexpected.

3. Currently, preventive measures such as weld cladding, spray coatings, and anodic
protection are being studied, and considerable information has been obtained. In the
meantime, the recommended guideline is to perform an annual examination.

6. SUMMARY OF SERVICE CRACKING EXPERIENCE.

1. The preceding discussion shows a strong influence of chemical environment on

cracking incidence. This is a factor that is not explicitly treated in most design codes.
Service experience is the best and often the only guide to in-service safety

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OSHA TECHNICAL MANUAL - SECTION IV: CHAPTER 3

assessment.

2. For vessels and tanks within the scope of this document, the service experience
indicates that the emphasis of the inspection and safety assessment should be on:

■ Vessels in deaerator, amine, wet H2S, ammonia and pulp digesting service.
■ Welds and adjacent regions.
■ Vessels that have not been thermally stress relieved (no PWHT of fabrication
welds).
■ Repaired vessels, especially those without PWHT after repair.

3. The evaluation of the severity of the detected cracks can be done by fracture
mechanics methods. This requires specific information about stresses, material
properties, and flaw indications. Generalized assessment guidelines are not easy to
formulate. However, fortunately, many vessels in the susceptible applications listed
above operate at relatively low stresses, and therefore, cracks have a relatively smaller
effect on structural integrity and continued safe operation.

7. NONDESTRUCTIVE EXAMINATION METHODS.

Of the various conventional and advanced nondestructive examination (NDE) methods, five
are widely used for the examination of pressure vessels and tanks by certified pressure vessel
inspectors. The names and acronyms of these common five methods are:

■ Visual Examination (VT)

■ Liquid Penetrant Test (PT)
■ Magnetic Particle Test (MT)
■ Gamma and X-ray Radiography (RT)
■ Ultrasonic Test (UT)

VT, PT, and MT can detect only those discontinuities and defects that are open to the surface
or are very near the surface. In contrast, RT and UT can detect conditions that are located
within the part. For these reasons, the first three are often referred to as "surface"
examination methods and the last two as "volumetric" methods. Table II of PUB 8-1.5
summarizes the main features of these five methods.

1. VISUAL EXAMINATION (VT) is easy to conduct and can cover a large area in a short
time. It is very useful for assessing the general condition of the equipment and for
detecting some specific problems such as severe instances of corrosion, erosion, and
hydrogen blistering. The obvious requirements for a meaningful visual examination are
a clean surface and good illumination.

2. LIQUID PENETRANT TEST (PT) depends on allowing a specially formulated liquid

(penetrant) to seep into an open discontinuity and then detecting the entrapped liquid
by a developing agent. When the penetrant is removed from the surface, some of it
remains entrapped in the discontinuities. Application of a developer draws out the
entrapped penetrant and magnifies the discontinuity. Chemicals which fluoresce under
black (ultraviolet) light can be added to the penetrant to aid the detectability and
visibility of the developed indications. The essential feature of PT is that the
discontinuity must be "open," which means a clean, undisturbed surface.

The PT method is independent of the type and composition of the metal alloy so it can
be used for the examination of austenitic stainless steels and nonferrous alloys where
the magnetic particle test is not applicable.

3. MAGNETIC PARTICLE TEST (MT).

1. This method depends on the fact that discontinuities in or near the surface
perturb magnetic flux lines induced into a ferromagnetic material. For a
component such as a pressure vessel where access is generally limited to one
surface at a time, the "prod" technique is widely used. The magnetic field is
produced in the region around and between the prods (contact probes) by an
electric current (either AC or DC) flowing between the prods. The ferromagnetic
material requirement basically limits the applicability of MT to carbon and low-
alloy steels.

2. The perturbations of the magnetic lines are revealed by applying fine particles
of a ferromagnetic material to the surface. The particles can be either a dry
powder or a wet suspension in a liquid. The particles can also be treated to
fluoresce under black light. These options lead to variations such as the "wet
fluorescent magnetic particle test" (WFMT). MT has some capability for
detecting subsurface defects. However, there is no easy way to determine the
limiting depth of sensitivity since it is highly dependent on magnetizing current,

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material, and geometry and size of the defect. A very crude approximation
would be a depth no more than 1.5 mm to 3 mm (1/16 in. to 1/8 in.).

3. A very important precaution in performing MT is that corners and surface

irregularities also perturb the magnetic field. Therefore, examining for defects in
corners and near or in welds must be performed with extra care. Another
precaution is that MT is most sensitive to discontinuities which are oriented
transverse to the magnetic flux lines and this characteristic needs to be taken
into account in determining the procedure for inducing the magnetic field.

4. RADIOGRAPHY (RT).

1. The basic principle of radiographic examination of metallic objects is the same

as in any other form of radiography such as medical radiography. Holes, voids,
and discontinuities decrease the attenuation of the X-ray and produce greater
exposure on the film (darker areas on the negative film).

2. Because RT depends on density differences, cracks with tightly closed surfaces

are much more difficult to detect than open voids. Also, defects located in an
area of an abrupt dimensional change are difficult to detect due to the
superimposed density difference. RT is effective in showing defect dimensions
on a plane normal to the beam direction but determination of the depth
dimension and location requires specialized techniques. Since ionizing radiation
is involved, field application of RT requires careful implementation to prevent
health hazards.

5. ULTRASONIC TESTING (UT).

1. The fundamental principles of ultrasonic testing of metallic materials are similar

to radar and related methods of using electromagnetic and acoustic waves for
detection of foreign objects. The distinctive aspect of UT for the inspection of
metallic parts is that the waves are mechanical, so the test equipment requires
three basic components:

■ Electronic system for generating electrical signal.

■ Transducer system to convert the electrical signal into mechanical
vibrations and vice versa and to inject the vibrations into and extract
them from the material.
■ Electronic system for amplifying, processing, and displaying the return
signal.

Very short signal pulses are induced into the material and waves reflected back
from discontinuities are detected during the "receive" mode. The transmitting
and detection can be done with one transducer or with two separate transducers
(the tandem technique).

2. Unlike radiography, UT in its basic form does not produce a permanent record of
the examination. However, more recent versions of UT equipment include
automated operation and electronic recording of the signals.

3. Ultrasonic techniques can also be used for the detection and measurement of
general material loss such as by corrosion and erosion. Since wave velocity is
constant for a specific material, the transit time between the initial pulse and
the back reflection is a measure of the travel distance and the thickness.

6. DETECTION PROBABILITIES AND FLAW SIZING.

1. The implementation of NDE (nondestructive examination) results for structural

integrity and safety assessment involves a detailed consideration of two
separate but interrelated factors.

■ Detecting the discontinuity.

■ Identifying the nature of the discontinuity and determining its size.

2. Much of the available information on detection and sizing capabilities has been

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developed for aircraft and nuclear power applications. This kind of information is
very specific to the nature of the flaw, the material, and the details of the test
technique, and direct transference to other situations is not always warranted.

3. The overall reliability of NDE is obviously an important factor in a safety and

hazard assessment. Failing to detect or undersizing existing discontinuities
reduces the safety margin while oversizing errors can result in unnecessary and
expensive outages. High reliability is achieved through a combination of factors,
as follows:

■ Validated procedures, equipment and test personnel.

■ Utilization of diverse methods and techniques.
■ Application of redundancy by repetitive and independent tests.

Finally, it is useful to note that safety assessment depends on evaluating the

"largest flaw that may be missed, not the smallest one that can be found."

8. INFORMATION FOR SAFETY ASSESSMENT.

This chapter and PUB 8-1.5 has a large amount of information on the design rules, inspection
requirements, and service experience, relevant to pressure vessels and low pressure storage
tanks used in general industrial applications. Though the compliance officer is not usually
qualified as a pressure vessel inspector, as a summary and a reminder, Appendix IV:3-1
outlines the information, data, and recordkeeping that are necessary, useful, or indicative of
safe management of operating vessels and tanks.

These records, in addition to the construction and maintenance logs, usually are kept by the
plant engineer, maintenance supervisor, or facility manager, and will be indicative of the
surveillance activities around safe operation of pressure vessels.

9. BIBLIOGRAPHY.

Chuse, R. 1984. Pressure Vessels: The ASME Code Simplified. 6th ed. McGraw-Hill: New York.

Forman, B. Fred. 1981. Local Stresses in Pressure Vessels. Pressure Vessel Handbook Publishing, Inc.:
Tulsa.

Hammer, W. 1981. Pressure Hazards in Occupational Safety Management and Engineering. 2nd ed. Prentice-
Hall: New York.

McMaster, R.C. and McIntire, P. (eds.) 1982-1987. Nondestructive Testing Handbook. 2nd ed., Vols. 1-3.
American Society for Metals/American Society of Nondestructive Testing: Columbus.

Megyesy, E.F. 1986. Pressure Vessel Handbook. 7th ed. Pressure Vessel Handbook Publishing Inc.: Tulsa.

OSHA Instruction Pub 8-1.5. 1989. Guidelines for Pressure Vessel Safety Assessment. Occupational Safety
and Health Administration: Washington, D.C.

Thielsch, H. 1975. Defects and Failures in Pressure Vessels and Piping. 2nd ed., Chaps. 16 and 17. Reinhold:
New York.

Yokell, S. 1986. Understanding the Pressure Vessel Code. Chemical Engineering 93(9):75-85.

APPENDIX IV:3-1. RECORDKEEPING DATA FOR STEEL VESSELS AND LOW PRESSURE STORAGE
TANKS

INTRODUCTION AND SCOPE. This outline summarizes information and data that will be helpful in
assessing the safety of steel pressure vessels and low pressure storage tanks that operate at temperatures
between -75° and 315 ° C (-100° and 600° F).

VESSEL IDENTIFICATION AND DOCUMENTATION. Information that identifies the specific vessel being
assessed and provides general information about it include the following items:

● Current owner of the vessel

● Vessel location
- Original location and current location if it has been moved
● Vessel identification
- Manufacturer's serial number
- National Board number if registered with NB
● Manufacturer identification
- Name and address of manufacturer

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- Authorization or identification number of the manufacturer

● Date of manufacture of the vessel
● Data report for the vessel
- ASME U-1 or U-2, API 620 form or other applicable report
● Date vessel was placed in service
● Interruption dates if not in continuous service.

DESIGN AND CONSTRUCTION INFORMATION.

Information that will identify the code or standard used for the design and construction of the vessel or tank
and the specific design values, materials, fabrication methods, and inspection methods used include the
following items:

● Design code
- ASME Code Section and Division, API Standard or other design code used
● Type of construction
- Shop or field fabricated or other fabrication method
● VIII, division 1 or 2 vessels
- Maximum allowable pressure and temperature
- Minimum design temperature
● API 620 vessels
- Design pressure at top and maximum fill
● Additional requirements included, such as
- Appendix Q (Low-Pressure Storage Tanks For Liquified Hydrocarbon Gases) and
- Appendix R (Low-Pressure Storage Tanks for Refrigerated Products)
● Other design code vessels
- Maximum design and allowable pressures
- Maximum and minimum operating temperatures
● Vessel materials
- ASME, ASTM, or other specification names and numbers for the major parts
● Design corrosion allowance
● Thermal stress relief (PWHT, postweld heat treatment)
- Design code requirements
- Type, extent, and conditions of PWHT performed
● Nondestructive examination (NDE) of welds
- Type and extent of examination performed
- Time when NDE was performed (before or after PWHT or hydrotest)

SERVICE HISTORY.

Information on the conditions of operating history of the vessel or tank that will be helpful in safety
assessment include the following items:

● Fluids handled
- Type and composition, temperature and pressures
● Type of service
- Continuous, intermittent or irregular
● Significant changes in service conditions
- Changes in pressures, temperatures, and fluid compositions and the dates of the changes
● Vessel history
- Alterations, reratings, and repairs performed
- Date(s) of changes or repairs

IN-SERVICE INSPECTION.

Information about inspections performed on the vessel or tank and the results obtained that will assist in
the safety assessment include the following items:

● Inspection(s) performed
- Type, extent, and dates
● Examination methods
- Preparation of surfaces and welds
- Techniques used (visual, magnetic particle, penetrant test, radiography, ultrasonic)
● Qualifications of personnel

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- ASNT (American Society for Nondestructive Testing) levels or equivalent of examining

and supervisory personnel
● Inspection results and report
- Report form used (NBIC NB-7, API 510 or other)
- Summary of type and extent of damage or cracking
- Disposition (no action, delayed action or repaired)

SPECIFIC APPLICATIONS.

Survey results indicate that a relatively high proportion of vessels in operations in several specific
applications have experienced in service-related damage and cracking. Information on the following items
can assist in assessing the safety of vessels in these applications:

● Service application
- Deaerator, amine, wet hydrogen sulfide, ammonia, or pulp digesting
● Industry bulletins and guidelines for this application
- Owner/operator awareness of information
● Type, extent, and results of examinations
- Procedures, guidelines and recommendations used
- Amount of damage and cracking
- Next examination schedule
● Participation in industry survey for this application
● Problem mitigation
- Written plans and actions

EVALUATION OF INFORMATION.

The information acquired for the above items is not adaptable to any kind of numerical ranking for
quantitative safety assessment purposes. However, the information can reveal the owner or user's apparent
attention to good practice, careful operation, regular maintenance, and adherence to the recommendations
and guidelines developed for susceptible applications. If the assessment indicated cracking and other serious
damage problems, it is important that the inspector obtain qualified technical advice and opinion.

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SECTION IV: CHAPTER 2

PETROLEUM REFINING PROCESSES

Contents:

I. Introduction
II. Overview of the Petroleum Industry
III. Petroleum Refining Operations
IV.
Description of Petroleum Refining Processes and
V. Related Health and Safety Considerations
VI. Other Refinery Operations
Bibliography

Appendix IV:2-1. Glossary

1. INTRODUCTION.

1. The petroleum industry began with the successful drilling of the first commercial oil well in 1859, and the opening of the
first refinery two years later to process the crude into kerosene. The evolution of petroleum refining from simple
distillation to today's sophisticated processes has created a need for health and safety management procedures and safe
work practices. To those unfamiliar with the industry, petroleum refineries may appear to be complex and confusing
places. Refining is the processing of one complex mixture of hydrocarbons into a number of other complex mixtures of
hydrocarbons. The safe and orderly processing of crude oil into flammable gases and liquids at high temperatures and
pressures using vessels, equipment, and piping subjected to stress and corrosion requires considerable knowledge,
control, and expertise.

2. Safety and health professionals, working with process, chemical, instrumentation, and metallurgical engineers, assure
that potential physical, mechanical, chemical, and health hazards are recognized and provisions are made for safe
operating practices and appropriate protective measures. These measures may include hard hats, safety glasses and
goggles, safety shoes, hearing protection, respiratory protection, and protective clothing such as fire resistant clothing
where required. In addition, procedures should be established to assure compliance with applicable regulations and
standards such as hazard communications, confined space entry, and process safety management.

3. This chapter of the technical manual covers the history of refinery processing, characteristics of crude oil, hydrocarbon
types and chemistry, and major refinery products and by-products. It presents information on technology as normally
practiced in present operations. It describes the more common refinery processes and includes relevant safety and
health information. Additional information covers refinery utilities and miscellaneous supporting activities related to
hydrocarbon processing. Field personnel will learn what to expect in various facilities regarding typical materials and
process methods, equipment, potential hazards, and exposures.

4. The information presented refers to fire prevention, industrial hygiene, and safe work practices, and is not intended to
provide comprehensive guidelines for protective measures and/or compliance with regulatory requirements. As some of
the terminology is industry-specific, a glossary is provided as an appendix. This chapter does not cover petrochemical
processing.

2. OVERVIEW OF THE PETROLEUM INDUSTRY.

1. BASIC REFINERY PROCESS: DESCRIPTION AND HISTORY. Petroleum refining has evolved continuously in response
to changing consumer demand for better and different products. The original requirement was to produce kerosene as a

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cheaper and better source of light than whale oil. The development of the internal combustion engine led to the
production of gasoline and diesel fuels. The evolution of the airplane created a need first for high-octane aviation
gasoline and then for jet fuel, a sophisticated form of the original product, kerosene. Present-day refineries produce a
variety of products including many required as feedstock for the petrochemical industry.

1. Distillation Processes. The first refinery, opened in 1861, produced kerosene by simple atmospheric distillation.
Its by-products included tar and naphtha. It was soon discovered that high-quality lubricating oils could be
produced by distilling petroleum under vacuum. However, for the next 30 years kerosene was the product
consumers wanted. Two significant events changed this situation: (1) invention of the electric light decreased the
demand for kerosene, and (2) invention of the internal combustion engine created a demand for diesel fuel and
gasoline (naphtha).

2. Thermal Cracking Processes. With the advent of mass production and World War I, the number of gasoline-
powered vehicles increased dramatically and the demand for gasoline grew accordingly. However, distillation
processes produced only a certain amount of gasoline from crude oil. In 1913, the thermal cracking process was
developed, which subjected heavy fuels to both pressure and intense heat, physically breaking the large
molecules into smaller ones to produce additional gasoline and distillate fuels. Visbreaking, another form of
thermal cracking, was developed in the late 1930's to produce more desirable and valuable products.

3. Catalytic Processes. Higher-compression gasoline engines required higher-octane gasoline with better
antiknock characteristics. The introduction of catalytic cracking and polymerization processes in the mid- to late
1930's met the demand by providing improved gasoline yields and higher octane numbers.

Alkylation, another catalytic process developed in the early 1940's, produced more high-octane aviation gasoline
and petrochemical feedstock for explosives and synthetic rubber. Subsequently, catalytic isomerization was
developed to convert hydrocarbons to produce increased quantities of alkylation feedstock. Improved catalysts
and process methods such as hydrocracking and reforming were developed throughout the 1960's to increase
gasoline yields and improve antiknock characteristics. These catalytic processes also produced hydrocarbon
molecules with a double bond (alkenes) and formed the basis of the modern petrochemical industry.

4. Treatment Processes. Throughout the history of refining, various treatment methods have been used to
remove nonhydrocarbons, impurities, and other constituents that adversely affect the properties of finished
products or reduce the efficiency of the conversion processes. Treating can involve chemical reaction and/or
physical separation. Typical examples of treating are chemical sweetening, acid treating, clay contacting, caustic
washing, hydrotreating, drying, solvent extraction, and solvent dewaxing. Sweetening compounds and acids
desulfurize crude oil before processing and treat products during and after processing.

Following the Second World War, various reforming processes improved gasoline quality and yield and produced
higher-quality products. Some of these involved the use of catalysts and/or hydrogen to change molecules and
remove sulfur. A number of the more commonly used treating and reforming processes are described in this
chapter of the manual.

TABLE IV: 2-1. HISTORY OF REFINING

Year Process name Purpose By-products, etc.

1862 Atmospheric distillation Produce kerosene Naphtha, tar, etc.

1870 Vacuum distillation Lubricants (original) Asphalt, residual
Cracking feedstocks (1930's) coker feedstocks
1913 Thermal cracking Increase gasoline Residual, bunker fuel
1916 Sweetening Reduce sulfur & odor Sulfur
1930 Thermal reforming Improve octane number Residual
1932 Hydrogenation Remove sulfur Sulfur
1932 Coking Produce gasoline basestocks Coke
1933 Solvent extraction Improve lubricant viscosity index Aromatics
1935 Solvent dewaxing Improve pour point Waxes
1935 Cat. polymerization Improve gasoline yield Petrochemical
& octane number feedstocks
1937 Catalytic cracking Higher octane gasoline Petrochemical
feedstocks
1939 Visbreaking Reduce viscosity Increased distillate,tar
1940 Alkylation Increase gasoline octane & yield High-octane aviation
gasoline

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1940 Isomerization Produce alkylation feedstock Naphtha

1942 Fluid catalytic cracking Increase gasoline yield & octane Petrochemical feedstocks
1950 Deasphalting Increase cracking feedstock Asphalt
1952 Catalytic reforming Convert low-quality naphtha Aromatics
1954 Hydrodesulfurization Remove sulfur Sulfur
1956 Inhibitor sweetening Remove mercaptan Disulfides
1957 Catalytic isomerization Convert to molecules with high Alkylation feedstocks
octane number
1960 Hydrocracking Improve quality and reduce Alkylation feedstocks
sulfur
1974 Catalytic dewaxing Improve pour point Wax
1975 Residual hydrocracking Increase gasoline yield from Heavy residuals
residual

2. BASICS OF CRUDE OIL.

1. Crude oils are complex mixtures containing many different hydrocarbon compounds that vary in appearance and
composition from one oil field to another. Crude oils range in consistency from water to tar-like solids, and in
color from clear to black. An "average" crude oil contains about 84% carbon, 14% hydrogen, 1%-3% sulfur, and
less than 1% each of nitrogen, oxygen, metals, and salts. Crude oils are generally classified as paraffinic,
naphthenic, or aromatic, based on the predominant proportion of similar hydrocarbon molecules. Mixed-base
crudes have varying amounts of each type of hydrocarbon. Refinery crude base stocks usually consist of mixtures
of two or more different crude oils.

2. Relatively simple crude oil assays are used to classify crude oils as paraffinic, naphthenic, aromatic, or mixed.
One assay method (United States Bureau of Mines) is based on distillation, and another method (UOP "K" factor)
is based on gravity and boiling points. More comprehensive crude assays determine the value of the crude (i.e.,
its yield and quality of useful products) and processing parameters. Crude oils are usually grouped according to
yield structure.

3. Crude oils are also defined in terms of API (American Petroleum Institute) gravity. The higher the API gravity, the
lighter the crude. For example, light crude oils have high API gravities and low specific gravities. Crude oils with
low carbon, high hydrogen, and high API gravity are usually rich in paraffins and tend to yield greater proportions
of gasoline and light petroleum products; those with high carbon, low hydrogen, and low API gravities are usually
rich in aromatics.

4. Crude oils that contain appreciable quantities of hydrogen sulfide or other reactive sulfur compounds are called
"sour." Those with less sulfur are called "sweet." Some exceptions to this rule are West Texas crudes, which are
always considered "sour" regardless of their H2S content, and Arabian high-sulfur crudes, which are not
considered "sour" because their sulfur compounds are not highly reactive.

TABLE IV: 2-2. TYPICAL APPROXIMATE CHARACTERISTICS AND

PROPERTIES AND GASOLINE POTENTIAL OF VARIOUS CRUDES
(Representative average numbers)

Paraffins Aromatics Naphthenes Sulfur API gravity Napht. yield Octane no

Crude source (% vol) (% vol) (% vol) (% wt) (approx.) (% vol) (typical)

Nigerian 37 9 54 0.2 36 28 60
-Light

Saudi 63 19 18 2 34 22 40
-Light

Saudi 60 15 25 2.1 28 23 35
-Heavy

Venezuela 35 12 53 2.3 30 2 60
-Heavy

Venezuela 52 14 34 1.5 24 18 50
-Light

USA - - - 0.4 40 - -
-Midcont. Sweet

USA 46 22 32 1.9 32 33 55
-W. Texas Sour

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North Sea 50 16 34 0.4 37 31 50

-Brent

3. BASICS OF HYDROCARBON CHEMISTRY. Crude oil is a mixture of hydrocarbon molecules, which are organic
compounds of carbon and hydrogen atoms that may include from one to 60 carbon atoms. The properties of
hydrocarbons depend on the number and arrangement of the carbon and hydrogen atoms in the molecules. The simplest
hydrocarbon molecule is one carbon atom linked with four hydrogen atoms: methane. All other variations of petroleum
hydrocarbons evolve from this molecule.

Hydrocarbons containing up to four carbon atoms are usually gases, those with 5 to 19 carbon atoms are usually liquids,
and those with 20 or more are solids. The refining process uses chemicals, catalysts, heat, and pressure to separate and
combine the basic types of hydrocarbon molecules naturally found in crude oil into groups of similar molecules. The
refining process also rearranges their structures and bonding patterns into different hydrocarbon molecules and
compounds. Therefore it is the type of hydrocarbon (paraffinic, naphthenic, or aromatic) rather than its specific chemical
compounds that is significant in the refining process.

1. Three Principal Groups or Series of Hydrocarbon Compounds that Occur Naturally in Crude Oil.

a. Paraffins. The paraffinic series of hydrocarbon compounds found in crude oil have the general formula CnH2n
+2 and can be either straight chains (normal) or branched chains (isomers) of carbon atoms. The lighter, straight-
chain paraffin molecules are found in gases and paraffin waxes. Examples of straight-chain molecules are
methane, ethane, propane, and butane (gases containing from one to four carbon atoms), and pentane and
hexane (liquids with five to six carbon atoms). The branched-chain (isomer) paraffins are usually found in heavier
fractions of crude oil and have higher octane numbers than normal paraffins. These compounds are saturated
hydrocarbons, with all carbon bonds satisfied, that is, the hydrocarbon chain carries the full complement of
hydrogen atoms.

FIGURE IV:2-1. TYPICAL PARAFFINS.

Example of simplest Examples of straight chain paraffin molecule (Butane) and branched
HC molecule (CH4): paraffin molecule (Isobutane) with same chemical formula (C4H10):

METHANE (CH4) BUTANE (C4H10) ISOBUTANE (C4H10)

b. Aromatics are unsaturated ring-type (cyclic) compounds which react readily because they have carbon atoms
that are deficient in hydrogen. All aromatics have at least one benzene ring (a single-ring compound
characterized by three double bonds alternating with three single bonds between six carbon atoms) as part of
their molecular structure. Naphthalenes are fused double-ring aromatic compounds. The most complex aromatics,
polynuclears (three or more fused aromatic rings), are found in heavier fractions of crude oil.

c. Naphthenes are saturated hydrocarbon groupings with the general formula CnH2n, arranged in the form of
closed rings (cyclic) and found in all fractions of crude oil except the very lightest. Single-ring naphthenes
(monocycloparaffins) with five and six carbon atoms predominate, with two-ring naphthenes (dicycloparaffins)
found in the heavier ends of naphtha.

2. Other Hydrocarbons.

a. Alkenes are mono-olefins with the general formula CnH2n and contain only one carbon-carbon double bond in
the chain. The simplest alkene is ethylene, with two carbon atoms joined by a double bond and four hydrogen
atoms. Olefins are usually formed by thermal and catalytic cracking and rarely occur naturally in unprocessed
crude oil.

FIGURE IV:2-2. TYPICAL AROMATICS.

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Example of simple aromatic compound: Examples of simple double-ring aromatic compound:

BENZENE (C6H6) NAPTHALENE (C10H8)

FIGURE IV:2-3. TYPICAL NAPHTHENES.

Example of typical single-ring Examples of naphthene with same chemical formula

naphthene: (C6H12) but different molecular structure:

CYCLOHEXANE (C6H12) METHYL CYCLOPENTANE (C6H12)

FIGURE IV:2-4. TYPICAL ALKENES.

Simplest Alkene (C2H4): Typical Alkenes with the same chemical formula (C4H8) but different
molecular structures:

ETHYLENE (C2H4) 1-BUTENE (C4H8) ISOBUTENE (C4H8)

b. Dienes and Alkynes. Dienes, also known as diolefins, have two carbon-carbon double bonds. The alkynes,
another class of unsaturated hydrocarbons, have a carbon-carbon triple bond within the molecule. Both these
series of hydrocarbons have the general formula CnH2n-2. Diolefins such as 1,2-butadiene and 1,3-butadiene, and
alkynes such as acetylene, occur in C5 and lighter fractions from cracking. The olefins, diolefins, and alkynes are

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said to be unsaturated because they contain less than the amount of hydrogen necessary to saturate all the
valences of the carbon atoms. These compounds are more reactive than paraffins or naphthenes and readily
combine with other elements such as hydrogen, chlorine, and bromine.

FIGURE IV:2-5. TYPICAL DIOLEFINS AND ALKYNES.

Simplest Alkyne: Typical Diolefins with the same chemical formula (C4H6) but different
(C2H2): molecular structures:

ACETYLENE (C2H2) 1,2-BUTADIENE (C4H6) 1,3-BUTADIENE (C4H6)

3. Nonhydrocarbons.

a. Sulfur Compounds. Sulfur may be present in crude oil as hydrogen sulfide (H2S), as compounds (e.g.
mercaptans, sulfides, disulfides, thiophenes, etc.) or as elemental sulfur. Each crude oil has different amounts
and types of sulfur compounds, but as a rule the proportion, stability, and complexity of the compounds are
greater in heavier crude-oil fractions. Hydrogen sulfide is a primary contributor to corrosion in refinery processing
units. Other corrosive substances are elemental sulfur and mercaptans. Moreover, the corrosive sulfur
compounds have an obnoxious odor.

Pyrophoric iron sulfide results from the corrosive action of sulfur compounds on the iron and steel used in refinery
process equipment, piping, and tanks. The combustion of petroleum products containing sulfur compounds
produces undesirables such as sulfuric acid and sulfur dioxide. Catalytic hydrotreating processes such as
hydrodesulfurization remove sulfur compounds from refinery product streams. Sweetening processes either
remove the obnoxious sulfur compounds or convert them to odorless disulfides, as in the case of mercaptans.

b. Oxygen Compounds. Oxygen compounds such as phenols, ketones, and carboxylic acids occur in crude oils
in varying amounts.

c. Nitrogen Compounds. Nitrogen is found in lighter fractions of crude oil as basic compounds, and more often
in heavier fractions of crude oil as nonbasic compounds that may also include trace metals such as copper,
vanadium, and/or nickel. Nitrogen oxides can form in process furnaces. The decomposition of nitrogen
compounds in catalytic cracking and hydrocracking processes forms ammonia and cyanides that can cause
corrosion.

d. Trace Metals. Metals, including nickel, iron, and vanadium are often found in crude oils in small quantities
and are removed during the refining process. Burning heavy fuel oils in refinery furnaces and boilers can leave
deposits of vanadium oxide and nickel oxide in furnace boxes, ducts, and tubes. It is also desirable to remove
trace amounts of arsenic, vanadium, and nickel prior to processing as they can poison certain catalysts.

e. Salts. Crude oils often contain inorganic salts such as sodium chloride, magnesium chloride, and calcium
chloride in suspension or dissolved in entrained water (brine). These salts must be removed or neutralized before
processing to prevent catalyst poisoning, equipment corrosion, and fouling. Salt corrosion is caused by the
hydrolysis of some metal chlorides to hydrogen chloride (HCl) and the subsequent formation of hydrochloric acid
when crude is heated. Hydrogen chloride may also combine with ammonia to form ammonium chloride (NH4Cl),
which causes fouling and corrosion.

f. Carbon Dioxide. Carbon dioxide may result from the decomposition of bicarbonates present in or added to
crude, or from steam used in the distillation process.

g. Naphthenic Acids. Some crude oils contain naphthenic (organic) acids, which may become corrosive at
temperatures above 450° F when the acid value of the crude is above a certain level.

4. MAJOR REFINERY PRODUCTS.

1. Gasoline. The most important refinery product is motor gasoline, a blend of hydrocarbons with boiling ranges
from ambient temperatures to about 400 °F. The important qualities for gasoline are octane number (antiknock),
volatility (starting and vapor lock), and vapor pressure (environmental control). Additives are often used to
enhance performance and provide protection against oxidation and rust formation.

2. Kerosene. Kerosene is a refined middle-distillate petroleum product that finds considerable use as a jet fuel and
around the world in cooking and space heating. When used as a jet fuel, some of the critical qualities are freeze
point, flash point, and smoke point. Commercial jet fuel has a boiling range of about 375°-525° F, and military

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jet fuel 130°-550° F. Kerosene, with less-critical specifications, is used for lighting, heating, solvents, and
blending into diesel fuel.

3. Liquified Petroleum Gas (LPG). LPG, which consists principally of propane and butane, is produced for use as
fuel and is an intermediate material in the manufacture of petrochemicals. The important specifications for proper
performance include vapor pressure and control of contaminants.

4. Distillate Fuels. Diesel fuels and domestic heating oils have boiling ranges of about 400°-700° F. The desirable
qualities required for distillate fuels include controlled flash and pour points, clean burning, no deposit formation
in storage tanks, and a proper diesel fuel cetane rating for good starting and combustion.

5. Residual Fuels. Many marine vessels, power plants, commercial buildings and industrial facilities use residual
fuels or combinations of residual and distillate fuels for heating and processing. The two most critical
specifications of residual fuels are viscosity and low sulfur content for environmental control.

6. Coke and Asphalt. Coke is almost pure carbon with a variety of uses from electrodes to charcoal briquets.
Asphalt, used for roads and roofing materials, must be inert to most chemicals and weather conditions.

7. Solvents. A variety of products, whose boiling points and hydrocarbon composition are closely controlled, are
produced for use as solvents. These include benzene, toluene, and xylene.

8. Petrochemicals. Many products derived from crude oil refining, such as ethylene, propylene, butylene, and
isobutylene, are primarily intended for use as petrochemical feedstock in the production of plastics, synthetic
fibers, synthetic rubbers, and other products.

9. Lubricants. Special refining processes produce lubricating oil base stocks. Additives such as demulsifiers,
antioxidants, and viscosity improvers are blended into the base stocks to provide the characteristics required for
motor oils, industrial greases, lubricants, and cutting oils. The most critical quality for lubricating-oil base stock is
a high viscosity index, which provides for greater consistency under varying temperatures.

5. COMMON REFINERY CHEMICALS.

1. Leaded Gasoline Additives. Tetraethyl lead (TEL) and tetramethyl lead (TML) are additives formerly used to
improve gasoline octane ratings but are no longer in common use except in aviation gasoline.

2. Oxygenates. Ethyl tertiary butyl ether (ETBE), methyl tertiary butyl ether (MTBE), tertiary amyl methyl ether
(TAME), and other oxygenates improve gasoline octane ratings and reduce carbon monoxide emissions.

3. Caustics. Caustics are added to desalting water to neutralize acids and reduce corrosion. They are also added to
desalted crude in order to reduce the amount of corrosive chlorides in the tower overheads. They are used in
some refinery treating processes to remove contaminants from hydrocarbon streams.

4. Sulfuric Acid and Hydrofluoric Acid. Sulfuric acid and hydrofluoric acid are used primarily as catalysts in
alkylation processes. Sulfuric acid is also used in some treatment processes.

3. PETROLEUM REFINING OPERATIONS.

1. INTRODUCTION. Petroleum refining begins with the distillation, or fractionation, of crude oils into separate
hydrocarbon groups. The resultant products are directly related to the characteristics of the crude processed. Most
distillation products are further converted into more usable products by changing the size and structure of the
hydrocarbon molecules through cracking, reforming, and other conversion processes as discussed in this chapter. These
converted products are then subjected to various treatment and separation processes such as extraction, hydrotreating,
and sweetening to remove undesirable constituents and improve product quality. Integrated refineries incorporate
fractionation, conversion, treatment, and blending operations and may also include petrochemical processing.

2. REFINING OPERATIONS. Petroleum refining processes and operations can be separated into five basic areas:

1. Fractionation (distillation) is the separation of crude oil in atmospheric and vacuum distillation towers into
groups of hydrocarbon compounds of differing boiling-point ranges called "fractions" or "cuts."

2. Conversion processes change the size and/or structure of hydrocarbon molecules. These processes include:

■ Decomposition (dividing) by thermal and catalytic cracking;

■ Unification (combining) through alkylation and polymerization; and
■ Alteration (rearranging) with isomerization and catalytic reforming.

3. Treatment processes are intended to prepare hydrocarbon streams for additional processing and to prepare
finished products. Treatment may include the removal or separation of aromatics and naphthenes as well as
impurities and undesirable contaminants. Treatment may involve chemical or physical separation such as
dissolving, absorption, or precipitation using a variety and combination of processes including desalting, drying,

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hydrodesulfurizing, solvent refining, sweetening, solvent extraction, and solvent dewaxing.

4. Formulating and Blending is the process of mixing and combining hydrocarbon fractions, additives, and other
components to produce finished products with specific performance properties.

5. Other Refining Operations include: light-ends recovery; sour-water stripping; solid waste and wastewater
treatment; process-water treatment and cooling; storage and handling; product movement; hydrogen
production; acid and tail-gas treatment; and sulfur recovery.

Auxiliary operations and facilities include: steam and power generation; process and fire water systems; flares
and relief systems; furnaces and heaters; pumps and valves; supply of steam, air, nitrogen, and other plant
gases; alarms and sensors; noise and pollution controls; sampling, testing, and inspecting; and laboratory,
control room, maintenance, and administrative facilities.

FIGURE IV:2-6. REFINERY PROCESS CHART.

TABLE IV:2-3. OVERVIEW OF PETROLEUM REFINING PROCESSES.

Process name Action Method Purpose Feedstock(s) Product(s)

FRACTIONATION PROCESSES
Atmospheric distillation Separation Thermal Separate fractions Desalted crude oil Gas, gas oil, distillate, residual
Vacuum distillation Separation Thermal Separate w/o cracking Atmospheric tower residual Gas oil, lube stock, residual
CONVERSION PROCESSED--DECOMPOSITION
Catalytic cracking Alteration Catalytic Upgrade gasoline Gas oil, coke distillate Gasoline, petrochemical
feedstock
Coking Polymerize Thermal Convert vacuum residuals Gas oil, coke distillate Gasoline, petrochemical
feedstock
Hydro-cracking Hydrogenate Catalytic Convert to lighter HC's Gas oil, cracked oil, residual Lighter, higher-quality products
*Hydrogen steam reforming Decompose Thermal/ Produce hydrogen Desulfurized gas, O2, steam Hydrogen, CO, CO2
catalytic
*Steam cracking Decompose Thermal Crack large molecules Atm tower hvy fuel/ distillate Cracked naphtha, coke,
residual

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Visbreaking Decompose Thermal Reduce viscosity Atmospheric tower residual Distillate, tar
CONVERSION PROCESSES--UNIFICATION
Alkylation Combining Catalytic Unite olefins & isoparaffins Tower isobutane/ cracker olefin Iso-octane (alkylate)
Grease compounding Combining Thermal Combine soaps & oils Lube oil, fatty acid, alky metal Lubricating grease
Polymerizing Polymerize Catalytic Unite 2 or more olefins Cracker olefins High-octane naphtha,
petrochemical stocks
CONVERSION PROCESSES--ALTERATION OR REARRANGEMENT
Catalytic reforming Alteration/ Catalytic Upgrade low-octane naphtha Coker/ hydro-cracker naphtha High oct. Reformate/ aromatic
dehydration
Isomerization Rearrange Catalytic Convert straight chain to Butane, pentane, hexane Isobutane/ pentane/ hexane
branch
TREATMENT PROCESSES
*Amine treating Treatment Absorption Remove acidic contaminants Sour gas, HCs w/CO2 & H2S Acid free gases & liquid HCs

Desalting Dehydration Absorption Remove contaminants Crude oil Desalted crude oil
Drying & sweetening Treatment Abspt/ therm Remove H2O & sulfur cmpds Liq Hcs, LPG, alky feedstk Sweet & dry hydrocarbons

*Furfural extraction Solvent extr. Absorption Upgrade mid distillate & lubes Cycle oils & lube feed-stocks High quality diesel & lube oil
Hydrodesulfurization Treatment Catalytic Remove sulfur, contaminants High-sulfur residual/ gas oil Desulfurized olefins
Hydrotreating Hydrogenation Catalytic Remove impurities, saturate Residuals, cracked HC's Cracker feed, distillate, lube
HC's
*Phenol extraction Solvent extr. Abspt/ therm Improve visc. index, color Lube oil base stocks High quality lube oils
Solvent deasphalting Treatment Absorption Remove asphalt Vac. tower residual, propane Heavy lube oil, asphalt
Solvent dewaxing Treatment Cool/ filter Remove wax from lube stocks Vac. tower lube oils Dewaxed lube basestock
Solvent extraction Solvent extr. Abspt/ precip. Separate unsat. oils Gas oil, reformate, distillate High-octane gasoline
Sweetening Treatment Catalytic Remv H2S, convert mercaptan Untreated distillate/gasoline High-quality distillate/gasoline

* Note: These processes are not depicted in the refinery process flow chart.

4. DESCRIPTION OF PETROLEUM REFINING PROCESSES AND RELATED HEALTH AND SAFETY CONSIDERATIONS.

1. CRUDE OIL PRETREATMENT (DESALTING).

1. Description.

a. Crude oil often contains water, inorganic salts, suspended solids, and water-soluble trace metals. As a first
step in the refining process, to reduce corrosion, plugging, and fouling of equipment and to prevent poisoning the
catalysts in processing units, these contaminants must be removed by desalting (dehydration).

b. The two most typical methods of crude-oil desalting, chemical and electrostatic separation, use hot water as
the extraction agent. In chemical desalting, water and chemical surfactant (demulsifiers) are added to the crude,
heated so that salts and other impurities dissolve into the water or attach to the water, and then held in a tank
where they settle out. Electrical desalting is the application of high-voltage electrostatic charges to concentrate
suspended water globules in the bottom of the settling tank. Surfactants are added only when the crude has a
large amount of suspended solids. Both methods of desalting are continuous. A third and less-common process
involves filtering heated crude using diatomaceous earth.

c. The feedstock crude oil is heated to between 150° and 350°F to reduce viscosity and surface tension for easier
mixing and separation of the water. The temperature is limited by the vapor pressure of the crude-oil feedstock.
In both methods other chemicals may be added. Ammonia is often used to reduce corrosion. Caustic or acid may
be added to adjust the pH of the water wash. Wastewater and contaminants are discharged from the bottom of
the settling tank to the wastewater treatment facility. The desalted crude is continuously drawn from the top of
the settling tanks and sent to the crude distillation (fractionating) tower.

TABLE IV:2-4. DESALTING PROCESS.

Feedstock From Process Typical products . . . To

Crude Storage Treating Desalted crude . . . Atmospheric distillation tower

Waste water . . . . . Treatment

FIGURE IV:2-7. ELECTROSTAITC DESALTING.

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2. Health and Safety Considerations.

a. Fire Prevention and Protection. The potential exists for a fire due to a leak or release of crude from
heaters in the crude desalting unit. Low boiling point components of crude may also be released if a leak occurs.

b. Safety. Inadequate desalting can cause fouling of heater tubes and heat exchangers throughout the refinery.
Fouling restricts product flow and heat transfer and leads to failures due to increased pressures and
temperatures. Corrosion, which occurs due to the presence of hydrogen sulfide, hydrogen chloride, naphthenic
(organic) acids, and other contaminants in the crude oil, also causes equipment failure. Neutralized salts
(ammonium chlorides and sulfides), when moistened by condensed water, can cause corrosion. Overpressuring
the unit is another potential hazard that causes failures.

c. Health. Because this is a closed process, there is little potential for exposure to crude oil unless a leak or
release occurs. Where elevated operating temperatures are used when desalting sour crudes, hydrogen sulfide
will be present. There is the possibility of exposure to ammonia, dry chemical demulsifiers, caustics, and/or acids
during this operation. Safe work practices and/or the use of appropriate personal protective equipment may be
needed for exposures to chemicals and other hazards such as heat, and during process sampling, inspection,
maintenance, and turnaround activities.

Depending on the crude feedstock and the treatment chemicals used, the wastewater will contain varying
amounts of chlorides, sulfides, bicarbonates, ammonia, hydrocarbons, phenol, and suspended solids. If
diatomaceous earth is used in filtration, exposures should be minimized or controlled. Diatomaceous earth can
contain silica in very fine particle size, making this a potential respiratory hazard.

2. CRUDE OIL DISTILLATION (FRACTIONATION).

1. Description. The first step in the refining process is the separation of crude oil into various fractions or straight-
run cuts by distillation in atmospheric and vacuum towers. The main fractions or "cuts" obtained have specific
boiling-point ranges and can be classified in order of decreasing volatility into gases, light distillates, middle
distillates, gas oils, and residuum.

2. Atmospheric Distillation Tower.

a. At the refinery, the desalted crude feedstock is preheated using recovered process heat. The feedstock then
flows to a direct-fired crude charge heater where it is fed into the vertical distillation column just above the
bottom, at pressures slightly above atmospheric and at temperatures ranging from 650° to 700° F (heating crude
oil above these temperatures may cause undesirable thermal cracking). All but the heaviest fractions flash into
vapor. As the hot vapor rises in the tower, its temperature is reduced. Heavy fuel oil or asphalt residue is taken
from the bottom. At successively higher points on the tower, the various major products including lubricating oil,
heating oil, kerosene, gasoline, and uncondensed gases (which condense at lower temperatures) are drawn off.

b. The fractionating tower, a steel cylinder about 120 feet high, contains horizontal steel trays for separating and
collecting the liquids. At each tray, vapors from below enter perforations and bubble caps. They permit the vapors
to bubble through the liquid on the tray, causing some condensation at the temperature of that tray. An overflow
pipe drains the condensed liquids from each tray back to the tray below, where the higher temperature causes re-
evaporation. The evaporation, condensing, and scrubbing operation is repeated many times until the desired
degree of product purity is reached. Then side streams from certain trays are taken off to obtain the desired
fractions. Products ranging from uncondensed fixed gases at the top to heavy fuel oils at the bottom can be taken
continuously from a fractionating tower. Steam is often used in towers to lower the vapor pressure and create a
partial vacuum. The distillation process separates the major constituents of crude oil into so-called straight-run
products. Sometimes crude oil is "topped" by distilling off only the lighter fractions, leaving a heavy residue that
is often distilled further under high vacuum.

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TABLE IV:2-5. ATMOSPHERIC DISTILLATION PROCESS

Feedstock From Process Typical products . . . . . . To

Crude Desalting Separation Gases . . . . . . . . . . . Atmospheric distillation tower
Naphthas. . . . . . . . . . . . Reforming or treating
Kerosene or distillates . . Treating
Gas oil . . . . . . . . . . . . . Catalytic cracking
Residual . . . . . . . . . Vacuum tower or visbreaker

FIGURE IV:2-8. ATMOSPHERIC DISTILLATION.

3. Vacuum Distillation Tower. In order to further distill the residuum or topped crude from the atmospheric tower
at higher temperatures, reduced pressure is required to prevent thermal cracking. The process takes place in one
or more vacuum distillation towers. The principles of vacuum distillation resemble those of fractional distillation
and, except that larger-diameter columns are used to maintain comparable vapor velocities at the reduced
pressures, the equipment is also similar. The internal designs of some vacuum towers are different from
atmospheric towers in that random packing and demister pads are used instead of trays. A typical first-phase
vacuum tower may produce gas oils, lubricating-oil base stocks, and heavy residual for propane deasphalting. A
second-phase tower operating at lower vacuum may distill surplus residuum from the atmospheric tower, which is
not used for lube-stock processing, and surplus residuum from the first vacuum tower not used for deasphalting.
Vacuum towers are typically used to separate catalytic cracking feedstock from surplus residuum.

4. Other Distillation Towers (Columns). Within refineries there are numerous other, smaller distillation towers
called columns, designed to separate specific and unique products. Columns all work on the same principles as
the towers described above. For example, a depropanizer is a small column designed to separate propane and
lighter gases from butane and heavier components. Another larger column is used to separate ethyl benzene and
xylene. Small "bubble" towers called strippers use steam to remove trace amounts of light products from heavier
product streams.

5. Health and Safety Considerations.

a. Fire Prevention and Protection. Even though these are closed processes, heaters and exchangers in the
atmospheric and vacuum distillation units could provide a source of ignition, and the potential for a fire exists
should a leak or release occur.

b. Safety. An excursion in pressure, temperature, or liquid levels may occur if automatic control devices fail.
Control of temperature, pressure, and reflux within operating parameters is needed to prevent thermal cracking
within the distillation towers. Relief systems should be provided for overpressure and operations monitored to
prevent crude from entering the reformer charge.

The sections of the process susceptible to corrosion include (but may not be limited to) preheat exchanger (HCl
and H2S), preheat furnace and bottoms exchanger (H2S and sulfur compounds), atmospheric tower and vacuum

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furnace (H2S, sulfur compounds, and organic acids), vacuum tower (H2S and organic acids), and overhead (H2S,
HCl, and water). Where sour crudes are processed, severe corrosion can occur in furnace tubing and in both
atmospheric and vacuum towers where metal temperatures exceed 450° F. Wet H2S also will cause cracks in
steel. When processing high-nitrogen crudes, nitrogen oxides can form in the flue gases of furnaces. Nitrogen
oxides are corrosive to steel when cooled to low temperatures in the presence of water.

Chemicals are used to control corrosion by hydrochloric acid produced in distillation units. Ammonia may be
injected into the overhead stream prior to initial condensation and/or an alkaline solution may be carefully
injected into the hot crude-oil feed. If sufficient wash-water is not injected, deposits of ammonium chloride can
form and cause serious corrosion. Crude feedstock may contain appreciable amounts of water in suspension
which can separate during startup and, along with water remaining in the tower from steam purging, settle in the
bottom of the tower. This water can be heated to the boiling point and create an instantaneous vaporization
explosion upon contact with the oil in the unit.

c. Health. Atmospheric and vacuum distillation are closed processes and exposures are expected to be minimal.
When sour (high-sulfur) crudes are processed, there is potential for exposure to hydrogen sulfide in the preheat
exchanger and furnace, tower flash zone and overhead system, vacuum furnace and tower, and bottoms
exchanger. Hydrogen chloride may be present in the preheat exchanger, tower top zones, and overheads.
Wastewater may contain water-soluble sulfides in high concentrations and other water-soluble compounds such
as ammonia, chlorides, phenol, mercaptans, etc., depending upon the crude feedstock and the treatment
chemicals. Safe work practices and/or the use of appropriate personal protective equipment may be needed for
exposures to chemicals and other hazards such as heat and noise, and during sampling, inspection, maintenance,
and turnaround activities.

TABLE IV:2-6. VACUUM DISTILLATION PROCESS

Feedstock From Process Typical products . . To

Residuals Atmospheric tower Separation Gas oils . . . . . . . . Catalytic cracker

Lubricants . . . Hydrotreating or solvent
Residual . . . Deasphalter, visbreaker, or coker

FIGURE IV:2-9. VACUUM DISTILLATION.

3. SOLVENT EXTRACTION AND DEWAXING.

1. Description. Solvent treating is a widely used method of refining lubricating oils as well as a host of other
refinery stocks. Since distillation (fractionation) separates petroleum products into groups only by their boiling-
point ranges, impurities may remain. These include organic compounds containing sulfur, nitrogen, and oxygen;
inorganic salts and dissolved metals; and soluble salts that were present in the crude feedstock. In addition,
kerosene and distillates may have trace amounts of aromatics and naphthenes, and lubricating oil base-stocks

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may contain wax. Solvent refining processes including solvent extraction and solvent dewaxing usually remove
these undesirables at intermediate refining stages or just before sending the product to storage.

2. Solvent Extraction.

a. The purpose of solvent extraction is to prevent corrosion, protect catalyst in subsequent processes, and
improve finished products by removing unsaturated, aromatic hydrocarbons from lubricant and grease stocks.
The solvent extraction process separates aromatics, naphthenes, and impurities from the product stream by
dissolving or precipitation. The feedstock is first dried and then treated using a continuous countercurrent solvent
treatment operation. In one type of process, the feedstock is washed with a liquid in which the substances to be
removed are more soluble than in the desired resultant product. In another process, selected solvents are added
to cause impurities to precipitate out of the product. In the adsorption process, highly porous solid materials
collect liquid molecules on their surfaces.

b. The solvent is separated from the product stream by heating, evaporation, or fractionation, and residual trace
amounts are subsequently removed from the raffinate by steam stripping or vacuum flashing. Electric
precipitation may be used for separation of inorganic compounds. The solvent is then regenerated to be used
again in the process.

c. The most widely used extraction solvents are phenol, furfural, and cresylic acid. Other solvents less frequently
used are liquid sulfur dioxide, nitrobenzene, and 2,2'-dichloroethyl ether. The selection of specific processes and
chemical agents depends on the nature of the feedstock being treated, the contaminants present, and the
finished product requirements.

TABLE IV:2-7. SOLVENT EXTRACTION PROCESS

Feedstock From Process Typical products . . . To

Naphthas, distillates, Atm. tower Treating/ High octane gasoline . . Storage

kerosene blending Refined fuels . . . . . . . . Treating and blending
Spent agents . . . . . . . . Treatment and blending

FIGURE IV:2-10. AROMATICS EXTRACTION.

3. Solvent Dewaxing. Solvent dewaxing is used to remove wax from either distillate or residual basestocks at any
stage in the refining process. There are several processes in use for solvent dewaxing, but all have the same
general steps, which are: (1) mixing the feedstock with a solvent, (2) precipitating the wax from the mixture by
chilling, and (3) recovering the solvent from the wax and dewaxed oil for recycling by distillation and steam
stripping. Usually two solvents are used: toluene, which dissolves the oil and maintains fluidity at low
temperatures, and methyl ethyl ketone (MEK), which dissolves little wax at low temperatures and acts as a wax
precipitating agent. Other solvents that are sometimes used include benzene, methyl isobutyl ketone, propane,
petroleum naphtha, ethylene dichloride, methylene chloride, and sulfur dioxide. In addition, there is a catalytic
process used as an alternate to solvent dewaxing.

TABLE IV:2-8. SOLVENT DEWAXING PROCESS

Feedstock From Process Typical products . . To

Lube basestock Vacuum tower Treating Dewaxed lubes . . . . . Hydrotreating
Wax . . . . . . . . . . . . . Hydrotreating
Spent agents . . . . . . Treatment or recycle

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FIGURE IV:2-11. SOLVENT DEWAXING.

Note: Diagrams in Figures IV:2-10, 11, 12, 13, 15, and 20 reproduced with permission
from Shell International Petroleum Company.

4. Health and Safety Considerations.

a. Fire Prevention and Protection. Solvent treatment is essentially a closed process and, although operating
pressures are relatively low, the potential exists for fire from a leak or spill contacting a source of ignition such as
the drier or extraction heater. In solvent dewaxing, disruption of the vacuum will create a potential fire hazard by
allowing air to enter the unit.

b. Health. Because solvent extraction is a closed process, exposures are expected to be minimal under normal
operating conditions. However, there is a potential for exposure to extraction solvents such as phenol, furfural,
glycols, methyl ethyl ketone, amines, and other process chemicals. Safe work practices and/or the use of
appropriate personal protective equipment may be needed for exposures to chemicals and other hazards such as
noise and heat, and during repair, inspection, maintenance, and turnaround activities.

4. THERMAL CRACKING.

1. Description.

a. Because the simple distillation of crude oil produces amounts and types of products that are not consistent
with those required by the marketplace, subsequent refinery processes change the product mix by altering the
molecular structure of the hydrocarbons. One of the ways of accomplishing this change is through "cracking," a
process that breaks or cracks the heavier, higher boiling-point petroleum fractions into more valuable products
such as gasoline, fuel oil, and gas oils. The two basic types of cracking are thermal cracking, using heat and
pressure, and catalytic cracking.

b. The first thermal cracking process was developed around 1913. Distillate fuels and heavy oils were heated
under pressure in large drums until they cracked into smaller molecules with better antiknock characteristics.
However, this method produced large amounts of solid, unwanted coke. This early process has evolved into the
following applications of thermal cracking: visbreaking, steam cracking, and coking.

2. Visbreaking Process. Visbreaking, a mild form of thermal cracking, significantly lowers the viscosity of heavy
crude-oil residue without affecting the boiling point range. Residual from the atmospheric distillation tower is
heated (800°-950° F) at atmospheric pressure and mildly cracked in a heater. It is then quenched with cool gas
oil to control overcracking, and flashed in a distillation tower. Visbreaking is used to reduce the pour point of
waxy residues and reduce the viscosity of residues used for blending with lighter fuel oils. Middle distillates may
also be produced, depending on product demand. The thermally cracked residue tar, which accumulates in the
bottom of the fractionation tower, is vacuum flashed in a stripper and the distillate recycled.

TABLE IV:2-9. VISBREAKING PROCESS.

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Feedstock From Process Typical products . . . . . To

Residual Atmospheric tower & Vacuum Decompose Gasoline or distillate . . Hydrotreating

tower Vapor . . . . . . . . . . . . Hydrotreater
Residue . . . . . . . . . . . Stripper or recycle
Gases . . . . . . . . . . . . Gas plant

FIGURE IV:2-12. VISBREAKING.

3. Steam Cracking Process. Steam cracking is a petrochemical process sometimes used in refineries to produce
olefinic raw materials (e.g., ethylene) from various feedstock for petrochemicals manufacture. The feedstock
range from ethane to vacuum gas oil, with heavier feeds giving higher yields of by-products such as naphtha. The
most common feeds are ethane, butane, and naphtha. Steam cracking is carried out at temperatures of 1,500°-
1,600° F, and at pressures slightly above atmospheric. Naphtha produced from steam cracking contains benzene,
which is extracted prior to hydrotreating. Residual from steam cracking is sometimes blended into heavy fuels.

4. Coking Processes. Coking is a severe method of thermal cracking used to upgrade heavy residuals into lighter
products or distillates. Coking produces straight-run gasoline (coker naphtha) and various middle-distillate
fractions used as catalytic cracking feedstock. The process so completely reduces hydrogen that the residue is a
form of carbon called "coke." The two most common processes are delayed coking and continuous (contact or
fluid) coking. Three typical types of coke are obtained (sponge coke, honeycomb coke, and needle coke)
depending upon the reaction mechanism, time, temperature, and the crude feedstock.

a. Delayed Coking. In delayed coking the heated charge (typically residuum from atmospheric distillation
towers) is transferred to large coke drums which provide the long residence time needed to allow the cracking
reactions to proceed to completion. Initially the heavy feedstock is fed to a furnace which heats the residuum to
high temperatures (900°-950° F) at low pressures (25-30 psi) and is designed and controlled to prevent
premature coking in the heater tubes. The mixture is passed from the heater to one or more coker drums where
the hot material is held approximately 24 hours (delayed) at pressures of 25-75 psi, until it cracks into lighter
products. Vapors from the drums are returned to a fractionator where gas, naphtha, and gas oils are separated
out. The heavier hydrocarbons produced in the fractionator are recycled through the furnace.

After the coke reaches a predetermined level in one drum, the flow is diverted to another drum to maintain
continuous operation. The full drum is steamed to strip out uncracked hydrocarbons, cooled by water injection,
and decoked by mechanical or hydraulic methods. The coke is mechanically removed by an auger rising from the
bottom of the drum. Hydraulic decoking consists of fracturing the coke bed with high-pressure water ejected from
a rotating cutter.

b. Continuous Coking. Continuous (contact or fluid) coking is a moving-bed process that operates at
temperatures higher than delayed coking. In continuous coking, thermal cracking occurs by using heat
transferred from hot, recycled coke particles to feedstock in a radial mixer, called a reactor, at a pressure of 50
psi. Gases and vapors are taken from the reactor, quenched to stop any further reaction, and fractionated. The
reacted coke enters a surge drum and is lifted to a feeder and classifier where the larger coke particles are
removed as product. The remaining coke is dropped into the preheater for recycling with feedstock. Coking
occurs both in the reactor and in the surge drum. The process is automatic in that there is a continuous flow of
coke and feedstock.

TABLE IV: 2-10. COKING PROCESSES.

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Feedstock From Process Typical products . . . To

Residual . . . . Atmospheric & Decompo- Naphtha, gasoline . . Distillation column,blending
vacuum sition
catalytic cracker
Clarified oil . . Catalytic cracker Coke . . . . . . . . . . . Shipping, recycle
Tars . . . . . . Various units Gas oil . . . . . . . . . . Catalytic cracking
Wasteater
(sour) . . . . . Treatment
Gases . . . . . Gas plant

FIGURE IV:2-13. DELAYED COKING.

5. Health and Safety Considerations.

a. Fire Protection and Prevention. Because thermal cracking is a closed process, the primary potential for fire
is from leaks or releases of liquids, gases, or vapors reaching an ignition source such as a heater. The potential
for fire is present in coking operations due to vapor or product leaks. Should coking temperatures get out of
control, an exothermic reaction could occur within the coker.

b. Safety. In thermal cracking when sour crudes are processed, corrosion can occur where metal temperatures
are between 450° and 900° F. Above 900° F coke forms a protective layer on the metal. The furnace, soaking
drums, lower part of the tower, and high-temperature exchangers are usually subject to corrosion. Hydrogen
sulfide corrosion in coking can also occur when temperatures are not properly controlled above 900° F.

Continuous thermal changes can lead to bulging and cracking of coke drum shells. In coking, temperature control
must often be held within a 10°-20° F range, as high temperatures will produce coke that is too hard to cut out of
the drum. Conversely, temperatures that are too low will result in a high asphaltic-content slurry. Water or steam
injection may be used to prevent buildup of coke in delayed coker furnace tubes. Water must be completely
drained from the coker, so as not to cause an explosion upon recharging with hot coke. Provisions for alternate
means of egress from the working platform on top of coke drums are important in the event of an emergency.

c. Health. The potential exists for exposure to hazardous gases such as hydrogen sulfide and carbon monoxide,
and trace polynuclear aromatics (PNA's) associated with coking operations. When coke is moved as a slurry,
oxygen depletion may occur within confined spaces such as storage silos, since wet carbon will adsorb oxygen.
Wastewater may be highly alkaline and contain oil, sulfides, ammonia, and/or phenol. The potential exists in the
coking process for exposure to burns when handling hot coke or in the event of a steam-line leak, or from steam,
hot water, hot coke, or hot slurry that may be expelled when opening cokers. Safe work practices and/or the use
of appropriate personal protective equipment may be needed for exposures to chemicals and other hazards such
as heat and noise, and during process sampling, inspection, maintenance, and turnaround activities. (Note: coke
produced from petroleum is a different product from that generated in the steel-industry coking process.)

5. CATALYTIC CRACKING.

1. Description.

a. Catalytic cracking breaks complex hydrocarbons into simpler molecules in order to increase the quality and
quantity of lighter, more desirable products and decrease the amount of residuals. This process rearranges the
molecular structure of hydrocarbon compounds to convert heavy hydrocarbon feedstock into lighter fractions such
as kerosene, gasoline, LPG, heating oil, and petrochemical feedstock.

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b. Catalytic cracking is similar to thermal cracking except that catalysts facilitate the conversion of the heavier
molecules into lighter products. Use of a catalyst (a material that assists a chemical reaction but does not take
part in it) in the cracking reaction increases the yield of improved-quality products under much less severe
operating conditions than in thermal cracking. Typical temperatures are from 850°-950° F at much lower
pressures of 10-20 psi. The catalysts used in refinery cracking units are typically solid materials (zeolite,
aluminum hydrosilicate, treated bentonite clay, fuller's earth, bauxite, and silica-alumina) that come in the form
of powders, beads, pellets or shaped materials called extrudites.

c. There are three basic functions in the catalytic cracking process:

■ Reaction: Feedstock reacts with catalyst and cracks into different hydrocarbons;
■ Regeneration: Catalyst is reactivated by burning off coke; and
■ Fractionation: Cracked hydrocarbon stream is separated into various products.

d. The three types of catalytic cracking processes are fluid catalytic cracking (FCC), moving-bed catalytic
cracking, and Thermofor catalytic cracking (TCC). The catalytic cracking process is very flexible, and operating
parameters can be adjusted to meet changing product demand. In addition to cracking, catalytic activities include
dehydrogenation, hydrogenation, and isomerization.

TABLE IV: 2-11. CATALYTIC CRACKING PROCESS

Feedstock From Process Typical products . . . . To

Gas oils . . . . Towers, coker Decomposition, Gasoline . . . . . . . . . Treater or blend

visbreaker alteration Gases . . . . . . . . . . Gas plant

Deasphalted Middle distillates . . .Hydrotreat, blend, or

oils . . . . . . . Deasphalter recycle
Petrochem feedstock . . Petrochem or other
Residue . . . . . . . . . . . Residual fuel blend

6. FLUID CATALYTIC CRACKING.

1. Description.

a. The most common process is FCC, in which the oil is cracked in the presence of a finely divided catalyst which
is maintained in an aerated or fluidized state by the oil vapors. The fluid cracker consists of a catalyst section and
a fractionating section that operate together as an integrated processing unit. The catalyst section contains the
reactor and regenerator, which, with the standpipe and riser, forms the catalyst circulation unit. The fluid catalyst
is continuously circulated between the reactor and the regenerator using air, oil vapors, and steam as the
conveying media.

b. A typical FCC process involves mixing a preheated hydrocarbon charge with hot, regenerated catalyst as it
enters the riser leading to the reactor. The charge is combined with a recycle stream within the riser, vaporized,
and raised to reactor temperature (900°-1,000° F) by the hot catalyst. As the mixture travels up the riser, the
charge is cracked at 10-30 psi. In the more modern FCC units, all cracking takes place in the riser. The "reactor"
no longer functions as a reactor; it merely serves as a holding vessel for the cyclones. This cracking continues
until the oil vapors are separated from the catalyst in the reactor cyclones. The resultant product stream (cracked
product) is then charged to a fractionating column where it is separated into fractions, and some of the heavy oil
is recycled to the riser.

c. Spent catalyst is regenerated to get rid of coke that collects on the catalyst during the process. Spent catalyst
flows through the catalyst stripper to the regenerator, where most of the coke deposits burn off at the bottom
where preheated air and spent catalyst are mixed. Fresh catalyst is added and worn-out catalyst removed to
optimize the cracking process.

FIGURE IV:2-14. FLUID CATALYTIC CRACKING.

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2. Moving Bed Catalytic Cracking. The moving-bed catalytic cracking process is similar to the FCC process. The
catalyst is in the form of pellets that are moved continuously to the top of the unit by conveyor or pneumatic lift
tubes to a storage hopper, then flow downward by gravity through the reactor, and finally to a regenerator. The
regenerator and hopper are isolated from the reactor by steam seals. The cracked product is separated into
recycle gas, oil, clarified oil, distillate, naphtha, and wet gas.

3. Thermofor Catalytic Cracking. In a typical thermofor catalytic cracking unit, the preheated feedstock flows by
gravity through the catalytic reactor bed. The vapors are separated from the catalyst and sent to a fractionating
tower. The spent catalyst is regenerated, cooled, and recycled. The flue gas from regeneration is sent to a carbon-
monoxide boiler for heat recovery.

4. Health and Safety Considerations.

a. Fire Prevention and Protection. Liquid hydrocarbons in the catalyst or entering the heated combustion air
stream should be controlled to avoid exothermic reactions. Because of the presence of heaters in catalytic
cracking units, the possibility exists for fire due to a leak or vapor release. Fire protection including concrete or
other insulation on columns and supports, or fixed water spray or fog systems where insulation is not feasible and
in areas where firewater hose streams cannot reach, should be considered.

In some processes, caution must be taken to prevent explosive concentrations of catalyst dust during recharge or
disposal. When unloading any coked catalyst, the possibility exists for iron sulfide fires. Iron sulfide will ignite
spontaneously when exposed to air and therefore must be wetted with water to prevent it from igniting vapors.
Coked catalyst may be either cooled below 120° F before it is dumped from the reactor, or dumped into
containers that have been purged and inerted with nitrogen and then cooled before further handling.

b. Safety. Regular sampling and testing of the feedstock, product, and recycle streams should be performed to
assure that the cracking process is working as intended and that no contaminants have entered the process
stream. Corrosives or deposits in the feedstock can foul gas compressors. Inspections of critical equipment
including pumps, compressors, furnaces, and heat exchangers should be conducted as needed. When processing
sour crude, corrosion may be expected where temperatures are below 900° F. Corrosion takes place where both
liquid and vapor phases exist, and at areas subject to local cooling such as nozzles and platform supports.

When processing high-nitrogen feedstock, exposure to ammonia and cyanide may occur, subjecting carbon steel
equipment in the FCC overhead system to corrosion, cracking, or hydrogen blistering. These effects may be
minimized by water wash or corrosion inhibitors. Water wash may also be used to protect overhead condensers in
the main column subjected to fouling from ammonium hydrosulfide. Inspections should include checking for leaks
due to erosion or other malfunctions such as catalyst buildup on the expanders, coking in the overhead feeder
lines from feedstock residues, and other unusual operating conditions.

c. Health. Because the catalytic cracker is a closed system, there is normally little opportunity for exposure to
hazardous substances during normal operations. The possibility exists of exposure to extremely hot (700° F)
hydrocarbon liquids or vapors during process sampling or if a leak or release occurs. In addition, exposure to
hydrogen sulfide and/or carbon monoxide gas may occur during a release of product or vapor.

Catalyst regeneration involves steam stripping and decoking, and produces fluid waste streams that may contain
varying amounts of hydrocarbon, phenol, ammonia, hydrogen sulfide, mercaptan, and other materials depending
upon the feedstock, crudes, and processes. Inadvertent formation of nickel carbonyl may occur in cracking
processes using nickel catalysts, with resultant potential for hazardous exposures. Safe work practices and/or the

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use of appropriate personal protective equipment may be needed for exposures to chemicals and other hazards
such as noise and heat; during process sampling, inspection, maintenance and turnaround activities; and when
handling spent catalyst, recharging catalyst, or if leaks or releases occur.

7. HYDROCRACKING.

1. Description.

a. Hydrocracking is a two-stage process combining catalytic cracking and hydrogenation, wherein heavier
feedstocks are cracked in the presence of hydrogen to produce more desirable products. The process employs
high pressure, high temperature, a catalyst, and hydrogen. Hydrocracking is used for feedstocks that are difficult
to process by either catalytic cracking or reforming, since these feedstocks are characterized usually by a high
polycyclic aromatic content and/or high concentrations of the two principal catalyst poisons, sulfur and nitrogen
compounds.

b. The hydrocracking process largely depends on the nature of the feedstock and the relative rates of the two
competing reactions, hydrogenation and cracking. Heavy aromatic feedstock is converted into lighter products
under a wide range of very high pressures (1,000-2,000 psi) and fairly high temperatures (750°-1,500° F), in the
presence of hydrogen and special catalysts. When the feedstock has a high paraffinic content, the primary
function of hydrogen is to prevent the formation of polycyclic aromatic compounds. Another important role of
hydrogen in the hydrocracking process is to reduce tar formation and prevent buildup of coke on the catalyst.
Hydrogenation also serves to convert sulfur and nitrogen compounds present in the feedstock to hydrogen sulfide
and ammonia.

c. Hydrocracking produces relatively large amounts of isobutane for alkylation feedstock. Hydrocracking also
performs isomerization for pour-point control and smoke-point control, both of which are important in high-
quality jet fuel.

2. Hydrocracking Process.

a. In the first stage, preheated feedstock is mixed with recycled hydrogen and sent to the first-stage reactor,
where catalysts convert sulfur and nitrogen compounds to hydrogen sulfide and ammonia. Limited hydrocracking
also occurs.

b. After the hydrocarbon leaves the first stage, it is cooled and liquefied and run through a hydrocarbon
separator. The hydrogen is recycled to the feedstock. The liquid is charged to a fractionator. Depending on the
products desired (gasoline components, jet fuel, and gas oil), the fractionator is run to cut out some portion of
the first stage reactor out-turn. Kerosene-range material can be taken as a separate side-draw product or
included in the fractionator bottoms with the gas oil.

c. The fractionator bottoms are again mixed with a hydrogen stream and charged to the second stage. Since this
material has already been subjected to some hydrogenation, cracking, and reforming in the first stage, the
operations of the second stage are more severe (higher temperatures and pressures). Like the outturn of the first
stage, the second stage product is separated from the hydrogen and charged to the fractionator.

TABLE IV: 2-12. HYDROCRACKING PROCESS.

Feedstock From Process Typical products . . . . To

High pour point Catalytic cracker, Decomposition, Kerosene, jet fuel . . . .Blending
atmospheric hydrogenation
& vacuum tower
Gas oil Vacuum tower, coker Gasoline, distillates . . Blending
Hydrogen Reformer Recycle, reformer gas . . Gas plant

FIGURE IV:2-15. TWO-STAGE HYDROCRACKING.

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3. Health and Safety Considerations.

a. Fire Prevention and Protection. Because this unit operates at very high pressures and temperatures,
control of both hydrocarbon leaks and hydrogen releases is important to prevent fires. In some processes, care is
needed to ensure that explosive concentrations of catalytic dust do not form during recharging.

b. Safety. Inspection and testing of safety relief devices are important due to the very high pressures in this
unit. Proper process control is needed to protect against plugging reactor beds. Unloading coked catalyst requires
special precautions to prevent iron sulfide-induced fires. The coked catalyst should either be cooled to below 120°
F before dumping, or be placed in nitrogen-inerted containers until cooled.

Because of the operating temperatures and presence of hydrogen, the hydrogen-sulfide content of the feedstock
must be strictly controlled to a minimum to reduce the possibility of severe corrosion. Corrosion by wet carbon
dioxide in areas of condensation also must be considered. When processing high-nitrogen feedstock, the
ammonia and hydrogen sulfide form ammonium hydrosulfide, which causes serious corrosion at temperatures
below the water dew point. Ammonium hydrosulfide is also present in sour water stripping.

c. Health. Because this is a closed process, exposures are expected to be minimal under normal operating
conditions. There is a potential for exposure to hydrocarbon gas and vapor emissions, hydrogen and hydrogen
sulfide gas due to high-pressure leaks. Large quantities of carbon monoxide may be released during catalyst
regeneration and changeover. Catalyst steam stripping and regeneration create waste streams containing sour
water and ammonia. Safe work practices and/or the use of appropriate personal protective equipment may be
needed for exposure to chemicals and other hazards such as noise and heat, during process sampling, inspection,
maintenance, and turnaround activities, and when handling spent catalyst.

8. CATALYTIC REFORMING.

1. Description.

a. Catalytic reforming is an important process used to convert low-octane naphthas into high-octane gasoline
blending components called reformates. Reforming represents the total effect of numerous reactions such as
cracking, polymerization, dehydrogenation, and isomerization taking place simultaneously. Depending on the
properties of the naphtha feedstock (as measured by the paraffin, olefin, naphthene, and aromatic content) and
catalysts used, reformates can be produced with very high concentrations of toluene, benzene, xylene, and other
aromatics useful in gasoline blending and petrochemical processing. Hydrogen, a significant by-product, is
separated from the reformate for recycling and use in other processes.

b. A catalytic reformer comprises a reactor section and a product-recovery section. More or less standard is a
feed preparation section in which, by combination of hydrotreatment and distillation, the feedstock is prepared to
specification. Most processes use platinum as the active catalyst. Sometimes platinum is combined with a second

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catalyst (bimetallic catalyst) such as rhenium or another noble metal.

c. There are many different commercial catalytic reforming processes including platforming, powerforming,
ultraforming, and Thermofor catalytic reforming. In the platforming process, the first step is preparation of the
naphtha feed to remove impurities from the naphtha and reduce catalyst degradation. The naphtha feedstock is
then mixed with hydrogen, vaporized, and passed through a series of alternating furnace and fixed-bed reactors
containing a platinum catalyst. The effluent from the last reactor is cooled and sent to a separator to permit
removal of the hydrogen-rich gas stream from the top of the separator for recycling. The liquid product from the
bottom of the separator is sent to a fractionator called a stabilizer (butanizer). It makes a bottom product called
reformate; butanes and lighter go overhead and are sent to the saturated gas plant.

d. Some catalytic reformers operate at low pressure (50-200 psi), and others operate at high pressures (up to
1,000 psi). Some catalytic reforming systems continuously regenerate the catalyst in other systems. One reactor
at a time is taken off-stream for catalyst regeneration, and some facilities regenerate all of the reactors during
turnarounds.

2. Health and Safety Considerations.

a. Fire Prevention and Protection. This is a closed system; however, the potential for fire exists should a leak
or release of reformate gas or hydrogen occur.

b. Safety. Operating procedures should be developed to ensure control of hot spots during start-up. Safe
catalyst handling is very important. Care must be taken not to break or crush the catalyst when loading the beds,
as the small fines will plug up the reformer screens. Precautions against dust when regenerating or replacing
catalyst should also be considered. Also, water wash should be considered where stabilizer fouling has occurred
due to the formation of ammonium chloride and iron salts. Ammonium chloride may form in pretreater
exchangers and cause corrosion and fouling. Hydrogen chloride from the hydrogenation of chlorine compounds
may form acid or ammonium chloride salt.

c. Health. Because this is a closed process, exposures are expected to be minimal under normal operating
conditions. There is potential for exposure to hydrogen sulfide and benzene should a leak or release occur.

Small emissions of carbon monoxide and hydrogen sulfide may occur during regeneration of catalyst. Safe work
practices and/or appropriate personal protective equipment may be needed for exposures to chemicals and other
hazards such as noise and heat during testing, inspecting, maintenance and turnaround activities, and when
handling regenerated or spent catalyst.

TABLE IV: 2-13. CATALYTIC REFORMING PROCESS

Feedstock From Process Typical products . . . . To

Desulfurized Coker Rearrange, High octane gasoline . . Blending
naphtha dehydrogenate
Aromatics . . . . Petrochemical
Naphthene- hydrocracker, Hydrogen . . . . Recycle, hydrotreat, etc.
rich fractions hydrodesulfur
Straight-run Atmospheric Gas . . . . . . . . Gas plant
naphtha fractionator

FIGURE IV:2-16. PLATFORMING PROCESS.

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9. CATALYTIC HYDROTREATING.

1. Description. Catalytic hydrotreating is a hydrogenation process used to remove about 90% of contaminants
such as nitrogen, sulfur, oxygen, and metals from liquid petroleum fractions. These contaminants, if not removed
from the petroleum fractions as they travel through the refinery processing units, can have detrimental effects on
the equipment, the catalysts, and the quality of the finished product. Typically, hydrotreating is done prior to
processes such as catalytic reforming so that the catalyst is not contaminated by untreated feedstock.
Hydrotreating is also used prior to catalytic cracking to reduce sulfur and improve product yields, and to upgrade
middle-distillate petroleum fractions into finished kerosene, diesel fuel, and heating fuel oils. In addition,
hydrotreating converts olefins and aromatics to saturated compounds.

2. Catalytic Hydrodesulfurization Process. Hydrotreating for sulfur removal is called hydrodesulfurization. In a

typical catalytic hydrodesulfurization unit, the feedstock is deaerated and mixed with hydrogen, preheated in a
fired heater (600°-800° F) and then charged under pressure (up to 1,000 psi) through a fixed-bed catalytic
reactor. In the reactor, the sulfur and nitrogen compounds in the feedstock are converted into H2S and NH3. The
reaction products leave the reactor and after cooling to a low temperature enter a liquid/gas separator. The
hydrogen-rich gas from the high-pressure separation is recycled to combine with the feedstock, and the low-
pressure gas stream rich in H2S is sent to a gas treating unit where H2S is removed. The clean gas is then
suitable as fuel for the refinery furnaces. The liquid stream is the product from hydrotreating and is normally sent
to a stripping column for removal of H2S and other undesirable components. In cases where steam is used for
stripping, the product is sent to a vacuum drier for removal of water. Hydrodesulfurized products are blended or
used as catalytic reforming feedstock.

3. Other Hydrotreating Processes.

a. Hydrotreating processes differ depending upon the feedstock available and catalysts used. Hydrotreating can
be used to improve the burning characteristics of distillates such as kerosene. Hydrotreatment of a kerosene
fraction can convert aromatics into naphthenes, which are cleaner-burning compounds.

b. Lube-oil hydrotreating uses catalytic treatment of the oil with hydrogen to improve product quality. The
objectives in mild lube hydrotreating include saturation of olefins and improvements in color, odor, and acid
nature of the oil. Mild lube hydrotreating also may be used following solvent processing. Operating temperatures
are usually below 600° F and operating pressures below 800 psi. Severe lube hydrotreating, at temperatures in
the 600°-750° F range and hydrogen pressures up to 3,000 psi, is capable of saturating aromatic rings, along
with sulfur and nitrogen removal, to impart specific properties not achieved at mild conditions.

c. Hydrotreating also can be employed to improve the quality of pyrolysis gasoline (pygas), a by-product from
the manufacture of ethylene. Traditionally, the outlet for pygas has been motor gasoline blending, a suitable
route in view of its high octane number. However, only small portions can be blended untreated owing to the
unacceptable odor, color, and gum-forming tendencies of this material. The quality of pygas, which is high in
diolefin content, can be satisfactorily improved by hydrotreating, whereby conversion of diolefins into mono-

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olefins provides an acceptable product for motor gas blending.

TABLE IV: 2-14 HYDRODESULFURIZATION PROCESS

Feedstock From Process Typical products . . To

Naphthas, Atmospheric & Treating, Naphtha . . . . . . . Blending
distillates sour vacuum tower, hydrogenation Hydrogen . . . . . . Recycle
gas oil, residuals catalytic & thermal Distillates . . . . . . Blending
cracker H2S, ammonia . . Sulfure plant,
treater
Gas . . . . . . . . . . Gas plant

FIGURE IV:2-17. DISTILLATE HYDRODESULFURIZATION.

4. Health and Safety Considerations.

a. Fire Prevention and Protection. The potential exists for fire in the event of a leak or release of product or
hydrogen gas.

b. Safety. Many processes require hydrogen generation to provide for a continuous supply. Because of the
operating temperatures and presence of hydrogen, the hydrogen sulfide content of the feedstock must be strictly
controlled to a minimum to reduce corrosion. Hydrogen chloride may form and condense as hydrochloric acid in
the lower-temperature parts of the unit. Ammonium hydrosulfide may form in high-temperature, high-pressure
units. Excessive contact time and/or temperature will create coking. Precautions need to be taken when
unloading coked catalyst from the unit to prevent iron sulfide fires. The coked catalyst should be cooled to below
120° F before removal, or dumped into nitrogen-inerted bins where it can be cooled before further handling.
Special antifoam additives may be used to prevent catalyst poisoning from silicone carryover in the coker
feedstock.

c. Health. Because this is a closed process, exposures are expected to be minimal under normal operating
conditions. There is a potential for exposure to hydrogen sulfide or hydrogen gas in the event of a release, or to
ammonia should a sour-water leak or spill occur. Phenol also may be present if high boiling-point feedstocks are
processed. Safe work practices and/or appropriate personal protective equipment may be needed for exposures
to chemicals and other hazards such as noise and heat; during process sampling, inspection, maintenance, and
turnaround activities; and when handling amine or exposed to catalyst.

10. ISOMERIZATION.

1. Description.

a. Isomerization converts n-butane, n-pentane and n-hexane into their respective isoparaffins of substantially
higher octane number. The straight-chain paraffins are converted to their branched-chain counterparts whose
component atoms are the same but are arranged in a different geometric structure. Isomerization is important
for the conversion of n-butane into isobutane, to provide additional feedstock for alkylation units, and the
conversion of normal pentanes and hexanes into higher branched isomers for gasoline blending. Isomerization is
similar to catalytic reforming in that the hydrocarbon molecules are rearranged, but unlike catalytic reforming,
isomerization just converts normal paraffins to isoparaffins.

b. There are two distinct isomerization processes, butane (C4) and pentane/hexane (C5/C6). Butane
isomerization produces feedstock for alkylation. Aluminum chloride catalyst plus hydrogen chloride are universally
used for the low-temperature processes. Platinum or another metal catalyst is used for the higher-temperature

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processes. In a typical low-temperature process, the feed to the isomerization plant is n-butane or mixed butanes
mixed with hydrogen (to inhibit olefin formation) and passed to the reactor at 230°-340° F and 200-300 psi.
Hydrogen is flashed off in a high-pressure separator and the hydrogen chloride removed in a stripper column. The
resultant butane mixture is sent to a fractionator (deisobutanizer) to separate n-butane from the isobutane
product.

c. Pentane/hexane isomerization increases the octane number of the light gasoline components n-pentane and n-
hexane, which are found in abundance in straight-run gasoline. In a typical C5/C6 isomerization process, dried
and desulfurized feedstock is mixed with a small amount of organic chloride and recycled hydrogen, and then
heated to reactor temperature. It is then passed over supported-metal catalyst in the first reactor where benzene
and olefins are hydrogenated. The feed next goes to the isomerization reactor where the paraffins are
catalytically isomerized to isoparaffins. The reactor effluent is then cooled and subsequently separated in the
product separator into two streams: a liquid product (isomerate) and a recycle hydrogen-gas stream. The
isomerate is washed (caustic and water), acid stripped, and stabilized before going to storage.

TABLE IV: 2-15. ISOMERIZATION PROCESSES

Feedstock From Process Typical products . . . To

n-Butane Various Processes Rearrangement Isobutane . . . . . . . Alkylation

n-Pentane Isopentane . . . . . . Blending
n-Hexane Isohexane . . . . . . . Blending
Gas . . . . . . . . . . . Gas Plant

FIGURE IV:2-18. C4 ISOMERIZATION.

FIGURE IV:2-19. C5 AND C6 ISOMERIZATION.

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2. Safety and Health Considerations.

a. Fire Protection and Prevention. Although this is a closed process, the potential for a fire exists should a
release or leak contact a source of ignition such as the heater.

b. Safety. If the feedstock is not completely dried and desulfurized, the potential exists for acid formation
leading to catalyst poisoning and metal corrosion. Water or steam must not be allowed to enter areas where
hydrogen chloride is present. Precautions are needed to prevent HCl from entering sewers and drains.

c. Health. Because this is a closed process, exposures are expected to be minimal during normal operating
conditions. There is a potential for exposure to hydrogen gas, hydrochloric acid, and hydrogen chloride and to
dust when solid catalyst is used. Safe work practices and/or appropriate personal protective equipment may be
needed for exposures to chemicals and other hazards such as heat and noise, and during process sampling,
inspection, maintenance, and turnaround activities.

11. POLYMERIZATION.

1. Description.

a. Polymerization in the petroleum industry is the process of converting light olefin gases including ethylene,
propylene, and butylene into hydrocarbons of higher molecular weight and higher octane number that can be
used as gasoline blending stocks. Polymerization combines two or more identical olefin molecules to form a single
molecule with the same elements in the same proportions as the original molecules. Polymerization may be
accomplished thermally or in the presence of a catalyst at lower temperatures.

b. The olefin feedstock is pretreated to remove sulfur and other undesirable compounds. In the catalytic process
the feedstock is either passed over a solid phosphoric acid catalyst or comes in contact with liquid phosphoric
acid, where an exothermic polymeric reaction occurs. This reaction requires cooling water and the injection of
cold feedstock into the reactor to control temperatures between 300° and 450° F at pressures from 200 psi to
1,200 psi. The reaction products leaving the reactor are sent to stabilization and/or fractionator systems to
separate saturated and unreacted gases from the polymer gasoline product.

NOTE: In the petroleum industry, polymerization is used to indicate the production of gasoline components,
hence the term "polymer" gasoline. Furthermore, it is not essential that only one type of monomer be involved. If
unlike olefin molecules are combined, the process is referred to as "copolymerization." Polymerization in the true
sense of the word is normally prevented, and all attempts are made to terminate the reaction at the dimer or
trimer (three monomers joined together) stage. However, in the petrochemical section of a refinery,
polymerization, which results in the production of, for instance, polyethylene, is allowed to proceed until materials
of the required high molecular weight have been produced.

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2. Safety and Health Considerations.

a. Fire prevention and Protection. Polymerization is a closed process where the potential for a fire exists due
to leaks or releases reaching a source of ignition.

b. Safety. The potential for an uncontrolled exothermic reaction exists should loss of cooling water occur. Severe
corrosion leading to equipment failure will occur should water make contact with the phosphoric acid, such as
during water washing at shutdowns. Corrosion may also occur in piping manifolds, reboilers, exchangers, and
other locations where acid may settle out.

c. Health. Because this is a closed system, exposures are expected to be minimal under normal operating
conditions. There is a potential for exposure to caustic wash (sodium hydroxide), to phosphoric acid used in the
process or washed out during turnarounds, and to catalyst dust. Safe work practices and/or appropriate personal
protective equipment may be needed for exposures to chemicals and other hazards such as noise and heat, and
during process sampling, inspection, maintenance, and turnaround activities.

TABLE IV: 2-16. POLYMERIZATION PROCESS

Feedstock From Process Typical products . . . . . . To

Olefins Cracking Unification High octane naphtha . . Gasoline
processes blending
Petrochem. feedstock . . Petrochemical
Liquefied petro. gas . . . Storage

FIGURE IV:2-20. POLYMERIZATION PROCESS.

12. ALKYLATION.

1. Description. Alkylation combines low-molecular-weight olefins (primarily a mixture of propylene and butylene)
with isobutene in the presence of a catalyst, either sulfuric acid or hydrofluoric acid. The product is called alkylate
and is composed of a mixture of high-octane, branched-chain paraffinic hydrocarbons. Alkylate is a premium
blending stock because it has exceptional antiknock properties and is clean burning. The octane number of the
alkylate depends mainly upon the kind of olefins used and upon operating conditions.

2. Sulfuric Acid Alkylation Process.

a. In cascade type sulfuric acid (H2SO4) alkylation units, the feedstock (propylene, butylene, amylene, and fresh
isobutane) enters the reactor and contacts the concentrated sulfuric acid catalyst (in concentrations of 85% to
95% for good operation and to minimize corrosion). The reactor is divided into zones, with olefins fed through
distributors to each zone, and the sulfuric acid and isobutanes flowing over baffles from zone to zone.

b. The reactor effluent is separated into hydrocarbon and acid phases in a settler, and the acid is returned to the
reactor. The hydrocarbon phase is hot-water washed with caustic for pH control before being successively
depropanized, deisobutanized, and debutanized. The alkylate obtained from the deisobutanizer can then go
directly to motor-fuel blending or be rerun to produce aviation-grade blending stock. The isobutane is recycled to
the feed.

3. Hydrofluoric Acid Alylation Process. Phillips and UOP are the two common types of hydrofluoric acid alkylation

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processes in use. In the Phillips process, olefin and isobutane feedstock are dried and fed to a combination
reactor/settler system. Upon leaving the reaction zone, the reactor effluent flows to a settler (separating vessel)
where the acid separates from the hydrocarbons. The acid layer at the bottom of the separating vessel is
recycled. The top layer of hydrocarbons (hydrocarbon phase), consisting of propane, normal butane, alkylate, and
excess (recycle) isobutane, is charged to the main fractionator, the bottom product of which is motor alkylate.
The main fractionator overhead, consisting mainly of propane, isobutane, and HF, goes to a depropanizer.
Propane with trace amount of HF goes to an HF stripper for HF removal and is then catalytically defluorinated,
treated, and sent to storage. Isobutane is withdrawn from the main fractionator and recycled to the reactor/
settler, and alkylate from the bottom of the main fractionator is sent to product blending.

4. The UOP process uses two reactors with separate settlers. Half of the dried feedstock is charged to the first
reactor, along with recycle and makeup isobutane. The reactor effluent then goes to its settler, where the acid is
recycled and the hydrocarbon charged to the second reactor. The other half of the feedstock also goes to the
second reactor, with the settler acid being recycled and the hydrocarbons charged to the main fractionator.
Subsequent processing is similar to the Phillips process. Overhead from the main fractionator goes to a
depropanizer. Isobutane is recycled to the reaction zone and alkylate is sent to product blending.

TABLE IV: 2-17. ALKYLATION PROCESS

Feedstock From Process Typical products . . . . To

Petroleum gas Distillation or cracking Unification High octane gasoline . . Blending
Olefins Cat. or hydro cracking n-Butane & propane . . . Stripper or
blender
Isobutane Isomerization

FIGURE IV:2-21. SULFURIC ACID ALKYLATION.

FIGURE IV:2-22. HYDROGEN FLUORIDE ALKYLATION.

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5. Health and Safety Considerations.

a. Fire Protection and Prevention. Alkylation units are closed processes; however, the potential exists for fire
should a leak or release occur that allows product or vapor to reach a source of ignition.

b. Safety. Sulfuric acid and hydrofluoric acid are potentially hazardous chemicals. Loss of coolant water, which is
needed to maintain process temperatures, could result in an upset. Precautions are necessary to ensure that
equipment and materials that have been in contact with acid are handled carefully and are thoroughly cleaned
before they leave the process area or refinery. Immersion wash vats are often provided for neutralization of
equipment that has come into contact with hydrofluoric acid. Hydrofluoric acid units should be thoroughly drained
and chemically cleaned prior to turnarounds and entry to remove all traces of iron fluoride and hydrofluoric acid.
Following shutdown, where water has been used the unit should be thoroughly dried before hydrofluoric acid is
introduced.

Leaks, spills, or releases involving hydrofluoric acid or hydrocarbons containing hydrofluoric acid can be
extremely hazardous. Care during delivery and unloading of acid is essential. Process unit containment by curbs,
drainage, and isolation so that effluent can be neutralized before release to the sewer system is considered.
Vents can be routed to soda-ash scrubbers to neutralize hydrogen fluoride gas or hydrofluoric acid vapors before
release. Pressure on the cooling water and steam side of exchangers should be kept below the minimum pressure
on the acid service side to prevent water contamination.

Some corrosion and fouling in sulfuric acid units may occur from the breakdown of sulfuric acid esters or where
caustic is added for neutralization. These esters can be removed by fresh acid treating and hot-water washing. To
prevent corrosion from hydrofluoric acid, the acid concentration inside the process unit should be maintained
above 65% and moisture below 4%.

c. Health. Because this is a closed process, exposures are expected to be minimal during normal operations.
There is a potential for exposure should leaks, spills, or releases occur. Sulfuric acid and (particularly)
hydrofluoric acid are potentially hazardous chemicals. Special precautionary emergency preparedness measures
and protection appropriate to the potential hazard and areas possibly affected need to be provided. Safe work
practices and appropriate skin and respiratory personal protective equipment are needed for potential exposures
to hydrofluoric and sulfuric acids during normal operations such as reading gauges, inspecting, and process
sampling, as well as during emergency response, maintenance, and turnaround activities. Procedures should be
in place to ensure that protective equipment and clothing worn in hydrofluoric acid activities are decontaminated
and inspected before reissue. Appropriate personal protection for exposure to heat and noise also may be
required.

13. SWEETENING AND TREATING PROCESSES.

1. Description.

a. Treating is a means by which contaminants such as organic compounds containing sulfur, nitrogen, and
oxygen; dissolved metals and inorganic salts; and soluble salts dissolved in emulsified water are removed from
petroleum fractions or streams. Petroleum refiners have a choice of several different treating processes, but the
primary purpose of the majority of them is the elimination of unwanted sulfur compounds. A variety of
intermediate and finished products, including middle distillates, gasoline, kerosene, jet fuel, and sour gases are
dried and sweetened. Sweetening, a major refinery treatment of gasoline, treats sulfur compounds (hydrogen
sulfide, thiophene and mercaptan) to improve color, odor, and oxidation stability. Sweetening also reduces
concentrations of carbon dioxide.

b. Treating can be accomplished at an intermediate stage in the refining process, or just before sending the
finished product to storage. Choices of a treating method depend on the nature of the petroleum fractions,
amount and type of impurities in the fractions to be treated, the extent to which the process removes the
impurities, and end-product specifications. Treating materials include acids, solvents, alkalis, oxidizing, and
adsorption agents.

2. Acid, Caustic, or Clay Treating. Sulfuric acid is the most commonly used acid treating process. Sulfuric acid
treating results in partial or complete removal of unsaturated hydrocarbons, sulfur, nitrogen, and oxygen
compounds, and resinous and asphaltic compounds. It is used to improve the odor, color, stability, carbon
residue, and other properties of the oil. Clay/lime treatment of acid-refined oil removes traces of asphaltic
materials and other compounds improving product color, odor, and stability. Caustic treating with sodium (or
potassium) hydroxide is used to improve odor and color by removing organic acids (naphthenic acids, phenols)
and sulfur compounds (mercaptans, H2S) by a caustic wash. By combining caustic soda solution with various
solubility promoters (e.g., methyl alcohol and cresols), up to 99% of all mercaptans as well as oxygen and
nitrogen compounds can be dissolved from petroleum fractions.

3. Drying and Sweetening. Feedstocks from various refinery units are sent to gas treating plants where butanes
and butenes are removed for use as alkylation feedstock, heavier components are sent to gasoline blending,
propane is recovered for LPG, and propylene is removed for use in petrochemicals. Some mercaptans are
removed by water-soluble chemicals that react with the mercaptans. Caustic liquid (sodium hydroxide), amine
compounds (diethanolamine) or fixed-bed catalyst sweetening also may be used. Drying is accomplished by the
use of water absorption or adsorption agents to remove water from the products. Some processes simultaneously

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dry and sweeten by adsorption on molecular sieves.

TABLE IV: 2-18. SWEETENING AND TREATING PROCESSES.

Feedstock From Process Typical products . . . . . . To

Gases, Various Treatment Butane & butene . . . . Alkylation

finished products, Propane, distillates . . . Storage
intermediates Gasoline . . . . . . . . . . . Blending
Propylene . . . . . . . . . . Petrochemical

FIGURE IV:2-23. MOLECULAR SIEVE DRYING AND SWEETENING.

4. Sulfur Recovery. Sulfur recovery converts hydrogen sulfide in sour gases and hydrocarbon streams to elemental
sulfur. The most widely used recovery system is the Claus process, which uses both thermal and catalytic-
conversion reactions. A typical process produces elemental sulfur by burning hydrogen sulfide under controlled
conditions. Knockout pots are used to remove water and hydrocarbons from feed gas streams. The gases are
then exposed to a catalyst to recover additional sulfur. Sulfur vapor from burning and conversion is condensed
and recovered.

5. Hydrogen Sulfide Scrubbing. Hydrogen sulfide scrubbing is a common treating process in which the
hydrocarbon feedstock is first scrubbed to prevent catalyst poisoning. Depending on the feedstock and the nature
of contaminants, desulfurization methods vary from ambient temperature-activated charcoal absorption to high-
temperature catalytic hydrogenation followed by zinc oxide treating.

6. Health and Safety Considerations.

a. Fire Protection and Prevention. The potential exists for fire from a leak or release of feedstock or product.
Sweetening processes use air or oxygen. If excess oxygen enters these processes, it is possible for a fire to occur
in the settler due to the generation of static electricity, which acts as the ignition source.

b. Health. Because these are closed processes, exposures are expected to be minimal under normal operating
conditions. There is a potential for exposure to hydrogen sulfide, caustic (sodium hydroxide), spent caustic, spent
catalyst (Merox), catalyst dust and sweetening agents (sodium carbonate and sodium bicarbonate). Safe work
practices and/or appropriate personal protective equipment may be needed for exposures to chemicals and other
hazards such as noise and heat, and during process sampling, inspection, maintenance, and turnaround activities.

14. UNSATURATED GAS PLANTS.

1. Description. Unsaturated (unsat) gas plants recover light hydrocarbons (C3 and C4 olefins) from wet gas
streams from the FCC, TCC, and delayed coker overhead accumulators or fractionation receivers. In a typical
unsat gas plant, the gases are compressed and treated with amine to remove hydrogen sulfide either before or
after they are sent to a fractionating absorber where they are mixed into a concurrent flow of debutanized
gasoline. The light fractions are separated by heat in a reboiler, the offgas is sent to a sponge absorber, and the
bottoms are sent to a debutanizer. A portion of the debutanized hydrocarbon is recycled, with the balance sent to
the splitter for separation. The overhead gases go to a depropanizer for use as alkylation unit feedstock.

TABLE IV: 2-19. UNSAT GAS PLANT PROCESS.

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Feedstock From Process Typical products . . . To

Gas Oils FCC, TCC, delayed Treatment Gasoline . . . . . . . . . Recycle or treating

coker Gases . . . . . . . . . . . Alkylation

2. Health and Safety Considerations.

a. Fire Prevention and Protection. The potential of a fire exists should spills, releases, or vapors reach a
source of ignition.

b. Safety. In unsat gas plants handling FCC feedstock, the potential exists for corrosion from moist hydrogen
sulfide and cyanides. When feedstocks are from the delayed coker or the TCC, corrosion from hydrogen sulfide
and deposits in the high pressure sections of gas compressors from ammonium compounds is possible.

c. Health. Because these are closed processes, exposures are expected to be minimal under normal operating
conditions. There is a potential for exposures to amine compounds such as monoethanolamine (MEA),
diethanolamine (DEA) and methyldiethanolamine (MDEA) and hydrocarbons. Safe work practices and/or
appropriate personal protective equipment may be needed for exposures to chemicals and other hazards such as
noise and heat, and during process sampling, inspection, maintenance, and turnaround activities.

15. AMINE PLANTS.

1. Description. Amine plants remove acid contaminants from sour gas and hydrocarbon streams. In amine plants,
gas and liquid hydrocarbon streams containing carbon dioxide and/or hydrogen sulfide are charged to a gas
absorption tower or liquid contactor where the acid contaminants are absorbed by counterflowing amine solutions
(i.e., MEA, DEA, MDEA). The stripped gas or liquid is removed overhead, and the amine is sent to a regenerator.
In the regenerator, the acidic components are stripped by heat and reboiling action and disposed of, and the
amine is recycled.

2. Health and Safety Considerations.

a. Fire Protection and Prevention. The potential for fire exists where a spill or leak could reach a source of
ignition.

b. Safety. To minimize corrosion, proper operating practices should be established and regenerator bottom and
reboiler temperatures controlled. Oxygen should be kept out of the system to prevent amine oxidation.

c. Health. Because this is a closed process, exposures are expected to be minimal during normal operations.
There is potential for exposure to amine compounds (i.e. monoethanolamine, diethanolamine,
methyldiethanolamine), hydrogen sulfide and carbon dioxide. Safe work practices and/or appropriate personal
protective equipment may be needed for exposures to chemicals and other hazards such as noise and heat, and
during process sampling, inspection, maintenance and turnaround activities.

16. SATURATE GAS PLANTS.

1. Description. Saturate (sat) gas plants separate refinery gas components including butanes for alkylation,
pentanes for gasoline blending, LPG's for fuel, and ethane for petrochemicals. Because sat gas processes depend
on the feedstock and product demand, each refinery uses different systems, usually absorption-fractionation or
straight fractionation. In absorption-fractionation, gases and liquids from various refinery units are fed to an
absorber-deethanizer where C2 and lighter fractions are separated from heavier fractions by lean oil absorption
and removed for use as fuel gas or petrochemical feed. The heavier fractions are stripped and sent to a
debutanizer, and the lean oil is recycled back to the absorber-deethanizer. C3/C4 is separated from pentanes in
the debutanizer, scrubbed to remove hydrogen sulfide, and fed to a splitter where propane and butane are
separated. In fractionation sat gas plants, the absorption stage is eliminated.

2. Health and Safety Considerations.

a. Fire Protection and Prevention. There is potential for fire if a leak or release reaches a source of ignition
such as the unit reboiler.

b. Safety. Corrosion could occur from the presence of hydrogen sulfide, carbon dioxide, and other compounds as
a result of prior treating. Streams containing ammonia should be dried before processing. Antifouling additives
may be used in absorption oil to protect heat exchangers. Corrosion inhibitors may be used to control corrosion in
overhead systems.

c. Health. Because this is a closed process, exposures are expected to be minimal during normal operations.
There is potential for exposure to hydrogen sulfide, carbon dioxide, and other products such as diethanolamine or
sodium hydroxide carried over from prior treating. Safe work practices and/or appropriate personal protective
equipment may be needed for exposures to chemicals and other hazards such as noise and heat, and during

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process sampling, inspection, maintenance, and turnaround activities.

17. ASPHALT PRODUCTION.

1. Description.

a. Asphalt is a portion of the residual fraction that remains after primary distillation operations. It is further
processed to impart characteristics required by its final use. In vacuum distillation, generally used to produce
road-tar asphalt, the residual is heated to about 750° F and charged to a column where vacuum is applied to
prevent cracking.

b. Asphalt for roofing materials is produced by air blowing. Residual is heated in a pipe still almost to its flash
point and charged to a blowing tower where hot air is injected for a predetermined time. The dehydrogenization
of the asphalt forms hydrogen sulfide, and the oxidation creates sulfur dioxide. Steam, used to blanket the top of
the tower to entrain the various contaminants, is then passed through a scrubber to condense the hydrocarbons.

c. A third process used to produce asphalt is solvent deasphalting. In this extraction process, which uses propane
(or hexane) as a solvent, heavy oil fractions are separated to produce heavy lubricating oil, catalytic cracking
feedstock, and asphalt. Feedstock and liquid propane are pumped to an extraction tower at precisely controlled
mixtures, temperatures (150°-250° F), and pressures of 350-600 psi. Separation occurs in a rotating disc
contactor, based on differences in solubility. The products are then evaporated and steam stripped to recover the
propane, which is recycled. Deasphalting also removes some sulfur and nitrogen compounds, metals, carbon
residues, and paraffins from the feedstock.

TABLE IV: 2-20. SOLVENT DEASPHALTING PROCESS.

Feedstock From Process Typical products . . To

Residual, Atmospheric tower Treatment Heavy lube oil . . . Treating or lube blending
reduced crude & Asphalt . . . . . . . . Storage of shipping
Vacuum tower Deasphalted oil . . Hydrotreat & catalytic cracker
Propane . . . . . . . Recycle

2. Safety and Health Considerations.

a. Fire Protection and Prevention. The potential for a fire exists if a product leak or release contacts a source
of ignition such as the process heater. Condensed steam from the various asphalt and deasphalting processes will
contain trace amounts of hydrocarbons. Any disruption of the vacuum can result in the entry of atmospheric air
and subsequent fire. In addition, raising the temperature of the vacuum tower bottom to improve efficiency can
generate methane by thermal cracking. This can create vapors in asphalt storage tanks that are not detectable by
flash testing but are high enough to be flammable.

b. Safety. Deasphalting requires exact temperature and pressure control. In addition, moisture, excess solvent,
or a drop in operating temperature may cause foaming, which affects the product temperature control and may
create an upset.

c. Health. Because these are closed processes, exposures are expected to be minimal during normal operations.
Should a spill or release occur, there is a potential for exposure to residuals and asphalt. Air blowing can create
some polynuclear aromatics. Condensed steam from the air-blowing asphalt process may also contain
contaminants. The potential for exposure to hydrogen sulfide and sulfur dioxide exists in the production of
asphalt. Safe work practices and/or appropriate personal protective equipment may be needed for exposures to
chemicals and other hazards such as noise and heat, and during process sampling, inspection, maintenance, and
turnaround activities.

18. HYDROGEN PRODUCTION.

1. Description.

a. High-purity hydrogen (95%-99%) is required for hydrodesulfurization, hydrogenation, hydrocracking, and

petrochemical processes. Hydrogen, produced as a by-product of refinery processes (principally hydrogen
recovery from catalytic reformer product gases), often is not enough to meet the total refinery requirements,
necessitating the manufacturing of additional hydrogen or obtaining supply from external sources.

b. In steam-methane reforming, desulfurized gases are mixed with superheated steam (1,100°-1,600° F) and
reformed in tubes containing a nickel base catalyst. The reformed gas, which consists of steam, hydrogen, carbon
monoxide, and carbon dioxide, is cooled and passed through converters containing an iron catalyst where the
carbon monoxide reacts with steam to form carbon dioxide and more hydrogen. The carbon dioxide is removed
by amine washing. Any remaining carbon monoxide in the product stream is converted to methane.

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c. Steam-naphtha reforming is a continuous process for the production of hydrogen from liquid hydrocarbons
and is, in fact, similar to steam-methane reforming. A variety of naphthas in the gasoline boiling range may be
employed, including fuel containing up to 35% aromatics. Following pretreatment to remove sulfur compounds,
the feedstock is mixed with steam and taken to the reforming furnace (1,250°-1,500° F) where hydrogen is
produced.

TABLE IV: 2-21. STEAM REFORMING PROCESS.

Feedstock From Process Typical products . . . To

Desufurized Various Decomposition Hydrogen . . . . . . . Processing
refinery gas treatment Carbon dioxide . . . Atmosphere
units Carbon monoxide . . Methane

2. Health and Safety Considerations.

a. Fire Protection and Prevention. The possibility of fire exists should a leak or release occur and reach an
ignition source.

b. Safety. The potential exists for burns from hot gases and superheated steam should a release occur.
Inspections and testing should be considered where the possibility exists for valve failure due to contaminants in
the hydrogen. Carryover from caustic scrubbers should be controlled to prevent corrosion in preheaters. Chlorides
from the feedstock or steam system should be prevented from entering reformer tubes and contaminating the
catalyst.

c. Health. Because these are closed processes, exposures are expected to be minimal during normal operating
conditions. There is a potential for exposure to excess hydrogen, carbon monoxide, and/or carbon dioxide.
Condensate can be contaminated by process materials such as caustics and amine compounds, with resultant
exposures. Depending on the specific process used, safe work practices and/or appropriate personal protective
equipment may be needed for exposures to chemicals and other hazards such as noise and heat, and during
process sampling, inspection, maintenance, and turnaround activities.

19. BLENDING.

1. Description. Blending is the physical mixture of a number of different liquid hydrocarbons to produce a finished
product with certain desired characteristics. Products can be blended in-line through a manifold system, or batch
blended in tanks and vessels. In-line blending of gasoline, distillates, jet fuel, and kerosene is accomplished by
injecting proportionate amounts of each component into the main stream where turbulence promotes thorough
mixing. Additives including octane enhancers, metal deactivators, anti-oxidants, anti-knock agents, gum and rust
inhibitors, detergents, etc. are added during and/or after blending to provide specific properties not inherent in
hydrocarbons.

2. Health and Safety Considerations.

a. Fire Prevention and Protection. Ignition sources in the area need to be controlled in the event of a leak or
release.

b. Health. Safe work practices and/or appropriate personal protective equipment may be needed for exposures
to chemicals and other hazards such as noise and heat; when handling additives; and during inspection,
maintenance, and turnaround activities.

20. LUBRICANT, WAX, AND GREASE MANUFACTURING PROCESSES.

1. Description. Lubricating oils and waxes are refined from the residual fractions of atmospheric and vacuum
distillation. The primary objective of the various lubricating oil refinery processes is to remove asphalts,
sulfonated aromatics, and paraffinic and isoparaffinic waxes from residual fractions. Reduced crude from the
vacuum unit is deasphalted and combined with straight-run lubricating oil feedstock, preheated, and solvent-
extracted (usually with phenol or furfural) to produce raffinate.

2. Wax Manufacturing Process. Raffinate from the extraction unit contains a considerable amount of wax that
must be removed by solvent extraction and crystallization. The raffinate is mixed with a solvent (propane) and
precooled in heat exchangers. The crystallization temperature is attained by the evaporation of propane in the
chiller and filter feed tanks. The wax is continuously removed by filters and cold solvent-washed to recover
retained oil. The solvent is recovered from the oil by flashing and steam stripping. The wax is then heated with
hot solvent, chilled, filtered, and given a final wash to remove all oil.

3. Lubricating Oil Process. The dewaxed raffinate is blended with other distillate fractions and further treated for
viscosity index, color, stability, carbon residue, sulfur, additive response, and oxidation stability in extremely
selective extraction processes using solvents (furfural, phenol, etc.). In a typical phenol unit, the raffinate is

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mixed with phenol in the treating section at temperatures below 400° F. Phenol is then separated from the
treated oil and recycled. The treated lube-oil base stocks are then mixed and/or compounded with additives to
meet the required physical and chemical characteristics of motor oils, industrial lubricants, and metal working
oils.

TABLE IV: 2-22. LUBRICATING OIL AND WAX

MANUFACTURING PROCESSES.
Feedstock From Process Typical products . . . . To
Lube Vacuum tower, Treatment Dewaxed raffinate . . Lube blend or
feedstock solvent dewaxing, compound,
and additives hydrotreating solvent grease compounding
extraction, etc. Wax . . . . . . . . . . . . Storage or shipping

4. Grease Compounding. Grease is made by blending metallic soaps (salts of long-chained fatty acids) and
additives into a lubricating oil medium at temperatures of 400°-600° F. Grease may be either batch-produced or
continuously compounded. The characteristics of the grease depend to a great extent on the metallic element
(calcium, sodium, aluminum, lithium, etc.) in the soap and the additives used.

5. Safety and Health Considerations.

a. Fire Protection and Prevention. The potential for fire exists if a product or vapor leak or release in the lube
blending and wax processing areas reaches a source of ignition. Storage of finished products, both bulk and
packaged, should be in accordance with recognized practices.

While the potential for fire is reduced in lube oil blending, care must be taken when making metal-working oils
and compounding greases due to the use of higher blending and compounding temperatures and lower flash point
products.

b. Safety. Control of treater temperature is important as phenol can cause corrosion above 400° F. Batch and in-
line blending operations require strict controls to maintain desired product quality. Spills should be cleaned and
leaks repaired to avoid slips and falls. Additives in drums and bags need to be handled properly to avoid strain.
Wax can clog sewer or oil drainage systems and interfere with wastewater treatment.

c. Health. When blending, sampling, and compounding, personal protection from steam, dusts, mists, vapors,
metallic salts, and other additives is appropriate. Skin contact with any formulated grease or lubricant should be
avoided. Safe work practices and/or appropriate personal protection may be needed for exposures to chemicals
and other hazards such as noise and heat; during inspection, maintenance, and turnaround activities; and while
sampling and handling hydrocarbons and chemicals during the production of lubricating oil and wax.

5. OTHER REFINERY OPERATIONS.

1. HEAT EXCHANGERS, COOLERS, AND PROCESS HEATERS.

1. Heating Operations. Process heaters and heat exchangers preheat feedstock in distillation towers and in
refinery processes to reaction temperatures. Heat exchangers use either steam or hot hydrocarbon transferred
from some other section of the process for heat input. The heaters are usually designed for specific process
operations, and most are of cylindrical vertical or box-type designs. The major portion of heat provided to process
units comes from fired heaters fueled by refinery or natural gas, distillate, and residual oils. Fired heaters are
found on crude and reformer preheaters, coker heaters, and large-column reboilers.

2. Cooling Operations. Heat also may be removed from some processes by air and water exchangers, fin fans, gas
and liquid coolers, and overhead condensers, or by transferring heat to other systems. The basic mechanical
vapor-compression refrigeration system, which may serve one or more process units, includes an evaporator,
compressor, condenser, controls, and piping. Common coolants are water, alcohol/water mixtures, or various
glycol solutions.

3. Health and Safety Considerations.

a. Fire Protection and Prevention. A means of providing adequate draft or steam purging is required to
reduce the chance of explosions when lighting fires in heater furnaces. Specific start-up and emergency
procedures are required for each type of unit. If fire impinges on fin fans, failure could occur due to overheating.
If flammable product escapes from a heat exchanger or cooler due to a leak, fire could occur.

b. Safety. Care must be taken to ensure that all pressure is removed from heater tubes before removing header
or fitting plugs. Consideration should be given to providing for pressure relief in heat-exchanger piping systems in
the event they are blocked off while full of liquid. If controls fail, variations of temperature and pressure could
occur on either side of the heat exchanger. If heat exchanger tubes fail and process pressure is greater than
heater pressure, product could enter the heater with downstream consequences. If the process pressure is less
than heater pressure, the heater stream could enter into the process fluid. If loss of circulation occurs in liquid or

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gas coolers, increased product temperature could affect downstream operations and require pressure relief.

c. Health. Because these are closed systems, exposures under normal operating conditions are expected to be
minimal. Depending on the fuel, process operation, and unit design, there is a potential for exposure to hydrogen
sulfide, carbon monoxide, hydrocarbons, steam boiler feed-water sludge, and water-treatment chemicals. Skin
contact should be avoided with boiler blowdown, which may contain phenolic compounds. Safe work practices
and/or appropriate personal protective equipment against hazards may be needed during process maintenance,
inspection, and turnaround activities and for protection from radiant heat, superheated steam, hot hydrocarbon,
and noise exposures.

2. STEAM GENERATION.

1. Heater and Boiler Operations. Steam is generated in main generation plants, and/or at various process units
using heat from flue gas or other sources. Heaters (furnaces) include burners and a combustion air system, the
boiler enclosure in which heat transfer takes place, a draft or pressure system to remove flue gas from the
furnace, soot blowers, and compressed-air systems that seal openings to prevent the escape of flue gas. Boilers
consist of a number of tubes that carry the water-steam mixture through the furnace for maximum heat transfer.
These tubes run between steam-distribution drums at the top of the boiler and water-collecting drums at the
bottom of the boiler. Steam flows from the steam drum to the superheater before entering the steam distribution
system.

2. Heater Fuel.

a. Heaters may use any one or combination of fuels including refinery gas, natural gas, fuel oil, and powdered
coal. Refinery off-gas is collected from process units and combined with natural gas and LPG in a fuel-gas balance
drum. The balance drum provides constant system pressure, fairly stable Btu-content fuel, and automatic
separation of suspended liquids in gas vapors, and it prevents carryover of large slugs of condensate into the
distribution system. Fuel oil is typically a mix of refinery crude oil with straight-run and cracked residues and
other products. The fuel-oil system delivers fuel to process-unit heaters and steam generators at required
temperatures and pressures. The fuel oil is heated to pumping temperature, sucked through a coarse suction
strainer, pumped to a temperature-control heater, and then pumped through a fine-mesh strainer before being
burned.

b. In one example of process-unit heat generation, carbon monoxide boilers recover heat in catalytic cracking
units as carbon monoxide in flue gas is burned to complete combustion. In other processes, waste-heat recovery
units use heat from the flue gas to make steam.

3. Steam Distribution. The distribution system consists of valves, fittings, piping, and connections suitable for the
pressure of the steam transported. Steam leaves the boilers at the highest pressure required by the process units
or electrical generation. The steam pressure is then reduced in turbines that drive process pumps and
compressors. Most steam used in the refinery is condensed to water in various types of heat exchangers. The
condensate is reused as boiler feedwater or discharged to wastewater treatment. When refinery steam is also
used to drive steam turbine generators to produce electricity, the steam must be produced at much higher
pressure than required for process steam. Steam typically is generated by heaters (furnaces) and boilers
combined in one unit.

4. Feedwater.

a. Feedwater supply is an important part of steam generation. There must always be as many pounds of water
entering the system as there are pounds of steam leaving it. Water used in steam generation must be free of
contaminants including minerals and dissolved impurities that can damage the system or affect its operation.
Suspended materials such as silt, sewage, and oil, which form scale and sludge, must be coagulated or filtered
out of the water. Dissolved gases, particularly carbon dioxide and oxygen, cause boiler corrosion and are
removed by deaeration and treatment. Dissolved minerals including metallic salts, calcium, carbonates, etc., that
cause scale, corrosion, and turbine blade deposits are treated with lime or soda ash to precipitate them from the
water. Recirculated cooling water must also be treated for hydrocarbons and other contaminants.

b. Depending on the characteristics of raw boiler feedwater, some or all of the following six stages of treatment
will be applicable:

■ Clarification;
■ Sedimentation;
■ Filtration;
■ Ion exchange;
■ Deaeration; and
■ Internal treatment.

5. Health and Safety Considerations.

a. Fire Protection and Prevention. The most potentially hazardous operation in steam generation is heater
startup. A flammable mixture of gas and air can build up as a result of loss of flame at one or more burners

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during light-off. Each type of unit requires specific startup and emergency procedures including purging before
lightoff and in the event of misfire or loss of burner flame.

b. Safety. If feedwater runs low and boilers are dry, the tubes will overheat and fail. Conversely, excess water
will be carried over into the steam distribution system and damage the turbines. Feedwater must be free of
contaminants that could affect operations. Boilers should have continuous or intermittent blowdown systems to
remove water from steam drums and limit buildup of scale on turbine blades and superheater tubes. Care must
be taken not to overheat the superheater during startup and shut-down. Alternate fuel sources should be
provided in the event of loss of gas due to refinery unit shutdown or emergency. Knockout pots provided at
process units remove liquids from fuel gas before burning.

c. Health. Safe work practices and/or appropriate personal protective equipment may be needed for potential
exposures to feedwater chemicals, steam, hot water, radiant heat, and noise, and during process sampling,
inspection, maintenance, and turnaround activities.

3. PRESSURE-RELIEF AND FLARE SYSTEMS.

1. Pressure-Relief Systems. Pressure-relief systems control vapors and liquids that are released by pressure-
relieving devices and blow-downs. Pressure relief is an automatic, planned release when operating pressure
reaches a predetermined level. Blowdown normally refers to the intentional release of material, such as
blowdowns from process unit startups, furnace blowdowns, shutdowns, and emergencies. Vapor depressuring is
the rapid removal of vapors from pressure vessels in case of fire. This may be accomplished by the use of a
rupture disc, usually set at a higher pressure than the relief valve.

2. Safety Relief Valve Operations. Safety relief valves, used for air, steam, and gas as well as for vapor and
liquid, allow the valve to open in proportion to the increase in pressure over the normal operating pressure.
Safety valves designed primarily to release high volumes of steam usually pop open to full capacity. The
overpressure needed to open liquid-relief valves where large-volume discharge is not required increases as the
valve lifts due to increased spring resistance. Pilot-operated safety relief valves, with up to six times the capacity
of normal relief valves, are used where tighter sealing and larger volume discharges are required. Nonvolatile
liquids are usually pumped to oil-water separation and recovery systems, and volatile liquids are sent to units
operating at a lower pressure.

3. Flare Systems. A typical closed pressure release and flare system includes relief valves and lines from process
units for collection of discharges, knockout drums to separate vapors and liquids, seals, and/or purge gas for
flashback protection, and a flare and igniter system which combusts vapors when discharging directly to the
atmosphere is not permitted. Steam may be injected into the flare tip to reduce visible smoke.

4. Pressure Relief Health and Safety Considerations.

a. Fire Protection and Prevention. Vapors and gases must not discharge where sources of ignition could be
present.

b. Safety. Liquids should not be discharged directly to a vapor disposal system. Flare knockout drums and flares
need to be large enough to handle emergency blowdowns. Drums should be provided with relief in the event of
overpressure.

Pressure relief valves must be provided where the potential exists for overpressure in refinery processes due to
the following causes:
■ Loss of cooling water, which may greatly reduce pressure in condensers and increase the pressure in the
process unit.
■ Loss of reflux volume, which may cause a pressure drop in condensers and a pressure rise in distillation
towers because the quantity of reflux affects the volume of vapors leaving the distillation tower.
■ Rapid vaporization and pressure increase from injection of a lower boiling-point liquid including water into
a process vessel operating at higher temperatures.
■ Expansion of vapor and resultant over-pressure due to overheated process steam, malfunctioning heaters,
or fire.
■ Failure of automatic controls, closed outlets, heat exchanger failure, etc.
■ Internal explosion, chemical reaction, thermal expansion, or accumulated gases.

Maintenance is important because valves are required to function properly. The most common operating
problems are listed below.
■ Failure to open at set pressure, because of plugging of the valve inlet or outlet, or because corrosion
prevents proper operation of the disc holder and guides.
■ Failure to reseat after popping open due to fouling, corrosion, or deposits on the seat or moving parts, or
because solids in the gas stream have cut the valve disc.
■ Chattering and premature opening, because operating pressure is too close to the set point.

c. Health. Safe work practices and/or appropriate personal protective equipment may be needed to protect
against hazards during inspection, maintenance, and turnaround activities.

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4. WASTEWATER TREATMENT.

1. Description. Wastewater treatment is used for process, runoff, and sewerage water prior to discharge or
recycling. Wastewater typically contains hydrocarbons, dissolved materials, suspended solids, phenols, ammonia,
sulfides, and other compounds. Wastewater includes condensed steam, stripping water, spent caustic solutions,
cooling tower and boiler blowdown, wash water, alkaline and acid waste neutralization water, and other process-
associated water.

2. Pretreatment Operations. Pretreatment is the separation of hydrocarbons and solids from wastewater. API
separators, interceptor plates, and settling ponds remove suspended hydrocarbons, oily sludge, and solids by
gravity separation, skimming, and filtration. Some oil-in-water emulsions must be heated to assist in separating
the oil and water. Gravity separation depends on the specific gravity differences between water and immiscible oil
globules and allows free oil to be skimmed off the surface of the wastewater. Acidic wastewater is neutralized
using ammonia, lime, or soda ash. Alkaline wastewater is treated with sulfuric acid, hydrochloric acid, carbon
dioxide-rich flue gas, or sulfur.

3. Secondary Treatment Operations. After pretreatment, suspended solids are removed by sedimentation or air
flotation. Wastewater with low levels of solids may be screened or filtered. Flocculation agents are sometimes
added to help separation. Secondary treatment processes biologically degrade and oxidize soluble organic matter
by the use of activated sludge, unaerated or aerated lagoons, trickling filter methods, or anaerobic treatments.
Materials with high adsorption characteristics are used in fixed-bed filters or added to the wastewater to form a
slurry which is removed by sedimentation or filtration. Additional treatment methods are used to remove oils and
chemicals from wastewater. Stripping is used on wastewater containing sulfides and/or ammonia, and solvent
extraction is used to remove phenols.

4. Tertiary Treatment Operations. Tertiary treatments remove specific pollutants to meet regulatory discharge
requirements. These treatments include chlorination, ozonation, ion exchange, reverse osmosis, activated carbon
adsorption, etc. Compressed oxygen is diffused into wastewater streams to oxidize certain chemicals or to satisfy
regulatory oxygen-content requirements. Wastewater that is to be recycled may require cooling to remove heat
and/or oxidation by spraying or air stripping to remove any remaining phenols, nitrates, and ammonia.

5. Health and Safety Considerations.

a. Fire Protection and Prevention. The potential for fire exists if vapors from wastewater containing
hydrocarbons reach a source of ignition during treatment.

b. Health. Safe work practices and/or appropriate personal protective equipment may be needed for exposures
to chemicals and waste products during process sampling, inspection, maintenance, and turnaround activities as
well as to noise, gases, and heat.

5. COOLING TOWERS.

1. Description. Cooling towers remove heat from process water by evaporation and latent heat transfer between
hot water and air. The two types of towers are crossflow and counterflow. Crossflow towers introduce the airflow
at right angles to the water flow throughout the structure. In counterflow cooling towers, hot process water is
pumped to the uppermost plenum and allowed to fall through the tower. Numerous slats or spray nozzles located
throughout the length of the tower disperse the water and help in cooling. Air enters at the tower bottom and
flows upward against the water. When the fans or blowers are at the air inlet, the air is considered to be forced
draft. Induced draft is when the fans are at the air outlet.

2. Cooling Water. Recirculated cooling water must be treated to remove impurities and dissolved hydrocarbons.
Because the water is saturated with oxygen from being cooled with air, the chances for corrosion are increased.
One means of corrosion prevention is the addition of a material to the cooling water that forms a protective film
on pipes and other metal surfaces.

3. Health and Safety Considerations.

a. Fire Prevention and Protection. When cooling water is contaminated by hydrocarbons, flammable vapors
can be evaporated into the discharge air. If a source of ignition is present, or if lightning occurs, a fire may start.
A potential fire hazard also exists where there are relatively dry areas in induced-draft cooling towers of
combustible construction.

b. Safety. Loss of power to cooling tower fans or water pumps could have serious consequences in the operation
of the refinery. Impurities in cooling water can corrode and foul pipes and heat exchangers, scale from dissolved
salts can deposit on pipes, and wooden cooling towers can be damaged by microorganisms.

c. Health. Cooling-tower water can be contaminated by process materials and by-products including sulfur
dioxide, hydrogen sulfide, and carbon dioxide, with resultant exposures. Safe work practices and/or appropriate
personal protective equipment may be needed during process sampling, inspection, maintenance, and turnaround
activities; and for exposure to hazards such as those related to noise, water-treatment chemicals, and hydrogen
sulfide when wastewater is treated in conjunction with cooling towers.

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6. ELECTRIC POWER.

1. Description. Refineries may receive electricity from outside sources or produce their own power with generators
driven by steam turbines or gas engines. Electrical substations receive power from the utility or power plant for
distribution throughout the facility. They are usually located in nonclassified areas, away from sources of vapor or
cooling-tower water spray. Transformers, circuit breakers, and feed-circuit switches are usually located in
substations. Substations feed power to distribution stations within the process unit areas. Distribution stations
can be located in classified areas, providing that classification requirements are met. Distribution stations usually
have a liquid-filled transformer and an oil-filled or air-break disconnect device.

2. Health and Safety Considerations.

a. Fire Protection and Prevention. Generators that are not properly classified and are located too close to
process units may be a source of ignition should a spill or release occur.

b. Safety. Normal electrical safety precautions including dry footing, high-voltage warning signs, and guarding
must be taken to protect against electrocution. Lockout/tagout and other appropriate safe work practices must be
established to prevent energization while work is being performed on high-voltage electrical equipment.

c. Health. Safe work practices and/or the use of appropriate personal protective equipment may be needed for
exposures to noise, for exposure to hazards during inspection and maintenance activities, and when working
around transformers and switches that may contain a dielectric fluid which requires special handling precautions.

7. GAS AND AIR COMPRESSORS.

1. Description. Both reciprocating and centrifugal compressors are used throughout the refinery for gas and
compressed air. Air compressor systems include compressors, coolers, air receivers, air dryers, controls, and
distribution piping. Blowers are used to provide air to certain processes. Plant air is provided for the operation of
air-powered tools, catalyst regeneration, process heaters, steam-air decoking, sour-water oxidation, gasoline
sweetening, asphalt blowing, and other uses. Instrument air is provided for use in pneumatic instruments and
controls, air motors and purge connections.

2. Health and Safety Considerations.

a. Fire Protection and Prevention. Air compressors should be located so that the suction does not take in
flammable vapors or corrosive gases. There is a potential for fire should a leak occur in gas compressors.

b. Safety. Knockout drums are needed to prevent liquid surges from entering gas compressors. If gases are
contaminated with solid materials, strainers are needed. Failure of automatic compressor controls will affect
processes. If maximum pressure could potentially be greater than compressor or process-equipment design
pressure, pressure relief should be provided. Guarding is needed for exposed moving parts on compressors.
Compressor buildings should be properly electrically classified, and provisions should be made for proper
ventilation.

Where plant air is used to back up instrument air, interconnections must be upstream of the instrument air drying
system to prevent contamination of instruments with moisture. Alternate sources of instrument air supply, such
as use of nitrogen, may be needed in the event of power outages or compressor failure.

c. Health. Safe work practices and/or appropriate personal protective equipment may be needed for exposure to
hazards such as noise and during inspection and maintenance activities. The use of appropriate safeguards must
be considered so that plant and instrument air is not used for breathing or pressuring potable water systems.

8. MARINE, TANK CAR, AND TANK TRUCK LOADING AND UNLOADING.

1. Description. Facilities for loading liquid hydrocarbons into tank cars, tank trucks, and marine vessels and barges
are usually part of the refinery operations. Product characteristics, distribution needs, shipping requirements, and
operating criteria are important when designing loading facilities. Tank trucks and rail tank cars are either top- or
bottom-loaded, and vapor-recovery systems may be provided where required. Loading and unloading liquefied
petroleum gas (LPG) require special considerations in addition to those for liquid hydrocarbons.

2. Health and Safety Considerations.

a. Fire Protection and Prevention. The potential for fire exists where flammable vapors from spills or releases
can reach a source of ignition. Where switch-loading is permitted, safe practices need to be established and
followed. Bonding is used to equalize the electrical charge between the loading rack and the tank truck or tank
car. Grounding is used at truck and rail loading facilities to prevent flow of stray currents. Insulating flanges are
used on marine dock piping connections to prevent static electricity buildup and discharge. Flame arrestors
should be installed in loading rack and marine vapor-recovery lines to prevent flashback.

b. Safety. Automatic or manual shutoff systems at supply headers are needed for top and bottom loading in the

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event of leaks or overfills. Fall protection such as railings are needed for top-loading racks where employees are
exposed to falls. Drainage and recovery systems may be provided for storm drainage and to handle spills and
leaks. Precautions must be taken at LPG loading facilities not to overload or overpressurize tank cars and trucks.

c. Health. The nature of the health hazards at loading and unloading facilities depends upon the products being
loaded and the products previously transported in the tank cars, tank trucks, or marine vessels. Safe work
practices and/or appropriate personal protective equipment may be needed to protect against hazardous
exposures when loading or unloading, cleaning up spills or leaks, or when gauging, inspecting, sampling, or
performing maintenance activities on loading facilities or vapor-recovery systems.

9. TURBINES.

1. Description. Turbines are usually gas- or steam-powered and are typically used to drive pumps, compressors,
blowers, and other refinery process equipment. Steam enters turbines at high temperatures and pressures,
expands across and drives rotating blades while directed by fixed blades.

2. Health and Safety Considerations.

a. Safety. Steam turbines used for exhaust operating under vacuum should have safety relief valves on the
discharge side, both for protection and to maintain steam in the event of vacuum failure. Where maximum
operating pressure could be greater than design pressure, steam turbines should be provided with relief devices.
Consideration should be given to providing governors and overspeed control devices on turbines.

b. Health. Safe work practices and/or appropriate personal protective equipment may be needed for noise,
steam and heat exposures, and during inspection and maintenance activities.

10. PUMPS, PIPING AND VALVES.

1. Description.

a. Centrifugal and positive-displacement (i.e., reciprocating) pumps are used to move hydrocarbons, process
water, fire water, and wastewater through piping within the refinery. Pumps are driven by electric motors, steam
turbines, or internal combustion engines. The pump type, capacity, and construction materials depend on the
service for which it is used.

b. Process and utility piping distribute hydrocarbons, steam, water, and other products throughout the facility.
Their size and construction depend on the type of service, pressure, temperature, and nature of the products.
Vent, drain, and sample connections are provided on piping, as well as provisions for blanking.

c. Different types of valves are used depending on their operating purpose. These include gate valves, bypass
valves, globe and ball valves, plug valves, block and bleed valves, and check valves. Valves can be manually or
automatically operated.

2. Health and Safety Considerations.

a. Fire Protection and Prevention. The potential for fire exists should hydrocarbon pumps, valves, or lines
develop leaks that could allow vapors to reach sources of ignition. Remote sensors, control valves, fire valves,
and isolation valves should be used to limit the release of hydrocarbons at pump suction lines in the event of
leakage and/or fire.

b. Safety. Depending on the product and service, backflow prevention from the discharge line may be needed.
The failure of automatic pump controls could cause a deviation in process pressure and temperature. Pumps
operated with reduced or no flow can overheat and rupture. Pressure relief in the discharge piping should be
provided where pumps can be overpressured. Provisions may be made for pipeline expansion, movement, and
temperature changes to avoid rupture. Valves and instruments that require servicing or other work should be
accessible at grade level or from an operating platform. Operating vent and drain connections should be provided
with double-block valves, a block valve and plug, or blind flange for protection against releases.

c. Health. Safe work practices and/or appropriate personal protective equipment may be needed for exposure to
hazards such as those related to liquids and vapors when opening or draining pumps, valves, and/or lines, and
during product sampling, inspection, and maintenance activities.

11. TANK STORAGE.

1. Description. Atmospheric storage tanks and pressure storage tanks are used throughout the refinery for storage
of crudes, intermediate hydrocarbons (during the process), and finished products. Tanks are also provided for fire
water, process and treatment water, acids, additives, and other chemicals. The type, construction, capacity and
location of tanks depends on their use and materials stored.

2. Health and Safety Considerations.

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a. Fire Prevention and Protection. The potential for fire exists should hydrocarbon storage tanks be overfilled
or develop leaks that allow vapors to escape and reach sources of ignition. Remote sensors, control valves,
isolation valves, and fire valves may be provided at tanks for pump-out or closure in the event of a fire in the
tank, or in the tank dike or storage area.

b. Safety. Tanks may be provided with automatic overflow control and alarm systems, or manual gauging and
checking procedures may be established to control overfills.

c. Health. Safe work practices and/or appropriate personal protective equipment may be needed for exposure to
hazards related to product sampling, manual gauging, inspection, and maintenance activities including confined
space entry where applicable.

6. BIBLIOGRAPHY.

American Petroleum Institute. 1971. Chemistry and Petroleum for Classroom Use in Chemistry Courses. Washington, D.C.:
American Petroleum Institute.

__________. 1973. Industrial Hygiene Monitoring Manual for Petroleum Refineries and Selected Petrochemical Operations.
Manual 2700-1/79-1M. Washington, D.C.: American Petroleum Institute.

__________. 1980. Facts About Oil. Manual 4200-10/80-25M. Washington, D.C.: American Petroleum Institute.

__________. 1990. Management of Process Hazards. RP 750. Washington, D.C.: American Petroleum Institute.

__________. 1990. Inspection of Piping, Tubing, Valves and Fittings. RP 574. Washington, D.C.: American Petroleum Institute.

__________. 1991. Inspection of Fired Boilers and Heaters. RP 573. Washington, D.C.: American Petroleum Institute.

__________. 1992. Inspection of Pressure Vessels. RP 572. Washington, D.C.: American Petroleum Institute.

__________. 1992. Inspection of Pressure Relieving Devices. RP 576. Washington, D.C.: American Petroleum Institute.

__________. 1994. Fire Protection in Refineries. Sixth Edition. RP 2001. Washington, D.C.: American Petroleum Institute.

Armistead, George, Jr. 1950. Safety in Petroleum Refining and Related Industries. New York: John G. Simmons & Co., Inc.

Exxon Company, USA. 1987. Encyclopedia for the User of Petroleum Products. Lubetext D400. Houston: Exxon Company, USA.

Hydrocarbon Processing. 1988. Refining Handbook. Houston: Gulf Publishing Co.

__________. 1992. Refining Handbook. Houston: Gulf Publishing Co.

IARC. [No date given.] Occupational Exposures in Petroleum Refining. IARC Monographs, Volume 45.

Kutler, A. A. 1969. "Crude distillation." Petro/Chem Engineering. New York: John G. Simmonds & Co., Inc.

Mobil Oil Corporation. 1972. "Light Products Refining, Fuels Manufacture." Mobil Technical Bulletin, 1972. Fairfax, Virginia: Mobil
Oil Corporation.

Parmeggiani, Luigi, Technical Editor. 1983. Encyclopaedia of Occupational Health and Safety. Third Edition. Geneva:
International Labour Organization.

Shell International Petroleum Company Limited. 1983. The Petroleum Handbook. Sixth Edition. Amsterdam: Elsevier Science
Publishers B.V.

Speight, James G. 1980. The Chemistry and Terminology of Petroleum. New York: Marcel Dekker, Inc.

Vervalin, Charles H., Editor. 1985. Fire Protection Manual for Hydrocarbon Processing Plants. Volume 1, Third edition. Houston:
Gulf Publishing Co.

APPENDIX IV: 2-1. GLOSSARY.

ABSORPTION The disappearance of one substance into another so that the absorbed substance loses its identifying
characteristics, while the absorbing substance retains most of its original physical aspects. Used in refining to selectively
remove specific components from process streams.

ACID TREATMENT A process in which unfinished petroleum products such as gasoline, kerosene, and lubricating oil stocks
are treated with sulfuric acid to improve color, odor, and other properties.

ADDITIVE Chemicals added to petroleum products in small amounts to improve quality or add special characteristics.

ADSORPTION Adhesion of the molecules of gases or liquids to the surface of solid materials.

AIR FIN COOLERS A radiator-like device used to cool or condense hot hydrocarbons; also called fin fans.

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ALICYCLIC HYDROCARBONS Cyclic (ringed) hydrocarbons in which the rings are made up only of carbon atoms.

ALIPHATIC HYDROCARBONS Hydrocarbons characterized by open-chain structures: ethane, butane, butene, acetylene, etc.

ALKYLATION A process using sulfuric or hydrofluoric acid as a catalyst to combine olefins (usually butylene) and isobutane to
produce a high-octane product known as alkylate.

API GRAVITY An arbitrary scale expressing the density of petroleum products.

AROMATIC Organic compounds with one or more benzene rings.

ASPHALTENES The asphalt compounds soluble in carbon disulfide but insoluble in paraffin naphthas.

ATMOSPHERIC TOWER A distillation unit operated at atmospheric pressure.

BENZENE An unsaturated, six-carbon ring, basic aromatic compound.

BLEEDER VALVE A small-flow valve connected to a fluid process vessel or line for the purpose of bleeding off small quantities
of contained fluid. It is installed with a block valve to determine if the block valve is closed tightly.

BLENDING The process of mixing two or more petroleum products with different properties to produce a finished product
with desired characteristics.

BLOCK VALVE A valve used to isolate equipment.

BLOWDOWN The removal of hydrocarbons from a process unit, vessel, or line on a scheduled or emergency basis by the use
of pressure through special piping and drums provided for this purpose.

BLOWER Equipment for moving large volumes of gas against low-pressure heads.

BOILING RANGE The range of temperature (usually at atmospheric pressure) at which the boiling (or distillation) of a
hydrocarbon liquid commences, proceeds, and finishes.

BOTTOMS Tower bottoms are residue remaining in a distillation unit after the highest boiling-point material to be distilled has
been removed. Tank bottoms are the heavy materials that accumulate in the bottom of storage tanks, usually comprised of oil,
water, and foreign matter.

BUBBLE TOWER A fractionating (distillation) tower in which the rising vapors pass through layers of condensate, bubbling
under caps on a series of plates.

CATALYST A material that aids or promotes a chemical reaction between other substances but does not react itself. Catalysts
increase reaction speeds and can provide control by increasing desirable reactions and decreasing undesirable reactions.

CATALYTIC CRACKING The process of breaking up heavier hydrocarbon molecules into lighter hydrocarbon fractions by use
of heat and catalysts.

CAUSTIC WASH A process in which distillate is treated with sodium hydroxide to remove acidic contaminants that contribute
to poor odor and stability.

CHD UNIT See Hydrodesulfurization.

COKE A high carbon-content residue remaining from the destructive distillation of petroleum residue.

COKING A process for thermally converting and upgrading heavy residual into lighter products and by-product petroleum
coke. Coking also is the removal of all lighter distillable hydrocarbons that leaves a residue of carbon in the bottom of units or
as buildup or deposits on equipment and catalysts.

CONDENSATE The liquid hydrocarbon resulting from cooling vapors.

CONDENSER A heat-transfer device that cools and condenses vapor by removing heat via a cooler medium such as water or
lower-temperature hydrocarbon streams.

CONDENSER REFLUX Condensate that is returned to the original unit to assist in giving increased conversion or recovery.

COOLER A heat exchanger in which hot liquid hydrocarbon is passed through pipes immersed in cool water to lower its
temperature.

CRACKING The breaking up of heavy molecular weight hydrocarbons into lighter hydrocarbon molecules by the application of
heat and pressure, with or without the use of catalysts.

CRUDE ASSAY A procedure for determining the general distillation and quality characteristics of crude oil.

CRUDE OIL A naturally occurring mixture of hydrocarbons that usually includes small quantities of sulfur, nitrogen, and
oxygen derivatives of hydrocarbons as well as trace metals.

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OSHA TECHNICAL MANUAL - SECTION IV: CHAPTER 2

CYCLE GAS OIL Cracked gas oil returned to a cracking unit.

DEASPHALTING Process of removing asphaltic materials from reduced crude using liquid propane to dissolve nonasphaltic
compounds.

DEBUTANIZER A fractionating column used to remove butane and lighter components from liquid streams.

DE-ETHANIZER A fractionating column designed to remove ethane and gases from heavier hydrocarbons.

DEHYDROGENATION A reaction in which hydrogen atoms are eliminated from a molecule. Dehydrogenation is used to
convert ethane, propane, and butane into olefins (ethylene, propylene, and butenes).

DEPENTANIZER A fractionating column used to remove pentane and lighter fractions from hydrocarbon streams.

DEPROPANIZER A fractionating column for removing propane and lighter components from liquid streams.

DESALTING Removal of mineral salts (most chlorides, e.g., magnesium chloride and sodium chloride) from crude oil.

DESULFURIZATION A chemical treatment to remove sulfur or sulfur compounds from hydrocarbons.

DEWAXING The removal of wax from petroleum products (usually lubricating oils and distillate fuels) by solvent absorption,
chilling, and filtering.

DIETHANOLAMINE A chemical (C4H11O2N) used to remove H2S from gas streams.

DISTILLATE The products of distillation formed by condensing vapors.

DOWNFLOW Process in which the hydrocarbon stream flows from top to bottom.

DRY GAS Natural gas with so little natural gas liquids that it is nearly all methane with some ethane.

FEEDSTOCK Stock from which material is taken to be fed (charged) into a processing unit.

FLASHING The process in which a heated oil under pressure is suddenly vaporized in a tower by reducing pressure.

FLASH POINT Lowest temperature at which a petroleum product will give off sufficient vapor so that the vapor-air mixture
above the surface of the liquid will propagate a flame away from the source of ignition.

FLUX Lighter petroleum used to fluidize heavier residual so that it can be pumped.

FOULING Accumulation of deposits in condensers, exchangers, etc.

FRACTION One of the portions of fractional distillation having a restricted boiling range.

FRACTIONATING COLUMN Process unit that separates various fractions of petroleum by simple distillation, with the column
tapped at various levels to separate and remove fractions according to their boiling ranges.

FUEL GAS Refinery gas used for heating.

GAS OIL Middle-distillate petroleum fraction with a boiling range of about 350°-750° F, usually includes diesel fuel, kerosene,
heating oil, and light fuel oil.

GASOLINE A blend of naphthas and other refinery products with sufficiently high octane and other desirable characteristics to
be suitable for use as fuel in internal combustion engines.

HEADER A manifold that distributes fluid from a series of smaller pipes or conduits.

HEAT As used in the Health Considerations paragraphs of this document, heat refers to thermal burns for contact with hot
surfaces, hot liquids and vapors, steam, etc.

HEAT EXCHANGER Equipment to transfer heat between two flowing streams of different temperatures. Heat is transferred
between liquids or liquids and gases through a tubular wall.

HIGH-LINE OR HIGH-PRESSURE GAS High-pressure (100 psi) gas from cracking unit distillate drums that is compressed
and combined with low-line gas as gas absorption feedstock.

HYDROCRACKING A process used to convert heavier feedstock into lower-boiling, higher-value products. The process
employs high pressure, high temperature, a catalyst, and hydrogen.

HYDRODESULFURIZATION A catalytic process in which the principal purpose is to remove sulfur from petroleum fractions in
the presence of hydrogen.

HYDROFINISHING A catalytic treating process carried out in the presence of hydrogen to improve the properties of low
viscosity-index naphthenic and medium viscosity-index naphthenic oils. It is also applied to paraffin waxes and microcrystalline

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OSHA TECHNICAL MANUAL - SECTION IV: CHAPTER 2

waxes for the removal of undesirable components. This process consumes hydrogen and is used in lieu of acid treating.

HYDROFORMING Catalytic reforming of naphtha at elevated temperatures and moderate pressures in the presence of
hydrogen to form high-octane BTX aromatics for motor fuel or chemical manufacture. This process results in a net production of
hydrogen and has rendered thermal reforming somewhat obsolete. It represents the total effect of numerous simultaneous
reactions such as cracking, polymerization, dehydrogenation, and isomerization.

HYDROGENATION The chemical addition of hydrogen to a material in the presence of a catalyst.

INHIBITOR Additive used to prevent or retard undesirable changes in the quality of the product, or in the condition of the
equipment in which the product is used.

ISOMERIZATION A reaction that catalytically converts straight-chain hydrocarbon molecules into branched-chain molecules
of substantially higher octane number. The reaction rearranges the carbon skeleton of a molecule without adding or removing
anything from the original material.

ISO-OCTANE A hydrocarbon molecule (2,2,4-trimethylpentane) with excellent antiknock characteristics on which the octane
number of 100 is based.

KNOCKOUT DRUM A vessel wherein suspended liquid is separated from gas or vapor.

LEAN OIL Absorbent oil fed to absorption towers in which gas is to be stripped. After absorbing the heavy ends from the gas,
it becomes fat oil. When the heavy ends are subsequently stripped, the solvent again becomes lean oil.

LOW-LINE or LOW-PRESSURE GAS Low-pressure (5 psi) gas from atmospheric and vacuum distillation recovery systems
that is collected in the gas plant for compression to higher pressures.

NAPHTHA A general term used for low boiling hydrocarbon fractions that are a major component of gasoline. Aliphatic
naphtha refers to those naphthas containing less than 0.1% benzene and with carbon numbers from C3 through C16. Aromatic
naphthas have carbon numbers from C6 through C16 and contain significant quantities of aromatic hydrocarbons such as
benzene (>0.1%), toluene, and xylene.

NAPHTHENES Hydrocarbons (cycloalkanes) with the general formula CnH2n, in which the carbon atoms are arranged to form
a ring.

OCTANE NUMBER A number indicating the relative antiknock characteristics of gasoline.

OLEFINS A family of unsaturated hydrocarbons with one carbon-carbon double bond and the general formula CnH2n.

PARAFFINS A family of saturated aliphatic hydrocarbons (alkanes) with the general formula CnH2n+2.

POLYFORMING The thermal conversion of naphtha and gas oils into high-quality gasoline at high temperatures and pressure
in the presence of recirculated hydrocarbon gases.

POLYMERIZATION The process of combining two or more unsaturated organic molecules to form a single (heavier) molecule
with the same elements in the same proportions as in the original molecule.

PREHEATER Exchanger used to heat hydrocarbons before they are fed to a unit.

PRESSURE-REGULATING VALVE A valve that releases or holds process-system pressure (that is, opens or closes) either by
preset spring tension or by actuation by a valve controller to assume any desired position between fully open and fully closed.

PYROLYSIS GASOLINE A by-product from the manufacture of ethylene by steam cracking of hydrocarbon fractions such as
naphtha or gas oil.

PYROPHORIC IRON SULFIDE A substance typically formed inside tanks and processing units by the corrosive interaction of
sulfur compounds in the hydrocarbons and the iron and steel in the equipment. On exposure to air (oxygen) it ignites
spontaneously.

QUENCH OIL Oil injected into a product leaving a cracking or reforming heater to lower the temperature and stop the
cracking process.

RAFFINATE The product resulting from a solvent extraction process and consisting mainly of those components that are least
soluble in the solvents. The product recovered from an extraction process is relatively free of aromatics, naphthenes, and other
constituents that adversely affect physical parameters.

REACTOR The vessel in which chemical reactions take place during a chemical conversion type of process.

REBOILER An auxiliary unit of a fractionating tower designed to supply additional heat to the lower portion of the tower.

RECYCLE GAS High hydrogen-content gas returned to a unit for reprocessing.

REDUCED CRUDE A residual product remaining after the removal by distillation of an appreciable quantity of the more
volatile components of crude oil.

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REFLUX The portion of the distillate returned to the fractionating column to assist in attaining better separation into desired
fractions.

REFORMATE An upgraded naphtha resulting from catalytic or thermal reforming.

REFORMING The thermal or catalytic conversion of petroleum naphtha into more volatile products of higher octane number.
It represents the total effect of numerous simultaneous reactions such as cracking, polymerization, dehydrogenation, and
isomerization.

REGENERATION In a catalytic process the reactivation of the catalyst, sometimes done by burning off the coke deposits
under carefully controlled conditions of temperature and oxygen content of the regeneration gas stream.

SCRUBBING Purification of a gas or liquid by washing it in a tower.

SOLVENT EXTRACTION The separation of materials of different chemical types and solubilities by selective solvent action.

SOUR GAS Natural gas that contains corrosive, sulfur-bearing compounds such as hydrogen sulfide and mercaptans.

STABILIZATION A process for separating the gaseous and more volatile liquid hydrocarbons from crude petroleum or
gasoline and leaving a stable (less-volatile) liquid so that it can be handled or stored with less change in composition.

STRAIGHT-RUN GASOLINE Gasoline produced by the primary distillation of crude oil. It contains no cracked, polymerized,
alkylated, reformed, or visbroken stock.

STRIPPING The removal (by steam-induced vaporization or flash evaporation) of the more volatile components from a cut or
fraction.

SULFURIC ACID TREATING A refining process in which unfinished petroleum products such as gasoline, kerosene, and
lubricating oil stocks are treated with sulfuric acid to improve their color, odor, and other characteristics.

SULFURIZATION Combining sulfur compounds with petroleum lubricants.

SWEETENING Processes that either remove obnoxious sulfur compounds (primarily hydrogen sulfide, mercaptans, and
thiophens) from petroleum fractions or streams, or convert them, as in the case of mercaptans, to odorless disulfides to
improve odor, color, and oxidation stability.

SWITCH LOADING The loading of a high static-charge retaining hydrocarbon (i.e., diesel fuel) into a tank truck, tank car, or
other vessel that has previously contained a low-flash hydrocarbon (gasoline) and may contain a flammable mixture of vapor
and air.

TAIL GAS The lightest hydrocarbon gas released from a refining process.

THERMAL CRACKING The breaking up of heavy oil molecules into lighter fractions by the use of high temperature without
the aid of catalysts.

TURNAROUND A planned complete shutdown of an entire process or section of a refinery, or of an entire refinery to perform
major maintenance, overhaul, and repair operations and to inspect, test, and replace process materials and equipment.

VACUUM DISTILLATION The distillation of petroleum under vacuum which reduces the boiling temperature sufficiently to
prevent cracking or decomposition of the feedstock.

VAPOR The gaseous phase of a substance that is a liquid at normal temperature and pressure.

VISBREAKING Viscosity breaking is a low-temperature cracking process used to reduce the viscosity or pour point of straight-
run residuum.

WET GAS A gas containing a relatively high proportion of hydrocarbons that are recoverable as liquids.

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 Revamping HDS Units
to Meet High Quality Diesel Specifications

by

F. Emmett Bingham
Haldor Topsoe, Inc., California, USA

and

Preben Christensen
Haldor Topsøe A/S, Lyngby, Denmark

Presented at the Asian Pacific Refining Technology Conference, 8 - 10 March, 2000,

Kuala Lumpur, Malaysia
 2

Abstract
Environmental legislation continues to favor more stringent global diesel quality
specifications. The U.S. and the European Union have adopted low sulfur diesel
specifications, and specifications with ultra low sulfur contents are either mandated or under
serious consideration after year 2005. Historically in Asia, air quality has tended to
deteriorate, perhaps as a consequence of very rapid industrial growth. However as Asia
becomes richer, the demand for future cleaner burning fuels is expected to become a
significant issue.

More stringent diesel quality criteria for properties other than sulfur content are also expected
in the future. The automobile industry, in their World Wide Fuel Charter issued in early 1999,
has recommended rather stringent limits on, aromatics, poly-aromatics, cetane number/index,
and back-end distillation temperatures. The European Union is expected to finalize their year
2005 fuel specifications during year 2000 and the specifications in the World Wide Fuel
Charter are anticipated to have a strong influence on the final criteria.

Depending on which of the above diesel quality specifications are adopted, it is likely that
meeting proposed diesel aromatics and burning quality criteria will be more challenging than
meeting the sulfur specifications. More sophisticated process technologies and catalysts will
be needed to meet these specifications.

Topsoe has been a pioneer in the design of units for diesel desulfurization and aromatics
saturation to produce cleaner burning fuels. This paper presents examples of Topsoe
commercial experience and provides a detailed illustration of the approach one refiner in the
United States will take to economically revamp an existing HDS unit to produce ultra low
sulfur, low aromatics diesel.
 3

Changes in Diesel Specifications and Demand

In recent years, the development and use of “environmentally friendly” fuels has been of high
priority throughout the world.

In Europe, the specifications adopted by EU for year 2000 concerns sulfur content, density,
Poly-Aromatic Hydrocarbons (PAH) content, cetane number and 95% ASTM distillation
point. Year 2005 concerns further reduction in the diesel sulfur specification. The remaining
specifications for the year 2005 are expected to be settled during year 2000 and these will
most likely be more stringent than those adopted for year 2000. This expectation is based on
the fact that, in January 1999, the automobile and engine manufacturing industry in USA,
Japan and Europe issued a World Wide Fuel Charter, that specified stringent requirements for
a large number of diesel fuel parameters.

The year 2000 diesel specifications in European Union and the World Wide Fuel Charter are
given in Table 1.

Table 1: Future Diesel Specifications

Specification EU Year 2000 Fuel Charter
Cetane Number 51 (min) 55 (min)
Cetane Index NA 52 (min)
Density @ 15°C, g/cm3 0.845 (max) 0.840 (max)
Distillation
90% Boiling Point, °C NA 320 (max)
95% Boiling Point, °C 360 (max) 340 (max)
Final Boiling Point, °C NA 350 (max)
Poly-aromatic Hydrocarbons, wt% 11 (max) 2.0 (max)
Total Aromatics Content, wt% NA 15 (max)
Sulfur Content, wppm 350 * (max) 30 (max)
*From Year 2005 the European Union has adopted a sulfur content of 50 wppm

The United States is currently debating diesel specifications with most recent speculation that
both EPA and CARB will slash allowable diesel fuel sulfur levels by year 2007 (EPA target:
5-40 ppm sulfur). Additionally, specifications for lower aromatics content and increased
diesel cetane number are expected.

Current diesel sulfur specifications in Asia vary over a wide range from 0.5 wt% to less than
0.05 wt% . The average sulfur level, estimated to be 0.19 wt% for year 2000, will continue to
drop due to the definite movement away from 0.5% grades toward 0.05% grades. Japan and
Korea, currently producing diesel at or below 0.05 wt% sulfur, are expected to tighten their
specifications to closely follow the specifications set by the EU and the United States.
Thailand, Malaysia, Singapore, and the Philippines will have 0.05% diesel sulfur
specifications by 2000 – 2004, and will most likely consider further future reduction. India
has recently introduced a sulfur specification of 0.2 wt% and is already now considering
further reductions to 0.05 wt% sulfur. Japan and Korea will also most likely follow the
 4

specifications for reduced aromatics and T95 promulgated in the EU and the United States.
However, there is less consensus among other Asian countries concerning the adoption of
these specifications.

At the same time diesel fuel specifications are tightening, the demand for diesel is growing in
Asia as well as in Europe. Although the growth rate for middle distillate in Asia was
relatively stagnant from 1997-1999, the forecast future demand, according to Trans-Energy
Research Associates, Inc. will be equivalent to the pre 1997 rate of approximately 5-8% per
year.

In Europe, the demand for home heating oil and fuel oil cutter stocks, two of the other
common uses of diesel boiling-range material, is decreasing. Consequently, more diesel
boiling-range material is becoming available for blending into the diesel pool. Unfortunately,
these lower quality stocks will require severe hydrotreatment to meet the future specifications
for diesel fuel.

The refining industry is therefore facing a double-edged challenge: to meet new, more
stringent specifications for diesel product, while simultaneously producing more diesel
product from lower quality feedstocks. The combination of these factors places a heavy
burden on the refiner’s hydroprocessing capabilities. As a result, increased new hydrotreating
capacity and revamp of existing facilities will be needed to meet the future diesel
specifications.

Meeting Future Specifications

Depending on the diesel specifications adopted, the issues facing the Asian refiner in the
future could be very similar to those faced by the EU refiner today. Properties being
considered for future diesel specifications in the EU can serve as basis of discussion for this
paper.

Sulfur Reduction

The current diesel sulfur specification in the EU is 350 wppm, and for year 2005 a further
reduction to 50 wppm will be mandatory. Germany is planning to offer tax incentives to
phase in diesel with 50 wppm sulfur from November 2001 and 10 wppm sulfur by November
2003. In Asia, the estimated average, year 2000, diesel sulfur content is 0.19 wt%. However,
the average diesel sulfur content is expected to rapidly approach 0.05 wt%, with lots of
pressure for continued reduction.

The impact of reducing diesel sulfur content from 2500 wppm to 10 wppm for the same feed
is illustrated in Table 2. The data, summarized in the table, was calculated using Topsøe’s
kinetic model, for a typical diesel hydrotreater using a CoMo type of catalyst. The left
column shows the target diesel sulfur content in wppm. The right-hand column shows how
much additional catalyst, having the same currently available activity, is required to reach the
lower sulfur targets while maintaining a constant cycle length. All other operating conditions
remain unchanged.
 5

Table 2: Impact of Diesel Sulfur Reduction

Diesel Sulfur Relative Catalyst Volume to
wppm Maintain Constant Cycle
2500 1.0
500 1.5
350 1.8
50 3.8
10 5.3

Table 2 can be used to predict the impact of future Asian regulations for reduced diesel sulfur.

New diesel HDS catalysts are being developed to provide sufficient additional activity to
allow the refiner to reduce the diesel sulfur content from 2500 wppm to 500 wppm with only
moderate impact on the cycle length. However it is very unlikely that catalyst HDS activities
will be increased sufficiently in the near future to meet the 10 – 50 wppm sulfur specifications
for Germany or the EU. Measures other than a simple catalyst replacement must therefore be
considered to meet these specifications.

Process parameters that can be modified to assist in meeting deeper desulfurization are
discussed in more detail in the process modification section.

Aromatics Reduction

The new EU diesel specifications for year 2000 and most likely also for year 2005 will
include a limit on the content of poly-aromatic hydrocarbons (PAH’s). The maximum content
of PAH’s for the year 2000 diesel has been set at 11 wt%. This can be met by most existing
hydrotreaters unless the feed being processed contains a very high amount of catalytically
cracked material. The PAH content increases significantly if LCO is blended in the feedstock.
The conversion of PAH’s is moderate in hydrotreaters having relatively low hydrogen partial
pressures. Therefore for units with hydrogen partial pressures in the range of 30 bar, the feed
blend should contain no more than approx. 30 wt% PAH’s in order to meet the year 2000
specification of 11% PAH.

No limit for PAH content for year 2005 diesel, has as yet been established by the EU.
However, it is likely that PAH content will be reduced considerably compared to the year
2000 specification and values as low as 1 wt% have been mentioned. Although PAH’s react
quite readily, their conversion is thermodynamically limited and can only be improved by
conversion of mono-aromatics compounds. As a result, it will be difficult for hydrotreaters
operating at typical design pressures with conventional base-metal catalysts, to meet the 2
wt% specification for PAH’s proposed by the automobile and engine manufacturers.

To demonstrate the thermodynamic limitations on PAH conversion, pilot plant tests at

operating conditions typical of many existing hydrotreaters, were conducted on a straight-run
gas oil containing 32 wt% total aromatics and 16 wt% PAH’s. The hydrogen partial pressure
was increased from 30 bars to 60 bars and the product PAH content was measured. The
results of these tests are shown in Figure 1.
 6

Figure 1: Saturation of PAH’s as Function of Temperature and Pressure

6
Increasing Hydrogen Pressure
5
Wt%, PAH

0
310 320 330 340 350 360 370 380 390 400 410
Temperature, °C

As seen in Figure 1, it is barely possible to reach 2 wt% PAH content in the diesel product,
even at the highest hydrogen partial pressure. Lowering space velocity lowers the required
reaction temperature and the thermodynamic equilibrium becomes more favorable for
aromatic saturation. Even so, as temperature is increased to compensate for catalyst
deactivation it becomes increasingly difficult to maintain the low PAH content. Achieving a
diesel product with 2 wt% PAH content will be even more demanding for the refiner
processing feeds containing cracked stocks.

In general, diesel hydrotreaters, operating at hydrogen partial pressures of 35-50 bars and 1.5-
2.0 hr-1space velocity, with base-metal catalysts, can produce diesel with a PAH content of 2-
6 wt% when processing 100% straight-run feed. However, at the same operating conditions,
the product PAH content will increase to 4-8 wt%, if the feed contains 20-30% LCO.

Many existing hydrotreaters may be capable of achieving a future PAH specification of 6 wt%
or higher. However, if a future PAH specification of 4 wt% or lower must be met, the
addition of a second stage hydrodearomatization reactor to the existing diesel hydrotreater or
the installation of a high pressure single stage unit will likely be required. When the final
specification for PAH content is settled, refiners will have to carefully evaluate their existing
hydrotreaters and determine how best to integrate new equipment with the existing unit.
Topsøe has the engineering experience to complete feasibility studies to help the refiner make
these evaluations.

In addition to reduction of the PAH content, the reduction in total aromatics content in diesel
may also become an issue in order to meet the anticipated new diesel specifications for
density, cetane number or cetane index. Eventually a specification on the total aromatics
content as suggested in the World Wide Fuel Charter may be introduced.

Reduction of the total aromatics content is much more difficult than reduction of the PAH
content, because saturation of mono-aromatics to naphthenes is much more difficult than
saturation of poly-aromatics to mono-aromatics. Processing a typical straight-run gas oil in a
single-stage hydrotreater, to reduce the total aromatics content to 15 wt% as required by the
World Wide Fuel Charter, will require rather severe reactor conditions including a hydrogen
 7

partial pressure greater than 70 bars. Processing feed blends containing cracked material, to
meet the 15 wt% total aromatics specification, will require even more severe conditions
making the single-stage approach less economically attractive.

For an existing moderate pressure diesel hydrotreater using base-metal catalyst (NiMo or
CoMo), the reduction in total aromatics content is very limited, due to the relatively low
hydrogenation activity of the base metal catalyst. Adding a second stage with a high activity
noble metal catalyst offers the best solution for achieving the required aromatics reduction at
moderate hydrogen partial pressures. A separate second stage is necessary because nitrogen
and sulfur containing compounds must be removed in the first stage, as they are temporary
poisons to the noble metal catalyst.

Improvement of Cetane Number/Cetane Index

Cetane number and cetane index are measures of the ignition quality of diesel fuel. Cetane
number is determined by running the fuel in a test engine and cetane index is calculated based
on measured values of the density and ASTM distillation.

Both the cetane number and the cetane index are strongly dependent on the amount and type
of aromatic components in the feedstock. Saturation of aromatics will therefore improve the
cetane number/index considerably. However the improvement attainable is also strongly
dependent on the amount of sulfur and nitrogen species in the feed and the nature of the
feedstock.

Deep hydrodesulfurization will in itself improve the cetane number/index, and further
increase in cetane number can be achieved through the use of additives. However, it is
possible that both cetane number and cetane index specifications will be adopted to limit the
use of additives to boost the cetane number. If this happens, the refiner will be forced to
reduce the total aromatics content of their diesel. Again a two-stage HDS/HDA process will
be an effective way of achieving this.

Reduction of 95% Distillation Point

Future diesel specifications may call for a reduction of up to 20°C below the year EU 2000
T95 specification of 360°C. Similarly future reductions in tail end distillation temperature
reductions are being considered in some parts of Asia. The simplest method of reducing the
95% distillation point of the diesel product is to adjust the draws of the crude distillation
tower to reduce the diesel endpoint. Reduction of the 95% distillation point will have the
following positive effects:

• Density will be reduced

• Most refractive sulfur components will be removed from the feedstock
• Concentration of heavy aromatics components will be reduced
• Cold flow properties of the diesel will be improved

However, this will leave the refiner with a heavy gas oil fraction boiling between say 340-
360°C representing 10–15 vol% of the diesel previously produced. One solution to bring this
material back in the future diesel pool would be to crack this narrow boiling fraction to lighter
material. The chemical nature of the hydrocarbons in this narrow boiling fraction, however,
limits what can be achieved by cracking this material. The molecules present in this narrow
boiling fraction will typically have between 20 and 23 carbon atoms. If these molecules are
 8

cracked in the middle, the product fractions will have between 10 and 12 carbon atoms, which
boil in the heavy naphtha/light kerosene range. If the molecules in the narrow boiling fraction
are cracked unsymmetrically, one fraction will boil in the diesel range and the other will be
either light naphtha or gas. Cracking the narrow boiling fraction will therefore result in a
major naphtha production as well as formation of significant amounts of hydrocarbon gasses
and only a limited portion will remain in the diesel range. Topsøe therefore believes that
means other than cracking of the narrow boiling fraction must be found.

Density Reduction

The extent to which the density is reduced during deep hydrodesulfurization depends on the
feedstock that is processed and on the operating conditions. For a typical EU refinery, the
density is reduced by 0.01-0.02 g/cm3 as a result of sulfur removal and saturation of di- and
tri-aromatic compounds. Together with the reduction in density of 0.003-0.005 g/cm3
obtained by the anticipated reduction in 95% distillation temperature, meeting the year 2005
density specification recommended by the automobile industry, will be possible for many
existing diesel hydrotreaters operating to meet 50 wppm of sulfur.

If additional density reduction is required, saturation of mono-aromatics compounds will be

necessary but difficult to achieve at the pressures normally employed in diesel hydrotreaters.
In this case a two-stage HDS/HDA unit is a good approach for achieving further density
reduction. Figure 2 shows the relationship between density reduction and aromatics saturation
obtained using Topsøe’s aromatic saturation catalysts TK-907 and TK-908. There is a spread
in the data, but roughly speaking a reduction in total aromatic content by 10% (absolute) will
result in approximately 0.007 g/cm3 reduction in density.

Figure 2: Density Improvement vs. Aromatics Removal

0
Density Improvement

-0.005
TK-907, TK-908
-0.01

-0.015

-0.02

-0.025

-0.03
0 10 20 30 40 50
Aromatics Removed, vol%
 9

Process Changes for Meeting Future Stringent Regulations

The primary process changes for meeting future sulfur regulations in an existing single stage
unit can be categorized into either reactor modifications, or recycle gas loop modifications.
Most units revamped to meet future product sulfur targets will employ one or both of these
modifications.

Reactor Modifications

Modifications associated with the reactor can include:

• Use of more active catalyst in the existing reactor

• Increasing the catalyst quantity by dense loading the existing reactor
• Increasing the catalyst quantity by addition of a new reactor
• Improving the performance of the existing reactor by installing state of the art reactor
internals.

Catalyst Options

Since the first introduction of the hydrotreating process in the 1950’s, catalyst manufacturers
have made significant improvements in catalyst activity. Today’s high activity hydrotreating
catalyst is about eight times more active than the first generation hydroprocessing catalyst.

Topsøe first introduced a full range of hydrotreating catalysts for refining applications in the
early 1980’s. Table 3 lists our Cobalt Moly catalyst typically used in hydrodesulfurization
applications and the relative improvement over the years. Topsøe’s most active CoMo
catalyst, TK-574, has about twice the activity of our first generation catalyst, TK-550,
introduced in early 1980’s. By applying TK-574 instead of older types of CoMo catalysts will
enable a typical Asian refiner to reduce the sulfur level from say 2500 wppm to 350-500
wppm with only minor influence on the cycle length.

Table 3: TK-Catalysts 500 Series: CoMo

Name HDS Activity
TK-550 Base
TK-554 Base * 1.5
TK-554+ Base * 1.7
TK-574 Base * 2.0

Catalyst Loading

Catalyst quantity can be increased by 15% by dense loading an existing reactor. This
corresponds to a 3°C to 4°C reduction in start of run temperatures. Hardware limitations such
as reactor internals design and compressor pressure drop limitations may prevent some refiners
from dense loading their catalyst.

LHSV

Adding a reactor or replacing an existing reactor is the normal approach taken when
revamping a unit to meet more stringent product requirements. Assuming all other operating
 10

conditions remain unchanged, doubling the catalyst volume results in a 20°C reduction in
average temperature. This reduction can be used to offset the increased severity required for
lower product sulfur. Increasing the catalyst volume has a double effect on performance:

• The start of run temperature is lower resulting in an increase in available temperature span
from start to end of run.
• The deactivation rate is lower due to the lower start of run temperature and larger quantity
of catalyst.

If an additional reactor is to be needed, the refiner has several choices as to the location of the
new reactor. The new reactor may be located in series, either up stream or down stream of the
existing reactor, or in parallel to the existing reactor. The decision will largely depend on the
limitations of the existing equipment. Reactor efficiencies based on mass velocity as well as
reactor pressure drop will play an important role in the choice of series or parallel configuration.
The recycle compressor capability should also be considered when making this decision. Other
factors that should be considered include quantity of catalyst to be added, reactor temperature
control requirements, heater requirements, etc.

Topsøe has revamped several hydrotreaters requiring the modification of, addition to, or
replacement of the existing reactor, and therefore has the experience to assist the refiner in
making the best decisions for modifying his particular unit.

Reactor Internals

Reactor internals play a key role in facilitating the contact of reactants with the catalyst. Poor
distribution of the reactants over the catalyst can contribute to channeling through the catalyst
bed resulting in inefficient utilization of the catalyst, development of hot spots, and catalyst
deactivation due to coke formation.

Figure 3 is a simple representation of the significant effect poor distribution can have on deep
diesel desulfurization.

Figure 3: Effect of liquid distribution on product sulfur

FEED Sulfur = 1.5 wt%

Unreated Liquid (wt ppm)
Sulfur Contribution from

350
300
250
200
150
100
50
0
0.5 1 2
% Untreated Liquid
Bypassing
Poor Flow
 11

The schematic on the left shows a poorly performing distributor causing catalyst bypassing.
The bar chart on the right shows the contribution of sulfur in the total product caused by
increasing levels of bypassing of the feed. As demonstrated by this figure, improved
distribution to eliminate bypassing will be mandatory as the product sulfur target gets lower
and lower. Reactor internals in most hydroprocessing units are out-dated and may not be able
to provide the high performance efficiency needed to fully utilize today’s high performance
catalysts.

Topsøe realized this dilemma and undertook and extensive development program in 1990
using several different pilot plants and test apparatus. From this test work computer models
were developed and used to design new state of the art reactor internals. These models can
also be used to evaluate the performance of existing reactor internals.

The most critical piece of hardware for evenly distributing reactants across the catalyst is the
liquid distribution tray. Topsøe’s “Vapor-Lift” distributor tray is a state-of-the-art design for
use at the top of the reactor or below a quench section. This design was first tested in our
pilot plants in 1996-1997 and the first commercial installation was completed in 1998. To
date we sold more than 40 trays. The commercial performance of the “Vapor Lift” Tray has
been excellent.

The replacement of the impingement mixers and bubble cap trays at the Syncrude Heavy
Coker Gas Oil Hydrotreater with Topsøe’s Vortex mixing chambers and “Vapor Lift” trays
demonstrates typical commercial results with Topsøe reactor internals. Reactor radial
temperature gradients are indicative of the performance of the reactor internals. Figure 4
shows reactor temperature profiles for the existing reactor internals. Figure 5 shows
temperature profiles after the impingement mixers and bubble cap trays were replaced with
the Topsøe Vortex mixers and “Vapor Lift” trays. These figures show the reactor
temperatures at a variety of radial locations from the top of the first bed (percent catalyst = 0)
to the bottom of the third bed (percent catalyst = 100). The radial temperature gradients were
reduced from over 40°C to an average radial temperature gradient of 2 °C with the Topsøe
internals, demonstrating the superiority of the Vortex mixers and “Vapor Lift” Trays.

It should be noted that at the time the trays were replaced, the reactor thermometry was also
changed to include Gayesco ‘Flexible-type’ thermocouples. This resulted in an increase in the
thermocouple coverage at any given elevation. The data from the dense pattern of flexible
thermocouples further establishes the temperature uniformity attained with the Topsøe tray.
 12

Figure 4: Temperature profile in existing 3 beds with bubble cap trays

+90

+80
Reactor Temperature, C

+70

+60

+50

+40
+30
+20
+10
Figure 5: Temperature profile in the 3 beds with Topsøe “Vapor Lift” trays
Base
0 10 20 30 40 50 60 70 80 90 100

Percent of Catalyst

+80
Reactor Temperature, C

+70
+60

+50
+40

+30
+20
+10

Base
0 10 20 30 40 50 60 70 80 90 100

Percent of Catalyst
 13

Recycle Gas Loop Modifications

Modifications to the recycle gas loop for meeting future regulations include:
• Increasing the hydrogen concentration of the gas
• Reducing the hydrogen sulfide concentration of the gas
• Increasing the recycle gas to oil ratio.

Recycle Gas H2 Concentration

Increasing the recycle gas H2 concentration will increase the reactor hydrogen partial pressure.
This is accomplished by purging recycle gas or by increasing the hydrogen concentration in
the make-up gas. Increasing the hydrogen partial pressure reduces the reactor start of run
temperature and also reduces the rate of catalyst deactivation. If purge of recycle gas is used,
the purge gas can be sent to a membrane separation unit or PSA unit in order to recover the
hydrogen, which can then be recycled.

In new units, higher hydrogen partial pressure can be achieved by designing the unit for
higher operating pressure. However, the effect of increasing hydrogen partial pressure by
increasing total pressure is not as pronounced as that achieved by increasing hydrogen purity.
The reason for the difference in response is that H2S partial pressure is also increased when
total pressure is increased where as increasing the recycle gas purity does not affect the H2S
partial pressure.

Figure 6: Impact of Increasing H2 Partial Pressure

00
WABT,°F
C

-2-4 Reactor Pressure

o
inWABT,

-8
-4
-12
-8
Reduction in

-16
Reduction

-10 Gas purity

-20
-12
-24
-14
-28
100 110 120 130 140 150 160
Hydrogen Partial Pressure (% Base)

Recycle Gas H2S Concentration

If the hydroprocessing unit does not have a recycle gas scrubber, H2S in the high-pressure
loop will build up to a high concentration and inhibit the desulfurization reaction. The reactor
temperature must then be increased to offset the hydrogen sulfide inhibition. This effect is
greater at higher total reactor pressure and more pronounced for CoMo catalysts than for
NiMo catalysts.
 14

Figure 7 shows the effect of recycle gas scrubbing on the average bed temperature at two
pressure levels.

Figure 7: Impact of Recycle Gas Scrubbing

C T,ºF

25
50
Pres sure = 30 Bar
W AoB

20
40
Pres sure = 60 Bar
e a s e inI nWABT,

15
30

10
20
I n c rIncrease

5
10

0
0
0 2 4 6
%H 2 S a t t h e R e a c t o r I n l e t

The effect on catalyst activity illustrated by this plot takes into account only the effect of the
hydrogen sulfide. Since the recycle gas hydrogen concentration also decreases when the
hydrogen sulfide concentration increases, there is an additional debit on the catalyst
deactivation rate due to the lower hydrogen partial pressure. An increase from zero to five
percent hydrogen sulfide in the recycle gas is equivalent to a required increase in average
reactor temperature of about 17°C at 30 bar total pressure and 22°C at 60 bar total pressure.

Recycle Gas/Oil Ratio

Increasing the recycle gas to oil ratio also decreases the average hydrogen sulfide partial
pressure in the reactor, and in turn increases the apparent catalyst activity for CoMo type
catalysts. However, a relatively large increase in the gas rate is required to have the same
effect as scrubbing the recycle gas. Even if the recycle gas is scrubbed, increasing the recycle
gas to oil ratio decreases the reactor hydrogen sulfide partial pressure and therefore reduces
the required reactor temperature. The effect is greater at higher pressure as shown in Figure 8.
 15

Figure 8: Impact of Increasing the Gas/Oil Ratio

Increase in WABT, °C
15

10 Pressure = 30 Bar

5 Pressure = 60 Bar
0
-5

-10

-15
50 100 150 200
Gas/Oil Ratio (% Base)
 16

Topsøe Two-Stage HDS/HDA Process

Process

Topsøe was one of the pioneers to design units for hydrodearomatization of diesel using noble
metal catalysts. The Topsøe two-stage HDS/HDA process is illustrated below in Figure 9.

Figure 9: Diesel Upgrading – Topsøe’s Two-stage Process

Recycle Gas
Compressor
Make-up
Hydrogen

M
Diesel
Feed
Amine Overhead
Wash Scrubber Vapor
Water

Reactor HDS
First Charge Stripper
Stage Heater
Water

HDS Wild
HDS Reactor
Separator Naphtha

Steam
Sour Product
Water Diesel
Stripper Diesel
Product

Diesel
Cooler
Second HDA
Stage Reactor

HDA Separator

The first stage is a conventional hydrotreating step performed over a base-metal catalyst to
reduce the feed sulfur and nitrogen contents to sufficiently low levels to allow the second-
stage noble-metal catalyst to perform the required degree of dearomatization at a high LHSV.
Following this first stage hydrotreating, the diesel leaving the separator contains a significant
quantity of dissolved hydrogen sulfide and ammonia, which is removed in an intermediate
stripper column, using recycle hydrogen as the stripping medium. The hydrogen sulfide and
ammonia containing off-gas is then purified in an amine scrubber.

The stripped diesel and scrubbed hydrogen are then fed to the second-stage,
hydrodearomatization reactor. The catalyst used for dearomatization is one of Topsøe’s noble
metal catalysts, TK-907, TK-908, or TK-915.

After the second dearomatization stage, the diesel is steam stripped to remove small amounts
of H2S present in the oil and to adjust the front-end distillation temperature to meet flash point
requirements for the product diesel.

The Topsøe two-stage diesel HDS/HDA process uses conventional hydrotreater technology at
moderate pressure which enables the potential maximum reuse of equipment during the
revamp of existing units.
 17

Commercial Experience, Topsøe’s HDS/HDA Process

Kuwait Petroleum (Denmark) Refining

One of the first units in Western Europe to produce Swedish Class I diesel (with less than 5
vol% aromatics and less than 10 wppm of sulfur) was the Kuwait Petroleum (Denmark)
Refining (KPDR) unit in Stigsnæs, Denmark, licensed from Topsøe in 1992.

The KPDR unit proved it was more practical to increase the severity in the HDS stage to
produce a better feed for the HDA stage than to increase the severity in the HDA stage to
compensate for insufficient pretreatment of the feedstock. KPDR therefore aimed at having a
feed to the HDA stage that only contained a few wt ppm of sulfur and nitrogen throughout the
life of the catalyst. Figures 10 shows the actual of feed and product aromatics content and
reactor WABT during operation of this unit.

Subsequent to the initial start-up in 1993, the unit was modified once in 1995 to increase the
plant capacity. However, the original charge of TK-908 from early 1992 was still in
operation, without any regeneration, when the KPDR refinery was closed in May 1997. The
analysis of operating data showed that the deactivation of the HDA activity was 15°C to 20ºC
from SOR until the unit was shut down.
 18

Figure 10: Performance of KPDR Diesel HDS/HDA Plant (After Revamp)

San Joaquin Refining, California

In September 1997, a Topsøe two-stage HDS/HDA unit was licensed to San Joaquin Refining
in California. The unit was designed to convert light vacuum gas oil to CARB diesel, i.e.
diesel with a total aromatics content of less than 10 wt%.

This Topsøe two-stage unit, having a hydrogen partial pressure of approximately 70 bar and a
HDA reactor average bed temperature of 287°C, produces diesel containing less than 4 wt%
aromatics. The specific gravity of the diesel is also reduced by about 5%, and the cetane
index of the diesel product is increased by 10 numbers. The properties of the feed,
intermediate first-stage product, and the final product from the HDA reactor are given in
Table 4. As can be seen, the severe pretreatment of the feed in the HDS stage allows a much
higher LHSV in the HDA stage resulting in a major saving in catalyst cost.
 19

Table 4: Performance of Topsøe Two-stage HDS/HDA Unit

Feed (1st Stage) Feed (2nd Stage) Product

TK-555 TK-907
Density @ 15°C, g/cm3 0.91 0.88 0.87
Sulfur, wt ppm 6515 9 1
Nitrogen, wt ppm 775 3 <1
Cetane Index (ASTM D976) 39 46 49
Total Aromatics, wt% 33 24 3.5
TBP Distillation, °C
10% 255 - -
50% 315 - -
95% 370 - -
Conditions
H2 Pressure, kg/cm2 71 70
Temperature, °C 367 287
LHSV, hr-1 Base Base x 3

Revamping Existing HDS Units to Topsøe HDS/HDA Technology

Ultramar/Diamond Shamrock – Wilmington, California

A good example of retrofitting Topsøe HDS/HDA technology is the revamp design of Unit 80
at the Ultramar/Diamond Shamrock (UDS) refinery in Wilmington, California. The existing
unit was originally designed as an FCC feed pretreater but was downgraded to diesel
hydrotreating service when UDS commissioned a new higher severity FCC pretreater. UDS
realized that Unit 80 had the potential for more than simple HDS service and thus after
evaluating currently available HDS/HDA technologies, selected Topsøe to convert this unit to
produce diesel meeting CARB specifications.

Pilot Plant Confirmation

The feed to the unit was a blend of straight-run light gas oil, with approximately 35% coker
light gas oil and FCC light cycle oil. Because the feed blend was very aromatic, UDS
required assurance that the proposed revamp modifications could accomplish the desired 10%
aromatic diesel product required by the CARB specifications. Therefore a pilot plant test was
undertaken to process the exact feed blend provided by UDS at the operating conditions
proposed for the revamped unit. The pilot plant test results in Table 5 summarize the
properties of the feedstock, product from the first HDS stage and final product from the
second HDA stage.
 20

Table 5: Ultramar/Diamond Shamrock Unit 80 Revamp – Pilot Plant Results

Feed 1st Stage Product 2nd Stage Product
Catalyst TK-555 TK-907
Temperature, °C Base Base – 75
-1
LHSV, h Base Base x 2
Sulfur, wppm 8220 11 2
3
Density @ 15°C, g/cm 0.871 0.859 0.842
Distillation, °C
IBP/50%/95% 90/269/353 211/264/352 201/258/348
Aromatics Content, wt%
Total 39.6 37.9 6.0
Mono 19.1 33.0 6.0
PAH 20.5 4.9 <0.4
Cetane Index 41.4 43.9 48.0
Cetane Number 36.2 38.8 47.4

The data in Table 5 shows that although approximately 99.9% desulfurization was achieved in
the first stage with high activity TK-555 NiMo catalyst, the overall reduction in aromatic
content at the 56 bar operating hydrogen partial pressure was quite small. Likewise, there was
only a moderate increase in cetane number and decrease in density due to the removal of
sulfur and nitrogen. The PAH content remained close to 5% because the removal of PAH’s
was constrained by thermodynamic equilibrium at the required 360oC operating temperature
of the first stage.

In the second HDA stage, using TK-907 aromatic saturation catalyst, more than 84% aromatic
saturation was obtained confirming a total product aromatics content of less than 6% was
achievable from the proposed unit. The cetane number was increased by more than 8
numbers compared to the first stage product and by more than 11 numbers compared to the
feed. Despite the much lower reaction temperature and higher LHSV used in the second
stage, the concentration of PAH was below the detection limit, confirming that the removal of
PAH was equilibrium controlled.

In addition to meeting CARB diesel specifications, UDS wanted to maximize the yield of
diesel while retaining the ability to extract 10 to 30% jet fuel boiling range material present in
the feed. The estimated product yields predicted by the pilot plant test work, based on the
218.6 m3/hr (33,000 bpd) design capacity, are shown in Table 6.

Table 6: Product Yields

Yields at 6% Product Aromatics
Rundown Product Start-of-Run End-of-Run
Fuel Gas, M Nm3/hr 2.3 2.8
Total Naphtha, m3/hr 11.1 13.4
Jet Fuel, m3/hr 33.1 33.1
CARB Diesel, m3/hr 186.9 185.2
 21

The fractionation section was designed so that UDS could consistently draw over 33 m3/hr of
jet fuel throughout the operating cycle and if desired they could draw up to 73 m3/hr. Since
the jet fuel fraction would be treated in the aromatic saturation section, the quality was very
good with a smoke point around 25 mm throughout the cycle. Furthermore, if jet fuel
demand decreased, UDS could reduce the jet draw and increase the production of CARB
diesel.

Revamp Procedure

Existing Configuration

Having proven that the proposed revamp design was capable of producing CARB quality
diesel; the next hurdle was to complete the revamp design while maximizing the use of the
existing equipment. The high pressure section of the existing Unit 80, as illustrated in Figure
11, is a conventional HDS single stage configuration with two single bed reactors in series, an
amine scrubber to remove H2S from the recycle gas, but no hot high pressure separator.

Figure 11: Unit 80 Existing Reaction Section Configuration

HDS Reactors
M
Reactor Recycle Gas
Charge Make-up Compressor
Heater Hydrogen

M Amine
Scrubber

E1A-D PC

V1
LSRGO
LCGO Water
P4
Wash Water
M
Surge Drum
Product
P1 Charge Pump E2 Separator

Sour Effluent to
Water Fractionation
Section

The process flow of the existing fractionation section is illustrated in Figure 12. The liquid
from the high-pressure separator is preheated by exchange with fractionator product streams
and light ends are removed in a high-pressure stripper. The stripped product then enters the
main fractionator which produces naphtha overhead and has side cut strippers to separate jet
and light diesel. Heavy diesel is produced as fractionator bottoms, and polished through a salt
dryer.
 22

Figure 12: Unit 80 Existing Fractionation Section Configuration

Sour Fuel Gas

Wild Naphtha

M
M Naphtha
S Fractionator
W
Wash Water
to Rx Section

Stripper
A
Jet
Stripper

Jet
M Fuel
Diesel
Stripper Diesel
A Fract
Feed Diesel
Heater
Salt
Dryer
From
Reactor M
Section

Process Feasibility Studies

One of UDS’ primary project objectives was to minimize the capital investment for the
revamp by reusing as much existing equipment as possible. Furthermore, the available plot
space was very tight, so the addition of new equipment had to be minimized.

To pursue this objective, the initial steps in the revamp approach included a series of
feasibility studies to determine the best options for utilizing the existing equipment.

Figure 13: Process Optimization Studies

• Evaluate Suitability of Existing Internals in HDS Reactors

• Evaluate Using Surplus Reactor for HDA Service
• Evaluate Liquid Quench versus Recycle Gas Quench
• Evaluate Options for Revamp of Recycle Gas Scrubber
• Optimize Conditions for HDS Stripper
 23

The optimization studies undertaken, summarized in Figure 13, included the following:

1. The top liquid distribution trays in the existing HDS reactors, originally designed for
vacuum gas oil service, were evaluated to determine their suitability for use under the
revamp conditions. The flow distribution model predicted the performance of the existing
trays would be adequate for the new service and they were not changed.

2. UDS had a new idle reactor. The use of this reactor was evaluated for second stage
aromatics saturation service as a possible means of minimizing the capital expenditure.
Although the catalyst volume of this reactor was acceptable, the reactor did not have
sufficient catalyst beds to handle the large heat release anticipated for HDA service, and
the diameter was too large to meet the mass flux criteria.

3. Various alternatives of quenching the large exotherm in the new HDA reactor were
studied. These included conventional gas quench options such as increasing the recycle
gas compressor capacity and adjusting the temperature profile through the reactor.
However, the most efficient and cost effective alternative proved to be the use of
hydrotreated diesel from the high-pressure separator bottoms as liquid quench. The highly
efficient Topsoe reactor internals in the new reactor made this choice a viable option.

4. The existing amine scrubber was re-piped to treat the overhead gas from the new high
pressure inter-stage stripper and had to be evaluated to insure it’s suitability for the new
service. The existing stripper was modified and the trays were replaced with packing to
enable reduction of the H2S content in the recycle gas below 5 ppm.

5. All of the existing exchangers in both the reaction and fractionation sections were
evaluated and reused. Some re-piping and minor re-tubing of some of the exchanger
bundles was required and some new exchangers were needed to complete the new
exchange scheme. However, the final layout ultimately provided an additional source of
heat to the fractionation section. This significantly reduced the fired heater duty of the
revamped unit by about 70% compared to SOR conditions and by about 50% compared to
the EOR conditions of the existing unit.

As illustrated in Figure 14, the fuel savings associated with the revamped heat integration
further helped defer some of the incremental capital expenditure.
 24

Figure 14: Fuel Savings

Fuel Gas (Absorbed Duty), MM BTU/hr

Before Revamp After Revamp

86 (Test Run) 27 (SOR)
42 (EOR)

Revamp Configuration

Upon completion of the process studies discussed above, the process configuration of the
revamped unit was finalized. The new process flow of the integrated HDS/HDA unit is
illustrated in Figure 15.
 25

Figure 15: Revamp Configuration of High Pressure Loop

HDS Reactors
M
Reactor Recycle Gas
Charge Compressor
Heater Amine
Make-up Scrubber
Hydrogen
M
HDS Strip
OH Drum

E1B-D HDS
Stripper
E102 Effluent to
PC Fractionation
Section
M
E101 Water
LSRGO V1
LCGO E1A P4 Wash
Sour
Water
LCO Surge Water
M Drum

Product
P1 HDA E2 Separator
Charge Pump Reactor

M
P101
E104
New HP Stripper
Preheat
Equipment
E103 E1A
Modified
Equipment

As seen, all existing equipment in the reaction section was reused, and only a few pieces of
equipment required modification. A new inter-stage high-pressure stripper, which used make-
up hydrogen to remove hydrogen sulfide and ammonia from the HDS stage reactor effluent,
was installed. A new second-stage HDA reactor and associated heat exchange bundles to
recover heat from the aromatic saturation reaction were also added. The existing high-
pressure separator was relocated to the second stage, and replaced by a new first stage
separator. Finally, a new low head pump was added to recycle treated diesel product for use
as liquid quench to the aromatic saturation reactor.

The combined liquid effluent from both first and second stage separators proceeds to the
fractionation section. Again, modifications to the existing process flow were minimized as
shown in Figure 16.
 26

Figure 16: Revamp Configuration of the Fractionation Section

New Sour Fuel Gas
Equipment
Wild Naphtha
Modified
Equipment
Disused M
Equipment M Naphtha
SW Fractionator

Wash Water
to Rx Section

Stripper
A
Jet
Stripper

Jet
M Fuel
Diesel
Stripper Diesel
A Fract Salt
Feed Dryer Diesel
Heater
Salt
Dryer
From
Reactor M
Section

Of all the existing equipment, only the diesel side-cut stripper was not reused. The existing
product and pump-around pumps were reused as spares in jet fuel service. The diesel product
heat exchangers were also all reused, by splitting the fractionator bottoms product into two
streams through the existing parallel run down lines formally used for the light and heavy
diesel products.

Revamp Summary

In summary, we believe that this was a very successful revamp project.

Figure 17: Revamp Summary

• All Product Quality Objectives were met or exceeded

• 24 Weeks for Engineering Design Package
• 9 Weeks for Prerelease of Reactor Shell Design
• The Total Installed Revamp Unit Cost was < $550/barrel
 27

All of the following initial objectives and requirements by UDS were met or exceeded:

1. All of the product quality objectives met or exceeded the specifications required for
CARB diesel.

2. UDS’s objectives of a ‘fast-track’ project were met, in as much as the early release of the
HDA reactor shell design was completed with 9 weeks and the final Engineering Design
Package was completed within 24 weeks from the kick-off date for the start-of-work
(including time for data collection and reconciliation for the existing Unit 80).

3. UDS’s objective to minimize capital expenditure was realized, in as much as the estimated
total incremental installed cost of the revamped unit was less than $550 US per barrel of
throughput capacity.

This is excellent ‘Bang for the Buck’.

Conclusions

New product specifications for diesel will be introduced in the European Union in year 2005.
Over time, the United States, and other nations of the world community, including Asia, are
expected to adopt similar specifications, concerning the sulfur content and diesel burning
quality. It is expected that the future specifications in Asia will also become more stringent
and the Asian refiner will be faced with perhaps even more difficult problems than the EU
refiner is faced with today.

Reducing the sulfur content from 2500 wppm to 350-500 wppm may be achieved in many
existing diesel hydrotreaters by a simple catalyst replacement or by a combination of catalyst
replacement and installation of low cost modification to the reactor internals. However,
meeting sulfur specifications of 10 wppm to 50 wppm is going to require more than these
simple changes.

Reduction of the aromatics content and diesel density, improvement of Cetane Number and
Cetane Index can be accomplished by hydrotreatment of the feed. The severity of the
treatment and the processing configuration will greatly depend on the specifications that are
mandated. A two stage unit utilizing base metal catalyst to clean the feed in the HDS stage,
followed by an HDA section using a noble metal catalyst offers a flexible cost effective way
to meet the most stringent specifications in these areas.

Topsoe has commercially proven HDS/HDA technology, and has demonstrated our
engineering capability and experience to efficiently apply this technology to either new units
or the cost-effective revamp of existing HDS units.
Sibneft - History of Oil In Russia

· What is Crude Oil? · History of Oil In Russia

· Oil And Gas Glossary

Although commercial oil production only began in the second half of the nineteenth century, for centuries oil was
gathered by peoples who lived in parts of the world where it seeped to the surface. In Russia, the first written
mention of the gathering of oil appeared in the sixteenth century. Travelers described how tribes living along the
banks of the river Ukhta in the far northern Timan Pechora region gathered oil from the surface of the river and
used it as a medicine and a lubricant. Oil gathered from the Ukhta river was delivered to Moscow for the first time
in 1597.

In 1702, Tsar Peter the First ordered the setting up of Russia's first regular newspaper, Vedomosti. The paper's first
issue carried a story about the discovery of oil on the surface of the river Sok in central Russia, while later issues
carried similar stories about oil seeps elsewhere in Russia. In 1745, Feodor Pryadunov received permission to
begin gathering oil seeping from the bed of the river Ukhta. Pryadunov also built a primitive refinery, delivering
some of the products to Moscow and St Petersburg.

Oil seeps had also been reported in the North Caucasus by various travelers who passing through the region. Local
people even gathered the oil using buckets to haul it up from wells up to one and a half metres deep. In 1823, the
Dubinin brothers opened a refinery in Mozdok to process oil gathered from the nearby Voznesensk oilfield.

Oil and gas seeps were recorded in Baku on the Western shores of the Caspian Sea by an Arab traveler and
historian as early as the tenth century. Marco Polo later wrote how people in Baku used oil for medicinal purposes
and to administer blessings. By the fourteenth century, oil gathered in Baku was already being exported to other
countries of the Middle East. The first oil well in the world was drilled at Bibi-Aybat near Baku in 1846, more than
a decade before the drilling of the first well in the US. This event marked the birth of the modern-day oil industry.

The Birth Of The Industry

The Baku region harbored many large fields which were very relatively easy to exploit, but transporting the oil to
market was difficult and expensive. The Nobel brothers and the Rothschild family played a major role in the
development of the oil industry in Baku, which was at that time part of the Russian Empire. The industry grew
rapidly, and by the turn of the century Russia accounted for over 30% of world oil production. Shell Transport &
Trading, which later became part of Royal Dutch/Shell, began life by ferrying oil produced by the Rothschilds to
Western Europe.

In the second half of the nineteenth century, Russia also began to discover oil fields in other parts of the country. In
1864, a well drilled in Krasnodar Krai produced the first gusher. Four years later, the first oil well was drilled on
the banks of the river Ukhta, while 1876 saw the start of commercial production on the Cheleken peninsula in
present-day Turkmenistan. The rapid development of oil production was accompanied by the construction of
various plants for processing crude oil, with a lubricants plant opening in 1879 near Yaroslavl and a similar facility
opening the same year in Nizhny Novgorod.

Oil production suffered as a result of the Russian revolution in 1917, and the situation worsened with

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Sibneft - History of Oil In Russia

nationalization of the oil fields by the Communists in 1920. The Nobels sold a significant part of their Russian
holdings to Standard Oil of New Jersey, which was later to become Exxon. Standard Oil protested the decision to
nationalize the oil fields and refused to cooperate with the new Soviet government. But other companies, including
Vacuum and Standard Oil of New York, which was later to become Mobil, invested in Russia. The continued
inflow of Western funds helped Russian oil production to recover, and by 1923 oil exports had climbed back to
their pre-revolutionary levels.

The Rise Of The Soviet Oil Industry

The Caspian and North Caucasus remained the center of the Soviet oil industry until the Second World War, with
rising production feeding the country's rapid drive to industrialize. Securing control over oil production in Baku
was a centrepiece of German strategy during the war, and for a time the Soviet Union found itself cut off from
access to its oil. Caspian oil production once again began to pick up after the end of the war, and reached a new
record high of some 850,000 b/d in 1951. Baku remained the centre of the industry and nearly two-thirds of Soviet
oil field equipment was manufactured in the area.

But at the same time, Soviet planners began to accelerate development of the Volga-Urals region, which had been
under development since the 1930s. Fields in the region were often close to existing transportation infrastructure,
and the geology was not particularly complex. By 1950, the new fields accounted for 45% of Soviet oil production.
The massive investments in the region paid off, allowing for a big hike in Soviet oil production. The extra barrels
went to feed a wave of new refineries which were brought on stream in the period between the 1930s and the
1950s. The Omsk refinery was opened in 1955, and later grew to become one of the largest refineries in the world.

The growth in production also allowed the Soviet Union to begin ramping up exports of oil. Moscow was keen to
maximize hard currency earnings from oil exports, and priced aggressively in order to boost its market share. By
the early 1960s, the Soviet Union had replaced Venezuela as the second largest oil producer in the world. The
arrival of lots of cheap Soviet barrels on the market forced many Western oil companies to cut their posted prices
for Middle Eastern oil, thus reducing royalty revenues for governments of the Middle East. This reduction in
revenues was one of the driving forces behind the formation of the Organization of Petroleum Exporting Countries
(OPEC).

Production from the Volga Urals region peaked at close to 4.5 million b/d in 1975 but later dropped back to less
than a third of that level. Just as the Soviet Union was thinking about how it could sustain production from
maturing fields in the Volga Urals, the first major discoveries in Western Siberia were announced. The early years
of the 1960s saw a series of discoveries in the region, culminating with the discovery of the super-giant Samotlor
field in 1965, home to recoverable reserves estimated at some 14 billion barrels.

The West Siberian basin presents a hostile environment in which to produce oil, with the territory ranging from
permafrost around the Arctic circle to extensive peat bogs in the south. But in spite of the difficulties, the Soviet
Union was able to ramp up production from the region at an astounding rate. Growth in West Siberian production
underpinned an increase in total Soviet production from 7.6 million b/d in 1971 to 9.9 million b/d by 1975. By the
middle of the 1970s, West Siberian production was filling the gap being left by the decline in Volga Urals output.

The Decline Of The Soviet Oil Industry

But in achieving phenomenal production from fields in Western Siberia, the Soviet oil industry had also sown the
seeds of its own decline. West Siberian fields were relatively cheap to develop and offered huge economies of
scale, and Soviet planners gave priority to maximizing short-term rather than long-term recovery. Production
associations tended to overproduce existing fields to meet production quotas without regard for proper reservoir
management practices, drilling too many wells and injecting too much water. There were also no incentives to
improve efficiency and scant investment in new technology. The problems soon began to manifest themselves in
the form of falling well productivity, low reservoir pressure and rising water cut.

By the middle of the 1970s, Moscow was already aware that a production decline was just around the corner. The
first decline hit in 1977, caused by chronic under-investment in exploration in Western Siberia, but authorities
managed to reverse the decline by boosting spending on drilling. The second fall happened in the period between

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1982 and 1986. This time too, Moscow managed to head off a crisis by injecting more cash.

In 1988, the Soviet Union hit a new record of some 11.4 million b/d. At this point, it was the largest producer in
the world, with output significantly higher than in either the US or Saudi Arabia. It was also this year that output
from Western Siberia peaked at 8.3 million b/d. But by that point, a sustained decline in production was inevitable
- thanks to poor reservoir management techniques, the Soviet Union only managed to lift production marginally
during the first part of the 1990s, despite a dramatic increase in capital expenditure. When it came, the collapse in
production was as dramatic as the rise had been -- Russian production fell continuously for a decade and ended up
at almost half its original level.

The slide was aided by the economic crisis which engulfed the region in the wake of the collapse of the Soviet
Union. The collapse of the economy resulted in a big drop in the domestic consumption of oil, but export capacity
restraints meant that companies were forced to continue selling a large portion of their output on the domestic
market, often to insolvent customers. The companies' financial difficulties forced a complete halt to all new
exploration and drilling activity, and even work-overs of existing wells, a situation which worsened the collapse in
production.

Future Development

Russian oil production finally halted its slide in 1997. An independent analysis suggests that Western Siberia could
still harbor over 150 billion barrels of reserves, three times the volume produced to date. But the picture is clouded
by the poor condition of the reservoirs at fields already under development, and by the fact that West Siberian
fields typically consist of a larger number of oil-bearing layers than fields in other regions, thus complicating
recovery.

Other provinces also offer significant potential. The Timan-Pechora basin stetches from the Urals in the east to the
Barents Sea in the north. The region suffers from a harsh climate, and a large part of the reserves are thought to
consist of heavy oil. Nevertheless, remaining discovered reserves are estimated at around nine billion barrels,
implying significant potential. East Siberia's remaining reserves are put at three billion barrels, but undiscovered
reserves could be several times larger. The region's main drawback is its distance from markets and the lack of
transport infrastructure. Reserves offshore Sakhalin island are also thought to be significant, but development to
date has been held back by high costs.

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 March 21, 2004

Succeed at Plant Debottlenecking

Critical evaluation, coupled with creativity, can enable you to turn up capacity and
improve performance
By Joseph C. Gentry, P.E.

March 10, 2004

You are a chemical engineer with the modest assignment to debottleneck your plant to boost
production. The direct approach seems obvious: assess current capacity, determine output
limitations, and then vigorously employ standard engineering techniques to upgrade or replace
equipment. Perhaps you find 20% reserve capacity, eliminating the need for major capital
expenditure. Even if you achieve good results, however, you can't necessarily be sure that your
straightforward analysis gives the optimum solution.
Now, consider a different assignment: expand the plant capacity by 75%, reduce the specific
energy consumption by 20%, increase the product yield by 5% – and do it on a shoestring budget.
This presents more of a challenge and definitely requires a different, more creative approach to the
problem.

Not every plant can be revamped to gain such improvements. However, many plants can operate
much more efficiently or at a substantially higher capacity than a routine analysis might suggest.
This article presents some ideas and case studies to help you critically evaluate your process to
see whether it offers exceptional revamp opportunities.

Beyond troubleshooting
Both revamp work and troubleshooting can be part of a plant debottlenecking effort. They share
many tactics, but differ in their intent.

The goal of a revamp is to improve some basic parameters such as capacity or processing
efficiency. In contrast, troubleshooting merely aims to solve a problem that hampers current
operation. Often, the problem exposes an opportunity for much greater gain in performance.

Both revamp and troubleshooting require open-minded thinking and proficiency with engineering
tools. Both often rely on test runs to diagnose problems and uncover design errors or inaccuracies
in equipment ratings. However, for a major improvement in the fundamental performance of a unit,
you must go beyond this first level of investigation and intentionally look for the greater gain.
The exceptional plant revamp will require, in one way or another, each of the following five
elements:

1. Thorough knowledge of process fundamentals;

2. A critique of the original design;
3. Proper approach to the problem;
4. Creative and talented individuals; and
5. A focus on developing the simplest solution.

Of course, you must also give due consideration to the specific problem and common sense.

Understand the process

A prerequisite for making a reasonable troubleshooting effort or having any hope of a successful
revamp is to have a thorough knowledge of the process. This must go beyond the mechanics of
operation and control points and include "what and why."
Let's review, for example, a traditional aromatics extraction plant. Most operators of Sulfolane or
UDEX plants do not fully understand the operation of the extractor/stripper and the various recycle
loops. What factors most affect the phase separation of the liquid-liquid extraction (LLE)? What
about water or other co-solvents? What is the loading profile in the towers? What is the stage
efficiency? What is the effect of process temperature? Do C5s in the feed help or hurt the
operation? How do you deal with three-phase distillation? There are a host of other questions to
answer before optimizing the process, let alone before proposing a revamp.
Figure 1: An LLE unit for aromatics recovery started to lose capacity because of flooding. Initially it
was assumed that the usual flooding mechanism (A) was occuring, but deeper analysis pointed to
the mechanism that was really responsible (B).

The following case study shows how an inadequate understanding of the process led to a
misdiagnosis and delay in resolving an operating constraint. Throughput of an LLE unit used for
aromatics recovery started to decrease. The tower design had the heavy-phase dispersed with the
interphase at the bottom. All parameters seemed normal, but the extractor would flood at high feed
rates. Typically, such flooding is due to restriction in the tower internals, as shown in Figure 1.
However, in this case, it occurred for a very a subtle and unexpected reason.

It was originally thought that the sieve trays were plugged because the tower would hold up solvent
as it became loaded. However, water-washing the debris from the tray decks did not restore the
capacity.

Pressure drop only measured the head of the continuous phase. To further debunk the theory of
holes being blocked, the tower could operate at virtually any flow rate of solvent. It only was
sensitive to higher flow rates of the continuous phase, or raffinate. Furthermore, an inspection
indicated that the tower was clean and mechanically intact, with no apparent restriction to the
solvent flow.

The troubleshooting effort then turned to a trial-and-error manipulation of various, unrelated

operating variables and external factors. Such efforts are costly and time consuming and, in this
case, were not based on fundamental process knowledge or all the phenomena occurring.

Hydraulic ratings showed that the sieve trays were operating in the normal range with no
entrainment of heavy phase into the upcomers. The overall aromatics recovery was as high as it
had always been, regardless of the feedstock or operating conditions, suggesting that back mixing
of the phases did not materially affect the efficiency.

At this point, experience and operation knowledge finally came in to play.

The trays were designed with 5/32-in. (4 mm.) nominal diameter sieve holes. Drill bits were used to
verify the hole size. Although the tray decks appeared clean, the 5/32-in. bits did not pass through
the perforations. There was a slight polymer buildup around the circumference of the holes. This
buildup was sufficient to shear the heavy phase into smaller droplets that became entrained in the
raffinate phase at higher continuous-phase rates. A simple chemical cleaning of the tower rrestored
the capacity to its prior level.

Eventually, the staff might have decided to chemically clean the tower and gained the improvement
in capacity. However, without understanding the cause of the problem, the plant would not have
been in a position for further improvements.

A key relationship
It is surprising how often the simple relationship of reflux-to-distillate (R/D) ratio versus the number
of theoretical stages (NTS) in a distillation application is overlooked. Yet, it can lead to a great return
on investment or a simple resolution to a problem.
In the post-fractionation section of one particular aromatics plant, three sets of xylene splitters
operated in parallel to produce ortho-xylene. The para- plus meta-xylene separation from ortho-
xylene typically requires about 120 trays at a moderate reflux-to-feed ratio. One of the xylene
splitters had only 51 trays. The tower was brought into xylene-splitting service several years earlier,
but had never produced much on-specification ortho-xylene. Higher reflux rates did not improve the
operation, as they did with the other towers. This was a classic case of operating at the far end of
the R/D versus NTS curve, near the point of minimum NTS (Figure 2). The tower was revamped by
replacing two trays for every original one. The result: a threefold increase in ortho-xylene production
at a lower energy consumption.

A different petrochemical application, this time involving structured packing, illustrates the same
oversight, but at the other end of the spectrum. Structured packing is commonly considered for low-
pressure or vacuum applications, but not those with high liquid loadings resulting from high vapor
density. Yet, the packing has been used in a successful revamp of a high-pressure depropanizer
application.

The next application of structured packing in the same service, however, did not perform nearly so
well. A closer examination of the first case showed the reason why. The first column was operating
at the edge of the R/D versus NTS curve, near the point of minimum reflux, as shown in Figure 3.
The packing gave substantially higher height equivalent of a theoretical plate (HETP) than claimed,
but this went unnoticed because the curve was steep in that region and the small loss in efficiency
was dismissed as inconsequential. The second case operated at a more normal position on the
curve. There, the loss of NTS was more pronounced, which exposed the true performance of the
packing.

The first revamp technically was not a failure because the depropanizer realized the claimed
capacity. The mistake, as shown in the second case, was in wrongly assuming that the revamp
would work in the same application without understanding the limitations of the technology.

Know where your towers operate on the R/D versus NTS curve. And understand how fluctuations in
operation over time will affect performance.

Critique the design

It is counterintuitive to question the integrity of a design that may have given years of adequate
service. However, the original design should be critically reviewed to address problems and to
understand the basis for significant improvement. The intent is not to denigrate the designers, but to
see where new methods and recently developed technologies may apply.

In some cases, the original designers used rules of thumb and practical guesswork. Therefore,
designs may appear to work well from the outset but may be on the verge of imminent failure. If
possible, contact the original designers to evaluate their methods. Do not automatically assume that
the design was done correctly.

Within a reasonable range, it is possible to exchange efficiency for capacity by adjusting the tray
design and process conditions. Like the high-pressure depropanizer application, the aromatics LLE
tower had far more efficiency than needed and great potential for capacity increase. Ultimately, the
column was debottlenecked with only minor modifications to approximately double the original
capacity at inconsequential loss of efficiency.

Some older plants may have been designed with rudimentary computer equipment, or perhaps
even slide rules. So, the designers may have missed opportunities for optimization due to the
difficulty of reviewing multiple design cases. In addition, all basic unit operations have benefited
from great improvements during the last several years. If you have a vintage plant that has not
been reviewed recently, chances are that such a review can uncover cost-effective revamp
opportunities.

Pay particular attention to any unit with an unusually large surplus capacity in one part. This often
means that the designer did not know what to do in that section and took the conservative approach
of oversizing the equipment. A mismatch in capacity within a unit presents a good opportunity to
find creative revamp solutions.

You may want to correct the imbalance first. Consider the example of some aromatics extraction
plants, where the capacity of the solvent stripping part of the unit substantially exceeds that of the
extraction part. An all-too-typical discussion during a plant walk-through goes like this:

"What is that tower?" a troubleshooter asks.

"It's our new extractor," the plant engineer replies.
"Why did you need a new extractor?"
"The first one couldn't handle the capacity we wanted."
"What other equipment did you have to replace to gain the higher processing rate?"
"Nothing else, only the extractor."

Adding a new extractor is a high price to pay for an improper design. If the original designer could
not balance the major pieces of equipment in the unit, you can be sure that the other parts are not
optimally designed either. Look for opportunities here.

Challenge assumptions
Sometimes, the evolution of knowledge and experience can lead to improvements. You should
question long-standing beliefs that govern important or peculiar operating practices.

At an olefins plant, troubleshooters were trying to discover the source of methanol found in polymer-
grade propylene. The cracker feedstock was immediately suspected, but the idea was dismissed as
unreasonable. A former staff member had read in a technical article that methanol would not
survive without decomposing in the cracking furnaces. An investigation found that the article,
written in 1929, described methanol decomposition after a few seconds at cracking furnace
temperatures. The cracking furnaces in this plant, however, provided only a fraction of a second
residence time and, therefore, could not decompose the methanol. The feed source was addressed
and the problem was solved. More reliable data in the beginning and closer scrutiny of what was
fact versus myth would have led to a quicker solution.

Figure 2: A xylene tower operated at the far end of the R/D versus NTS curve, near the point of
minimum NTS. Revamping, by replacing each tray with two at closer spacing, moved the operating
point and led to increased production and lower energy consumption.

Likewise, don't assume that popular technology represents the current state of the art. Process
engineers tend to be a conservative lot. For the most part, this is prudent because the business
risks of failure outweigh potential gains from minor improvements. But, consider the example of
aromatics recovery technology where the LLE process had remained unchallenged for 40 years.
Now, extractive distillation (ED) with modern solvents and design techniques is accepted as the
superior process and current state of the art.
Figure 3: Structured packing gave substantially higher HETP than expected, but this went
unnoticed because the curve was steep in that region. Revamping moved the operating point to a
more favorable location.

Approach the problem right

Before you can embark on solving a problem, you have to recognize that you have a problem and,
perhaps, a significant opportunity. The xylene-splitting column in the example above had been in
operation for many years before anyone saw the revamp opportunity. Countless other cases
throughout the industry await discovery and process engineers need to take the initiative to find and
make improvements to their units.

Ensure that the problem to be resolved is not defined too narrowly. Before you try to increase the
capacity of a unit, make sure that you are effectively using its current capacity. For example, one
plant wanted to revamp its toluene product column to gain more capacity. While that could have
been done, a broader investigation revealed five distillation columns within the complex that were
already separating C7 from C8 components. The streams were re-mixed after an intermediate
processing step, only to be separated again. A proper redesign of the process configuration would
bring this down to two separations and free up the desired capacity.

Remember: solving the problem in the traditional manner will get you whatever spare capacity the
original designer left in the unit, at about the same yield and specific energy consumption. Tackling
the same problem with a creative approach could gain much more capacity and reduce operating
cost at the same time.

Perhaps you have heard, "I've already revamped my unit," from someone who is proud of
increasing capacity by, say, 80% more than the original design. The implication is that the
maximum capacity possible has been reached. Don't blithely accept that. My company re-
engineered a Sulfolane unit to boost its capacity by another 80% (to 320% of the original design) by
adding only one new distillation column (Figure 4). The solution was to convert the LLE unit
equipment to ED service. The revamp cost only one-third as much as that of building a completely
new unit, which is the conventional approach. Lesson learned: make sure that you do not kill the
project before reviewing it from more than one viewpoint.

Find creative people

If you need technology experts or process licensors to help with the revamp design, ask some basic
questions. What are their skills with this type of revamp? How do they approach the problem? Are
they proficient with the process engineering? Certainly, major firms have qualified engineers – but
that doesn't mean these particular engineers will work on your project. Exercise your right to select
the team.

Figure 4: Capacity of an aromatics extraction unit was boosted by an additional 80% (to 320% of
original design) by converting the LLE unit equipment to ED service at one-third the cost of building
a completely new unit.

When assessing licensors, remember that a long list of licensed units does not necessarily mean
the technology is state of the art or the units were properly designed. Neither of these may be true.

Many engineers don't take the trouble to understand the distinctions among available processes
and instead believe the safe selection is the most widely known licensor. Maybe this is the best
choice, but it is essential to carry out a thorough investigation. Otherwise, you will not know what
has been missed and may find yourself in a competitively difficult situation against others who have
chosen a more creative approach.

Strive to put together a team with a variety of skill sets. Engineers who have worked in regions
where energy costs are high will likely keep energy conservation and process integration clearly in
mind. Engineers from relatively "technology-poor" regions usually have well-developed basic skills
and practical insight. A variety of backgrounds often foster fresh ideas.

Here is an example of a revamp solution that came from a solid understanding of the original
process, coupled with out-of-the-box thinking. A producer of terephthalic acid (PTA) needed to
increase the capacity of its dehydration column, which separates water from acetic acid using
traditional distillation. The component separation profile had a pinch near the point of maximum
water concentration at the top of the column, as depicted in Figure 5. It was prohibitively expensive
to debottleneck the vessel because of the titanium construction material.

The solution to this problem was to decrease the reflux ratio on the column and feed the overhead
stream into an LLE unit, which is more effective than distillation in removing water from dilute
streams. Furthermore, part of the feedstock is diverted to this new extraction unit, which increases
the capacity even more. Because the separation is moved away from the pinch point, the unit will
simultaneously reduce the specific energy consumption.

These kinds of ideas come from creative and talented individuals who go beyond the requirements
of typical process engineering assignments.

Figure 5: Reducing reflux ratio of a dehydration column, to move it away from a pinch point, led to
increased capacity and lower specific energy consumption. The overhead feed now goes to an LLE
unit.

Simple is best
Implementing a clever revamp solution requires selling it to others, either inside or outside the
company. Engineers are proud of their technical accomplishments, but many of these brilliant ideas
go unrealized because of poor selling. The reason may be as straightforward as a manager
rejecting a proposal because it looks too complicated. In general, the revamp idea should be
packaged to be no more complex than what the least technically capable person in the decision-
making chain can readily understand. The solution itself does not have to be simplistic – but the
presentation must be simple and clear. Look at the big picture. The project will never happen if you
can't convince people of its value.

Seize the opportunity

Revamping and troubleshooting a chemical plant is a complex task aimed at achieving a simple
and measurable objective. So, make sure you approach the effort properly. Know your process;
question and then take responsibility for its design; think creatively about solution alternatives; use
reliable engineering practices to implement the work. Start with conventional wisdom but check out
recent developments which may open up some opportunities. Choose the right people to be
involved in your design. Simplicity is always best, even if the means of achieving that simplicity is
complicated.

Competition in the marketplace is fierce. Everyone wants to become the "low-cost" producer.
Achieving that is tough if your plant uses the same technology as everyone else. A competitor
 taking a more creative approach may leapfrog you. So, don't be complacent. Spend the time
aggressively reviewing the opportunities for improving your plant's operation.

Joseph C. Gentry, P.E., is petrochemical business manager for GTC Technology Inc., Houston, a
firm that provides technology licenses to the petrochemical and refining industries. E-mail him at
jgentry@gtchouston.com.

Chemical Processing © 2004 Putman Media

Atmospheric/Vacuum Distillation Units - Atmospheric Distillation

SPM-2700 Atmospheric/Vacuum Distillation Units

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Copyright © 2002 by Simtronics Corporation. All rights reserved worldwide.

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Atmospheric/Vacuum Distillation Units

SPM-2700 Atmospheric/Vacuum Distillation Units

Process Description

The ADU (Atmospheric Distillation

Unit) separates most of the lighter
end products such as gas, gasoline,
naphtha, kerosene, and gas oil from
the crude oil. The bottoms of the ADU
is then sent to the VDU (Vacuum
Distillation Unit).

Crude oil is preheated by the bottoms

feed exchanger, further preheated
and partially vaporized in the feed
furnace and passed into the
atmospheric tower where it is
separated into off gas, gasoline,
naphtha, kerosene, gas oil, and
bottoms.
Click here to view the schematic display A.
This tower contains 20 fractionation Click here to view the schematic display B.
trays, is equipped with one top pump
around, an overhead reflux system,
and three side strippers (for naphtha,
kerosene, and gas oil products).

The liquid from the feed furnace

enters the tower bottoms, where it is
collected and sent for further
processing to the VDU. Steam is
injected into the base of the tower to
reduce the hydrocarbon partial
pressure by stripping some light
boiling components from the bottoms
liquid. The vapors from the feed
heater enter the tower below tray 20.

At tray 19, a draw pan is located from

which gas oil product is drawn. The
gas oil product flows by gravity to the
top of the gas oil stripper. Stripping
steam is used to remove the light
ends, improving the flash point. The
stripped gas oil product is pumped to
storage.

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Atmospheric/Vacuum Distillation Units

The next product draw is located at

tray 12, where the kerosene product
is drawn. The kerosene product flows
by gravity to the top of the kerosene
stripper. Stripping steam is used to
remove the light ends, improving the
flash point. The stripped kerosene
product is pumped to storage.

The last product side draw is located

at tray 5, where the naphtha product
is drawn. The naphtha product flows
by gravity to the top of the naphtha
stripper. Stripping steam is used to
remove the light ends, improving the
flash point. The stripped naphtha
product is pumped to storage.

A pump around liquid stream is drawn

from tray 6, cooled and returned to
tray 3.

The condensed gasoline and water are

separated by gravity in the reflux
drum. Part of the gasoline is pumped
back to the tower as reflux, with the
rest going to storage. The water is
drained to disposal and the vapor
from the ADU overhead is passed to
an untreated fuel gas system.

The VDU (Vacuum Distillation Unit)

takes the residuum from the ADU
(Atmospheric Distillation Unit) and
separates the heavier end products
such as vacuum gas oil, vacuum
distillate, slop wax, and residue.

Heavy crude oil is preheated by the

bottoms feed exchanger, further
preheated and partially vaporized in
the feed furnace, and passed into the
vacuum tower where it is separated
into slop oil, vacuum gas oil, vacuum
distillate, slop wax, and bottoms
residue.

This tower contains a combination of

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Atmospheric/Vacuum Distillation Units

14 fractionation trays and beds. It is

equipped with three side draws and
pump around sections for vacuum gas
oil, vacuum distillate, and slop wax
products.

The liquid from the feed furnace

enters the tower bottoms, where it is
collected and sent for further
processing. Steam is injected into the
base of the tower to reduce the
hydrocarbon partial pressure by
stripping some light boiling
components from the bottoms liquid.
The vapors from the feed heater enter
the tower below tray 14.

At tray 14, a draw pan is located from

which slop wax product is drawn. The
slop wax product and pump around
are cooled, with the slop wax product
going to storage, while the pump
around is returned to the tower at
tray 11.

The next product draw is located at

tray 8, where the draw for vacuum
distillate product is located. The
vacuum distillate draw tray is a total
draw tray, where the reflux from the
tray is pumped under flow control to
the tray below. The product and
pump around are cooled, with the
vacuum distillate product going to
storage, while the pump around is
returned to the tower at tray 7.

The last product draw is located at

tray 4, where the draw for vacuum
gas oil product is located. The
vacuum gas oil draw tray is also a
total draw tray, where the reflux from
the tray is pumped under flow control
to the tray below. The product and
pump around are cooled with the
vacuum gas oil product going to
storage, while the pump around is
returned to the tower at tray 1.

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Atmospheric/Vacuum Distillation Units

The overhead from the VDU is

condensed and combined with the
vacuum steam. The slop oil and water
are separated by gravity in the
vacuum drum. The water is drained to
disposal, while the slop oil is
accumulated and occasionally drained
to slop collection.

Process Specifications

The ADU fractionates 18.40 MBPD of Crude Oil to produce 2.19 MSCFD of Gas, 1.34
MBPD of Gasoline, 1.92 MBPD of Naphtha, 1.19 MBPD of Kerosene, 0.80 MBPD of Gas Oil,
and 10.56 MBPD of Atmospheric Residuum.

The ADU feed is heated to 690 Deg F before entering the tower which is maintained at
2.70 PSIG. The top temperature is controlled at 280 Deg F which maintains the Gasoline
quality, and draw temperatures of 355 Deg F for the Naphtha, 529 Deg F for the
Kerosene and 583 Deg F for the Gas Oil.

The VDU fractionates 10.56 MBPD of Atmospheric Residuum to produce 0.39 MBPD of
Vacuum Gas Oil, 1.22 MBPD of Vacuum Distillate, 1.44 MBPD of Slop Wax and 6.07 MBPD
of Vacuum Residuum.

The VDU feed is heated to 750 Deg F before entering the tower which is maintained at
2.00 inHg. The top draw temperature is controlled at 310 Deg F which maintains the
Vacuum Gas Oil quality, and draw temperatures of 607 Deg F for the Vacuum Distillate,
and 668 Deg F for the Slop Wax.

Instrumentation

The ADU feed is pumped by P-100 (HS-100) and controlled by FIC-100. It is preheated in
the bottoms feed exchanger (E-100) before entering the Feed Furnace (F-100). TIC-100
controls the crude oil temperature entering the ADU (T-100) by adjusting fuel gas flow to
the furnace.

Bottoms liquid is collected and sent to the VDU by LIC-114 through the Bottoms Pump P-
114 (HS-114). This flow is indicated by FI-124. Stripping steam is injected into the ADU
bottoms by FIC-134.

Hot gas oil flows by gravity to the Gas Oil Stripper (T-113) through FIC-113. The gas oil
enters the stripper at the top and flows downward over six trays. Stripping steam is
introduced into the bottom of the stripper through FIC-133. The gas oil product is
pumped from the base of the stripper by the Gas Oil Product Pump P-113 (HS-113) to
storage. The gas oil product flow is controlled by LIC-113 and the flow rate is indicated by
FI-123. The gas oil product's 95% point is monitored by AI-123.

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Atmospheric/Vacuum Distillation Units

Hot kerosene flows by gravity to the Gas Oil Stripper (T-112) through FIC-112. The
kerosene enters the stripper at the top and flows downward over six trays. Stripping
steam is introduced into the bottom of the stripper through FIC-132. The kerosene
product is pumped from the base of the stripper by the Gas Oil Product Pump P-112 (HS-
112) to storage. The gas oil product flow is controlled by LIC-112 and the flow rate is
indicated by FI-122. The kerosene product's 95% point is monitored by AI-122.

Hot naphtha flows by gravity to the Naphtha Stripper (T-111) through FIC-111. The
naphtha enters the stripper at the top and flows downward over six trays. Stripping
steam is introduced into the bottom of the stripper through FIC-131. The naphtha product
is pumped from the base of the stripper by the Naphtha Product Pump P-111 (HS-111) to
storage. The naphtha product flow is controlled by LIC-111 and the flow rate is indicated
by FI-121. The naphtha product's 95% point is monitored by AI-121.

A naphtha pump around is drawn from tray 6, pumped through P-115 (HS-115) and
controlled by FIC-115. The pump around return temperature is controlled by TIC-115
which modulates cooling water flow to E-115.

The ADU overhead vapor flows through the overhead condenser E-110 (HV-110), whose
outlet temperature is indicated by TI-120, into the Overhead Reflux Drum D-111. The
hydrocarbons are partially condensed and the two phases (vapor and liquid) enter the
overhead reflux drum where the condensed water separates from the hydrocarbon liquid
by gravity.

The reflux is returned to tray 1 of the tower from the reflux drum via pump P-110 (HS-
110). The reflux flow is controlled by FIC-110 which is reset by TIC-110 to control the
tower overhead temperature. The level of the overhead drum is maintained by LIC-110
which sends the gasoline product to storage, whose rate is indicated by FI-120.

Water collects in the boot of the overhead reflux drum, and is transferred to the water
system for treating. LIC-120 maintains a constant sour water level. The sour water is
sent to treatment through pump P-120 (HS-120).

The uncondensed gas (FI-130) is sent to fuel gas through PIC-120, which maintains the
ADU back pressure. Analyzers are present to monitor the C3 composition of the off gas
(AI-130) and vapor pressure (AI-120) of the gasoline.

ADU overhead pressure is indicated by PI-110 and bottoms pressure is indicated by PI-
114.

The ADU tower temperature profile is indicated by TI-120 (overhead), TI-110 (gasoline),
TI-111 (naphtha), TI-112 (kerosene), TI-113 (Gas Oil), and TI-114 (ADU residuum).

The VDU feed is pumped by P-114 (HS-114) controlled by LIC-114 and indicated by FI-
124. It is preheated by the bottoms feed exchanger E-200 before entering the Feed
Furnace (F-200). TIC-200 controls the temperature of the feed entering the VDU (T-200)
by adjusting fuel gas flow to the furnace.

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Atmospheric/Vacuum Distillation Units

Bottoms liquid is collected and sent to storage through pump P-214 (HS-214), controlled
by LIC-214, and indicated by FI-224. This residue's 95% point is monitored by AI-224.
Stripping steam is injected into the VDU bottoms by FIC-234.

Hot slop wax is pumped from the tower by pump P-213 (HS-213). The slop wax product
flow to storage (FI-223) is controlled by LIC-213, and it's 95% point is monitored by AI-
123. Cooled pump around is controlled by FIC-213 and returned to the tower above the
slop wax draw tray.

Hot vacuum distillate is pumped from the tower by pump P-212 (HS-212). The vacuum
distillate product flow to storage (FI-222) is controlled by LIC-212, and it's 95% point is
monitored by AI-122. Cooled pump around is controlled by FIC-212 and returned to the
tower above the vacuum distillate draw tray. Vacuum distillate reflux is controlled by FIC-
232 and returned to the tower below the vacuum distillate draw tray.

Hot vacuum gas oil is pumped from the tower by pump P-211 (HS-211). The vacuum gas
oil product flow to storage (FI-221) is controlled by LIC-211, and it's 95% point is
monitored by AI-121. Cooled pump around is controlled by FIC-211 and returned to the
tower above the vacuum gas oil draw tray. Vacuum gas oil reflux is controlled by FIC-231
and returned to the tower below the vacuum gas oil draw tray.

The VDU overhead vapor flows through the overhead condenser E-210 (HV-212) into the
Overhead Vacuum Drum D-211. The hydrocarbons are fully condensed and mixed with
the vacuum condensate flow from E-211.

The water separates from the hydrocarbon liquid by gravity, where the drum's water level
is maintained by LIC-210 which sends the water to treatment via P-210 (HS-210). The
hydrocarbon phase overflows from the water phase, and slowly accumulates (LI-209).
When a sufficient level of slop oil has accumulated, the level can be drained via P-209
(HS-209).

The VDU vacuum pressure is maintained by the steam to the vacuum ejector (HV-211),
the cooling water (HV-212) to the steam condenser E-211, and the hydrocarbon
condenser E-210. The pressure is regulated by PIC-210, which reduces the vacuum by
circulating water to the vacuum ejector. The VDU bottoms pressure is indicated by PI-
214.

The VDU tower temperature profile is indicated by TI-210 (overhead), TI-211 (gas oil), TI-
212 (vacuum distillate), TI-213 (slop wax), and TI-214 (VDU residuum).

Advanced Controls

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Atmospheric/Vacuum Distillation Units

The Burner Ignitor Switch (HS-103) serves as both a burner ignitor and a flame indicator.
The fuel gas to the Feed Heater (F-100) will trip (TIC-100 mode=manual and output=0.0)
if the FD Fan (HS-102) is not on, the flame goes out, or the tube flow (FIC-100) is less
than it’s low alarm. This trip logic can be bypassed with HS-101.

The Burner Ignitor Switch (HS-203) serves as both a burner ignitor and a flame indicator.
The fuel gas to the Feed Heater (F-200) will trip (TIC-200 mode=manual and output=0.0)
if the FD Fan (HS-202) is not on, the flame goes out, or the tube flow (FI-124) is less
than it’s low alarm. This trip logic can be bypassed with HS-201.

Faults

All faults can be failed high or low to any degree with any of 8 fault function generators
(step change, square wave, staircase, stairs, ramp, sawtooth, slope, or sine wave). Faults
can be programmed to start and/or stop at various times during a simulation exercise.

● Fault 1: Pump P-100 ● Fault 26: Pump P-209

● Fault 2: Pump P-110 ● Fault 27: Pump P-210
● Fault 3: Pump P-111 ● Fault 28: Pump P-211
● Fault 4: Pump P-112 ● Fault 29: Pump P-212
● Fault 5: Pump P-113 ● Fault 30: Pump P-213
● Fault 6: Pump P-114 ● Fault 31: Pump P-214
● Fault 7: Pump P-115 ● Fault 32: E-200 Hex
● Fault 8: Pump P-120 ● Fault 33: Vacuum Efficiency
● Fault 9: E-100 Hex ● Fault 34: Gas Oil Hex
● Fault 10: E-110 Hex ● Fault 35: Vacuum Distillate Hex
● Fault 11: E-115 Hex ● Fault 36: Slop Wax Hex
● Fault 12: LIC-110 Valve ● Fault 37: PIC-210 Valve
● Fault 13: FIC-111 Valve ● Fault 38: FIC-211 Valve
● Fault 14: FIC-112 Valve ● Fault 39: LIC-212 Valve
● Fault 15: FIC-113 Valve ● Fault 40: FIC-232 Valve
● Fault 16: TIC-100 Valve ● Fault 41: FIC-213 Valve
● Fault 17: FIC-131 Valve ● Fault 42: FIC-234 Valve
● Fault 18: FIC-132 Valve ● Fault 43: FI-221 Transmitter
● Fault 19: FIC-133 Valve ● Fault 44: FI-223 Transmitter
● Fault 20: FIC-134 Valve ● Fault 45: PI-214 Transmitter
● Fault 21: TI-120 Transmitter ● Fault 46: TI-211 Transmitter
● Fault 22: TIC-110 Transmitter ● Fault 47: TI-213 Transmitter
● Fault 23: TI-111 Transmitter ● Fault 48: AI-221 Transmitter
● Fault 24: F-100 FD Fan ● Fault 49: F-200 FD Fan
● Fault 25: F-100 Flame ● Fault 50: F-200 Flame

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Atmospheric/Vacuum Distillation Units

Training Exercises

You may create a virtually unlimited number of scenarios and training exercises by
programming the faults described in the previous section. You can then establish
performance standards for each one of those exercises. Simtronics provides a number of
exercises with established performance standards for each process simulation. The
objective, time to complete the exercise, cause, effect, solution, and procedure for each
exercise is documented. You may modify these procedures to more closely reflect your
particular process plant operating procedures.

● Exercise 1: Design
● Exercise 2: Cold Start
● Exercise 3: TIC-110 Reads Low

Copyright © 2002 by Simtronics Corporation. All rights reserved worldwide.

For questions or comments concerning this web site, please contact the Webmaster.

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Atmospheric/Vacuum Distillation Units - Vacuum Distillation

SPM-2700 Atmospheric/Vacuum Distillation Units

Click Here to Return

Copyright © 2002 by Simtronics Corporation. All rights reserved worldwide.

For questions or comments concerning this web site, please contact the Webmaster

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Vacuum Distillation Unit - Schematic

SPM-2600 Fluidized Catalytic Cracking Unit

Click Here to Return

Copyright © 2002 by Simtronics Corporation. All rights reserved worldwide.

For questions or comments concerning this web site, please contact the Webmaster

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Vacuum Distillation Unit

SPM-2600 Vacuum Distillation Unit

Process Description

The VDU (Vacuum Distillation Unit)

takes the residuum from the ADU
(Atmospheric Distillation Unit) and
separates the heavier end products
such as vacuum gas oil, vacuum
distillate, slop wax, and residue.

Heavy crude oil is preheated by the

bottoms feed exchanger, further
preheated and partially vaporized in
the feed furnace, and passed into the
vacuum tower where it is separated
into slop oil, vacuum gas oil, vacuum
distillate, slop wax, and bottoms
residue.

Click here to view the schematic display.

This tower contains a combination of
14 fractionation trays and beds. It is
equipped with three side draws and
pump around sections for vacuum gas
oil, vacuum distillate, and slop wax
products.

The liquid from the feed furnace

enters the tower bottoms, where it is
collected and sent for further
processing. Steam is injected into the
base of the tower to reduce the
hydrocarbon partial pressure by
stripping some light boiling
components from the bottoms liquid.
The vapors from the feed heater enter
the tower below tray 14.

At tray 14, a draw pan is located from

which slop wax product is drawn. The
slop wax product and pump around
are cooled, with the slop wax product
going to storage, while the pump
around is returned to the tower at
tray 11.

The next product draw is located at

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Vacuum Distillation Unit

tray 8, where the draw for vacuum

distillate product is located. The
vacuum distillate draw tray is a total
draw tray, where the reflux from the
tray is pumped under flow control to
the tray below. The product and
pump around are cooled, with the
vacuum distillate product going to
storage, while the pump around is
returned to the tower at tray 7.

The last product draw is located at

tray 4, where the draw for vacuum
gas oil product is located. The
vacuum gas oil draw tray is also a
total draw tray, where the reflux from
the tray is pumped under flow control
to the tray below. The product and
pump around are cooled with the
vacuum gas oil product going to
storage, while the pump around is
returned to the tower at tray 1.

The overhead from the VDU is

condensed and combined with the
vacuum steam. The slop oil and water
are separated by gravity in the
vacuum drum. The water is drained to
disposal, while the slop oil is
accumulated and occasionally drained
to slop collection.

Process Specifications

The VDU fractionates 10.56 MBPD of Atmospheric Residuum to produce 0.39 MBPD of
Vacuum Gas Oil, 1.22 MBPD of Vacuum Distillate, 1.44 MBPD of Slop Wax and 6.07 MBPD
of Vacuum Residuum.

The VDU feed is heated to 750 Deg F before entering the tower which is maintained at
2.00 inHg. The top draw temperature is controlled at 310 Deg F which maintains the
Vacuum Gas Oil quality, and draw temperatures of 607 Deg F for the Vacuum Distillate,
and 668 Deg F for the Slop Wax.

Instrumentation

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Vacuum Distillation Unit

The VDU feed is pumped by P-114 (HS-114) controlled by FIC-124. It is preheated by the
bottoms feed exchanger E-200 before entering the Feed Furnace (F-200). TIC-200
controls the temperature of the feed entering the VDU (T-200) by adjusting fuel gas flow
to the furnace.

Bottoms liquid is collected and sent to storage through pump P-214 (HS-214), controlled
by LIC-214, and indicated by FI-224. This residue's 95% point is monitored by AI-224.
Stripping steam is injected into the VDU bottoms by FIC-234.

Hot slop wax is pumped from the tower by pump P-213 (HS-213). The slop wax product
flow to storage (FI-223) is controlled by LIC-213, and it's 95% point is monitored by AI-
123. Cooled pump around is controlled by FIC-213 and returned to the tower above the
slop wax draw tray.

Hot vacuum distillate is pumped from the tower by pump P-212 (HS-212). The vacuum
distillate product flow to storage (FI-222) is controlled by LIC-212, and it's 95% point is
monitored by AI-122. Cooled pump around is controlled by FIC-212 and returned to the
tower above the vacuum distillate draw tray. Vacuum distillate reflux is controlled by FIC-
232 and returned to the tower below the vacuum distillate draw tray.

Hot vacuum gas oil is pumped from the tower by pump P-211 (HS-211). The vacuum gas
oil product flow to storage (FI-221) is controlled by LIC-211, and it's 95% point is
monitored by AI-121. Cooled pump around is controlled by FIC-211 and returned to the
tower above the vacuum gas oil draw tray. Vacuum gas oil reflux is controlled by FIC-231
and returned to the tower below the vacuum gas oil draw tray.

The VDU overhead vapor flows through the overhead condenser E-210 (HV-212) into the
Overhead Vacuum Drum D-211. The hydrocarbons are fully condensed and mixed with
the vacuum condensate flow from E-211.

The water separates from the hydrocarbon liquid by gravity, where the drum's water level
is maintained by LIC-210 which sends the water to treatment via P-210 (HS-210). The
hydrocarbon phase overflows from the water phase, and slowly accumulates (LI-209).
When a sufficient level of slop oil has accumulated, the level can be drained via P-209
(HS-209).

The VDU vacuum pressure is maintained by the steam to the vacuum ejector (HV-211),
the cooling water (HV-212) to the steam condenser E-211, and the hydrocarbon
condenser E-210. The pressure is regulated by PIC-210, which reduces the vacuum by
circulating water to the vacuum ejector. The VDU bottoms pressure is indicated by PI-
214.

The VDU tower temperature profile is indicated by TI-210 (overhead), TI-211 (gas oil), TI-
212 (vacuum distillate), TI-213 (slop wax), and TI-214 (VDU residuum).

Advanced Controls

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Vacuum Distillation Unit

The Burner Ignitor Switch (HS-203) serves as both a burner ignitor and a flame indicator.
The fuel gas to the Feed Heater (F-200) will trip (TIC-200 mode=manual and output=0.0)
if the FD Fan (HS-202) is not on, the flame goes out, or the tube flow (FIC-124) is less
than it’s low alarm. This trip logic can be bypassed with HS-201.

Faults

All faults can be failed high or low to any degree with any of 8 fault function generators
(step change, square wave, staircase, stairs, ramp, sawtooth, slope, or sine wave). Faults
can be programmed to start and/or stop at various times during a simulation exercise.

● Fault 1: Pump P-114 ● Fault 14: FIC-211 Valve

● Fault 2: Pump P-209 ● Fault 15: LIC-212 Valve
● Fault 3: Pump P-210 ● Fault 16: FIC-232 Valve
● Fault 4: Pump P-211 ● Fault 17: FIC-213 Valve
● Fault 5: Pump P-212 ● Fault 18: FIC-234 Valve
● Fault 6: Pump P-213 ● Fault 19: FI-221 Transmitter
● Fault 7: Pump P-214 ● Fault 20: FI-223 Transmitter
● Fault 8: E-200 Hex ● Fault 21: PI-214 Transmitter
● Fault 9: Vacuum Efficiency ● Fault 22: TI-211 Transmitter
● Fault 10: Gas Oil Hex ● Fault 23: TI-213 Transmitter
● Fault 11: Vacuum Distillate Hex ● Fault 24: AI-221 Transmitter
● Fault 12: Slop Wax Hex ● Fault 25: F-200 FD Fan
● Fault 13: PIC-210 Valve ● Fault 26: F-200 Flame

Training Exercises

You may create a virtually unlimited number of scenarios and training exercises by
programming the faults described in the previous section. You can then establish
performance standards for each one of those exercises. Simtronics provides a number of
exercises with established performance standards for each process simulation. The
objective, time to complete the exercise, cause, effect, solution, and procedure for each
exercise is documented. You may modify these procedures to more closely reflect your
particular process plant operating procedures.

● Exercise 1: Design
● Exercise 2: Cold Start
● Exercise 3: Water Pump P-210 Fails

Copyright © 2002 by Simtronics Corporation. All rights reserved worldwide.

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Vacuum Distillation Unit

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 Technical Program Menu

[31e] - Novel Hightech Mass-Transfer Trays: GESIPHighspeedTray

and GESIPSieveTray

Presented at: [31] - Distillation Equipment and Applications # 1

For schedule information click here
Author Information:

Klaus R Hartmann (speaker)

GESIP Process Technologies
Rudower Chaussee 29
Berlin, 12489
Germany
Phone: #49 30 6789 2160
Fax: #49 30 6789 2161
Email: khartmann@gesip.de

Abstract:

Novel Hightech Mass-Transfer Trays:

GESIPHighspeedTray and GESIPSieveTray

Klaus Hartmann
GESIP Gesellschaft fuer Informations- und Prozesstechnik mbH, Berlin,
Technical University Berlin, Germany

The mass-transfer trays GESIPHighspeedTray and GESIPSieveTray represent a

new generation of mass transfer trays that enable increase the volume-specific
efficiency and the capacity of tray columns significantly.
This increase is achieved for the vortex - jet based GESIPHighspeedTray (HST)
by a special vapor and fluid distribution and contact system within the contact
and separation elements (CSE) installed on this tray and the special construction
and design of the tray itself.
The CSE consists of a short vertical cylinder with perforation and a special
constructed internal swirler on the inlet and a second, larger cylinder as a
separator on the outlet. The GESIPHighspeedTray increases the productivity /
capacity of mass transfer equipment of up to 3 to 5 times. The velocity factor F of
vapor in total section column with tray spacing 400-600 mm is from 3.5 to 11, the
liquid load up to 75 m3 /m2 h. The HST is self-regulating, the tray performance is
indifferent to deviations from the tray levelness. The new construction also offers
miniaturization of the columns, which results in significant savings of equipment
costs. The column can be preassembled for operation by the factory. The
preferred application fields are the natural gas and petroleum industries,
especially for the dehydration of natural gas using glycol but also for sweetening
(alkanolamine processes), absorption, desorption, as well as rectification of
hydrocarbon mixtures or other mixtures, and the utilization in environmental
technologies for gas scrubbing and in the power generation industry for the
separation of water droplets from high-pressure steam. The
GESIPHighspeedTray may also be used as a mist eleminator(demister).
The GESIPSieveTray represents a sieve tray with a special stabilizing froth
retainer in the form of a grid. This grid guarantees an uniform distribution for the
liquid phase as well as for the gas phase increasing the froth height up to 2 and
the gas velocity up to 4 m/s. The froth height on the tray is between 100 to 200
mm, so that a tray spacing of 250 to 300 mm and F-factors of 1.8 to 2.5 and
liquid loads from 0.1 to 50 m3 /m2 h are possible in columns with large diameters.
These columns operate very stable and efficient with low liquid loads. This new
type of construction increases the performance of frothing equipment to double
compared to the traditional construction. At the same time, it also improves
operating characteristics and the mass exchange efficiency (the surface area
between the phases is about 400 m2 /m3 ). GESIPSieveTray may be used in
distillation columns, in absorbers and gas washers with a high throughput and
performance, e.g., in the natural gas treatment industry, petrochemical industry,
and the petroleum processing industry, as well as in smaller and medium-sized
columns in separation plants and/or gas purification plants and as in gas
scrubbers for washing dust-containing and toxic gases.
Both novel internals are developed to meet the special needs and requirements
for offshore installations.
 MASS TRANSFER: DISTILLATION

Trouble-free design of
refinery fractionators
A review of factors most frequently the cause of distillation towers falling short
of design objectives. Analysis of case histories provides guidelines for identifying
potential troublespots in the most important fractionators
Henry Z Kister
Fluor Corporation

A
two-phase survey was recently refinery fractionators,
completed by Fluor of all the case and these form the
histories related to malfunctions in basis for the current
refinery towers that have been document- analysis.
ed over the last 50 years. Altogether, 400 As with other Fluor
case histories were found in the literature. surveys, certain
The first phase identified the most com- ground rules were
mon root causes of problems in refinery applied to limit the
fractionators (towers), but did not exam- scope. Only specific
ine the troublespots in each specific ser- incidents were
vice. This phase yielded general included. For exam-
guidelines for trouble-free design, but did ple, a statement such
not address issues related to each specific as “leakage from
fractionator. chimney trays in
In the second phase, case histories of refinery vacuum tow-
tower malfunctions are analysed specifi- ers can be reduced by
cally for each of the major refinery frac- seal-welding” does
tionators. Each case history teaches a not constitute a case Figure 1 Uplifted packing in wash section of a vacuum tower
lesson. Together, these lessons are the best history. On the other
tool for understanding the potential trou- hand, a statement such as “one vacuum It clearly shows that the vacuum tower is
blespots in each service, and for drawing tower experienced severe chimney tray by far the most troublesome refinery ser-
guidelines for trouble-free design of each leakage at low-rate operation. Seal weld- vice, which is where the survey begins.
service. ing tray sections reduced leakage to
I have previously described the Fluor acceptable levels” does. Vacuum tower malfunctions
survey methodology in Distillation Opera- Also, incidents of corrosion and foul- The 86 case histories reported for the vac-
tion (McGraw-Hill, New York,1990). All ing were included only if a feature unique uum tower is almost double the number
the case histories used as a basis for the to the column design, operation, or con- reported for the atmospheric crude tower,
survey were extracted from the published trol contributed to their occurrence. For which is the next most troublesome refin-
literature. There were 900 total cases, of instance, an incident where the wrong ery tower. When a vacuum tower per-
which 400 were for refinery towers. In corrosion inhibitor or antifoulant was forms poorly, valuable distillate is lost to
about one quarter of these, the specific applied does not qualify as a case history the resid, and poor distillate quality poi-
service was not stated or the service was in this survey. A case where fouling was sons FCC catalyst. The wash section of
one that did not have enough cases caused by insufficient liquid flow, maldis- the fractionator is the most critical sec-
reported on to permit detailed analysis. tribution, or poor process control, does. tion and also one where most of the mal-
This left about 300 cases for the main Finally, optimisation case studies
(where capacity was raised or pressure Top causes of vacuum tower
Fractionator malfunctions drop lowered by replacing trays by pack- malfunctions
ings) are outside of the scope of the sur-
Number of vey. The objective of the current survey is No. Description Cases
cases to identify the issues that make towers fall
1 Damage 27
1. Vacuum towers 86 short of achieving these design capacities.
2 Coking 21
2. Atmospheric crude fractionators 45 There is some overlap in the tabulation
3 Intermediate draws 17
3. Debutanisers 37 of cases for each fractionator. For 4 Misleading measurements 10
4. FCC main fractionators 33 instance, a coked chimney tray case study 5 Plugging 9
5. Deethanisers 23 will be listed once under “coking” and – Installation mishaps 9
6. Depropanisers, C3/C4 splitters 22 another time under “intermediate draws”. – Abnormal operation (startup,
7. Alky main fractionators/isostrippers 17 This means that adding the individual shutdown, commissioning) 9
8. Coker main fractionators 15 malfunctions may yield a number greater 8 Maldistribution 6
9. Naphtha splitters 11
than the number of malfunctions report- – Weeping 6
10. Deisobutanisers 8
ed for the service. Table 1 lists the main 10 Condenser 4
11. Amine towers 8
fractionators surveyed and the concise
Table 1 number of cases reported for each service. Table 2

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functions were reported. The wash sec-

tion of the vacuum tower is therefore the
most troublesome tower section in the
refinery (Figure 1, on previous page).
Table 2 shows the common malfunc-
tions reported in vacuum towers. With 27
case histories, damage tops the list. Most
of this damage can be readily prevented.
Table 3 shows the most common causes.
Foremost are water-induced pressure
surges, which account for one third of the
reported damage incidents. In three of the
nine cases reported, the source of water
was poor draining of stripping steam
lines. In another three, pockets of water
lying in the piping of spare pumps
entered the hot tower when these pumps
were connected to the hot tower.
A lesson in this case is that many, pos-
sibly most, vacuum tower damage inci-
dents can be prevented by design and
operating procedures that adequately
drain the tower steam lines in wet towers, Figure 2 Broken, non-standard flange in the spray header supplying wash oil to the
and that positively prevent water from wash bed of a vacuum tower
spare pump piping from entering the
tower. A joint designer/refiner “hazop” flanges and gaskets and properly allowing sory Committee, San Antonio, Texas, Nov
should focus on these troublespots. for thermal expansion in the header 2001].
The next source of damage in Table 3, design. Coking of the wash section (Figure 3) is
insufficient mechanical strength, is also Damage due to high base level con- a close second in Table 2 with 21 reported
readily preventable. It should be recog- tributed three out of the 27 damage case case studies. Table 4 gives a breakdown of
nised (as can be readily seen from Table 2), histories. This again is an issue that can be the causes. Excess stages and vaporisation
that damage is a major issue in a vacuum at least alleviated by good level monitor- occur in wash beds that are either too tall
tower, and that heavy duty internals ing, alarms, and well-designed trip sys- or contain packings that are too efficient.
design should be used. Although the tems. Another three damage-related In either case, the additional stages inten-
heavy duty design would not be able to case-histories were caused by packing sify the vaporisation of the wash oil, leav-
withstand a major pressure surge, it would fires. This type of damage is more difficult ing little liquid to reach and wet the lower
weather the smaller pressure surges. Some to prevent due to the difficulty of clean- sections of the bed. These lower sections
good heavy duty design practices have ing the packings, especially when coked. of the bed dry and coke. Poor modelling
been described by Shieveler [Shieveler G H, Nonetheless, much progress has been and simulation is another cause of coking.
Use heavy-duty trays for severe services; Chem reported in developing preventive mea- Golden et al stress that the heavy ends of
Eng Progr, Aug 1995]. sures, and is discussed in two excellent the crude must be correctly characterised
Special attention should be paid to grid papers by Bouck and Markeloff [Bouck D in the simulation and that the feed entry
installation and tightening. In two of the S, Vacuum Tower Packing Fires; API Operat- to the tower must be modelled by a series
five cases, poorly fastened grids disinte- ing Practices Symposium, 27 April 1999. of flash steps that correctly represent the
grated in service. Through-bolting has Markeloff R, Packing fires; FRI Technical Advi- physical sequence of steps between the
been far more effective than J-bolting for
keeping grid together, and should be rou-
tinely specified.
Spray distributors and their headers are
prone to damage (Table 3). Again, this
damage can be easily prevented by sound
design, good installation, and thorough
inspection and testing. Water testing
spray nozzles and headers can readily
detect damage (Figure 2). Header damage
can be prevented by using standard

Causes of damage in vacuum

fractionators

No. Description Cases

1 Water-induced pressure surges 9
2 Insufficient mechanical strength 5
3 Broken nozzles or headers of
spray distributors 4
4 High bottom liquid level 3
– Packing fires 3

Table 3 Figure 3 Coking of fouling resistant grid in the wash section of a vacuum tower

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chimney tray that caused entrainment to be troublesome in three cases. With

Causes of coking in vacuum into the wash bed. Three cases were trapout trays being used only in tray tow-
fractionators reported where maldistribution of vapour ers, this issue is experienced mainly in
or liquid led to coking. vacuum lube towers that have trays and
No. Description Cases Intermediate draw malfunctions is a not packing. The remaining case histories
1 Excess stages and vaporisation 4 close third of the most common vacuum describe coking and excessive hydraulic
– Poor modelling and simulation 4 tower malfunctions, with 17 reported gradients.
3 Insufficient wash, reason not case histories (Table 5). Foremost is leak- With 10 case histories, misleading mea-
reported 3 age from total draw chimney trays. Leak- surements are in the 4th spot in Table 2.
– Misleading measurement 3 age of the HVGO chimney tray represents Three of these 10 are the troublesome
– Liquid, vapour maldistribution 3 good distillate degraded into resid with chimney tray level measurements previ-
Not reported 4 no beneficial effects whatsoever, as leak- ously mentioned. Other troublesome
ing liquid poorly distributes in the wash cases have been reported with short coil
Table 4 bed and does little washing. Leakage of outlet thermocouples (two cases), ambi-
LVGO into the HVGO section lowers the ent changes affecting vacuum measure-
heater outlet and flash zone [Golden S W, HVGO boiling point and can reduce heat ments with ordinary gauges (two cases),
Vacuum Tower Troubleshooting; AIChE Spring transfer, even limiting vacuum on the bottom level, heater fuel flow rate, and
Meeting, 1994]. tower. reflux to a packed bed distributor. The
When these principles are overlooked, This leakage needs to be avoided with lessons from these case histories to instru-
the simulation underestimates wash oil totally seal-welded chimney trays. Special ment specifications are self-explanatory.
vaporisation, leading to the drying up and techniques, as recommended by Lieber- With nine case histories, plugging (as
coking reported in four cases. man, are effective and need to be incor- distinct from coking) is in the 5th spot in
In three other reported coking inci- porated to avoid tray buckling due to Table 2. Of the nine cases, five were plug-
dents, it was stated that the wash flowrate thermal expansion [Lieberman N P, Process ging of spray headers. One case was
was insufficient but no specific reason Design for Reliable Operation, 2nd ed; Gulf Pub- reported of plugged packing, plugged
was given. It is likely that in those cases lishing, Houston, Texas, 1988]. quench pipe, plugged instrument line
too, either the number of stages was Level measurement on chimney trays and plugged ejector. In two cases, the
excessive, or the modelling/simulation has been troublesome in three reported plugging was by corrosion products. Mea-
were poor, or both. In three other cases, cases. This may lead to overflow or sures found effective for alleviating plug-
coking was produced by either a mislead- entrainment. While overflow is equiva- ging in the wash spray headers, which is
ingly low coil outlet temperature signal lent to leakage, entrainment from the one of the most common troublespots,
that caused excessive firing, or a faulty overflash chimney tray can induce cok- are to provide good wash oil filtration and
level measurement on the overflash ing. Leaking trapout trays were reported to specify an all-stainless-steel wash oil

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fluid dynamics (CFD) is essential. Improv-

Intermediate draw malfunctions ing the vapour horn design has eliminat- Top causes of FCC main
in vacuum towers ed operating problems and improved the fractionator malfunctions
performance of several vacuum towers.
No. Description Cases A surprisingly high number of case
histories place tray weeping in the equal No. Description Cases
1 Leaking total draw chimney trays 6 1 Plugging 9
2 Level measurement on total draw 8th spot in Table 2. This is not an issue
with most vacuum towers that are all – Abnormal operation 9
chimney trays 3
3 Liquid maldistribution to
– Leaking trapout tray 3 packed, but appears to be a major issue packing liquid 6
4 Chimney tray coking 2 with the trayed vacuum towers, mostly
4 Intermediate draws 5
– Excessive hydraulic gradients on in lube service. In two of these, leakage
5 Water-induced pressure surges 4
chimney trays 2 at a draw tray made it impossible to draw
6 Pressure control 3
6 Others 1 sufficient product. In two others, weep- – Vapour maldistribution 3
age at pumparound trays starved the
Table 5 pump and reduced heat transfer. In two
more, poor separation between side-cuts Table 7
line downstream of the wash oil filter. resulted. Blanking valves, using high
Sharing the 5th spot in Table 2 is installa- turndown valves, and in the case of the and air premature upon shutdown,
tion mishaps. Three case histories were side cuts, drawing from seal-welded another from a packing fire at the
associated with spray header installation, chimney trays, was the solution. turnaround. The cause of damage in the
and another three with grid or packing Condenser issues complete Table 2. other three cases was not reported.
assembly. Other cases describe problems Condenser problems raise pressure in the Seven abnormal operation (startup/
with tower out-of-roundness, using car- tower and thus reduce distillate recovery. shutdown/commissioning) incidents
bon steel bolts where stainless steel was Two of the four cases reported dealt with were reported. Four of these resulted in
specified and poor installation of strip- excess lights, one with ejector plugging, four of the damage incidents listed, one
ping trays. and one with flash equilibrium at the pre- led to an explosion, another to a fire and
Also sharing the 5th spot in Table 2 is condenser. one to a chemical release. Three of the
abnormal operation incidents. Five of pressure surges listed in Table 6 under
these describe incidents during startup Crude tower malfunctions Damage resulted from poor dehydration
where a pocket of water entered the hot The three most common atmospheric during startup or pump switchover. Poor
tower and created a pressure surge. Poor crude tower malfunctions (Table 6) are blinding and unblinding led to one
blinding/unblinding contributed two case plugging, intermediate draw malfunc- reported case of explosion and another of
histories, one case is related to pressur- tions and damage. chemicals release.
ing/depressuring and one to flushing. There were nine plugging incidents The next four entries in Table 6 are well
Maldistribution problems, other than reported in atmospheric crude towers: below the top four, and have three to
those attributed to coking, plugging or Four in the wash section or gasoil four reported malfunctions. Four installa-
damage, are in the 8th spot in Table 2 pumparound, three in the top section or tion mishaps were reported, all involving
with a surprisingly low number of case top pumparound, and two in the strip- trays or chimney trays. Leaks of a
histories (six). Of the reported six, four ping section. No cases of plugging were pumparound exchanger, a pump seal,
were vapour maldistribution, three of reported in the middle of the tower. In the and resid to atmosphere were the three
these originating in the flash zone and wash section, the most common cause of reported leak cases. Controlling liquid
one in the previously mentioned tray plugging was entrainment from the flash flow to the wash section has been a spe-
chimney. The other two cases reported zone or from the vapour overhead of a cial challenge, contributing three more
liquid maldistribution problems. Fluor’s preflash drum. In the top of the tower, the case histories.
experience has been that maldistribution, plugging was by scale and corrosion prod- Trouble-free designs properly dis-
especially of vapour from the flash zone, ucts, corrosion inhibitors, and salting out. entrain the vapour in the flash zone and
has been far more troublesome than sug- Five of the nine incidents resulted in preflash drum and properly desalt the
gested by the low spot of this item in plugged trays, two in plugged downcom- crude in order to minimise plugging in
Table 2. ers. In one case, a packed bed plugged, in the fractionator. Specifying fouling-resis-
Vapour horn design and good distribu- another, a liquid distributor to the pack- tant hardware in the wash zone and
tion of liquid to the wash bed are central ing plugged. upper trays is good practice. Downcomer
for achieving trouble-free performance of The large number of intermediate draw trapouts and chimney trays are the most
the wash bed. An expert hydraulic analy- incidents is well in line with Fluor’s expe- important internals for ensuring trouble-
sis, often with the aid of computational rience: chimney trays and downcomer free operation. They need to be designed
trapouts make or break fractionators. Of and inspected carefully, not just left to
Top causes of atmospheric crude the nine, seven took place with down-
tower malfunctions comer trapouts, two with chimney trays. Top causes of malfunctions in
Four of these involved choking or restric- debutanisers, incl stabilisers and
tion in the outlet liquid line, while in two depentanisers
No. Description Cases
others, leakage at the drawoff restricted
1 Plugging 9
the recovery of a side cut. No. Description Cases
– Intermediate draws 9
Four of the nine damage incidents 1 Control 10
– Damage 9
4 Abnormal operations 7 reported were due to water-induced pres- 2 Vapour cloud release 5
5 Installation mishaps 4 sure surges. Two of these were caused by – Installation Mishaps 5
6 Condenser Problems 3 undrained stripping steam lines, one by a 4 Feed arrangement, tray towers 4
– Poor control of wash 3 water pocket in a spare pump, and one by – Reboiler draw arrangements 4
– Leaks 3 plugged drainholes in the bottom seal
pan. One case of damage resulted from
Table 6 exposing column internals to cold water Table 8

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incidents.There were nine plugging inci-

Top causes of malfunctions in dents reported in FCC main fractiona- Top causes of malfunctions in
C3/C4 splitters and depropanisers tors. Of these, four were salting-out deethanisers (absorbers and
(excl those in alky units) incidents that plugged trays near the top strippers)
of the tower, and were overcome by
online water washes. Three were inci- No. Description Cases
No. Description Case dents in which grid in the slurry section 1 Reboiler draw and return
1 Reboiler draw and return coked up due to vapour or liquid-maldis- arrangements 6
arrangements 7 tribution. One incident was plugging due 2 Excessive tower base level 4
2 Tower flooding by excess to catalyst carryover into upper packed – Control 4
base level 6
sections, and another was plugging of a – Component accumulation 4
3 Vapour cloud release 5 line draining the main feed line, both – Side draw arrangements 4
Table 10 during startup.
There were also nine abnormal opera- Table 9
others, as these will make or break the tion (startup/shutdown/commissioning)
fractionator. incidents reported, two of which were FCC main fractionators than in atmo-
Prevention of water entry, by ensuring previously described. These two inci- spheric crude fractionators, so it comes as
adequate drainage on stripping steam lines dents, plus two others, occurred during little surprise to find liquid maldistribu-
and eliminating dead pockets inside the liquid circulation and dehydration. In tion to packings in a prominent spot in
tower, is central for damage prevention. three of these four, a pressure surge and Table 7. Of the six reported incidents, two
Proper inspection of equipment is a must major damage resulted, the other was the involved liquid maldistribution to the
to prevent the installation mishaps. Proper catalyst carryover. The remaining five slurry pumparound section, the others to
control of the liquid flow rate to the wash incidents include poor unblinding caus- various fractionation sections. Intermedi-
section is the prime control consideration ing a toxic release; switching over oxygen ate draws in FCC main fractionator have
in the tower. and nitrogen purge gas causing explo- been troublesome in five reported case
sions; trip failure on the reflux drum caus- histories, more in chimney trays than in
Main column malfunctions ing liquid carryover and major downcomer trapouts. Finally, four cases
There are some similarities with regard to compressor damage; a major leak due to of water-induced pressure surges were
FCC main fractionator malfunctions the thermal shock while opening or clos- reported, three of which led to major
(Table 7) and atmospheric crude tower ing the valve in the tower inlet; and a damage.
malfunctions, but there are also major startup pressure control problem resulting Two other malfunctions are also shown
differences. Two malfunctions top the from steam condensation. in Table 7: Vapour maldistribution, all
list: plugging and abnormal operation Packings are used more frequently in cases dealing with grid in the slurry sec-

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tion, and pressure control. Plugging and There is a distinct link between the or pan causing liquid to bypass a once-
coking can be alleviated by providing ade- feeds and reboiler draw arrangements. through thermosiphon reboiler (two
quate on-line wash facilities near the top Both constitute “points of transition”, ie, cases); a reboiler tube leak and slug flow at
of the fractionator, by using plugging- where a stream enters or leaves the tower. the reboiler outlet pipe.
resistant trays there, and by preferring These points of transition are some of the Slightly behind, with six case histories, is
shed decks or disk and donut trays to grid major troublespots in a tower. The lesson tower flooding by excess base level. Two of
in the wash section. Shed decks and disk for debutanisers is that all points of transi- these resulted from the type of reboiler
and donut trays are far less sensitive to tion need to be critically examined for problems previously discussed. False level
vapour or liquid maldistribution than potential bottlenecks, both at the design indica tions led to two others, and frothing
grid, and therefore far less prone to cok- (or debottleneck) and when trouble- or foaming at the tower base led to the
ing during upsets. shooting. remaining two. Clearly, the lessons learned
Startups and shutdowns are major are that troubleshooting and trouble-free
issues in FCC main fractionators, and a Deethaniser malfunctions debottlenecks of C3/C4 splitters should focus
good design of these needs to hazop what The strippers and absorbers are included in on the reboiler piping and the bottom
can go wrong and take preventive mea- the deethaniser malfunctions. Topping the sump.
sures. Intermediate draws and liquid dis- list (Table 9) with six case histories are Similar to debutanisers, depropanisers
tributors are the weakest links in the reboiler draw and return arrangements. and C3/C4 splitters have experienced a high
internals design, and need to be designed Three of the six cases report excessive pres- number of vapour cloud releases, mostly
and inspected carefully, not just left to sure drop in the process inlet or outlet due to line rupture (three cases), but also
others. Finally, pressure controls as well as pipes of a kettle reboiler. The high-pressure due to poor blinding or plugging/freeze ups
liquid flow control to the wash section drop either caused the tower base liquid of valves. Some major blasts resulted. The
are major considerations in these frac- level to rise above the reboiler return inlet, vapour cloud lessons described under
tionators. or back liquid up on the chimney tray feed- debutanisers extend to depropanisers and
ing the reboiler to the top of the chimneys. C3/C4 splitters.
Debutaniser malfunctions Insufficient heat during coke drum
Due to similar functions, stabilisers and switchover was reported in two cases, one Other fractionators
depentenisers have been lumped together of them due to weeping from the draw tray For other refinery fractionators, the num-
with debutanisers. Over 70% of the cases, to a once-through thermosiphon reboiler. ber of case studies reported was less than
however, were contributed by debutanis- Four case histories were reported of base 20, a sample too small for a detailed anal-
ers. level exceeding the reboiler return. Two of ysis. Nonetheless, some observations are
Table 8 shows that the most common these were due to high-pressure drop in significant and require more detail, includ-
malfunctions experienced in debutanisers the kettle piping (those previously men- ing coker main fractionators, alky unit
are widely different from those experi- tioned), the other two due to absence of or main fractionators/isostrippers, naphtha
enced in the vacuum, crude and FCC frac- to poor level indication. As with debu- splitters, deisobutanisers and amine
tionators. Topping the list with 10 case tanisers, control issues are also important absorbers/regenerators.
histories is controls, an item that showed in deethanisers, and account for four case A total of 15 malfunction case histories
low down (if at all) on the main fraction- histories. Also, with four case histories, of coker fractionators have been report-
ator malfunctions list. Of the 10 cases, five component accumulation in deethanisers ed. Of the 15, seven described fouling by
reported difficulties with pressure and con- is a problem. coking or carryover of coke, while five
denser controls. In all five, a total con- Either ethane or water or both accumu- others described damage due to water-
denser was used with partial flooding of late and can lead to cycling, capacity bot- induced pressure surges. There is no
the condenser. In two of the five, the prob- tlenecks, and in the case of water, also doubt that coking and water-induced
lem was induced by presence of non-con- corrosion. Finally, choking of side draws pressure surges are the major issues with
densables. Composition control or the with entrained gas bubbles has been a these fractionators.
assembly of a control system contributed problem in four case histories. A total of 17 malfunctions have been
the other control case histories. The lessons learned from this documen- reported for alky unit main
Vapour cloud release and installation tation are that the points of transition in fractionators/isostrippers. Of these, four
mishaps share the second spot in Table 8. deethanisers (the side draws as well as the described plugging, mostly by scale or cor-
Three of the five case histories of vapour region below the bottom tray, including rosion products; three described explo-
clouds ended in explosions, and one more the reboiler draw and return lines) require sions, either due to vapour cloud release or
in a fire. Some of these were accompanied thorough design, review, and inspection, due to HF carryover in the hydrocarbons
by injuries and heavy damage. Line frac- and must not just be left to others. Preven- and a violent reaction in a caustic bed
ture (two incidents), poor blinding (two tion of component accumulation and care- downstream; and three others described
incidents), and freeze-ups in leaking ful review of the control systems are also accumulation of either ethane or water in
valves (two incidents) were some of the prime considerations that make the differ- the overhead system.
contributing factors. Hazops of debutanis- ence between a troublesome and trouble- A total of 11 naphtha splitter malfunc-
ers should consider some of the lessons free deethaniser. tions have been reported. Four of these
learned from previous vapour cloud reported plugging, mainly by scale and cor-
releases to positively eliminate further Splitter malfunctions rosion products; three reported reboiler
accidents. Malfunctions in C3/C4 splitters also include issues; two were a result of poor installa-
With four case histories, poor feed depropanisers other than those in alky tion; and two reported control problems.
arrangements closely follow, leading to a units, which are uniquely different (Table Control problems with other refinery
capacity bottleneck or an efficiency loss in 10). Similar to deethanisers, reboiler draw fractionators have also been reported. For
the feed region. Also with four case histories and return arrangements lead the list with example, of the eight total deisobutaniser
are reboiler draw arrangements, including seven reported cases. Again, the main prob- malfunctions that have been reported, six
vapour entrainment choking the reboiler lems have been excess pressure drop in involved control problems. Five of these
draw lines and liquid leaking from a trapout inlet and outlet lines of a reboiler causing were temperature control issues that can
tray to a once-through thermosiphon base liquid level to exceed the reboiler be particularly troublesome with narrow-
reboiler, thus “starving” the reboiler. return inlet (two cases); leaking draw tray boiling mixtures.

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Eight case histories were specifically Plugging, abnormal operation issues,

reported for refinery amine absorbers. Of liquid maldistribution to packings and
these, six reported foaming, and three grid, intermediate draws and water-
reported scale and corrosion products induced pressure surges are the key trou-
catalysing the foam and causing plugging. ble spots in FCC main fractionators. The
Seven other case histories of amine main causes of plugging in these fraction-
absorbers and/or regenerators were report- ators are salting out near the top and cok-
ed without stating whether they came ing of grid beds in the slurry section.
from refineries or natural gas plants. Of These can be alleviated by online water
these, foaming was the issue in five. There washes and by avoiding grid in the slurry
is no doubt that foaming, and to a lesser section, respectively. Tower dehydration
degree fouling with scale and and corro- and liquid circulation are important start-
sion products, is the prime issue in amine up operations that can turn troublesome
towers. and lead to pressure surges or catalyst car-
ryover into the less fouling-resistant
Lessons learned regions.
The vacuum tower is by far the most trou- Control issues, especially pressure con-
blesome refinery fractionator. Damage, trols, are the primary source of problems
wash bed coking, and intermediate draws in debutanisers. Vapour cloud releases
are prime trouble spots. Water-induced have led to explosions and major damage
pressure surges are the leading cause of in debutanisers. Installation mishaps,
damage. Most of the damage incidents in tower feed entry arrangements, and
the vacuum towers are preventable. reboiler draw arrangements have also
Hazoping the possibility of water entry, been major trouble spots. Critical review
using heavy duty mechanical designs, of the control system, especially the pres-
inspecting, testing and correctly design- sure/ condenser controls, learning from
ing spray headers, paying attention to past vapour-cloud accidents, using
level measurement, and applying special hazops to minimise the possibility of
designs and procedures to prevent pack- vapour cloud releases, and sound design
ing fires can drastically reduce damage of feed entry piping and of tower base
incidents. arrangements are the key for trouble-free
Wash bed coking can be alleviated by debutanisers.
using short beds of relatively inefficient In deethanisers, the points of transi-
packing by correctly simulating the wash tion (the side draw arrangements and the
zone and by avoiding insufficient wash. region below the bottom tray, including
Intermediate draw malfunctions can be the reboiler draw and return lines) are
alleviated by seal-welding draw trays and prime trouble spots and are key to trou-
water-testing them at the turnaround and ble-free design and operation. Prevention
by paying attention to reliable level mea- of water and ethane accumulation in this
surement on these trays. Other measures tower is also important.
that promote trouble-free operation are Reboiler draw and return arrange-
good instrument specifications, good fil- ments, and tower base level are the key to
tration of the wash oil followed by stain- trouble-free depropanisers and C3/C4
less steel piping downstream of the filters, splitters. Depropanisers also are prone to
and good distribution of vapour and of vapour cloud releases, and lessons
the wash oil to the wash bed. learned from past vapour cloud incidents
Plugging, intermediate draws, damage should be incorporated in the design and
and abnormal operation incidents are the operation of depropanisers and C3/C4
most troublesome malfunctions in atmo- splitters.
spheric crude towers. Plugging is most Trouble-free operation of coker frac-
common near the tower top, at the wash tionators focuses on preventing coking
zone, or in the stripping section. Choking and water-induced pressure surges; of
and leakage are the prime intermediate alky main fractionators focuses on plug-
draw issues. Water-induced pressure ging, vapour cloud and component accu-
surges are the most common cause of mulation prevention; of deisobutanisers
damage. Plugging problems can be allevi- on composition control; and of amine
ated by eliminating the source of fouling absorbers and regenerators on foaming
and/or by using plugging-resistant inter- and plugging prevention.
nals. Downcomer trapouts and chimney
trays need to be designed and inspected
carefully, not just left to others.
Prevention of water entry by proper Henry Z Kister is a Fluor Corporation
draining of stripping steam lines and by fellow and director of fractionation
sound dehydration procedures at startup technology at Aliso Viejo, California, USA.
is critical. Proper inspection of equip- He has over 25 years’ experience in design,
ment and adequate control of liquid control and startup of frationation processes
flowrate to the wash section are also and equipment. He obtained his BE and ME
important for promoting trouble-free degrees from the University of New South
operation. Wales, Australia.

115
P T Q AUTUIMN 2003
Troubleshooting crude vacuum tower overhead ejector systems
Use these guidelines to improve performance and product quality

J. R. LINES AND L. L. FRENS, GRAHAM MANUFACTURING CO. INC., BATAVIA, NEW YORK

R outinely surveying tower overhead vacuum systems can

improve performance and product quality. These vacuum sys-
tems normally provide reliable and consistent operation.
However, process conditions, supplied utilities, corrosion, erosion
and fouling all have an impact on ejector system performance.

Refinery vacuum distillation towers use ejector systems to main-

tain tower top pressure and remove overhead gases (Fig. 1).
However, as with virtually all refinery equipment, performance
may be affected by a number of variables. These variables may act
independently or concurrently. It is important to understand
basic operating principles of vacuum systems and how perform-
ance is affected by:

• Utilities Fig. 1. Twin-element, three-stage ejector system on crude vacuum

tower.
• Corrosion and erosion
• Fouling An ejector may operate unstably if it is not supplied with enough
energy to allow compression to its design discharge pressure. If
• Process conditions.
the actual motive steam pressure is below design or its tempera-
Reputable vacuum-system suppliers have service engineers that ture above design, then, within limits, an ejector’s nozzle can be
will come to a refinery to survey the system and troubleshoot per- rebored to a larger diameter. The larger nozzle diameter allows
formance or offer suggestions for improvement. A skilled more steam to flow through and expand across the nozzle. This
vacuum-system engineer may be needed to diagnose and remedy increases the energy available for compression.
system problems.
If motive steam supply pressure is more than 10% to 20% above
design, then too much steam expands across the nozzle. This
UTILITIES tends to choke the diffuser. When this occurs, less suction load is
handled by the ejector and vacuum tower top pressure tends to
When a vacuum system is initially designed, utilities are estab- rise. If an increase in tower top pressure is not desired, then ejec-
lished and the most extreme conditions are usually used for the tor nozzles must be replaced with ones with smaller throat
design basis. Once operating, actual utility supply conditions can diameters.
be different than those set at the design stage and vary occasional-
ly. Important utilities for ejector systems are motive steam and Steam quality is important. Wet steam can be damaging to an
cooling water. Motive steam pressure, quality and temperature are ejector system. Moisture droplets in motive steam lines are accel-
critical variables. Flowrate and inlet temperature are important for erated to supersonic velocities and become very erosive. Moisture
cooling water. in motive steam is noticeable when inspecting ejector nozzles.
Rapidly accelerated moisture droplets erode nozzle internals. They
Motive steam conditions. These are very important and have a etch a striated pattern on the nozzle’s diverging section and may
direct impact on an ejector’s operation. If motive steam supply actually wear out the nozzle mouth. Also, the inlet diffuser tapers
pressure falls below design, then the nozzle will pass less steam. and throat will have signs of erosion. The exhaust elbow at the
When this happens, the ejector is not provided with enough ener- ejector’s discharge can erode completely through. Severe tube
gy to compress the suction load to the design discharge pressure. impingement in the intercondenser can also occur depending
The same problem occurs when the supply motive steam temper- upon ejector orientation. To solve wet steam problems, all lines
ature rises above its design value. Result: Increased specific up to the ejector should be well insulated. Also, a steam separator
volume and, therefore, less steam passes through the nozzle.

Hydrocarbon Processing®, March 1995 1

 EJECTOR FUNDAMENTALS
The basic operating principle of an ejector is to convert pressure
energy into velocity. This occurs with adiabatic expansion of
motive steam across a converging/diverging nozzle from motive
pressure to suction load operating pressure. Supersonic velocity
from the nozzle mouth results. Typically, velocities of mach 3 to 4
are achieved.
In operation, motive steam expands to a pressure below the suction
pressure. This creates a driving force to bring the suction load into
the ejector. High-velocity motive steam entrains and mixes with
the suction load gas. The resulting mixture is still supersonic. As
this mixture enters the converging/diverging diffuser, high velocity
is reconverted into pressure. A diffuser’s converging section reduces
velocity as crossflow area is reduced. The diffuser’s throat is
designed to create a normal shock wave. A dramatic increase in
pressure occurs as the flow across the shock wave goes from super-
sonic to sonic to subsonic after the shock wave. In the diffuser’s
diverging section, cross-sectional flow area is increased and velocity
is further converted to pressure. Fig. 2 details ejector components
Fig. 2. Ejector components and pressure profile. and a pressure profile for an ejector having a compression ratio in
excess of 2:1.
Ejector systems are required to operate over a wide range of condi-
with a trap should be installed immediately before an ejector’s tions—from very light loads to loads above design. An ejector
motive steam inlet connection. In some cases, a steam superheater system must stably adapt to all anticipated operating conditions.
may be required. Determining the design non-condensable and light-end hydrocar-
bon loading is essential for stable operation. Furthermore, an
Wet steam can also cause performance problems. When water accurate understanding of system back pressure is important.
droplets pass through an ejector nozzle, they decrease the energy Ejector systems may be configured a number of different ways to
available for compression. The effect is a decrease in load han- offer flexibility in handling various feedstocks and differing refin-
dling ability. With extremely wet steam, the ejector may even ery operations. A single vacuum train with one set of ejectors and
condensers has the lowest initial capital cost, but flexibility is limit-
break operation. ed. Often, parallel ejector trains are installed for each stage. Each
parallel ejector will handle a percentage of the total loading. For
Cooling water. Ejector system intercondensers and intercon- example:
densers are designed to condense steam and condensible • Twin element ejectors, each designed for 50% of total load
hydrocarbons, and cool non-condensible gases. This occurs at a
pressure corresponding to the preceding ejector’s design discharge • Triple element ejectors, each designed for 40% of total load-
ing for 120% capacity
pressure and the following ejector’s design suction pressure. When
the cooling water supply temperature rises above its design value, • Twin element, 1⁄3:2⁄3 ejector trains
ejector system performance is penalized. A rise in cooling water • Other configurations.
temperature drives down a condenser’s available log-mean temper-
Parallel ejector trains allow one train to be shut down for mainte-
ature difference (LMTD). The condenser does not condense nance while the column operates at reduced conditions. Also, at
enough and more vapors are carried out with the non-condensible light loadings, a train may be shut down to conserve refinery oper-
gases as saturation components. A pressure drop increase across ating costs. Fig. 3 shows a typical vacuum tower ejector system
the condenser is noticeable. The ejector following this condenser with a triple element ejector and first intercondenser. The second
cannot handle the increased load at this pressure. Pressure rises intercondenser and aftercondenser are a single element.
and the preceding ejector does not have enough energy to dis-
charge to the higher pressure. Result: The preceding ejector
breaks operation and the system may become unstable. CORROSION AND EROSION

This also occurs if the cooling water flowrate falls below design. Corrosion may occur in ejectors, condensers or vacuum piping.
At lower-than-design cooling water flowrates, there is a greater Extreme corrosion can cause holes and air leaks into the system.
water temperature rise across a condenser. This also lowers This destroys vacuum system performance.
LMTD and the above situation occurs.
Erosion may occur within the ejectors. Poor steam quality and
Problems with cooling water normally occur during summer high velocities erode diffuser and motive nozzle internals. An
months. This is when the water is at its warmest and demands on ejector manufacturer will provide certified information that gives
refinery equipment are highest. If the cooling water flowrate or the motive nozzle and diffuser throat design diameters. If a rou-
temperature is off design then new ejectors or condensers may be tine inspection of these parts indicates an increase in
required to provide satisfactory operation. cross-sectional area over 7%, then performance may be compro-
mised and replacement parts will be necessary.

Hydrocarbon Processing®, March 1995 2

Corrosion is a result of improperly selected metallurgy. Ensure that
the most appropriate materials are used before replacing parts. A
common corrosion problem occurs when carbon steel tubing is
used in condensers. Although carbon steel may be suitable for the
crude feedstock handled, it is not always the best practical choice.
It does offer the initial advantage of lower capital cost. However,
operating problems far outweigh modest up-front savings.

Vacuum towers undergo periods of extended shutdown for routine

maintenance, revamp or other reasons. During this period, a con-
denser with carbon steel tubing will be exposed to air and will rust
and develop a scale buildup. When the system starts up, the con-
densers are severely fouled. They will not operate as designed and
vacuum system operation is compromised. Modest savings in ini-
tial investment for steel tubing is quickly lost with less-than-
optimal tower operations due to rusted and scaled tubing. Vacuum
system manufacturers often caution against using carbon steel tub-
ing.

Fouling. Intercondensers and aftercondensers are subject to foul-

Fig 3. Typical multi-stage system.
ing like all other refinery heat exchangers. This may occur on the
tubeside, shellside or both. Fouling deters heat transfer and, at
some point, may compromise system performance. Process conditions. These are very important for reliable vacuum
system operation. Process conditions used in the design stage are
Cooling tower water is most often used as the cooling fluid for vac- rarely experienced during operation. Vacuum system performance
uum condensers. This water is normally on the tubeside. Typical may be affected by the following process condition variables that
fouling deposits on tubing internals cause a resistance to heat trans- may act independently or concurrently:
fer. Over a prolonged period of time, actual fouling may exceed the
design value and condenser performance falls short of design. • Non-condensible gas loading, either air leaks or light-end
hydrocarbons
Vacuum tower overhead gases, vapors and motive steam are nor-
• Condensible hydrocarbons
mally on the condenser’s shellside. Depending on tower
fractionation and the type of crude processed, a hydrocarbon film • Vacuum tower loading
may develop on the tube’s outside surface. This film is a resistance
• Vacuum system back pressure
to heat transfer, and over time, this fouling will exceed design.
Once this occurs, condenser performance falls short. • Condenser condensate barometric leg.

Routine refinery procedures should include periodic cleaning of

condenser bundles. Cleaning procedures must be for the con-
denser’s tubeside and shellside. To facilitate shellside cleaning, a
common practice is to specify removable tube bundles (i.e.,
TEMA designation BXU, AXS or AXT).

Fig 4. Typical first-stage ejector operating curve. Fig. 5. Preferred recycle control scheme to maintain tower pres-
sure at design when handling overheads below design.

Hydrocarbon Processing®, March 1995 3

Non-condensible loading. Vacuum systems are susceptible to Non-condensible loadings must be accurately stated. If not, any
poor performance when non-condensible loading increases above vacuum system is subject to performance shortcomings. If non-
design. Non-condensible loading to a vacuum system consists of condensible loadings are consistently above design, then new
air leaking into the system, light-end hydrocarbons and cracked ejectors are required. New condensers may be required depending
gases from the fired heater. The impact of higher-than-design on severity.
non-condensible loading is severe. As non-condensible loading
increases, the amount of saturated vapors discharging from the Condensible hydrocarbons. Tower overhead loading consists of
condenser increases. The ejector following a condenser may not steam, condensible hydrocarbons and noncondensibles. As differ-
handle increased loading at the condenser’s design operating pres- ent crude oils are processed or refinery operations change, the
sure. The ejector before the condenser is not designed for a higher composition and amount of condensible hydrocarbons handled
discharge pressure. This discontinuity in pressure causes the first by the vacuum system vary. A situation may occur where the con-
ejector to break operation. When this occurs, the system will densible hydrocarbon loading is so different from design that
operate unstably and tower pressure may rapidly rise above design condenser or ejector performance is adversely affected.
values.

Hydrocarbon Processing®, March 1995 4

This may occur a couple of different ways. If the
condensing profile is such that condensible hydro-
carbons are not condensed as they were designed
to, then the amount vapor leaving the condenser
increases. Ejectors may not tolerate this situation,
resulting in unstable operation.

Another possible effect of increased condensible

hydrocarbon loading is an increased oil film on the
tubes. This reduces the heat-transfer coefficient.
Again, this situation may result in increased vapor
and gas discharge from the condenser. Unstable
operation of the entire system may also result.

To remedy performance shortcomings, new con-

densers or ejectors may be necessary.
Readings that should be taken include:
Tower overhead loading. In general, a vacuum system will track
tower overhead loading as long as noncondensible loading does • Motive steam pressure and temperature measured as close as
not increase above design. Tower top pressure follows the first- possible to each ejector’s inlet
stage ejector’s performance curve. Fig. 4 shows a typical
• Suction and discharge pressure and temperature of each ejector
performance curve. At light tower overhead loads, the vacuum
system will pull tower top operating pressure down below design. • Cooling water inlet and outlet temperature of each condenser
This may adversely affect tower operating dynamics and pressure
• Tubeside pressure drop across each condenser
control may be necessary.
• Condensate temperature from each condenser, if available.
Tower pressure control is possible with multiple element trains.
At reduced overhead loading, one or more parallel elements may With these readings, a step-by-step comparison against the origi-
be shut off. This reduces handling capacity, permitting tower nal design criteria should allow a determination of the cause for a
pressure to rise to a satisfactory level. If multiple trains are not deficiency in the system.
used, recycle control is another possible solution. Here, the dis-
charge of an ejector is recycled to the system suction. This acts as Tables 1 and 2 will help in analyzing the system to determine any
an artificial load, driving the suction pressure up. With a multi- problems. Table 1 pertains to ejector performance and the areas
ple-stage ejector system, recycle control should be configured to that effect an ejector, namely, motive steam pressure and quality,
recycle the load from before the first condenser back to system suction load and discharge pressure. Table 2 relates to condenser
suction (Fig. 5). This way, noncondensible loading is not allowed performance.
to accumulate and negatively impact downstream ejectors.
If there are still questions as to the exact problem after complet-
System back pressure. Vacuum system back pressure may have ing the above evaluation, it is advisable to contact the original
an overwhelming influence on satisfactory performance. Ejectors manufacturer or have a qualified vacuum service engineer visit the
are designed to compress to a design discharge pressure. If the site to help analyze the system.
actual discharge pressure rises above design, the ejector will not
have enough energy to reach the higher pressure. When this
occurs, the ejector breaks operation and there is a sharp increase CASE HISTORY
in suction pressure.
Operating survey of a vacuum system on a crude tower in a
When back pressure is above design, possible corrective actions are South American refinery. Problem: The refiner was dissatisfied
to lower the system back pressure, rebore the steam nozzle to permit with the vacuum tower operating pressure. Tower top pressure
the use of more motive steam or install a completely new ejector. exceeded its design value. Table 3 shows the results of the system
operating survey that was conducted.
Operating survey. The most practical way to troubleshoot a vacu-
um system is to perform a pressure and temperature profile while The system was shut down and equipment inspected.
the system is in normal operation. Then compare the readings Dimensionally, the ejectors were in satisfactory condition, but the
with the origin design criteria. Before performing a survey, it is second-stage ejector showed signs of erosion. The second-stage
critical to have accurate instrumentation like calibrated ther- ejector diffuser throat was 4% to 5% larger in diameter than
mometers, pressure gauges and an absolute pressure gauge for design. Each ejector had a heavy hydrocarbon film on the motive
vacuum readings. steam nozzle exterior and air chamber.

Hydrocarbon Processing®, March 1995 5

The first intercondenser was heavily fouled on the tube and shell- THE AUTHORS
side. The tubeside had an excessive white scale that caused the
high tubeside pressure drop. The shellside had a heavy black film James R. Lines is vice president of engineering
coating the tubes. This was similar to what was noticed in the for Graham Manufacturing Co., Inc., Batavia,
ejectors. The secondary condensers had similar fouling deposits. New York. Since joining Graham in 1984, he
has held positions as an application engineer,
Solutions: The higher-than-design cooling water temperature product supervisor and sales engineer focusing
and excessive fouling of the condensers affected condenser per- on vacuum and heat transfer processes. Mr
formance to the extent that inadequate condensation was taking Lines holds a BS degree in aerospace engineer-
place. The impact of this was with the ejectors. Higher vapor ing from the University of Buffalo.
loads exiting the first intercondenser could not be handled by the
second-stage ejector at 86 mmHg Abs. This forced the first-stage Lance L. Frens is a senior contract engineer/
ejector to break operation resulting in a substantial increase in technical services supervisor for Graham
tower operating pressure. Manufacturing Co., Inc., Batavia, New York.
Since joining Graham in 1967, he has held posi-
Quotes were made to replace existing condensers with new ones tions as a test technician and application engineer.
having higher design fouling factors and based on higher cooling Mr Frens has 27 years of hands-on experience
water temperature. Once installed, tower top pressure was at surveying, revamping and troubleshooting vacu-
design and operation was stable. um and heat transfer systems worldwide.

Hydrocarbon Processing®, March 1995 6

 What Caused Tower Malfunctions
in the Last 50 Years?

Henry Z. Kister
Fluor Corporation, Aliso Viejo, California, USA

Presented at the Spring Distillation Symposium

South Texas Section of the AIChE - 3 April 2003 - Houston, Texas

AV\20030099005.ppt 1
 Are Malfunctions Repetitious?

Q:
Q: HOW
HOW MANY
MANY CASE
CASE HISTORIES
HISTORIES WERE
WERE REPORTED
REPORTED WITH
WITH
THE
THE FOLLOWING
FOLLOWING SEQUENCE
SEQUENCE OF
OF EVENTS?
EVENTS? (1986
(1986 -- 2001)
2001)

Faulty bottom level indication

Level rises above reboiler return

Tower prematurely floods

Possibly followed by tray/packing damage, excess

level in reflux drum, liquid discharges?

CHOICES: (a) 1 (b) 3 (c) 6 (d) 12

AV\20030099005.ppt 2
 Are Malfunctions Repetitious?

ANSWER: (d) 12

AV\20030099005.ppt 3
 2002 Malfunctions Survey

900 CASE HISTORIES

OPEN LITERATURE, LAST 50 YEARS

ABOUT 50% REPORTED IN THE LAST DECADE

40% CHEMICAL, 40% REFINERY, 20% OLEF / GAS

EXCLUDES “INHERENT” ISSUES

(e.g., fouling because inhibitor not injected)

SOME OVERLAP
(e.g., distributor plugging shows under distributors & plugging)

AV\20030099005.ppt 4
 Top 20 Malfunctions
CASES 1992+
11 FOAMING 51 26
12 SIMULATIONS 47 35
13 LEAKS 41 22
14 COMPOSITION CONTROL 33 16
15 CONDENSERS 31 12
16 CONTROL ASSEMBLY 29 16

17 PRESSURE & CONDENSER CONTROLS 29 14

18 OVERPRESSURE RELIEF 24 10
19 FEED INLETS TO TRAYS 18 11
20 FIRES (NO EXPLOSIONS) 18 8
AV\20030099005.ppt 5
 Top 10 Malfunctions
CASES
CASES 1992+
1992+ REF
REF CHEM
CHEM

10 CHEMICAL EXPLOSIONS 53 17 11 34

AV\20030099005.ppt 6
 Top 10 Malfunctions

CASES 1992+
CASES 1992+ REF
REF CHEM
CHEM

9 REBOILERS 62 28 28 13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53
53 17
17 11
11 34
34

AV\20030099005.ppt 7
 Top 10 Malfunctions

CASES 1992+
CASES 1992+ REF
REF CHEM
CHEM

8 MISLEADING MEASUREMENTS 64 31 31 9
99 REBOILERS
REBOILERS 62
62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53
53 17
17 11
11 34
34

AV\20030099005.ppt 8
 Top 10 Malfunctions

CASES
CASES 1992+
1992+ REF
REF CHEM
CHEM

7 INTERMEDIATE DRAWS 68 45 50 10
88 MISLEADING
MISLEADING MEASUREMENTS
MEASUREMENTS 64
64 31
31 31
31 99
99 REBOILERS
REBOILERS 62
62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53
53 17
17 11
11 34
34

AV\20030099005.ppt 9
 Top 10 Malfunctions

CASES 1992+
CASES 1992+ REF
REF CHEM
CHEM

6 PACKING LIQUID DISTRIBUTORS 74 48 18 40

77 INTERMEDIATE
INTERMEDIATE DRAWS
DRAWS 68
68 45
45 50
50 10
10
88 MISLEADING
MISLEADING MEASUREMENTS
MEASUREMENTS 64
64 31
31 31
31 99
99 REBOILERS
REBOILERS 62
62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53
53 17
17 11
11 34
34

AV\20030099005.ppt 10
 Why Liquid Distributors Don’t Work
CASES
1 PLUGGING 20
2 OVERFLOW 17
3 POOR IRRIGATION 14
4 FABRICATION / INSTALLATION 13
5 FEED ENTRY 11
6 HOLE PATTERN 8
7 DAMAGE 8

AV\20030099005.ppt 11
 Case 956 - The Overlong Feed Pipe

FEED PIPE
PARTING
BOX

1.5 mm

Based on OLSSON, R.F., CHEM ENG PROG., Oct 1999

AV\20030099005.ppt 12
 Top 10 Malfunctions

CASES
CASES 1992+
1992+ REF
REF CHEM
CHEM

5 ASSEMBLY MISHAPS 75 36 23 16
66 PACKING
PACKING LIQUID
LIQUID DISTRIBUTORS
DISTRIBUTORS 74
74 48
48 18
18 40
40
77 INTERMEDIATE
INTERMEDIATE DRAWS
DRAWS 68
68 45
45 50
50 10
10
88 MISLEADING
MISLEADING MEASUREMENTS
MEASUREMENTS 64
64 31
31 31
31 99
99 REBOILERS
REBOILERS 62
62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53
53 17
17 11
11 34
34

AV\20030099005.ppt 13
 Assembly Mishaps
CASES
1 PACKING LIQUID DISTRIBUTORS 13
2 POOR PACKING ASSEMBLY 13
3 UNTIGHTENED NUTS, BOLTS, CLAMPS 9
4 POOR TRAY PANEL ASSEMBLY 8
5 OBSTRUCTION OR MISORIENTATION AT FEEDS / 7
DRAWS
6 LEAKING COLLECTORS & LOW LIQUID LOAD TRAYS 7
7 DC CLEARANCES / INLET WEIR 5
8 DEBRIS LEFT IN COLUMN 5
9 TRAY MANWAYS LEFT UNBOLTED 4
10 INFERIOR MATERIALS OF CONSTRUCTION 4
AV\20030099005.ppt 14
 Top 10 Malfunctions

CASES 1992+
CASES 1992+ REF
REF CHEM
CHEM

4 ABNORMAL OPERATION INCIDENTS 84 25 35 31

55 ASSEMBLY
ASSEMBLY MISHAPS
MISHAPS 75
75 36
36 23
23 16
16
66 PACKING
PACKING LIQUID
LIQUID DISTRIBUTORS
DISTRIBUTORS 74
74 48
48 18
18 40
40
77 INTERMEDIATE
INTERMEDIATE DRAWS
DRAWS 68
68 45
45 50
50 10
10
88 MISLEADING
MISLEADING MEASUREMENTS
MEASUREMENTS 64
64 31
31 31
31 99
99 REBOILERS
REBOILERS 62
62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53
53 17
17 11
11 34
34

AV\20030099005.ppt 15
 Abnormal Operation Gone Wrong

CASES
1 BACKFLOW 15
2 BLINDING / UNBLINDING 15
3 WATER REMOVAL (REFINERY ONLY) 15
4 WASHING 12
5 STEAM & WATER OPERATION 10
6 OVERHEATING 7
7 PRESSURING / DEPRESSURING 6
8 OVERCHILLING (OLEF / GAS PLANTS) 6

AV\20030099005.ppt 16
 Cases 1108, 1109 - Reverse Flow Thru
Unblinded Valves Blowdown header
contained toxic gas

Valve closed,
but not blinded.
Valve leaked at
shutdown

Product
storage

Column end open Toxic gas came

out here and killed
Product came for maintenance
the operator Column was
out here
drained here

AV\20030099005.ppt (Based on Kletz, T.A. “What Went Wrong”, 2nd Ed, Gulf, 1988) 17
 Top 10 Malfunctions

CASES 1992+
CASES 1992+ REF
REF CHEM
CHEM

3 INTERNALS DAMAGE (Excl.

(Excl. Explosion, Fire) 84
Explosion, Fire) 35 35 33
4 ABNORMAL OPERATION INCIDENTS 84 25 35 31
55 ASSEMBLY
ASSEMBLY MISHAPS
MISHAPS 75 36
36 23
23 16
16
66 PACKING
PACKING LIQUID
LIQUID DISTRIBUTORS
DISTRIBUTORS 74 48
48 18
18 40
40
77 INTERMEDIATE
INTERMEDIATE DRAWS
DRAWS 68 45
45 50
50 10
10
88 MISLEADING
MISLEADING MEASUREMENTS
MEASUREMENTS 64 31
31 31
31 99

99 REBOILERS
REBOILERS 62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53 17
17 11
11 34
34

AV\20030099005.ppt 18
 What Damaged Tower Internals*
CASES
CASES
1 WATER- INDUCED PRESSURE SURGES 26
(90% Refinery)
22 INSUFFICIENT
INSUFFICIENT MECHANICAL
MECHANICAL RESISTANCE
RESISTANCE 10
10
33 HIGH
HIGH BASE
BASE LIQUID
LIQUID LEVEL
LEVEL 88
44 VALVE
VALVE TRAY
TRAY DOWNFLOW
DOWNFLOW 88
55 UPLIFT
UPLIFT DUE
DUE TO
TO RAPID
RAPID UPFLOW
UPFLOW 77
66 BREAKAGE
BREAKAGE OF
OF PACKING
PACKING 66
77 MELTING
MELTING // SOFTENING
SOFTENING OF
OF PACKING
PACKING 55
88 POOR
POOR ASSEMBLY
ASSEMBLY // FABRICATION
FABRICATION 55

* (EXCLUDES FIRE, EXPLOSION, IMPLOSION)

AV\20030099005.ppt 19
 Top 10 Malfunctions
CASES
CASES 1992+
1992+ REF
REF CHEM
CHEM

2 TOWER BASE 103 51 51 22

3 INTERNALS DAMAGE (Excl.
(Excl. Explosion,
Explosion, Fire)
Fire) 84 35 35 33
4 ABNORMAL OPERATION INCIDENTS 84 25 35 31
55 ASSEMBLY
ASSEMBLY MISHAPS
MISHAPS 75 36
36 23
23 16
16
66 PACKING
PACKING LIQUID
LIQUID DISTRIBUTORS
DISTRIBUTORS 74 48
48 18
18 40
40
77 INTERMEDIATE
INTERMEDIATE DRAWS
DRAWS 68 45
45 50
50 10
10
88 MISLEADING
MISLEADING MEASUREMENTS
MEASUREMENTS 64 31
31 31
31 99

99 REBOILERS
REBOILERS 62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53 17
17 11
11 34
34

AV\20030099005.ppt 20
 Tower Base: #1 Troublesome Internal

CASES
1 HIGH BASE LEVEL 49
2 VAPOR MALDISTRIBUTION 13

3 IMPINGEMENT 10

4 WATER-INDUCED PRESSURE SURGES 8

5 LEAKING DRAW TO ONCE-THRU THERMOSIPHON 7

6 LOW BASE LEVEL 7

7 GAS ENTRAINMENT 6

AV\20030099005.ppt 21
 Top 10 Malfunctions
CASES
CASES 1992+
1992+ REF
REF CHEM
CHEM

1 PLUGGING, COKING 121 70 68 32

2 TOWER BASE 103 51 51 22
3 INTERNALS DAMAGE (Excl.
(Excl. Explosion,
Explosion, Fire)
Fire) 84 35 35 33
4 ABNORMAL OPERATION INCIDENTS 84 25 35 31
55 ASSEMBLY
ASSEMBLY MISHAPS
MISHAPS 75 36
36 23
23 16
16
66 PACKING
PACKING LIQUID
LIQUID DISTRIBUTORS
DISTRIBUTORS 74 48
48 18
18 40
40
77 INTERMEDIATE
INTERMEDIATE DRAWS
DRAWS 68 45
45 50
50 10
10
88 MISLEADING
MISLEADING MEASUREMENTS
MEASUREMENTS 64 31
31 31
31 99

99 REBOILERS
REBOILERS 62 28
28 28
28 13
13
10
10 CHEMICAL
CHEMICAL EXPLOSIONS
EXPLOSIONS 53 17
17 11
11 34
34

AV\20030099005.ppt 22
 Sources of Solids in Towers

CASES
CASES 1992+
1992+ REF
REF CHEM
CHEM O/G
O/G

1 COKING 26* 22 25 1 –
2 CORROSION PRODUCTS 22 11 14 4 4
3 PRECIPITATION, SALTING OUT 17 13 9 8 –
44 SOLIDS
SOLIDS IN
IN FEED
FEED 10
10 44 33 66 11

55 POLYMER
POLYMER 99 55 –– 44 44

* 17 of these in Refinery Vacuum Towers

AV\20030099005.ppt 23
 Where Do Towers Plug?
CASES
CASES

1 PACKING 50
• BEDS 32*
• LIQ DISTRIBUTORS 20

2 TRAYS 39
• ACTIVE AREAS 32
•• DOWNCOMERS
DOWNCOMERS 99

33 DRAW
DRAW LINES
LINES 16
16
44 INSTRUMENT
INSTRUMENT LINES
LINES 10
10
55 FEED
FEED LINES
LINES 66
* 17 of these in Refinery Vacuum Towers
AV\20030099005.ppt 24
 Conclusions

1. PLUGGING / COKING UNDISPUTED #1 MALFUNCTION

Refineries: Coking, corrosion products, salting out
Chem: Precipitation, solids in feed, polymer, corrosion products
Packings, distributors & tray active areas plug most

2. TOWER BASE IS #1 TROUBLESOME INTERNAL

~ 50% of cases bottom level above vapor inlet
vapor maldistribution, impingement also troublesome

3. WATER- INDUCED SURGES #1 CAUSE OF INTERNALS DAMAGE

Low mech. resistance, high base level, rapid upflow, valve tray downflow

AV\20030099005.ppt 25
 Conclusions (continued)

4. ABNORMAL OPERATION MISHAPS ARE ON THE WAY DOWN

Backflow, blinding, water removal, washing, steam/water lead

5. ASSEMBLY MISHAPS LEVELED OFF

Distributor, tray assemblies and untightened clamps / bolts lead

6. PACKING LIQUID DISTRIBUTORS #1 MALFUNCTION IN CHEMICALS

Plugging, overflow, poor irrigation, fabrication / assembly, feed entry lead

7. INTERMEDIATE DRAWS ARE TROUBLESOME

Leakage, vapor choke dominate

8. MISLEADING MEASUREMENTS, REBOILERS, CHEM EXPLOSIONS

COMPLETE TOP 10
AV\20030099005.ppt 26
 Refinery Technology Online

INDEX

Students Guide to Refinery Processes

Isomerization

Bitumen Blowing

Cat Cracking

Crude Distillation

Gas Treating and Sulfur Recovery

Hydrotreating

Thermal Cracking

Vacuum Distillation

Process Workshop

A better Way to Clean

New Cleaning Technology Used on Texas Towers

Gamma scanning for troubleshooting of distillation columns

★
New industrial distillation technology - Linas distillation technology

Refinery configuration - how is it done

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Source: Petroleum Refinery Processes - OSHA Technical Manual

The isomerisation process involves the transformation of one molecular structure into another (isomer)
whose component atoms are the same but arranged in a different geometrical structure. Since isomers
Search may differ greatly in physical and chemical properties, isomerisation offers the possibility of converting
less desirable compounds into isomers with desirable properties, in particular to convert n-paraffins into
iso-paraffins, thereby increasing the octane of the hydrocarbon stream. The main fields of application of
isomerisation are:

Extra Resources ISOMERISATION of normal butane into isobutane

ISOMERISATION of pentanes and hexanes into higher- branched isomers
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Since branched isomers have a higher antiknock quality than the corresponding linear paraffins, this form
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Calculation xylene into para- xylene, used for the manufacturing of polester fibres.
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Isomerisation of low molecular weight paraffins has been commercially applied for many years. After
extensive laboratory work had been carried out during the 1930s, World War 2 prompted the development
RTOL?s Internet of the laboratory processes into full- scale commercial units in order to meet the demand for isobutane
Resources necessary for the manufacture of large amounts of alkylate. While the first butane isomerisation unit went
on stream in late 1941, by the end of the war nearly 40 butane isomerisation units were in operation in the
All Resources USA and the Caribbean. Two pentane and two light naphtha isomerisation units also came on stream
Articles towards the end of the war to provide an additional source of blending aviation gasoline.
Best Practises
Catalyst Catalogs Though butane isomerisation has maintained its importance, present day interest isomerisation is specially
Discussion Groups focussed on the upgrading of fractions containing C5 Pentane and C6 Hexane for use as motor gasoline
Events components. This application has been prompted by the world drive to remove the lead additives gradually
Glossaries from motor gasoline in order to reduce air pollution. The octane loss caused by the removal or reduction of
Jobs lead antiknock additives can be compensated for by isomerisation of pentane/hexane paraffin fraction of
Learning the light gasoline fractions.
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Software order to achieve the low temperature necessary to obtain an acceptable yield of isomers, the Catalyst
Success Stories systems used in the early units were based on aluminium chloride in some form. These catalyst systems,
Training however, had the drawback of being highly corrosive and difficult to handle. In recent years, catalyst of a
Websites different type have come in use. These are solid catalysts consisting of a support having an acidic carrier
and a hydrogenation function, frequently a noble metal. Modern isomerisation units utilise these dual-
function catalysts and operate in the vapour phase and the presence of hydogen. For these reasons, these
process are called hydro- isomerisation processes.

The first hydro- isomerisation unit was introduced nin 1953 by UOP, followed in 1965 by the first BP one,
while in 1970 the first Shell hydro-isomerisation (HYSOMER) unit was started up. At present the following
hydro-isomerisation processes are commercially available:

UOP BUTAMER for butane isomerisation

UOP PENEX for pentane/hexane isomerisation
BP C4 isomerisation for butane isomerisation
BP C5/C6 isomerisation for pentane/hexane isomerisation
SHELL Hysomer for pentane/hexane isomerisation

All these processes takes place in the vapour phase on a fixed bed catalyst containing platinum on a solid
carrier.

As an example, the Shell Hysomer process will be briefly described. The liquid feedstock is pentane/hexane
from light naphtha. naphtha splitters are widely used to split light naphtha, heavy naphtha and also LPG.
The light naphtha (C5/C6) is combined with the recycle gas/ fresh gas mixture. The resultant combined
reactor feed is routed to a feed/ effluent heat exchanger, where it is heated and completely vaporised by
the effluent of the reactor. The vapourised combined reactor feed is further heated to the desired reactor
inlet temperature in the reactor charge heater. The hot charge enters the Hysomer reactor at the top and
flows downwards through the catalyst bed, where a portion of normal and mono- branched paraffins is
converted into higher branched (high octane) components. Temperature rise from the heat of reaction
release is controlled by a cold quench gas injection into the reactor. Reactor effluent is cooled and
subsequently separated in the product separator into two streams: a liquid product (isomerate) and a
recycle gas stream returning to the reactor via the recycle gas compressor.

The catalyst is a dual function catalyst consisting of platinum on a zeolite basis, highly stable and
regenerable.

Temperatures and pressure vary in a range of 230 - 285 degrees C and 13-30 bar, C5/C6 content in
product relative to that in feed is 97% or better, and octane upgrading ranges between 8 and 10 points,
depending on feedstock quality. <>The Hysomer process can be integrated with catalytic reformer,
resulting in substantial equipment savings, or with iso-normal separation processes which allows for a
complete conversion of pentane/hexane mixtures into isoparaffin mixtures. An interesting application in
this field is the total isomerisation process (TIP) in which the isomerisation is completely integrated with a
Union Carbide molecular sieve separation process or the naphtha IsoSiv Process by UOP.

Highlights of TIP

The following are some of the highlights of the TIP process:

A. TIP has been in commercial operation since 1975

B. UOP manufacturers both the zeolite isomerisation catalyst and the IsoSiv Grade Molecular Sieve
adsorbent.
C. UOP's zeolite catalyst will tolerate sulfur and/or water upsets, the effects of which are usually reversible,
either with time or by in situ regeneration (which minimises any down time).
D. The expected life of the catalyst and adsorbent is 10 years or more.
F. The combination of zeolite isomerisation and IsoSiv molecular separation is possible because each
station has similar operating conditions of temperature, pressure and environment. This eliminates the
need for a second compressor, intermediate stabilisation and the costs associated with cooling, purifying
and reheating the recylce normal paraffins.
G. TIP and IsoSiv separation permits maximum flexibility in changing the C5/C6 ratio and iso/normal ratio
of the feed.

Conclusion:

Nowadays many refiners are looking into the isomerisation processes to add potential extra value and
complimentary to the platforming process. Directly both the platforming and isomerisation process work
hand in hand in several ways. C5 paraffins tend to crack away in the platformer, but give high upgrading in
the isomerisation unit. C6 components convert nicely to benzene in the platformer, but nowadays the
specs on aromatics and benzene are tightening, which makes conversion of these components to C6
isomers preferred. Furthermore, benzene is hydrogenated in the isomerisation unit. By adjusting the
cutpoint between the light and heavy naphtha, i.e. the cutpoint between the feed to the isomerisation feed
and the platformer feed, the refiner has the flexibility to control the benzene content of its gasoline pool.
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Asphaltic bitumen, normally called " bitumen" is obtained by vacuum distillation or vacuum flashing
of an atmospheric residue. This is " straight run" bitumen. An alternative method of bitumen production is
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Search withdrawn continuously from the surge vessel under level control and pumped to storage through feed/
product heat exchangers.
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Source: Petroleum Refinery Processes - OSHA Technical Manual
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INTRODUCTION
Send to a Friend
Already in the 30's it was found that when heavy oil fractions are heated over clay type materials, cracking
Search reactions occur, which lead to significant yields of lighter hydrocarbons. While the search was going on for
suitable cracking catalysts based on natural clays, some companies concentrated their efforts on the
development of synthetic catalyst. This resulted in the synthetic amorphous silica-alumina catalyst, which
was commonly used until 1960, when it was slightly modified by incorporation of some crystalline material
(zeolite catalyst). When the success of the Houdry fixed bed process was announced in the late 1930s, the
Search companies that had developed the synthetic catalyst decided to try to develop a process using finely
powdered catalyst. Subsequent work finally led to the development of the fluidised bed catalytic cracking
(FCC) process, which has become the most important catalytic cracking process.

Originally, the finely powdered catalyst was obtained by grinding the catalyst material, but nowadays, it is
Extra Resources produced by spray-drying a slurry of silica gel and aluminium hydroxide in a stream of hot flue gases.
Under the right conditions, the catalyst is obtained in the form of small spheres with particles in the range
of 1-50 microns.
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When heavy oil fractions are passed in gas phase through a bed of powdered catalyst at a suitable velocity
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(0.1-0.7m/s), the catalyst and the gas form a system that behaves like liquid, i.e. it can flow from one
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vessel to another under the influence of a hydrostatic pressure. If the gas velocity is too low, the powder
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does not fluidise and it behaves like a solid. If velocity is too high, the powder will just be carried away with
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the gas. When the catalyst is properly fluidised, it can be continously transported from a reactor vessel,
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where the carcking reactions take place and where it is fluidised by the hydrocarbon vapour, to a
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regenerator vessel, where it is fluidised by the air and the products of combustion, and then back to the
reactor. In this way the proces is truly continous.
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The first FCC unit went on stream in Standard Oil of New Jersey's refinery in Baton Rounge, Louisiana in
May 1942. Since that time, many companies have developed their own FCC process and there are
All Resources numerous varieties in unit configuration.
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Catalyst Catalogs FCC Process Configuration
Discussion Groups
Events : Hot feed, together with some steam, is introduced at the bottom of the riser via special distribution
Glossaries nozzles. Here it meets a stream of hot regenerated catalyst from the regenerator flowing down the
Jobs inclined regenerator standpipe. The oil is heated and vaporised by the hot catalyst and the cracking
Learning reactions commence. The vapour, initially formed by vaporisation and successively by cracking, carries the
Licensed Processes catalyst up the riser at 10-20 m/s in a dilute phase. At the outlet of the riser the catalyst and hydrocarbons
Magazines are quickly separated in a special device. The catalyst (now partly deactivated by deposited coke) and the
News Sources vapour then enter the reactor. The vapour passes overhead via cyclone separator for removal of entrained
Software catalyst before it enters the fractionator and further downstream equipment for product separation. The
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Training catalyst then descends into the stripper where entrained hydrocarbons are removed by injection of steam,
Websites before it flows via the inclined stripper standpipe into the fluidised catalyst bed in the regenerator.

Air is supplied to the regenerator by an air blower and distributed throughout the catalyst bed. The coke
deposited is burnt off and the regenerated catalyst passes down the regenerator standpipe to the bottom
of the riser, where it joins the fresh feed and the cycle recommences.

The flue gas (the combustion products) leaving the regenerator catalyst bed entrains catalyst particles. In
particular, it entrains "fines", a fine dust formed by mechanical rubbing of catalyst particles taking place in
the catalyst bed. Before leaving the regenerator, the flue gas therefore passes through cyclone separators
where the bulk of this entrained catalyst is collected and returned to the catalyst bed.

Normally modern FCC is driven by an expansion turbine to mimimise energy consumption. In this
expansion turbine, the current of flue gas at a pressure of about 2 barg drives a wheel by striking impellers
fitted on this wheel. The power is then transferred to the air blower via a common shaft. This system is
usually referred to as a "power recovery system". To reduce the wear caused by the impact of catalyst
particles on the impellers (erosion), the flue gas must be virtually free of catalyst particles. The flue gas is
therefore passed through a vessel containing a whole battery of small, highly efficient cyclone separators,
where the remaining catalyst fines are collected for disposal.

Before being disposed of via a stack, the flue gas is passed through a waste heat boiler, where its
remaining heat is recovered by steam generation.

In the version of the FCC process described here, the heat released by burning the coke in the regenerator
is just sufficient to supply the heat required for the riser to heat up, vaporise and crack the hydrocarbon
feed. The units where this balance occurs are called " heat balanced" units. Some feeds caused excessive
amounts of coke to be deposited on the catalyst, i.e. much more than is required for burning in the
regenerator and to have a "heat balanced" unit. In such cases, heat must be removed from the
regenerator, e.g. by passing water through coils in the regenerator bed to generate steam. Some feeds
cause so little coke to be deposited on the catalyst that heat has to be supplied to the system. This is done
by preheating the hydrocarbon feed in a furnace before contacting it with the catalyst.

Main Characteristics

A special device in the bottom of the riser to enhance contacting of catalyst and hydrocarbon feed.
The cracking takes place during a short time (2-4 seconds) in a riser ("short-contact time riser") at
high temperatures ( 500-540 C at riser outlet).
The catalyst used is so active that a special device for quick separation of catalyst and
hydrocarbons at the outlet of the riser is required to avoid undesirable cracking after the mixture
has left the riser. Since, no cracking in thereactor is required, the reactor no longer functions as a
reactor; it merely serves as a holding vessel for cyclones.
The regenerator takes place at 680-720 C. With the use of special catalysts, all the carbon
monoxide (CO) in the flue gas is combusted to carbon dioxide (CO2) in the regenerator.
Modern FCC includes a power recovery system for driving the air blower.

Equipment in FCC

Large storage vessels for catalyst (fresh and equilibrium)

Regenerator
Reactor
Main Fractionator
Product Work Up section (several distillation columns in series
Product treating facilities

Feedstock & Yield

Before the introduction of residues, vacumn distillates were used as feedstock to load the Catalytic Cracker
fully. These days, even residues are used to load the cracker. The term used for this type of configuation is
Long Residue Catalytic Cracking Complex. The only modification or addition needed are a residue desalter
and a bigger and more heat resistent reactor.

The yield pattern of an FCC unit is typically as follows:

Product % wt on fresh feed

C3 & C4 15
Gasoline 40-50
Heavy Gas Oil 10
Coke 5

Conclusion
The FCC Unit can a real margin improver for many refineries. It is able to convert the residues into high
value products like LPG , Butylene, Propylene and Mogas together with Gasoil. The FCC is also a start for
chemical production (poly propylene). Many FCC's have 2 modes: a Mogas mode and a Gasoil mode and
FCC's can be adapted to cater for the 2 modes depending on favourabale economic conditions. The only
disavantage of an FCC is that the products produced need to be treated (sulfur removal) to be on
specification. Normally Residue FCCs act together with Residue Hydroconversion Processes and
Hydrocrackers in order to minimise the product quality give away and get a yield pattern that better
matches the market specifications. Via product blending, expensive treating steps can be avoided and the
units prepare excellent feedstock for eachother: desulfurised residue or hydrowax is excellent FCC feed,
while the FCC cycle oils are excellent Hydrocracker feed.

In the near future, many refiners will phase the challenge how to desulfurise cat cracked gasoline without
destroying its octane value. Catalytic destillation appears to be one of the most promising candidate
processes for that purpose.
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Source: Petroleum Refinery Processes - OSHA Technical Manual

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Distillation is the first step in the processing of crude oil and it takes place in a tall steel tower called a Send to a Friend
Search fractionation column. The inside of the column is divided at intervals by horizontal trays. The column is
kept very hot at the bottom (the column is insulated) but as different hydrocarbons boil at different
temperatures, the temperature gradually reduces towards the top, so that each tray is a little cooler than
the one below.

Search The crude needs to be heated up before entering the fractionation column and this is done at first in a
series of heat exchangers where heat is taken from other process streams which require cooling before
being sent to rundown. Heat is also exchanged against condensing streams from the main column.
Typically, the crude will be heated up in this way upto a temperature of 200 - 280 deg. C, before entering a
furnace.
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As the raw crude oil arriving contains quite a bit of water and salt, it is normally sent for salt removing first,
RTOL?s Useful Stuff in a piece of equipment called a desalter. Upstream the desalter, the crude is mixed with a water stream,
typically about 4 - 6% on feed. Intense mixing takes place over a mixing valve and (optionally) as static
mixer. The desalter, a large liquid full vessel, uses an electric field to separate the crude from the water
RTOL Catalyst Vendor droplets. It operates best at 120 - 150 deg C, hence it is conveniently placed somewhere in the middle of
Database the preheat train.
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Oil Refineries in the World Part of the salts contained in the crude oil, particularly magnesium chloride, are hydrolysable at
Asia Pacific Refinery Index temperatures above 120 C. Upon hydrolysis, the chlorides get converted into hydrochloric acid, which will
find its way to the distillation column's overhead where it will corrode the overhead condensers. A good
performing desalter can remove about 90% of the salt in raw crude.
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Downstream the desalter, crude is further heated up with heat exchangers, and starts vaporising, which
will increase the system pressure drop. At about 170 -200 deg. C, the crude will enter a 'pre-flashvessel',
All Resources operating at about 2 - 5 barg, where the vapours are separated from the remaining liquid. Vapours are
Articles directly sent to the fractionation column, and by doing so, the hydraulic load on the remainder of the crude
Best Practises preheat train and furnace is reduced (smaller piping and pumps).
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Events Just upstream the preflash vessel, a small caustic stream is mixed with the crude, in order to neutralise
Glossaries any hydrochloric acid formed by hydrolysis. The sodium chloride formed will leave the fractionation column
Jobs via the bottom residue stream. The dosing rate of caustic is adjusted based on chloride measurements in
Learning the overhead vessel (typically 10 - 20 ppm).
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Magazines At about 200 - 280 C the crude enters the furnace where it is heated up further to about 330 -370 C. The
News Sources furnace outlet stream is sent directly to the fractionation column. Here, it is separated into a number of
Software fractions, each having a particular boiling range.
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Training At 350 degrees Celsius, and about 1 barg, most of the fractions in the crude oil vapourise and rise up the
Websites column through perforations in the trays, losing heat as they rise. When each fraction reaches the tray
where the temperature is just below its own boiling point, it condenses and changes back into liquid phase.
A continuous liquid phase is flowing by gravity through 'downcomers' from tray to tray downwards. In this
way, the different fractions are gradually separated from each other on the trays of the fractionation
column. The heaviest fractions condense on the lower trays and the lighter fractions condense on the trays
higher up in the column. At different elevations in the column, with special trays called draw-off trays,
fractions can be drawn out on gravity through pipes, for further processing in the refinery.

At top of the column, vapours leave through a pipe and are routed to an overhead condenser, typically
cooled by air fin-fans. At the outlet of the overhead condensers, at temperature about 40 C, a mixture of
gas, and liquid naphtha exists, which is falling into an overhead accumulator. Gases are routed to a
compressor for further recovery of LPG (C3/C4), while the liquids (gasoline) are pumped to a hydrotreater
unit for sulfur removal.

A fractionation column needs a flow of condensing liquid downwards in order to provide a driving force for
separation between light and heavy fractions. At the top of the column this liquid flow is provided by
pumping a stream back from the overhead accumulator into the column. Unfortunately, a lot of the heat
provided by the furnace to vaporise hydrocarbons is lost against ambient air in the overhead fin-fan
coolers. A clever way of preventing this heat lost of condensing hydrocarbons is done via the circulating
refluxes of the column. In a circulating reflux, a hot side draw-off from the column is pumped through a
series of heat exchangers (against crude for instance), where the stream is cooled down. The cool stream
is sent back into the column at a higher elevation, where it is been brought in contact with hotter rising
vapours. This provides an internal condensing mechanism inside the column, in a similar way as the top
reflux does which is sent back from the overhead accumulator. The main objective of a circulating reflux
therefore is to recover heat from condensing vapours. A fractionating column will have several (typically
three) of such refluxes, each providing sufficient liquid flow down the corresponding section of the column.
An additional advantage of having circulating refluxes is that it will reduce the vapour load when going
upwards in the column. This provided the opportunity to have a smaller column diameter for top sections
of the tower. Such a reduction in diameter is called a 'swage'.

The lightest side draw-off from the fractionating column is a fraction called kerosene, boiling in the range
160 - 280 C, which falls down through a pipe into a smaller column called 'side-stripper'. The purpose of
the side stripper is to remove very light hydrocarbons by using steam injection or an external heater called
'reboiler'. The stripping steam rate, or reboiled duty is controlled such as to meet the flashpoint
specification of the product. Similarly to the atmospheric column, the side stripper has fractionating trays
for providing contact between vapour and liquid. The vapours produced from the top of the side stripper
are routed back via pipe into the fractionating column.

The second and third (optional) side draw-offs from the main fractionating column are gasoil fractions,
boiling in the range 200 - 400 C, which are ultimately used for blending the final diesel product. Similar as
with the kerosene product, the gasoil fractions (light and heavy gasoil) are first sent to a side stripper
before being routed to further treating units.

At the bottom of the fractionation column a heavy, brown/black coloured fraction called residue is drawn
off. In order to strip all light hydrocarbons from this fraction properly, the bottom section of the column is
equipped with a set of stripping trays, which are operated by injecting some stripping steam (1 - 3% on
bottom product) into the bottom of the column.
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RTOL Navigation Process Practice: RTOL Students' Guide: Gas Treating and Sulfur Recovery Related Links

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Removal of CO2 and H2S from gas streams is normally done via absorption in regenerable solvents.
Sulfur is recovered from the concentrated off-gas via the so called 'Claus' reaction, which involves partial
Community oxydation over catalyst at very high temperatures. Most read story about
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Calendar Crude oils contain up to 4% wt. Sulphur, and the gas streams produced by distillation or conversion
Top10 (cracking) contain significant quantities of hydrogen sulphide (H2S). This highly poisonous and corrosive
and odourous compound must normally be removed from the gas before it can be burnt in the furnaces.
Downloads & Links Apart from H2S, LPG (propane and butane) also contain carbonyl sulphide and mercaptans, and these too
may have to be removed, because upon combustion they get converted in SO2/SO3, which has a major Please take a second and vote
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Gasification of residual fuel oils (or, in the future, of coal) a raw gas consisting of is used to make carbon
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monoxide and hydrogen. This raw gas an be used to produce pure hydrogen, methanol or ammonia. The
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gas from the gasifier has to be treated to remove H2S and carbonyl sulphide (sometimes to very low levels)
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and in many applications large quantities of CO2 have to be removed.
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Quizz Natural gas streams may contain sulphur in the form of H2S, carbonyl sulphides or mercaptans, and
significant amounts of CO2 may be also be present. Because of the great variety of combination of
contaminants in naturally occurring gases, these projects usually pose the most problems to the process
Services designer. Typical specifications for natural gas for domestic use are 4ppm (vol.) H2S and about 100ppm Cast my Vote!
Store Front (vol.) for other sulphur components. The CO2 specification may be set by calorific value of product, or by
Advertising secondary processing, such as cryogenic plants. Complete liquefaction, for example, requires a CO2
Products and Services specification of about 100 ppm vol.

The water content often needs to be reduced to avoid formation of hydrates, which will plug the pipeline
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The removal of the contaminants discussed above is almost always carried out by absorption in
regenerable solvents. Occasional exceptions, for small amounts of contaminants, are adsorbents or caustic Send to a Friend
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soda on a disposable basis.

The regenerable solvents can be classified as chemical, physical and mixtures of physical and chemical. The
choice of solvent depends on pressure and type of feed gas, amount and combination of contaminants in
Search the feed, and treated gas specification.

Another relevant factor in choice of solvent is the composition of the sulphur-rich stream removed when
the solvent is regenerated, in cases in which a process for recovering sulphur from this stream is required.
The " CLAUS" process is by far the most common sulphur recovery process, with limitations on the amount
Extra Resources of CO2 and hydrocarbons in its feed.

RTOL?s Useful Stuff There may be therefore be specifications on selectivity of the treating process for absorption of H2S as
opposed to absorption of CO2 or hydrocarbons.
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The two most important contaminants we are concerned with, H2S and CO2, are both acidic in aqueous
Oil Refineries in the World
solution, and early in the development of gas-treating processes a weakly basic water-soluble solvent was
Asia Pacific Refinery Index
sought which would react reversibly with these components. Many alkanolamines have the correct
combination of properties, and already in the early thirties di-ethanolamine (DEA) gained popularity, under
RTOL?s Internet the name "GIRBOTOL PROCESS". Other alkanolamines such as mono-ethanolamine (MEA) and di-
Resources isopropanolamine (DIPA) have also found wide application.

All Resources A typical high pressure gas treating process is always accompanied with a sulphur recovery unit. In the gas
Articles treating unit, the H2S and all or part of the CO2 are removed by countercurrently contacting the gas with
Best Practises DEA solution in a column with trays or random packing. The amount of solution and number of trays (or
Catalyst Catalogs packing height) is chosen to meet specification on H2S and if relevant, CO2. The DEA solution leaving the
Discussion Groups absorber is let down in pressure to allow dissolved entrained hydrocarbons to escape. These gases are
Events usually sent to the fuel gas system. After picking up heat from the hot regenerated solvent, the DEA
Glossaries solution enters the regenerator, where it is contacted countercurrently with steam. The solvent is raised to
Jobs its boiling point of about 110 deg. C and stripped by the steam. The regenerated solvent, after giving up
Learning heat to the loaded solvent, is cooled to about 40 deg. C before re-entering the absorber. Typically, the
Licensed Processes solvent inventory is circulated 50 times per hour between regenerator and absorber. A single regenerator
Magazines may serve several gas absorbers and also LPG extractors (see below).
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Software
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Training The steam in the H2S/CO2 stream from the regenerator is largely condensed by cooling to 40 deg. C and
Websites the water is returned to the top of the regenerator or sent to the sour water stripper.

The H2S/CO2 gases are then fed to the sulphur recovery unit to remove 99.9% of the sulphur.

Chemical absorption processes consume appreciable amounts of steam, and in an effort to reduce costs,
there is a move towards higher concentrations of amine and higher H2S/CO2 loadings. This tends to go
with higher degradation and corrosion rates, so a compromise must be found. Among the various
alkanolamines, those which are most selective towards H2S in the presence of CO2 are gaining ground,
since the sulphur recovery unit becomes considerably smaller if the CO2 content in the feed to this unit is
low, and also the overall treating costs are much lower if the CO2 content in the feed to this unit is low, and
also the overall treating costs are much reduced. DEA has a poor selectivity and is losing some ground for
this reason.

A key task for the process engineer looking after the gas treating units is to optimise the solvent circulation
rate. A lower circulation rate can give significant steam savings in the regenerator, but may also lead to the
risk of H2S breakthrough when the H2S quantity in the feedgas goes up. With regular testruns, the process
engineer can determine the "safe" minimum circulation rate.

It should be noted the deep removal of carbonyl sulphide from gases is not possible with chemical
solvents, owing to slow reaction rate, and the mercaptans are hardly removed at all.

Physical Solvents:

The solubility of H2S, CO2 and Carbonyl sulphide is much higher than that of methane, carbon monoxide
and hydrogen in many liquids, methanol for example. Since the heat of solution is much lower than the
heat of reaction in chemical solvents, desorption can be achieved by pressure reduction, possibly combined
with moderate heating or inert gas stripping.

Such solvents are only attractive, however, when sufficiently high loadings of contaminants in the solvent
can be achieved, avoiding excess solvent requirements. Such high loadings are possible when the partial
pressure of the contaminant is very high (typically above 10 bar) or when refrigeration is used to cool
solvent to increase solubility.

Methanol, under the trade name " Rectisol", is widely applied for removal H2S, CO2 and COS from
synthesis gas. The lean Rectisol is cooled to -40 dC. Very deep sulphur removal is achieved and the solvent
is selective for H2S. Such units are complex, with high capital cost.

The application of other physical solvents for natural gas treating is limited to feeds with low concentration
of propane and heavier hydrocarbons, since the solubility of these components is too high.

Mixed Physical/Chemical Solvents:

The Sulfinol solvent, developed in the early sixties, combines many of the attractive properties of physical
and chemical solvents. It is a mixture of Sulfolane, an alkanolamine and water. It has found wide
application in treating natural and syngases. Its main features are deep removal of H2S, carbonyl sulphide,
mercaptans and when required CO2. Operating cost are in general significantly lower than for purely
chemical processes.

LPG Treating:

LPG may contain H2S, COS and mercaptans, the quantity depending on origin. The H2S is usually removed
by applying a chemical absorption solvent in a packed column.

COS may be removed if sufficient residence time is created in a mixer-settler also using a chemical
solvent, with a propeller mixer as contacting device. A solvent regeneration system combined with gas
absorbers can be used.

For mercaptans removal, alkali hydroxide solution is generally used, also in a packed column. The caustic
soda may be regenerable by oxidising the mercaptides to disulphides, which are then separated from the
solution by contacting with gasoline.

Both gas treating and liquid treating columns are prone to fouling, since the solvent picks up contaminants
that are not removed in the regenerator column. The contaminants deposit themselves on the packing or
trays of the treating column, which is normally detected by an increase in pressure drop over the column.
When the pressure drop becomes excessive, the solvent gets carried over with the gas or LPG stream and
H2S can breakthrough. Early detection of this fouling is therefore essential, after which the column can be
taken off-line for washing with soda ash and hot condensate.

Because LPG is commonly used in the household as cooking gas, the contaminant specifications are very
stringent and it has to be free of H2S. Running down off-spec LPG to storage can be a very costly affair for
a refinery, because the off-spec LPG will have to be burnt in its own furnaces. To prevent this from
happening, normally a vessel with caustic is installed downstream the LPG extractor. The caustic will
remove residual H2S and act as a buffer for LPG extractor operating problems.

Sulphur Recovery & Tail Gas Treating:

The " CLAUS" process is by far the most widely applied means of sulphur recovery, and is based on partial
combustion of H2S to SO2 (at 1200-1400 deg. C) and the second step "Claus" reaction of H2S and SO2 to
form elemental sulphur in accordance with

2H2S + SO2 <-------> 3S + 2H2O

Thermodynamic equilibrium limits sulphur recovery to about 95%, favoured by lower operating
temperatures. In order to achieve a reasonable speed of reaction, the Claus reaction takes place in a
reactor over a Sodium Oxyde/Alumina catalyst.

The main parts of the Claus unit consist of:

Furnace (partial combustion of H2S)

Reactor (Claus reaction to produce sulfur)
Heat recovery section (steam production)
 Sulfure condensation and pump-out

The cost of the process and sulphur emission increases with CO2 concentrations in feed, and excessive
hydrocarbons in the feed can result in black and unsaleable sulphur. These set the two main requirements
Languages of the upstream treating process.

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The process is exothermic and if little CO2 is present, the combined gas treating and sulphur recovery units
are at least self- sufficient in terms of energy. The heat generated in the Claus process is normally used to
produce steam, which is sent to the refinery's steam network.
English

The so-called "tail gas" from a sulphur recovery unit is a mixture of sulphur gases (SO2, H2S etc)
remaining after condensing the sulphur. It is usually incinerated to SO2 and vented. More stringent
environmental requirements have led to the development of several tailgas treating processes, which
remove the remaining sulfur compounds. The best known tailgas treating processes are Superclaus
(Comprimo) and Scot (Shell). The combined claus and tailgas treating process normally has a sulfur
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recovery of about 99 %.
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chapter deals with another workhorse of the refinery processes: the hydrotreating process, a sulphur
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The objective of the Hydrotreating prococess is to remove suplur as well as other unwanted compunds,
e.g. unsaturated hydrocarbons, nitrogen from refinery process streams.
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Until the end of World War 2, there was little incentive for the oil industry to pay significant attention to Options
improving product quality by hydrogen treatment. However, soon after the war the production of high
sulphur crudes increased significantly, which gave a more stringent demand on the product blending
flexibility of refineries, and the marketing specifications for the products became tighter, largely due to Printer Friendly Page
environmental considerations. Furthermore, the catalyst used in the Platforming process can only handle
Search sulfur in the very low ppm level, so hydrotreating of naphtha became a must. The necessity for Send to a Friend
hydrotreating of middle distillates (kerosene/gasoil) originates from pressure to reduce sulfur emissions
into the environment. Overall, this situation resulted in an increased necessity for high sulphur removal
capability in many refineries.

Search As catalytic reforming gives hydrogen as a byproduct, it gave additional momentum to the development of
sulphur removal process by hydrogen treatment. In this treatment, the sulphur compounds are removed
by converting them into hydrogen sulphide by reaction with hydrogen in the presence of a catalyst. This
results in high liquid product yields, since only sulphur is removed. Furthermore, the hydrogen sulphide
produced can be easily removed from the product gas stream, for example by an amine wash. In this way,
Extra Resources hydrogen sulphide is recovered as a higly concentrated stream and can be further converted into elemental
sulphur via the "Claus" process.
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Hydrodesulphursiation has been extensively used commercially for treating naphtha as feedstock for
catalytic reformers to meet the very stringent sulphuir specification of less than 1 ppm wt to protect the
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platinum catalyst. It has also been widely used for removal of sulphur compounds from kerosine and
Database
gasoils to make them suitable as blending components. In cases where products are from catalytic or
Water Dewpoint
thermal crackers, hydrogen treatment is used to improve product quality specifications like colour, smoke
Calculation
point, cetane index, etc.
Oil Refineries in the World
Asia Pacific Refinery Index
For Hydrotreating, two basic processes are applied, the liquid phase (or trickle flow) process for kerosine
and heavier straight-run and cracked distillates up to vacumn gas oil and the vapour phase process for light
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straight-run and cracked fractions.
Resources

Both processes use the same basic configuration: the feedstock is mixed with hydrogen-rich make up gas
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and recycle gas. The mixture is heated by heat exchange with reactor effluent and by a furnace and enters
Articles
a reactor loaded with catalyst. In the reactor, the sulphur amd nitrogen compounds present in the
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feedstock are converted into hydrogen sulphide and ammonia respectively. The olefins present are
Catalyst Catalogs
saturated with hydrogen to become di-olefins and part of the aromatics will be hydrogenated. If all
Discussion Groups
aromatics needs to be hydrogenated, a higher pressure is needed in the reactor compared to the
Events
conventional operating mode.
Glossaries
Jobs
Learning The reactor operates at temperatures in the range of 300-380 degrees C and at a pressure of 10-20 bar for
Licensed Processes naphta and kero, as compared with 30-50 bar for gasoil, with excess hydrogen supplied. The temperature
Magazines should not exceed 380 degrees C, as above this temperature cracking reactions can occur, which
News Sources deteriorates the colour of the final product. The reaction products leave the reactor and, after having been
Software cooled to a low temperature, typically 40-50 degrees C, enter a liquid/gas separation stage. The hydrogen-
Success Stories rich gas from the high pressure separation is recycled to combine with the feedstock, and the low pressure
Training off-gas stream rich in hydrogen sulphide is sent to a gas-treating unit, where hydrogen sulphide is
Websites removed. The clean gas is then suitable as fuel for the refinery furnaces. The liquid stream is the product
from hydotreating. It is normally sent to a stripping column where H2S and other undesirable components
are removed, and finally, in cases where steam is used for stripping, the product is sent to a vacumn drier
for removal of water. Some refiners use a salt dryer in stead of a vacuum drier to remove the water.

The catalyst used is normally cobalt, molybdenum and nickel finely distributed on alumina extrudates. It
slowly becomes choked by coke and must be renewed at regular intervals (typically 2-3 years). It can be
regenerated (by burning off the coke) and reused typically once or twice before the breakdown of the
support's porous structure unacceptably reduces its activity. Catayst regeneration is, nowadays, mainly
carried out ex- situ by specialised firms. Other catalysts have also been developed for applications where
denitrification is the predominant reaction required or where high stauration of olefins is necessary.

A more recent development is the application of Hydrotreating for pretreatment of feedstcok for the
catalytic cracking process. By utilisation of a suitable hydrogenation-promoting catalyst for conversion of
aromatics and nitrogen in potential feedstocks, and selection of severe operating conditions, hydrogen is
taken up by the aromatic molecules. The increased hydrogen content of the feedstock obtained by this
treatment leads to significant conversion advantages in subsequent catalytic cracking, and higher yield of
light products can be achieved.

Hydrotreatment can also be used for kerosine smoke point improvement (SPI). It closely resembles the
conventional Hydrotreating Process however an aromatic hydrogenation catalyst consisting of noble metals
on a special carrier is used. The reactor operates at pressure range of 50-70 bar and temperatures of 260-
320 degrees C. To restrict temperature rise due to the highly exothermic aromatics conversion reactions,
quench oil is applied between the catalysts beds. The catalyst used is very sensitive to traces of sulphur
and nitrogen in the feedstock and therefore pretreatment is normally applied in a conventional
hydrotreater before kerosine is introduced into the SPI unit. The main objective of Smoke Point
Improvement is improvement in burning characteristics as the kerosine aromatics are converted to
naphthenes.

Hydrotreatment is also used for production of feedstocks for isomersiation unit from pyrolysis gasoline
(pygas) which is one of the byproducts of steam cracking of hydrocarbon fractions such as naphtha and
gasoil.

A hyrotreater and a hydrodesulphuriser are basically the same process but a hydotreater termed is used
for treating kerosene or lighter feedstock, while a hydodesulhuriser mainly refers to gasoil treating. The
hydrotreatment process is used in every major refinery and is therefore also termed as the work horse of
the refinery as it is the hydrotreater unit that ensures several significant product quality specifications. In
most countries the Diesel produced is hydrodesulhurised before its sold. Sulphur specifications are getting
more and more stringent. In Asia, countries such as Thailand, Singapore and Hong Kong already have a
0.05%S specification and large hydrodesulphurisation units are required to meet such specs.

The by-products obtained from HDT/HDS are light ends formed from a small amounts of cracking and these
products are used in the refinery fuelgas pool. The other main by-product is Hydrogen Sulphide which is
oxidized to sulphur and sold to the chemical industry for further processing

In combination with temperature, the pressure level (or rather the partial pressure of hydrogen) generally
determines the types of components that can be removed and also determines the working life of the
catalyst. At higher (partial) pressures, the desulphurisation process is 'easier', however, the unit becomes
more expensive for instance due to larger compressors and heavier reactors. Also, at higher pressure, the
hydrogen consumption of the unit increases, which can be a signficant cost factor for the refinery. The
minimum pressure required typically goes up with the required severity of the unit, i.e. the heavier the
feedstock, or the lower levels of sulphur in product required.
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Source: Petroleum Refinery Processes - OSHA Technical Manual

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INTRODUCTION
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Thermal cracking is the oldest and, in a way, the simplest cracking process. It basically aims at the
reduction of molecular size by application of heat without any additional sophistication such as catalyst or Send to a Friend
Search hydrogen. At a temperature level of 450-500 C, the larger hydrocarbon molecules become unstable and
tend to break spontaneously into smaller molecules of all possible sizes and types. By varying the time,
temperature and pressure under which a particular feedstock remains under cracking conditions, the
desired degree of cracking (conversion) can be controlled. Temperature and time (residence time) are
important process variables pressure plays a secondary role.
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Obviously, the cracking conditions to be applied and the amount and type of cracked products will depend
largely on the type of feedstock. In practice, the feedstock for thermal cracking is a mixture of complex
heavy hydrocarbon molecules left over from atmospheric and/or vacuum distillation of crude. The nature
of these heavy, high molecular weight fractions is extremely complex and much fundamental research has
been carried out on their behaviour under thermal cracking conditions. However, a complete and
Extra Resources
satisfactory explanation of these reactions that take place cannot be given, except for relatively simple and
well-defined types of products. For instance, long chain paraffinic hydrocarbon molecules break down into
RTOL?s Useful Stuff a number of smaller ones by rupture of a carbon-to-carbon bond (the smaller molecules so formed may
break down further). When this occurs, the number of hydrogen atoms present in the parent molecule is
RTOL Catalyst Vendor insufficient to provide the full complement for each carbon atom, so that olefins or "unsaturated"
Database compounds are formed. The rupturing can take place in many ways, usually a free radical mechanism for
Water Dewpoint the bond rupture is assumed.
Calculation
Oil Refineries in the World However, paraffinic hydrocarbons are usually only a small part of the heavy petroleum residues, the rest
Asia Pacific Refinery Index being cyclic hydrocarbons, either aromatic or naphthenic in character. In these, the rupture takes place in
the paraffinic side-chain and not in the ring. Other side reactions also take place. In particular, the
RTOL?s Internet condensation and polymerisation reactions of olefins and of the aromatics are of considerable practical
Resources importance, since they can lead to undesirable product properties, such as an increase in the sludge or tar
content. Hence, in practice, it is very difficult to assess the crackability of various feedstocks without plant
trials. The final products consist of gas, light hydrocarbons in the gasoline and gasoil range and heavier
All Resources products. By selection of the type of unit, feedstock and operating conditions, the yields and quality of the
Articles various products can, within limits be controlled to meet market requirements.
Best Practises
Catalyst Catalogs
Discussion Groups The maximum conversion that can be obtained will be determined by the quality of the bottom product of
Events the thermal cracker, thermally cracked residue. This stream is normally routed to the fuel oil blending
Glossaries pool. When the cracking has taken place at a too high severity, the fuel can become 'unstable' upon
Jobs blending with diluent streams (see below). Normally, the refinery scheduler will assess what the maximum
Learning severity is that the thermal cracking unit can operate on, without impacting on the stability of the refinery
Licensed Processes fuel blending pool.
Magazines
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Software production of gasoline. The units were relatively small (even applying batch processing), were inefficient
Success Stories and had a very high fuel consumption. However, in the twenties and thirties a tremendous increase in
Training thermal cracking capacity took place, largely in the version of the famous DUBBS process, invented by
Websites UOP. Nevertheless, thermal cracking lost ground quickly to catalytic cracking (which produces gasoline of
higher octane number) for processing heavy distillates with the onset of the latter process during World
War II. Since then and up to the present day, thermal cracking has mostly been applied for other
purposes : cracking long residue to middle distillates (gasoil), short residue for viscosity reduction
(visbreaking), short residue to produce bitumen, wax to olefins for the manufacture of chemicals, naphtha
to ethylene gas (also for the manufacturing of chemicals), selected feedstocks to coke for use as fuel or for
the manufacture of electrodes.

In modern oil refineries there are three major applications of the thermal cracking process:

VISBREAKING
THERMAL GASOIL PRODUCTION
COKING

VISBREAKING
Visbreaking (i.e. viscosity reduction or breaking) is an important application of thermal cracking because it
reduces the viscosity of residue substantially, thereby lessening the diluent requirements and the amount
of fuel oil produced in a refinery. The feed, after appropriate preheat, is sent to a furnace for heating to the
cracking temperature, at about 450-460 degrees C. The cracking takes place to a small extent in the
furnace and largely in a soaker (reaction chamber) just downstream of the furnace. At the soaker outlet,
the temperature is lower than at the furnace outlet (soaker inlet) because the cracking reactions are
endothermic. The products are quenched at the soaker outlet to stop the cracking reaction (to prevent
excessive coke formation). After that, the products enter the fractionator at a temperature level of 300-
400 degrees C and from here onward the processing is similar to any normal distillation process. The
products are separated into gas, gasoline, kero, gasoil and residue. The residue so obtained has a lower
viscosity that the feed (visbreaking), which leads to a lower diluent requirement to make the fuel on
specification for viscosity. The up-flow soaker provides for a prolonged residence time and therefore
permits a lower cracking temperature than if the soaker was not used. This is advantageous as regards
cost in furnace and fuel. Modern soakers are equipped with internals so as to reduce back mixing- effects ,
thus maximising the viscosity reduction. Since only one cracking stage is involved, this layout is also
named one-stage cracking. The cracking temperature applied is about 440-450 degree C at a pressure of
5-10 barg in the soaker. The fractionator can be operated at 2-5 barg, depending on furnace constraints,
condenser constraints and fuel cost.

THERMAL GASOIL PRODUCTION

This is a more elaborate and sophisticated application of thermal cracking as compared with visbreaking.
Its aim is not only to reduce viscosity of the feedstock but also to produce and recover a maximum amount
of gasoil. Altogether, it can mean that the viscosity of residue (excluding gasoil) run down from the unit is
higher than that of the feed.
In the typical lay out is the first part of the unit quite similar to a visbreaking unit. The visbroken residue is
vacuum-flashed to recover heavy distillates, which are then sent back to a thermal cracking stage,
together with heavy distillate recovered from the fractionator, in a second furnace under more severe
cracking conditions ( temperature 500 degrees C; pressure 20-25 barg) . More severe conditions are
necessary because the feedstock has a smaller molecular size and is therefore more difficult to crack than
the larger residue molecules in the first stage. This layout is referred to as tow-stage cracking.

DELAYED COKING
This is an even more severe thermal cracking application than the previous one. The goal is to make a
maximum of cracking products - distillates - whereby the heavy residue becomes so impoverished in
hydrogen that it forms coke. The term "delayed" is intended to indicate that the coke formation does not
take place in the furnace (which would lead to a plant shutdown) but in the large coke drums after the
furnace. These drums are filled/emptied batch-wise (once every 24 hours), though all the rest of the plant
operates continuously. A plant usually has two coke drums, which have adequate capacity for one day's
coke production (500-1500 m2). The process conditions in the coke drum are 450-500 degrees C and 20 -
30 bar. Only one coke drum is on-line; the other is off line, being emptied or standing by. Only the vapour
passes from the top of the coke drums to the fractionator, where the products are separated into the
desired fractions. The residue remains in the coke drum to crack further until only the coke is left. Often
the heaviest part of the fractionator products is recycled to feed.

PRODUCT QUALITY
Thermally cracked products - distillates - are not suitable for commercial use as produced in other units;
they require further refinement or treatment in order to improve their quality, particularly sulfur and
olefins content. Formerly, wet treating processes, for example treatment with caustic or an other
extraction medium, were applied to remove or "sweeten" the smelly sulfur products, but nowadays the
catalytic hydrotreating is employed almost without exception, both for gasoline and for gas oil range
products. Of course, the gases too have to be desulphurised before being used as fuel gas within the
refinery.
The residual products from thermal cracking are normally not treated any further, except for coke, which
may be calcined if the specifications require it to be treated. The cracked residue is normally disposed of as
refinery or commercial fuel. Here a very important aspect of the process is the stability of the cracked
residues or of the final fuels after blending with suitable diluents. Residue contains asphaltenes,which are
colloidally dispersed uniformly in the oil in a natural way. In the cracking process, the character of the
asphaltenes as well as of the oil changes, and if the cracking is too severe the natural balance of the
colloidal system can be affected to the extent that part of the asphaltenes precipitates in the equipment or
in the storage tanks, forming sludge. If the sludge formation is excessive, i.e. above a certain specified
limit, the product (fuel) is considered to be unstable.

PLANT OPERATIONS/DECOKING
A practical aspect of operation of thermal cracking units is that, in spite of good design and operating
practice, furnaces, and sometimes also other equipment, gradually coke up, so that the unit has to be shut
down and decoked. Furnaces can be decoked by " turbining" (using special rotary tools to remove coke
from inside furnace pipes) or by steam-air decoking process. In the latter case, the coke is burnt off in a
carefully controlled decoking process in which air and steam are passed through the tubes at elevated
temperatures. Air serves to burn coke, where as the steam serves to keep the burning temperatures low
so that they do not exceed the maximum tolerable temperatures.
More recently, a new decoking method using studded 'pigs' propelled with water, is getting more popular.
The plastic pigs have a size slightly smaller than the tube inside diameter and are equipped with metal
studs. When the pigs are pumped through the furnace pipes, they move around in a rotating fashion, thus
scraping the cokes from the inside of the furnace tubes.
Other coked equipment is usually cleaned by hydrojetting techniques. Owing to these unavoidable stops for
decoking, the on-stream time i.e. on stream days per annum, for thermal cracking units is slightly shorter
than for most other oil processes.
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recover additional distillates from long residue, distillation at reduced pressure and high temperature
has to be applied. This vacuum distillation process has become an important chain in maximising the
Community upgrading of crude oil. As distillates, vacuum gas oil, lubricating oils and/or conversion feedstocks are Most read story about
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Search
INTRODUCTION
To recover additional distillates from long residue, distillation at reduced pressure and high temperature
has to be applied. This vacuum distillation process has become an important chain in maximising the
Search
upgrading of crude oil. As distillates, vacuum gas oil, lubricating oils and/or conversion feedstocks are
generally produced. The residue from vacuum distillation - short residue - can be used as feedstock for
further upgrading, as bitumen feedstock or as fuel component. The technology of vacuum distillation has
developed considerably in recent decades. The main objectives have been to maximise the recovery of
valuable distillates and to reduce the energy consumption of the units.
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At the place where the heated feed is introduced in the vacuum column - called the flash zone - the
RTOL?s Useful Stuff temperature should be high and the pressure as low as possible to obtain maximum distillate yield. The
flash temperature is restricted to about 420 C, however, in view of the cracking tendency of high-
RTOL Catalyst Vendor molecular-weight hydrocarbons. Vacuum is maintained with vacuum ejectors and lately also with liquid ring
Database pumps. Lowest achievable vacuum in the flash zone is in the order of 10 mbar.
Water Dewpoint
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Oil Refineries in the World not be achieved without the use of "lifting" steam. The steam acts in a similar manner as the stripping
Asia Pacific Refinery Index steam of crude distillation units. This type of units is called "wet" units. One of the latest developments in
vacuum distillation has been the deep vacuum flashers, in which no steam is required. These "dry" units
RTOL?s Internet operate at very low flash zone pressures and low pressure drops over the column internals. For that reason
Resources the conventional reflux sections with fractionation trays have been replaced by low pressure- drop spray
sections. Cooled reflux is sprayed via a number of specially designed spray nozzles in the column
countercurrent to the up-flowing vapour. This spray of small droplets comes into close contact with the hot
All Resources vapour, resulting in good heat and mass transfer between the liquid and vapour phase.
Articles
Best Practises
Catalyst Catalogs To achieve low energy consumption, heat from the circulating refluxes and rundown streams is used to
Discussion Groups heat up the long residue feed. Surplus heat is used to produce medium and/or low-pressure steam or is
Events exported to another process unit (via heat integration). The direct fuel consumption of a modern high-
Glossaries vacuum unit is approximately 1% on intake, depending on the quality of the feed. The steam consumption
Jobs of the dry high-vacuum units is significantly lower than that of the "wet" units. They have become net
Learning producers of steam instead of steam consumers.
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News Sources application, viz.:
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Training FEED PREPARATION UNITS
Websites LUBOIL HIGH- VACUUM UNITS
HIGH - VACUUM UNITS FOR BITUMEN PRODUCTION

FEED PREPARATION UNITS

These units make a major contribution to deep conversion upgrading ("cutting deep in the barrel"). They
produce distillate feedstocks for further upgrading in catalytic crackers, hydrocrackers and thermal
crackers. To obtain an optimum waxy distillate quality a wash oil section is installed between feed flash
zone and waxy distillate draw-off. The wash oil produced is used as fuel component or recycled to feed. The
flashed residue (short residue) is cooled by heat exchange against long residue feed. A slipstream of this
cooled short residue is returned to the bottom of the high-vacuum column as quench to minimise cracking
(maintain low bottom temperature).

LUBOIL HIGH- VACUUM UNITS

Luboil high vacuum units are specifically designed to produce high-quality distillate fractions for luboil
manufacturing. Special precautions are therefore taken to prevent thermal degradation of the distillates
produced. The units are of the "wet" type. Normally, three sharply fractionated distillates are produced
(spindle oil, light machine oil and medium machine oil). Cutpoints between those fractions are typically
controlled on their viscosity quality. Spindle oil and light machine oil are subsequently steam- stripped in
dedicated strippers. The distillates are further processed to produce lubricating base oil. Short residue is
normally used as feedstock for the solvent de-asphalting process to produce deasphalted oil, an
intermediate for bright stock manufacturing.

HIGH- VACUUM UNITS FOR BITUMEN PRODUCTION

Special vacuum flashers have been designed to produce straight-run bitumen and/or feedstocks for
bitumen blowing. In principle, these units are designed on the same basis as the previously discussed feed
preparation units, which may also be used to provide feedstocks for bitumen manufacturing.
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In this article by AIMM technologies, not only the innovative heat exchanger cleaning process is Average Score: 4.33
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Your_Account A patented tube cleaning process called Hydrokinetics has helped AIMM Technologies build a business
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By Carol Brzozowski-Gardner
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Hydroblasting and chemical cleaning are the dominant methods for industrial cleaning. But AIMM
Technologies, a service company in LaMarque, Texas, has used a patented ultrasonic-based cleaning
process to make headway in the United States market as well as internationally.
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AIMM manufactures the Hydrokinetic equipment for its own use. The process uses sonic resonance with
water to clean pipes and tubes. Most clients are in the petrochemical field; the balance are in government
services and food processing.
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Search Brooks Bradford, AIMM president, bought the company five years ago from Ralph and Pat Garcia, who Send to a Friend
started it in 1991 and patented the Hydrokinetic process (Ralph was the inventor). Since Brooks took over,
the company has expanded the process, received a second patent, and has a third patent pending.
Company revenues have tripled, and AIMM has gone global. Besides operating offices in Texas, Louisiana
and Monterrey, Mexico (where Miguel Morett is country manager), AIMM partners with companies in
Search Norway, the United Kingdom and Saudi Arabia.

Another partner, The Atlantic Group, provides service to the U.S. Navy, which added Hydrokinetics to its
technical manual this year, opening the doors for servicing Navel vessels worldwide. “The Hydrokinetics
process is impressive from several viewpoints, but the thing that impressed me initially was its safety,”
Brooks says. “It’s inherently safe, and in the industrial cleaning industry, safety is paramount. There is no
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AIMM employee Americo Almeida operates a Hydrokinetic ram assembly. Hydrokinetics is a

new tube cleaning technology developed by AIMM, based in LaMarque, Texas.

In Hydrokinetics, the tube or pipe to be cleaned is first filled with an oscillating water stream based on
resonant frequency. Once flow is achieved, a bullet is introduced in the tube to maintain the water column.
The tube is then refilled with water. The water (or sonic) resonance is transferred through the water
column.

The fouling and the tube resonate at different frequencies, breaking the bond between them, allowing the
fouling to be easily expelled out of the opposite end of the tube. The fouling is expelled in a snakelike
fashion rather than in particles. Additionally, the material comes out the end opposite the operator.

Upper photo, heat exchanger tubes before Hydrokinetic cleaning; lower photo, entry to tubes
after Hydrokinetic cleaning.

The equipment consists of a hydraulic monitoring device in a sealed cabinet. With the availability of a
conventional plunger pump (10,000 psi at 20 gpm) and 120 psi service air, technician and unit can perform
cleaning anywhere.

Brooks, who owned a rental company providing refrigeration services before he acquired AIMM, bought the
cleaning company because he was impressed with the Hydrokinetics method and its low ratio of employees
to income. He also likes the fact that the method produces immediate results. Despite its emphasis on
Hydrokinetics, AIMM still does traditional cleaning. The company offers hydroblasting up to 36,000 psi for
industrial cleaning or surface preparation.

While safety is a prime concern in any industry, it is even more so in chemical plants, where AIMM does
much of its work. AIMM counts on Hydrokinetics to be safe but also has developed a safety program that is
 approved by all clients and is adhered to daily, says Gary Dunn, health, safety and environment (HSE)
manager.

No lost-time accidents
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A Hydrokinetics nozzle about to enter a tube.

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Last 10 page visits Refinery

Technology On Line

01 Guest
---> News
02 Guest
---> News
03 Guest
04 Guest
---> News
05 Guest
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06 Guest Pieces of scale removed from tubes.
---> News
07 Guest
In AIMM’s 10 years, the company has never had a lost time accident. Clients’ industrial plant safety
---> News
engineers and supervisors appreciate that record. Ronald Cotton, AIMM operations supervisor, attributes
08 Guest the safety record to the fact that Hydrokinetics does not expose the operator to high-pressure water or
---> News waste materials. In addition to safety, AIMM vice-president Ralph Garcia favors Hydrokinetics for:
09 Guest
---> News Time savings. The method is 40 to 50 percent faster than others, reducing the hours needed for
pipefitting and saving on downtime for clients. “It gets them back into production faster, and that’s
10 Guest
where they make their money,” Ralph points out. “These plants work on several hundred thousand
---> News
dollars an hour, and if you can bring them back into production 5 to 10 hours faster, then you’ve
done a good job for them.”
View MS-Analysis Water savings. Hydrokinetics uses 90 percent less water than conventional industrial cleaning
methods.
Ecological advantages. The method generates 75 percent less wastewater than traditional
methods.

Ephemerids So what’s the down side of Hydrokinetics? Some question whether Hydrokinetics stresses tubes and
piping. The company says the maximum frequency of 11,250 vibrations per minute is far below the
Love is a matter of chemistry, number of cycles that would cause fatigue in even the softest metals.
sex is a matter of physics.
For some companies, cost may be a factor. Antone Belcher, operations manager, concedes that
Hydrokinetics costs more than traditional methods. “But it’s probably less expensive when you consider
the entire cost,” he adds. “We get them back into service faster. And we clean much cleaner, so that the
tube will stay cleaner longer and allow them to remain in production for a longer period. When you consider
Web Search all those things, we’re probably cheaper.”

90 percent success
 Search at Google
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Visit Us Again

· Bookmark Us Brooks Bradford, president of AIMM Technologies, shows the equipment used for the
· Set Your Home Page company's patented Hydrokinetics process.

Technicians Domingo Blanco (left) and Mark Guidry display foulant removed from tubes.

Company officials say Hydrokinetics cleaning is successful in more than 90 percent of applications. It often
works best where other traditional methods fail, such as in U-shaped piping or tubing, confined
workspaces, and lines that are completely blocked.

“Anybody who says they’re 100 percent on any job might be hedging a little bit,” says Brooks. “We’re not
100 percent; I don’t know that anyone is. But with things that give us a problem, they’re impossible for
other people.” Ask Brooks to cite his company’s most difficult job and, like most company presidents, he
laughs and says they’re all difficult. But in fact, AIMM’s files bulge with examples of challenging projects.
For instance, AIMM used Hydrokinetics to clean 14 heat exchangers in a Houston-area petrochemical plant
with a total of 7,000 tubes that were severely fouled with polymers — and did it in onefifth the time
traditional cleaning had taken.

In another case, the company saved a Houston syngas plant half a million dollars in potential capital
investment by cleaning nickel and sulphite from an 840-tube exchanger the plant was planning to scrap.
While only 35 tubes in the bundle had been open after sand jetting, the AIMM Tech crew opened up 66
percent of the tubes.

Image

Foulant exits a pipe during Hydrokinetics cleaning.

Foulant comes out in a snakelikefashion.

Prestigious clientele

The company’s clients include such prestigious firms as Dow, Chevron, Exxon Mobil, Fina, Bayer, and
Union Carbide. Brooks is not intent on amassing a large number of clients — he prefers doing quality work
for his present base of loyal customers. AIMM focuses on tube and pipe cleaning. “Everything else we leave
to other companies,” Brooks adds. “As a result, we know what we’re doing. We get in and do it in a safe,
fast, profitable manner. And we’re constantly trying to be in front of the customer, trying to determine
what his needs are, what his problems are, and helping him to solve those.”
 AIMM is now considering license agreements or joint venture offers from Brazil, Holland and India and
expects these to be in place by 2003.

Ralph says he’s happy with the company’s growth — each year, revenues have mushroomed by 30
percent. “A lot of that goes back to a combination of the process and the individuals we have going out and
representing the company.”

Employees key to success

AIMM has 30 employees. Antone says that when he looks for an employee, he looks for someone who is a
“cut above the rest. We try to find someone who has some formal education. We don’t necessarily look for
experience. We have them trained in our process rather than repeating some mistakes they may have
made somewhere else.”

Bo Davenport, shop manager, says he also looks to hire someone with potential to become a supervisor.
“We’re not looking for laborers,” he says. “We’re looking for people who can lead and represent the
company in the manner that we want it to be shown.”

Brooks attributes his company’s success to making the right hires. He offers employees incentives such as
a percentage of profits from a job, health benefits, and a newly introduced 401k program. “Some of these
things are very expensive for a small company,” he concedes. “But in order to attract the right people,
those who we would like to have carry this company to the next level, those are the things that we will have
to offer.”

Hiring and retaining with such a focus is essential. Brooks says five years from now, his company may be
at the point of going public. So he’s been developing the management in-house to ensure the company’s
continued success long after his own retirement.

Going Global

There are cultural considerations when a company takes a global presence. Out of respect for the local
cultures overseas, AIMM Technologies educates its employees on the local customs through in-house or
outsourced training.

“There’s a vast difference between going to England or Norway versus going to Saudi Arabia,” says
Brooks Bradford, AIMM’s president. The amount of instruction employees receive depends on where they
are going and how long they will stay.

For example, Antone Belcher, operations manager, recently completed a fivemonth project in Central
Africa in an area with few roads, where most travel was by air or water. “As a result, you’ve got to learn
to be patient and make do with what you have,” he says. “You learn to work with the local population,
not against it. Their workdays and their attitudes are different.”

“You have to be careful who you send on a project like that, as to whether they can stand the type of
pressure that comes about both daily and after they’ve been there for 30 days. If they’re off in the
jungles, can they handle that isolation? Not many people can do it.”

Productivity takes a nosedive after about three weeks. Consequently, managers rotate employees in and
out of international projects on a three- or four-week basis. While Brooks tries to employ local people, he
sends an AIMM supervisor to every job. If the job requires more U.S. employees, he’ll send them. But
his goal is to have his alliance partners act independently, with little input from the U.S. AIMM office.

For more information you can contact Antone Belcher at belcher@aimmtechnologies.com

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Downloads & Links What is the best way to quickly clean an exchanger that has 3,870 plugged tubes that are 70 feet long?
Web_Links That is the question that presented itself to the turnaround team a few weeks before PTR4 was shut down.
Downloads That's a lot of tube length, especially when you take into account that there were two exchangers to clean.
If placed end to end, that would be over 100 miles of tubes. AIMM Technologies (AIMMTECH) had just the Please take a second and vote
right answer, "HydrokineticsTM". What is "HydrokineticsTM", you say? Well, its a cleaning process for this article:
Infos developed by AIMM Technologies that uses the induction of sonic resonance into a cleaning water stream
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Site stats material, they resonate at different frequencies. When a sonic resonance is introduced through a water
Your_Account medium, the bond is broken between the tube and the fouling material, thus allowing it to be expelled
WebMail easily. AIMM Technologies uses a two phase system to accomplish this. First, the tube is pumped full of
Quizz water and the sonic resonance introduced. After the material is broken up, a bullet made of a hardened
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Software The process not only proved to be quicker than conventional lancing, but also more compact and cleaner.
Success Stories Instead of having to build a long platform to accommodate the use of a 70 foot plus stiff lance, a small
Training scaffold was only required. The water was injected using a small gun mounted on a track, which didn't
Websites spray water back on the technicians. All of which proved to be a safer, cleaner, and a quicker method!

Image

For more information you can contact Antone Belcher at belcher@aimmtechnologies.com

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Gamma scanning is troubleshooting technique for distillation columns that can be used to detect coloumn
Infos malfunctioning due to overloading, damage to internals or severe fouling. The technique effectively measures
Members_List density through the column shell, clearly differentiating betweena areas of liquid space, vapour space and the
Site stats metal of the column internals. In this article by Towerscan, a company specialised in this technique, the basics
Your_Account of gammascanning is explained and some real life examples are given how this technique is used on trayed
and packed distillation columns.
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Gamma scanning technology for troubleshooting of
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distilation columns
Search Introduction

When a process column develops a problem, your engineering and operations personnel need to know what is
happening with the tower. Any kind of difficulty within a distillation column is very expensive when it involves
Search lower yields, off-spec product, or having to shut down. The more quickly the problem is corrected, the lower
your blood pressure, and the higher your company’s profits.

Gamma scanning provides a fast, efficient means of finding out what is happening inside your column, while it
is still operating.
Extra Resources
Gamma scans are a very effective, well established, troubleshooting technique. As illustrated in the below
RTOL?s Useful Stuff chart, the information from a scan, can help diagnose and solve approximately 70% of the column problems
your staff will encounter. In the other 30% of the cases, a gamma scan often eliminates possible scenarios,
allowing resources to be focused on finding the true source(s) of the problem.
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Figure 1

Gamma scanning of trayed towers

Gamma scanning service increases your profitability and provides a fast, efficient means of determining:

the location of damaged and/or missing trays

the severity and extent of flooding, entrainment, foaming, or weeping
the location of process bottlenecks
the depth and relative densities of the aerated liquid on the trays, and the liquid level in the base of
the column
and preventing unnecessary shutdowns

Column scans, a well established troubleshooting tool, are quick, safe, and easy to set up. They require no
support from process personnel other than drawings of the column, a work permit, and access to the top of
the column.

Figure 2: scan orientations for a trayed column scan

A gamma scan is performed by placing a small radioactive source, and a sensitive radiation detector on
opposite sides of an operating distillation column. As seen in Figure 2, the source and detector are aligned so
that the path between them is across the tray active area, but avoids intersecting the tray downcomers.
Maintaining a constant geometry, any variation in the signal is due to density differences within the column,
as the scan proceeds down the tower.

Image

Figure 3: Gamma scan results

In Figure 3, overlying a baseline scan (red), is a scan profile (blue) of a column that had experienced damage
to tray 6, the debris from which has caused flooding from tray 9 upwards. The second scan (green), on the
lower half of the chart, shows the profile typically seen when a column experiences entrainment problems.

The effectiveness of gamma scanning is always increased when the opportunity arises to compare two or
more scans, whether they be at different operating conditions, or a current scan vs. a baseline scan taken
after a turnaround. These situations offer great process optimization opportunities.

Trayed tower customer example

Pre Job Info:

Under normal operating conditions, this tower was experiencing a high pressure differential across the column
with liquid being carried overhead.
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· 1: Energy Best Practices (US

Department of Energy)
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articles Figure 4
· 3: APC-network.com
· 4: Jardine Solutions TowerScan Results:
· 5: Site remediation The rates were reduced in order to obtain a scan under stable operating conditions. The first scan showed
technologies overview trays in the middle of the tower (11-17) to be flooded, with evidence of entrainment on the trays higher up the
· 6: Lubricants and Lube Oil tower. The unusual aspect of the scan profile was the extremely high density seen for the liquid on tray 11.
Tutorial (Shell Global Given that tray 11 was holding liquid, and therefore mechanically sound, TowerScan personnel were confident
Solutions) in predicting that the vapor was bypassing the tray via the downcomer, as there was no aeration of the liquid.
· 7: HPLC Description of
Refinery Processes The mechanism causing the vapor bypassing wasn't 100% certain from the profile. Trays 10 and 9
· 8: Sulphur Review immediately below 11 showed reduced liquid loadings on the tray. Typically this is due to damage, however
(Environment Directorate the possibility of the trays from 11 upwards loading up with liquid and then dumping, temporarily starving the
General) trays below of liquid remained as an alternative explanation.
· 9: Joshua Corbin's Oil
Refinery Process Description Subsequent inspection of the tower revealed that the tray deck panel immediately underneath the downcomer
· 10: CIA Delayed Coking from tray 11 was damaged. As such it allowed the vapor to bypass tray 11, and most of the liquid to bypass
Technical Papers trays 10 and 9 even though the remainder of that decking was mechanically sound.

Gamma scanning of packed towers

Stats summary When a packed bed column is scanned, the two most common scenarios are where a single scan is performed,
or scans on a 2 x 2 grid as shown in Figure 5.
Last 10 page visits Refinery
Technology On Line Image

01 tmura214 Figure 5: scan line orientations for a packed bed "grid" scan
---> Web_Links
02 tmura214 A single “TowerScan” on a packed bed column will determine:
---> News
03 Guest whether all the packed beds are present in the column
04 Guest if the beds have experienced any damage or settling of the packing
---> News the liquid level on chimney and collector trays
if any of the beds have experienced flooding or fouling
05 tmura214 whether the demister pads and distributors are at their proper elevations
---> News
06 Guest By performing and analyzing a full 2 x 2 grid scan, the evaluation will reveal:
07 tmura214
---> News the extent and location of liquid maldistributions within the packed beds
08 tmura214 if tower internals such as demister pads, distributors, trays, and the top of beds are level

09 Guest
---> CCart In addition to using gamma scans as a troubleshooting tool, some companies make excellent use of the
information to schedule shutdowns. In processes where they know the packing gradually becomes fouled,
10 Guest periodic scans monitor the progress of the fouling. They can then accurately schedule their turnarounds, as
---> News opposed to having to incur sudden, unexpected, and costly shutdowns. Examples of “before” and “after”
scans performed on a tower are shown in Figure 6.
View MS-Analysis
Image

Figure 6: Gamma scan results

Ephemerids
The initial scan (red) showed that the demister pad, distributor, packed bed, and chimney tray were at the
Love is a matter of chemistry, proper elevations. The bed however had experienced severe fouling, especially in the bottom two thirds of the
sex is a matter of physics. packing. The column was shut down and the packing replaced.

PThe scan taken after the turnaround (blue) exhibited a far more uniform bed density. Some higher densities
were seen at the top and bottom of the bed due to the hold down plate and bed support, and part way down
the bed due to external mechanical interference.
Web Search

Packed tower customer example

Pre Job Info:

Historically, this tower with multiple packed bed was relatively stable in its operation but running over several
years between shutdowns it was prone to fouling. As such the customer had TowerScan to perform baseline
grid scans so that they could minimize in the future time spent on turnarounds as well as aiding with
troubleshooting.

Image

Figure 7
 TowerScan Results:
The adjoining scan profile shows the grid scan performed two years after the baseline scan, as well the profile
(black) for one of the four baseline profiles. Each of the six packed beds consisted of a shorter section of
structured packing in the top of the bed, with the majority of the bed consisting of dumped packing.

The greatest contrast was between the two beds shown, beds C and D. While all of the beds showed some
Search at Google fouling in the smaller structured packing element at the top, the remainder of the beds, consisting of the
dumped packing, were seen to be still operating with a uniform density profile, clear of significant fouling. The
Search! exception was bed D below the feed inlet, which showed appreciable fouling throughout the bed. Armed with
this knowledge going into the shutdown, the unit engineer was able to minimize the downtime by planning to
clean the distributor above each bed, as well as the top structured packing element, while leaving in place all
of the dumped packing except for that in bed D which was replaced.

Towerscan gamma scanning services

Visit Us Again TowerScan was established in order to provide the highest level of quality and service in the process
troubleshooting and optimization industry.
· Bookmark Us
· Set Your Home Page Our personnel have been working in the industry since 1989, and over the years have developed many
technical innovations. We have performed hundreds of jobs across the U.S., Canada, and around the world.

TowerScan's process diagnostic services are fast, safe, and economical. We service the refining,
petrochemical, gas processing, oil & gas production, pulp & paper, and pipeline industries. We help our
customers come up with the solutions that:

Increase product yield, throughput, and profit

Avoid unnecessary shutdowns
Plan for turnarounds, minimizing downtime

We offer a complete range of process diagnostic services including:

Gamma scans on distillation columns

Neutron backscatter level / interface measurements on process equipment
Pipe scanning to locate blockages and the buildup of deposits in process lines

Please feel free to contact Ian Blackmore should you have any questions or would like additional
information. Visit our web site soon.

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RTOL Navigation New industrial distillation technology - Linas distillation technology Related Links

Linas technology is a revolutionary distillation technology that could revolutionarise the refining · More about Distillation
business. It is based on film distillation, which requires significantly less vapour/liquid traffic in the
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column to achieve the separation targets. Because of that characteristic, Linas distillation towers can
be drastically smaller than conventional towers (3 to 10 times), and also energy consumption is significantly
Community lower (at least 10%). In this article, the basics of the technology is explained, illustrated by practical Most read story about
industry applications. Distillation:
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Linas-Tekhno Inc., Novosibirsk, Russia Votes: 4
Stories_Archive G.A.Nesterov - Linas Technology International Corporation, New York, USA For further information or
Calendar enquiries, contact G.A. Nesterov at nesterovg@cs.com.
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1. Importance of distillation
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Web_Links for this article:
Downloads Modern industry requires more and more very pure chemicals. Therefore separation technologies become
more and more important, complex and more expensive . The most common separation technology is
distillation. Modern industrial distillation was established some 40 to 50 years ago. Distillation consumes
Infos huge amounts of energy and it can generate more than 50% of plant operating cost. Significant efforts by
Members_List thousands of researchers and developers around the world during the second half of the twentieth century
Site stats did not bring a sufficient improvement in industrial distillation. Main efforts were concentrated on local
Your_Account improvements and the main principles of the modern distillation were not changed for a long time. In
WebMail general distillation is a very conservative process. Huge rectification towers at chemical and petrochemical
Quizz plants are a symbol of modern industry. Very often the height of rectification towers exceeds 50 m. Costs of
design, manufacturing, transport, operation and repairs become enormous. Rectification towers looks like
dinosaurs and clash with the technological image of the 21st century. Cast my Vote!
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For more than 50 years of an existence of modern distillation the opinion was gradually created, that
significant improvement of industrial distillation is impossible. However the careful and detailed analysis of
· Disclaimer principles of modern industrial distillation shows several fundamental contradictions and disadvantages. Options
· Recommend_Us These contradictions are concentrated in an arrangement of heat and mass exchange processes and reflux.

2.1. Arrangement of heat and mass exchange processes. Printer Friendly Page

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Indeed the distillation is based on heat and mass exchange processes between two or more compounds. In
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the modern industrial distillation the basic and actually the main attention is given to the mass exchange
process. All efforts of technologists and engineers are directed on improvements mass exchange processes
between substances. Complex and expensive packing and plates of the most complex forms are created.
Thus to the heat exchange processes between substances it is not given due attention.
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From our point of view the basic problem of modern distillation consists in absence of the control over
energy of condensation of substances. Moreover, the control over energy of evaporation and condensation
is primary process, and it can operate and rule the mass exchange process.

Extra Resources Let's consider the elementary case of water and alcohol distillation (Fig.1). After evaporation of both
substances, there is their condensation on surface of packing or plates. Thus it generates large amounts of
the heat of condensation.
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Training Fig.1. Scheme of elementary step of a water-alcohol distillation
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Modern distillation practically does not pay an attention to the heat of condensation and does not operate
with him. It results to that a part of the condensed water can again evaporate and to rise above on a column.
In an ideal step after the first act of evaporation and condensation all amount of water should be removed at
a bottom of the column, and all amount of alcohol should be evaporated again and to be removed at the top
of the column. It means, that the energy of water condensation is necessary to remove from a distillation
space or to make a heat transfer to a process of alcohol evaporation. However in modern columns there are
no devices or technology which would supervise this process.

Thus, the energy of a condensation is not supervised and it results in an efficiency reduction of a separation,
increase in height of columns and expenses of energy. It is necessary to organize a certain device or
technology which would supervise the energy of condensation of separated substances. Introduction of such
mechanism would lead to sharp decrease in a height of a column and an efficiency of distillation.

2.2. Arrangement of reflux.

The second weak point of the modern industrial distillation is the arrangement of reflux and its distribution
along height of a rectification tower. As the reflux a pure distillate is usually used. The reflux is entered
through a special device at a top of the column (Fig.2). Simple analysis of mass exchange processes in a
column show, that huge surplus of pure distillate on top of a column is not required. Moreover, the reflux at
the top of a column is absolutely not necessary. The amount and structure of a reflux should be various on
the height of a column. The bottom part of a column needs a plenty of reflux with a large amount of a higher
boiling component. With increase in the height of a column a content of higher boiling component in reflux
should decrease and an amount of the reflux should decrease too (Fig.2. Desirable distribution of reflux).
This arrangement of the reflux reduces a total amount of the reflux and makes the reflux more effective.

Fig.2. Scheme of reflux distributions for a conventional and Linas distillations (V is vapor).

Modern industrial distillation uses a different arrangement of reflux. There is no the arrangement of the
reflux composition and an amount of the reflux along the height of a column. It results in a sharp increase in
expenses of energy and increase in height of columns.

Thus, from our point of view the basic lacks of modern industrial rectification are the following:

1. Modern design of rectification towers does not pay attention to real microbalance of heat and mass
exchange processes.
2. Mass and heat exchange processes in every point of conventional columns are not correlated to each
others.
3. Modern arrangement of a composition and an amount of reflux does not correlated to real
distribution of low and high boiling compounds along rectification towers.

This results in columns being very tall and requires more energy for the distillation process. Finally this
results in a high cost of an operation.

3. Linas distillation technology

A small group of enthusiastic and highly professional researchers and engineers united in companies known
as Linas-Tekhno Inc., and Linas Technology International Corporation made an attempt to develop a new
breakthrough distillation technology. This attempt was successful. They have developed a totally new
solution for industrial distillation and have developed the new industrial distillation technology called Linas
technology.

Linas technology is based on a modified very much film distillation. Indeed, the conventional film distillation
has several attractive advantages such as simple construction, a very low flow resistance and good
separation ability. The film distillation has the lowest height of theoretical tray (equal 5 mm) among all
distillation technologies, but only if vapor velocity is around 1 cm/s. Therefore applications of the film
distillation are very limited and not really applicable in a large scale industrial distillation. At high velocity of
vapor along vertical surfaces a film ceases to be uniform and heat and mass exchange processes become
unstable and all advantages of film rectification are not realized.

This main problem of the conventional film distillation was solved by Linas technology. Linas technology is
based on vertical tubes with a length from 0.5 m till 3 m and a diameter 6-25 mm. Linas technology
concentrates the main attention on the energy of a condensation of separated compounds. Temperature of
Linas tube?s wall (Fig.3, temperature TW) is fixed on certain level along a height of tube between TA and TB.
Under distillation conditions the compound B with higher boiling point (TB) condenses always on the walls of
tubes and removed down in a liquid form. The compound A with lower boiling point (TA) is evaporated
always from a surface of the walls and leave tube?s space as a final distillate. Indeed the energy and an
energy barrier are the moving force of Linas technology.
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Fig.3. Moving force of Linas technology (TA and TB are boiling temperatures of compounds A and
· 3: APC-network.com B).
· 4: Jardine Solutions
· 5: Site remediation Industrial applications of Linas technology give:
technologies overview
· 6: Lubricants and Lube Oil
Tutorial (Shell Global 1. Stable distillation film under a velocity of a vapor stream inside of Linas?s column up to 1.5-2.0 m/s.
2. Adaptation of a heat and mass exchange processes inside the Linas column to the physical
Solutions)
properties of separated compounds.
· 7: HPLC Description of 3. New arrangement of the reflux process. All evaporated distillate is a final product according to the
Refinery Processes scheme below. Linas reflux process takes place inside the Linas distillation towers.
· 8: Sulphur Review 4. Three to ten fold reduction of the height of rectification towers and 50 to 100 time reduction in the
(Environment Directorate amount of separated compounds inside the Linas column compared with conventional rectification
General) towers. Technologic scheme of Linas rectification tower is little bit different from the conventional
· 9: Joshua Corbin's Oil one (Fig.4). There is no a reflux line back to the tower from the distillate drum.
Refinery Process Description
· 10: CIA Delayed Coking
Technical Papers

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Fig.4. Technologic scheme of Linas and conventional rectification towers.
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The Linas tower contains an additional device (the condenser) compare to the conventional tower.
08 tmura214
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4. Industrial applications of Linas technology
09 Guest
10 Guest
---> News Because of commercial reasons only several industrial applications of Linas technology were developed for
last four years. These applications are presented below.
View MS-Analysis
4.1. Distillation of trifluoride-methyl sulfonic acid.

The first industrial application of Linas technology was accomplished four years ago for a distillation of
trifluoride-methyl sulfonic acid at the Angarsk chemical plant. The height of the column is only 1.4 meters.
Ephemerids After four years of a continuous operation the rectification column did not demonstrate any real distillation
problems.
Love is a matter of chemistry,
sex is a matter of physics. 4.2. Linas technology in oil refinery applications.

Linas distillation technology was applied for an oil refinery two years ago. First oil refinery SMR-8(10) (8000-
10.000 MTY of crude oil) is in operation in West Siberia (Russia). Certified gasoline, diesel oil and fuel oil
were produced by the oil refinery. A height of rectification part of the tower is 1.5 m. In operation, the
Web Search rectification tower contains 1.5 kg of compounds only. Industrial operation of SMR-8(10) confirms the
advantages of the new technology. A general view of SMR-8(10) is presented on a photo below. The
rectification tower is very small and it is even difficult to find the tower in the photo.
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Technical data of oil refinery SMR-8(10) (Linas-Tekhno) based on Linas technology and a conventional oil
refinery presented in Table. Indeed the technical parameters of Linas rectification tower are superior to
conventional rectification tower.

Type of plant SMR-8(10) conventional

Capacity 8000-10000 MTY 7000 MTY
Type of feed Crude oil Crude oil
Yield of useful products >99.9 % wt. >99.9 % wt.
Pressure 0.11 MPa 0.17 MPa
Weight of unit 4.2 t (column + furnace + reboiler) 5.6 t (column only)
Height of column 1.5 m 10 m
Diameter of column 500 mm 520 mm
Consumption of furnace 13-24 kg/h 12.6-22 kg/h
Electric power 8 KW 27 KW
Operation conditions -55C ? (+55C) -30C ? (+55)
Type of operation continuous continuous

Next generation of Linas oil refinery SMR-50 (50,000 MTY of crude oil) will be in an industrial operation in
July 2004. The oil refinery contains a rectification tower with a diameter 1.0 meter. A height of the
rectification part is 2.0 meter and a height of the stripping is 1 meter. Total height of the column is around 5
meters. General scheme of the column of the oil refinery SMR-50 is presented on Fig.5.

Image

Fig.5. General scheme of Linas rectification tower including in oil refinery SMR-50.

Several SMR-50 will be built up and start up at the end of 2004 and the beginning of 2005 in Russia. Current
commercial price of the SMR-50 begins from $ 1.2 million. Linas rectification tower for the oil refinery with a
capacity 50,000 MTY is much smaller than an average size of conventional rectification tower with the same
capacity 50,000 MTY (Fig.6).

Image

Fig.6. Comparable size of Linas and conventional rectification towers for oil refinery with the
capacity 50,000 MTY.

Companies Linas-Tekhno Inc., and Linas Technology International Corporation plan to develop and built up
in a nearest future the Linas oil refinery towers with a capacity 150 MTY and 500 MTY. Heights of
rectification towers for SMR-150 and SMR-500 will be around 6 meters.

Several Linas distillation processes on a pilot plant level were developed.

4.3. Distillation of waste lube oil.

Very compact pilot distillation unit was built up for a regeneration of waste lube oil in 2002. Height of a
rectification part is 1.5 meter only. Total height of the rectification tower is 2.5 meters. Around 75-85 % wt.
of waste lube oil is regenerated and could be used again as lube oil for ship diesel engines. Based on the pilot
plant data the design of an industrial plant with a capacity 75.000 MTY was done.

4.4. Distillation of ethanol-water mixture.

Pilot plant for alcohol-water distillation was designed and built up at beginning of 2003. The pilot plant was
tested in the Netherlands in company Zeton BV and presented in Achema (Germany) in May 2003. The pilot
plant contains a rectification part with a height 1.5 meter, stripping part with a height 0.5 meter. Diameter
of the rectification tower is 0.2 meter. Standard feeding is 15 kg of alcohol-water mixture per hour. Feed
contains 6-20 % wt. of alcohol. Concentration of alcohol in a distillate is around 90 % wt. In September
2003 the technology was improved and a distillate with the alcohol concentration 94 wt. % was obtained.
Photo of the pilot plant is presented below.
5. Advantages of Linas technology

Real industrial operation of rectification towers based on the Linas technology confirms several advantages
of the technology such as:

1. Heights of rectification towers were reduced 3-10 times compared to heights of conventional tray
and packed columns.
2. Amount of separated compounds inside of the columns is reduced 50-100 times compared to
conventional rectification towers.
3. The residue time of the rectification process is between 2 and 60 seconds only.
4. Total reduction of energy consumption is 10 % at least.
5. It is possible to separate thermo unstable compounds.
6. Operation, cleaning and repair costs are at least 50% less than for conventional columns.
7. Dramatic reduction of manufacturing, transport and assembly costs.
8. Totally simple scale up process. For example the design of Linas rectification part included in Linas
rectification tower for crude oil refinery with a capacity from 70 MTY till 50,000 MTY is presented
below.

Image

The basic element of above described rectification sectors is the single Linas distillation tube and the
capacity of columns can be increased by a number of the tubes. Indeed there is no difference in parameters
of distillation processes between one tube with a capacity 70 MTY and 721 tubes with a total capacity
50,000 MTY. It means that the design of any industrial unit can be done very fast if there is data for one
distillation tube. A conventional scale up process is very complex and does not give any warranty for a
successful result. Increasing of a capacity always gives larger diameters and heights of columns.

All advantages of the Linas technology together mean a real breakthrough in industrial distillation.

6. Future of Linas technology

Linas technology has one further advantage. The technology saves money for every step in a long chain of
an industrial distillation from designers to users of rectification towers as is presented below (data are
compared to conventional technology):

1. Design - 2-3 times less cost

2. Scale up process ~ 3 times less cost
3. Manufacturing - 1.5-2 times less cost
4. Transport and assembly (for large rectification units) - 2-5 times less cost
5. Operation, repair and service are at least 50 % less.
6. Energy saving is around 10 % at least.

Linas technology will change the landscape of the chemical, petrochemical and oil refining process. Huge
conventional rectification towers will be replaced by compact Linas columns. It will take many years. First,
the rectification industry is relatively conservative and needs huge investments to built new rectification
units. Second, we need time to convince the rectification world in a high economic and technical efficiency of
Linas technology. The only commercial Linas rectification towers in permanent operation can convince
engineers and investors. Therefore we try to speed up this process and concentrate our main efforts on oil
refinery to start up several Linas rectification tower as soon as possible.

7. The Authors

Albert F. Saifutdinov is a deputy general director of Linas-Tekhno Inc. and a leader of Linas Technology
R&D. He has many years of engineering experience in the nuclear industry of the former Soviet Union.
 Mr.Saifutdinov holds MsD in mechanical engineering from Novosibirsk Technical University and he is a TRIZ
expert for the last 10 years. Mr.Saifutdinov is the founder, inventor and co-owner of Linas technology.

Oleg E.Beketov is a general director of Linas-Tekhno Inc. He has many years of research and development
experience in the nuclear industry of the former Soviet Union. Mr.Beketov holds an MsD in mechanical
engineering from Altai Technical University. Some eight years ago, together with partners, he founded the
company, Linas-Techno Inc. Mr.Beketov is an inventor and a co-owner of Linas technology.

Image
Viktor S.Ladoushkin is CEO of Linas-Tekhno Inc. He has ten years experience in technology development
for the nuclear industry of the former Soviet Union. Mr.Ladoushkin holds a MsD in technology of machinery
from Novosibirsk Technical University (Novosibirsk). Eight years ago, along with his partners, he founded
the company Linas-Techno. Mr.Ladoushkin is an inventor and a co-owner of Linas technology.

Dr. Guennadi A. Nesterov is a president of Linas Technology International Corporation (New York, USA).
He has over 28 years of varied experience involving invention and development of new industrial chemical
technologies mainly focusing on catalysis, chemical technology and energy saving processes. Dr. Nesterov
holds a PhD degree in chemistry from Institute of Catalysis (Novosibirsk). During the last 10 years he has
been a consultant for several German, Dutch, Korean and Japanese companies. Dr. Nesterov is an inventor
and co-owner of Linas technology.

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RTOL - November 2003

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RTOL Navigation Process Practice: Refinery configuration - how is it done? Related Links

Many of you must have wondered at some point in time how the configuration for a greenfield · More about Refinery
refinery is chosen. Is there a 'one size fits all' model or logical groups of process unit that you can Processes
Home build from? In this article, Craig Hesser, RTOL's Process Expert explains the logic behind the choices
· News by Craig
for a refinery configuration and gives configuration examples for three archetypes: the gasoline refinery,
Community the distillate fuel refinery and the LPG refinery.
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Refinery configuration - how is it done? for this article:

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Recently, I received a question from a chemical engineer. It was interesting, because it focused on one of
the basic decisions involved in a new refinery: "What is the best oil refining process pattern if we want to
Services produce a maximum amount of gasoline when refining a heavy crude oil." We all know that crude
Store Front distillation is the first major process step that happens when crude is introduced into the refining process, Cast my Vote!
Advertising but how do you decide what comes next?
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The various licensing companies (UOP, IFP, Exxon Research, etc.) as well as most of the engineering and
construction companies and many engineering consultants all have their own ideas, which (naturally)
· Disclaimer feature their own processing strengths and prejudices, and tend to ignore their weaknesses.
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What you can do with the crude naturally depends upon what sort of crude you have. Responding to the
question above, I assumed that it is not only heavy, but also contains a medium amount of sulphur (2-3%)
and middle- to high-level of metals (particularly nickel and vanadium), and not an excessive level of organic
nitrogen. In other words, nothing very exotic. Printer Friendly Page
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A more-or-less traditional flow scheme from the USA (which is where the gasoline refinery concept was
All Resources born) for a medium to large gasoline refinery is shown below.
Articles
Best Practises
Catalyst Catalogs Atmospheric distillation block
Discussion Groups Atmospheric distillation
Events Full range naphtha HT à splitter à heavy naphtha to catalytic reforming and light naphtha
Glossaries directly or via C5/C6 isomerisation to gasoline
Jobs Light distillate (kerosene) HT à burning kerosene and/or jet fuel
Learning Mid-distillate (gasoil) HT à Diesel fuel and light heating oil
Licensed Processes Vacuum distillation block
Magazines Vacuum distillation
News Sources Delayed coker (a large one, so that essentially no vacuum residue remains)
Software Conversion block
Success Stories
Training FCCU – feeding medium VGO and heavy VGO with a feed hydrotreater to reduce the
Websites sulphur and nitrogen contents
The C3/C4 streams go to alkylation, catalytic polymerisation, and MTBE (except in the USA,
where MTBE now is being removed from gasoline)
An AGO/LVGO hydrocracker (depends on the needs of the gasoline pool and the cost of
hydrogen production) as a combined FCCU feed HT and a gasoline component producer
(the economics should decide whether this is necessary or not)
Hydrogen production
There will be a large volume of refinery gas and C3-C4's coming from the cracking units;
these can be used for hydrogen production (steam-HC reforming), perhaps balancing
gasoline component production against hydrogen requirements Another option, if there is
not enough light material after fuel and LPG product needs for steam-HC reforming, is a
partial oxidation unit, using vacuum residue as the feedstock (this is effective but capital-
intensive) instead of or in addition to the steam-HC reformer, and partially replacing the
coker
Other process and utility units
Steam production for various usages, including utility, process, and power generation
Sulphur recovery – there will be large amounts of sulphur, produced as H2S, that will have
to be converted to sulphur for safe handling
Coke handling – if the refinery coke is not burned in the power plant, it will need to be
handled – one option is to calcine the coke to make it a more valuable product

What you will note above is that virtually all products are hydrotreated at some stage of the process. This is
to make the distillate fuels (gasoline, kerosene for jet or burning fuel, gasoil for heating oil and Diesel fuel)
compatible with today's low sulphur environmental requirements. The product slate should be roughly 55-
65% gasoline by weight, 25-15% light distillates by weight, with the remaining ±20% going to coke,
sulphur, maybe propane/butane, and own consumption (including catalytic coke produced and consumed
in the FCCU). The volume yield of gasoline can be almost 100% on crude, depending on specifications and
the degree of cracking utilised.

If the refinery is to be a (relatively) small one, then the best solution may be to build the entire
atmospheric distillation train, but only the vacuum distillation, the delayed coker, and the hydrocracker in
the vacuum distillation train as a high-conversion unit (primarily gasoline components), and leave out the
FCCU and its downstream units. The hydrogen production facilities will still be necessary unless H2 is
available from outside the refinery. In the future, if the refinery is expanded, an FCCU plus its downstream
units can be added, and the hydrocracker can be converted to an FCC feed pretreater with lower
conversion (50-60 %) but with a much greater throughput (1,5 –2 times original).

Distillate fuel refinery

But what to do if you want mostly distillate fuels instead of gasoline? Well, this is the "European solution" as
opposed to the "American solution" for gasoline above. Typically, the distillate solution (high cetane
number Diesel fuel, light heating oil, kerosene for household fuel and aviation turbine fuel) uses different
choices of conversion equipment, in order to provide more middle distillates and less gasoline.

The earlier solution (1960's) was to use low severity FCC units, which produced more than 50% light cycle
oil instead of a large volume of gasoline. Today, the light cycle oil is a problem because of it's aromatic and
sulphur contents, not to mention stability.

The more current solution is a hydrocracker using (typically) one conversion stage to produce large
amounts of desulphurised, low aromatic distillates (kerosene and Diesel fuel). Depending on catalyst type
and feedstocks, the volumetric yield of distillates can be over 100%. Conversion of the vacuum residue to
hydrogen and energy (steam and/or electricity) may take on a more important aspect, since the large
hydrocracking process requires a lot of hydrogen. However, the steam reforming straight run naphtha to
hydrogen is another H2 source, if gasoline is really unimportant. The refinery is somewhat less complicated
than a gasoline refinery of the same size. The blocks below give an idea of what might be used in a
distillate fuels refinery. The common components that remain more or less unchanged are shown in bold
print.

Atmospheric distillation block

Atmospheric distillation
Medium + light naphtha HT à medium range naphtha to catalytic reforming and light
naphtha directly or via C5/C6 isomerisation to gasoline
Heavy naphtha / light distillate (kerosene) HT à burning kerosene and/or jet fuel
Mid-distillate (gasoil) HT à Diesel fuel and light heating oil
Vacuum distillation block
Vacuum distillation
Delayed coker (a large one, so that essentially no vacuum residue remains)
Conversion block
AGO/VGO hydrocracker optimised for distillate production
HC medium naphtha to catalytic reforming (hydrogen production and gasoline octane), HC
light naphtha to gasoline blending, naphtha sales, or H2 feedstock
Hydrogen production
There will be a respectable volume of refinery gas and C3-C4's coming from the cracking
units; these can be used for hydrogen production (steam-HC reforming), perhaps balancing
LPG component production against hydrogen requirements Another option, if there is not
enough light material for steam-HC reforming, is a partial oxidation unit, using vacuum
residue as the feedstock (this is effective but capital-intensive) instead of or in addition to
the steam-HC reformer, and partially replacing the coker
Other process and utility units (sizes may be different, but functions are essentially identical)
Steam production for various usages, including utility, process, and power generation
Sulphur recovery – there will be large amounts of sulphur, produced as H2S, that will have
to be converted to sulphur for safe handling
Coke handling – if the refinery coke is not burned in the power plant, it will need to be
handled – one option is to calcine the coke to make it a more valuable product
 LPG Refinery

There is a third type variant, although it is relatively rare, and that is the LPG refinery. The objective of the
Languages LPG refinery is to optimise the production of propane and butanes in order to satisfy local market needs for
LPG as fuel and as petrochemical feedstocks. Usually, these refineries are based on light crudes or
condensates as a starting point, since there is a large potential for wasted energy when converting heavy
Select Interface Language:
oils to LPG.

English The special process unit configurations that can be used here are

hydrocracking units that are optimised for LPG production,

catalytic reformers that are in what normally would be a stable but overchlorided catalyst condition
(excessive hydrocracking), and
FCC units trimmed for maximum LPG production (with ZSM-5 cracking additive).
Latest Links
Naturally, in every case, there will be a number of options to consider before the refinery's final process
· 1: Energy Best Practices (US design is set, and special solutions for unusual problems or local conditions are the rule, rather than the
Department of Energy) exception, in refinery basic process flow design.
· 2: Beville Engineering
Alarm/Display technical The selection of the optimum process flow scheme generally takes place by making runs with a refinery
articles economics model. The objective of the model is to product the optimum profit, taking into consideration
· 3: APC-network.com the base feedstock (crude), additional feedstocks, the product requirements, and the prices for feedstocks,
· 4: Jardine Solutions products, and utilities. The first runs are with most, if not all of the process units in the model "open", i.e.
the economic model is allowed to optimise without considering which of the process units will be built. This
· 5: Site remediation will probably give a refinery with a large number of process units, some of which will be uneconomically
technologies overview small from an operating point of view. Later, the least attractive processes will be eliminated, and the
· 6: Lubricants and Lube Oil capital cost will be factored into the calculation (this may have already been done in the first step). Finally,
Tutorial (Shell Global other considerations (site requirements and limitations, transport considerations for feedstocks and
Solutions) products, transport considerations for refinery equipment delivery to the refinery, market limitations,
· 7: HPLC Description of process complexity, licensing considerations, political considerations, etc.) will finalise the refinery process
Refinery Processes design.
· 8: Sulphur Review
(Environment Directorate This should give you an overview of the process that is termed "refinery process design". Flexibility and an
General) open mind are necessary in order to come up with a refinery configuration that fits all needs.
· 9: Joshua Corbin's Oil
Refinery Process Description
· 10: CIA Delayed Coking
Technical Papers Your feedback and questions are welcome! You can reach me at craig@r-t-o-l.com or on the RTOL Forum
at "Questions to the Editor of RTOL Process Q&A".

About Craig Hesser

Stats summary
Craig Hesser is a chemical engineer, operating specialist, and manager with more than 35 years
experience in refining and petrochemicals, marketing, refining technology, and engineering and
Last 10 page visits Refinery
construction. Craig has a BSChE from Kansas University and an MBA from Pepperdine University. His
Technology On Line
international experience ranges from major to independent organisations in Europe, Africa, both Americas,
and Asia, and he is today a consultant to the refining industry worldwide. He is also a certified SAP
01 tmura214 consultant and a certified CRM (Customer Relationship Management) consultant. Craig lives in the beautiful
---> News mountains of Switzerland.
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