Engineering Encyclopedia

Saudi Aramco DeskTop Standards

DISTILLATION PROCESS

Note: The source of the technical material in this volume is the Professional
Engineering Development Program (PEDP) of Engineering Services.
Warning: The material contained in this document was developed for Saudi
Aramco and is intended for the exclusive use of Saudi Aramco’s employees.
Any material contained in this document which is not already in the public
domain may not be copied, reproduced, sold, given, or disclosed to third
parties, or otherwise used in whole, or in part, without the written permission
of the Vice President, Engineering Services, Saudi Aramco.

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Section Page

INFORMATION ............................................................................................................... 4
INTRODUCTION............................................................................................................. 4
CALCULATION OF VAPOR/ LIQUID EQUILIBRIUM (VLE) ........................................... 5
Ideal and Nonideal Gases .................................................................................... 5
Vapor Pressure..................................................................................................... 6
Ideal Mixtures - Dalton's, Raoult's Laws ............................................................... 6
Mixtures Approximated as Ideal ........................................................................... 7
Two-Component Example .................................................................................... 8
Fugacity (Nonideal Gas Mixtures).............................................................. 9
Equilibrium K-Values .......................................................................................... 10
Relative Volatility ................................................................................................ 12
Nonideal Liquids ................................................................................................. 12
Equations Of State ............................................................................................. 13
SRK Equation of State ............................................................................. 13
PR Equation of State ............................................................................... 13
CALCULATION OF EQUILIBRIUM SEPARATION ....................................................... 14
Equilibrium Diagram ........................................................................................... 14
Vapor/Liquid Phase Diagrams ............................................................................ 16
Bubble Point and Dew Point ............................................................................... 17
Equilibrium Flash Separation.............................................................................. 18
One-Stage Flash...................................................................................... 19
Representation of Petroleum with Pseudo-Components .................................... 21
DETERMINATION OF KEY DISTILLATION PROCESS PARAMETERS ..................... 23
Conventional Distillation Column ........................................................................ 23
Reflux ................................................................................................................. 25
Major Equipment ................................................................................................ 25
McCabe-Thiele Diagram..................................................................................... 25
Effect of Reflux on Required Stages................................................................... 27
Minimum Reflux.................................................................................................. 28
Total Reflux - Minimum Stages........................................................................... 29
Stages Versus Reflux ......................................................................................... 30

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Simulations with PROCESS/PRO II ................................................................... 31

Steps in Column Design ..................................................................................... 33
Column Operating Pressure, Temperature......................................................... 33
Tray Efficiency .................................................................................................... 34
DETERMINATION OF REQUIRED OPERATIONAL CHANGES IN SAUDI
ARAMCO DISTILLATION PROCESSES............................................................ 36
Crude Stabilization and Sweetening................................................................... 36
Condensate Stripping ......................................................................................... 38
Crude Distillation ................................................................................................ 39
Atmospheric Unit ..................................................................................... 41
Vacuum Unit ............................................................................................ 43
NGL Fractionation .............................................................................................. 45
NGL Fractionators ................................................................................... 47
Hydrocracker Fractionator .................................................................................. 51
Summary ............................................................................................................ 54
GLOSSARY .................................................................................................................. 55
ADDENDUM ................................................................................................................. 56
References ......................................................................................................... 56

List of Figures

Figure 1. Propane - n - Butane System Pressure........................................................... 9

Figure 2. De Priester Nomograph ................................................................................ 11
Figure 3. Vapor/Liquid Equilibrium ................................................................................ 14
Figure 4. Equilibrium Diagram ...................................................................................... 15
Figure 5. Phase Diagram ............................................................................................. 17
Figure 6. Bubble Point and Dew Point.......................................................................... 18
Figure 7. One-Stage Flash ........................................................................................... 19
Figure 8. One-Stage Flash on a Phase Diagram ......................................................... 20
Figure 9. One-Stage Flash on an Equilibrium Diagram ................................................ 21
Figure 10. Pseudo-Component Breakdown ................................................................. 22

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Figure 11. Conventional Distillation Column................................................................. 24

Figure 12. McCabe- Thiele Diagram ............................................................................. 26
Figure 13. Effect of Reflux on Required Stages ........................................................... 27
Figure 14. Minimum Reflux .......................................................................................... 28
Figure 15. Total Reflux - Minimum Stages ................................................................... 29
Figure 16. Stages Versus Reflux.................................................................................. 30
Figure 17. Simulation Tower Profile ............................................................................. 32
Figure 18. Abqaiq Crude Stabilizer No. 17 ................................................................... 37
Figure 19. Abqaiq Condensate Stripper ....................................................................... 39
Figure 20. Plant 15 Crude Unit - Ras Tanura Refinery................................................. 40
Figure 21. Plant 15 Atmospheric Crude Unit - Ras Tanura Refinery ............................ 42
Figure 22. Plant 15 Vacuum Crude Unit - Ras Tanura Refinery................................... 44
Figure 23. YANBU Fractionation and Treating: Deethanizer....................................... 49
Figure 24. Ras Tanura NGL Fractionation Plant 10 ..................................................... 50
Figure 25. Ras Tanura Plant 10 - NGL Fractionation ................................................... 50
Figure 26. Hydrocracker Fractionation .......................................................................... 52

List of Tables

Table 1. Typical Overall Efficiencies, % ........................................................................ 35

Table 2. Hydrocarbon Components............................................................................... 46
Table 3. NGL (Light-Ends) Fractionators....................................................................... 48
Table 4. Hydrocracker Production ................................................................................. 53

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INFORMATION

INTRODUCTION
This module reviews the concept of Vapor/Liquid Equilibrium
(VLE), the basic relationships that apply to VLE, and the
process side of the multistage distillation unit operation. It also
reviews the main Saudi Aramco distillation processes. The
module is divided into four parts:
1. Vapor/Liquid Equilibrium: Examines the basic
relationships that apply to vapor-liquid equilibrium.
2. Equilibrium Separation: Introduces concepts such as the
equilibrium diagram and reviews simple separation
calculations.
3. Distillation Process Parameters: Introduces basic
distillation process parameters such as reflux, minimum
reflux, and minimum number of stages.
4. Saudi Aramco Distillation Processes: Reviews the main
Saudi Aramco distillation processes: crude stabilization
and sweetening, condensate stripping, crude distillation,
and NGL fractionation.
Distillation hardware is covered in Module ChE 104.02.
Throughout this module, we will be referring to a number of
equations. The following listing is the nomenclature appropriate
to those equations.

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CALCULATION OF VAPOR/ LIQUID EQUILIBRIUM (VLE)

A vapor/liquid system is considered to be in equilibrium when
there are no longer any detectable changes occurring in the
system. Generally, a system is assumed to be in equilibrium
when the mass, energy, and composition of each phase remain
constant with time. An example of a system in equilibrium is a
mixture of water and air in a closed vessel. After some time,
there will be no change in temperature, in the amount of water
in the vapor phase, or in the number of gas molecules dissolved
in the water. The system is in equilibrium.
Equilibrium also applies to systems that are not static. We may
have equilibrium in an overhead condenser separator of a
distillation column. The vapor and liquid leaving the separator
are in equilibrium, and their compositions can be described by
relationships for systems in equilibrium.

Ideal and Nonideal Gases

Ideal gases are those whose behavior can be described by the
ideal gas law, which is stated mathematically as:
PV
PV = nRT or = 1.0
nRT
The ideal gas law indicates that the product of pressure, (P),
times volume, (V), is proportional to the number of moles (n) of
the component times the absolute temperature, (T). R is an
ideal gas proportionality constant.
Gases tend to behave as ideal gases at temperatures higher
than their critical temperature and pressures well below their
critical pressure.
When the ideal gas law does not apply, the gases are called
nonideal or real.
PV
≠1
nRT
The compressibility factor, (Z), expresses the deviation from the
ideal gas equation. It can be used to predict real gas properties.
The compressibility factor is the ratio of the real gas volume to
that of the ideal gas at the same temperature and pressure:

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PV
PV = ZnRT or =Z
nRT
For an ideal gas, the compressibility factor is 1.0. The
compressibility factor, Z, can be obtained from generalized
graphs such as those in Maxwell, pages 148-153, or the GPSA
Engineering Data Book, Chapter 16.

Vapor Pressure
The vapor pressure of a pure component at a given temperature
is the pressure that is exerted by the component when it is in
the liquid phase. Vapor pressure is a unique property, and it is a
direct function of temperature. A material having a higher vapor
pressure at the same temperature than another is said to be
more volatile.
Vapor pressure (VP) and temperature (T) are often related by
means of the Antoine equation:
B
Log (VP ) = A −
T+C
where A, B, and C are constants for a particular compound over
a relatively narrow temperature range, usually not over 100°C.
Values of these constants for various compounds and the
temperature ranges for which the constants apply appear in a
number of references. The Antoine equation is often plotted in
charts with the horizontal axis in a inverse absolute temperature
scale and a vertical axis in a logarithmic scale.

Ideal Mixtures - Dalton's, Raoult's Laws

Ideal mixtures, gas or liquid, consist of components that do not
interact with each other chemically or physically. The concept of
ideal mixtures has formed the basis for many quantitative
relationships describing equilibrium. Of particular interest are
Dalton's law of partial pressures and Raoult's law relating the
pressure exerted by a component in the vapor phase to its
concentration in the liquid phase.

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Dalton's law states that the total pressure of a mixture (PT) of

gases is equal to the sum of the partial pressures (pp) of the
mixed gases. Thus,
∑PPi = PT = PP1 + PP2 + PP3 + …

Dalton also postulated that the partial pressure of an ideal gas

in a gas mixture is proportional to its vapor mole fraction (y), that
is, the relative number of moles of that gas in the mixture. Thus,
PPi = yi PT

Raoult's law, relating the partial pressure in the vapor phase to

the liquid phase mole fraction (x), is expressed as:
PPi = xi VPi

Combining Dalton's and Raoult's laws results in an expression

describing mixtures of ideal vapors and liquids in equilibrium.
PT = ∑PPi= ∑yi PT = ∑xi VPi

and for component i,

 VP 
y i = x i  i 
 PT 

Mixtures Approximated as Ideal

The mixtures that can be approximated as ideal must satisfy the
following requirements:
• Total pressure of the system must be below 200 psia.
• The components must be chemically similar, for example,
butane and pentane, both of which are paraffins. A mixture
of an aromatic component and a paraffin such as benzene
and hexane cannot be approximated as ideal.
• The components must be close boiling; that is, they must
have similar boiling points.
• The system pressure and temperature must not be near the
critical pressure and temperature of the mixture.
Using ideal mixture correlations in calculations results in
approximate compositions or conditions (P, T). The error may
be acceptable for a simple operation, such as a flash drum
separation. The same correlations used in a superfractionator,

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where tray-to-tray calculations compound the error, may

produce unacceptable results.

Two-Component Example

Let's assume that we have an ideal mixture of propane and n-

butane at 100ºF. The vapor pressures of the two components at
100ºF are:
• Propane 13 atm = 191 psia (Component 1).
• n-Butane 3.5 atm = 52 psia (Component 2).
The total pressure, PT, of the mixture can be calculated from:
PT = PP1 + PP2 = x1VP1 + x2VP2

PT = 191x1 + 52(1-x1) = 52 + 139x1

This last equation indicates that the total pressure of an ideal

binary mixture is a linear function of the composition. This
relationship is illustrated in Figure 1, which shows that the total
pressure is the sum of partial pressures and is a straight line
between the vapor pressure of n-butane (x1 = 0) and propane
(x1 = 1.0).

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Figure 1. Propane - n - Butane System Pressure

Fugacity (Nonideal
Gas Mixtures)

The vapor-liquid equilibrium of an ideal mixture can be

described by:
PPi = yiPT = xiVPi

To improve the accuracy of prediction, we can replace the

pressures by analogous fugacities:
fPPi= yifPT = xifVPi

where: fPPi = Fugacity of i in either phase of the system.

fPT = Fugacity of i as a pure saturated liquid (or

vapor) at its vapor pressure corresponding
to the equilibrium temperature of the
system.
fVPi = Fugacity of i as a pure vapor at the
equilibrium temperature and pressure of the
system.
Generalized correlations have been developed for the ratio of
fugacity to pressure for pure hydrocarbons as a function of

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reduced temperature and reduced pressure. A correlation of this

type was used in conjunction with the vapor pressure charts to
develop the fugacity function charts for individual hydrocarbons.
The fugacity function given by these charts is defined as:
f
VPP y
f i = f i T = PT i
PT xi

The fugacity function, (Fi), may be considered a corrected vapor

pressure and used in place of vapor pressure in any equation
pertaining to liquid-vapor equilibrium.
Values for fugacity functions can be obtained from sources such
as Maxwell's Data Book on Hydrocarbons (pp 49-60). Fugacities
for petroleum fractions can be obtained from a generalized
graph in Maxwell.
Equilibrium K-Values

The definition of equilibrium K-value, also called K-factor or

distribution coefficient, of component i in a mixture is given in
the following equation:
y
Ki = i
xi

The K-value is simply the ratio of the vapor to the liquid mole
fraction of i. This ratio has no special thermodynamic
significance, but has found extensive use in high-pressure VLE
work. For ideal systems where Raoult's law applies, it can be
expressed as:
y VPi
Ki = i =
xi PT

It can also be expressed in terms of fugacities as:

y f
K i = i = VPi
xi fP T

Equilibrium K-values are useful in simple VLE hand calculations.

K-values can be obtained from graphs or nomographs like the
De Priester nomograph, Figure 2. K-values are a function of
temperature and pressure. For nonideal mixtures, K-values are
also a function of composition.

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Figure 2. De Priester Nomograph

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Relative Volatility

Relative volatility (α) is a relation widely used in distillation. It is

defined by:
y /x K
αij = i i = i
y j /x j K j

Relative volatility is a measure of separability. The larger the

value of αij, the easier the separation. For close boiling
components, such as pentane and isopentane, the relative
volatility approaches 1.0.
Because the value of relative volatility is not as sensitive to
temperature as other measures of equilibrium, it is used in a
number of shortcut distillation calculations. For ideal mixtures
(Raoult's law applies), the relative volatility of two components is
equal to the ratio of their vapor pressures.
VPi
αij =
VPj

The small effect of temperature on relative volatility is the

reason for using relative volatility in shortcut distillation
calculations. Relative volatility data for only two or three points
in the column provide results of acceptable accuracy.
Nonideal Liquids

In liquids and liquid mixtures, the distances between molecules

are much smaller than in gases, and the forces attracting
molecules to each other are much greater. Nonideal behavior of
liquids is indicated by heat of mixing and non-additivity of
volumes when two liquids are mixed. The deviation from ideality
is greater for chemically dissimilar substances. The activity
coefficient, γ, measures the deviation from ideal liquid solution
behavior. Using the coefficient in Raoult's law results in:
y iPT = PPi = γi xi VPi

Activity coefficients are used in a number of VLE methods such

as the Chao-Seader and the Grayson-Streed correlations. The
Chao-Seader correlation requires relatively short computing
times. It was used extensively in the '60s and '70s when
computing was costly. Hydrocarbon VLE methods using activity
coefficients have been replaced by the more rigorous equations
of state.

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Equations Of State
Equations of State (EOS) predict the PVT behavior of gases
and liquids. The simplest EOS is the one for ideal gases. Per
mole of gas:
RT
P=
V
In general, real fluids deviate from ideal fluids in two ways: there
are variations in the sizes and shapes of the molecules, and
specific interactions between molecules, such as polarity or
hydrogen bonding, must be considered. The large variations in
size and shape of molecules have a great effect on PVT
behavior.
The Soave-Redlich-Kwong (SRK) and the Peng-Robinson (PR)
equations of state are among the best known.

SRK Equation of
State

The Soave-Redlich-Kwong EOS is a two-parameter equation of

the following form:
RT a
P= −
V − b V (V + b )
where: ( )
a = Σ Σ(1 − c ij ) aia j (x i x j )
i j

b = Σ xibi
i

The parameters a and b must be specified for each component

in a mixture and then combined as a function of composition.
The a parameter is temperature-dependent. In addition, a binary
interaction parameter, cij, is used to calculate the aij term for
mixtures, to improve vapor-liquid equilibrium calculations.

PR Equation of State

The Peng-Robinson EOS is similar to the SRK equation of

state, except that it has an expanded volume term:
RT a
P= −
V − b V (V + b ) + b(V − b )
The a parameter varies with temperature. Both constants use
the same mixing equations as the SRK equation of state.

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CALCULATION OF EQUILIBRIUM SEPARATION

Equilibrium Diagram
Figure 3 depicts a simple flash separation. The feed consists of
two components: propane and n-butane. The feed temperature
and composition vary. Figure 3 lists vapor and liquid
concentrations of propane, and the distribution coefficients (K1
and K2) for propane and n-butane for five temperatures.
Pressure is fixed at 100 psia. The feed composition is not the
same for all temperatures.
At 70ºF, the mole fraction of propane in the liquid phase is
0.746. Its mole fraction in the vapor phase is higher, 0.907,
since propane is the more volatile of the two components. The
distribution coefficient, K1, for propane is equal to the ratio
y1/x1 = 0.907/0.746 = 1.22.

Note: y2 = 1 – y1, x2 = 1 – x1

Figure 3. Vapor/Liquid Equilibrium

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Figure 4 is an equilibrium diagram for the propane/n-butane

system using the data from Figure 3 at 100 psia. The horizontal
axis indicates the liquid mole fraction of the more volatile
component, propane. The vertical axis indicates its vapor mole
fraction. The equilibrium line connects all the (x1, y1) points.
Given the mole fraction in the liquid phase, you can use the
equilibrium line to obtain the mole fraction in the vapor phase.
Given the mole fraction in the vapor phase, you can find the
mole fraction in the liquid phase.

Figure 4. Equilibrium Diagram

Figure 4 contains a second line, the reference line. It is simply

the diagonal of the diagram: for all the points on the reference
line, y = x. The reference line makes it easier to see the
differences between the vapor and the liquid phase
compositions. Since by convention the horizontal axis
represents the composition of the more volatile component, y1 is
larger than x1. Therefore, the equilibrium line is above the
reference line. Large differences in y1, x1 mole fractions

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indicate large differences in the volatility of the two components.

Accordingly, equilibrium lines bulging away from the reference
line are indicative of mixtures that are easy to separate by
successive vaporization and condensation steps, that is, by
multistage distillation.

For some mixtures, there is a reversal in relative volatilities and

the equilibrium line intersects the y = x reference line. Because
the vapor and liquid fractions at that point are equal, these
mixtures cannot be separated by distillation. Such mixtures are
called azeotropes.

Vapor/Liquid Phase Diagrams

Phase diagrams are used to describe two-phase systems by
plotting two of the three independent variables (composition,
temperature, and pressure) at a constant value of the third
variable. Figure 5 is a phase diagram at constant pressure for
the binary mixture of propane and n-butane. The two lines
indicate the temperatures at which a phase change takes place.
The temperatures and concentrations (at the diagram pressure)
below the two lines correspond to an all-liquid mixture. In the
region between the two lines, the vapor and liquid phases are
present. Above the lines, there is only a vapor phase.
The phase lines in Figure 5 were drawn from data in Figure 3.
For example, at 120ºF, the point on the liquid phase line
corresponds to x1 = 0.191 and the point on the vapor line to
y1 = 0.398 (see Figure 3, 120ºF, x1, y1 data). The phase
diagram can be used to determine the compositions of the
vapor and liquid phases from the pressure and temperature at
equilibrium.

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Figure 5. Phase Diagram

Bubble Point and Dew Point

The phase diagram can also be used to determine the phase

transition points. Figure 6 is an example for a mixture of 45%
propane and 55% n-butane at 70ºF. The phase diagram in
Figure 6 indicates that the mixture at 70ºF is in the liquid phase.
If the temperature is increased at constant pressure, 100 psia,
the mixture will be liquid up to 92ºF, at which point vaporization
begins. This is the bubble point of the mixture, the temperature
and pressure at which a liquid is in equilibrium with an
infinitesimal amount of vapor.

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Figure 6. Bubble Point and Dew Point

Between 92ºF and 117ºF, the mixture is in two phases, vapor
and liquid. At 117ºF, all of the liquid is vaporized. This is the dew
point, the temperature and pressure at which vapor is in
equilibrium with an infinitesimal amount of liquid. At
temperatures above the dew point, there is only a vapor phase.
The liquid phase line is the bubble point curve; the vapor phase
line is the dew point curve.
Equilibrium Flash Separation
The equilibrium flash separator is the simplest equilibrium
process for engineers to consider. The process involves the
separation of a two-phase feed into vapor and liquid in a vessel.
The feed is at a desired temperature and pressure, achieved by
heating, cooling, pumping, or letting down with a control valve.
Calculations of the compositions and the relative amounts of the
liquid and vapor phases at any given pressure and temperature
involve a tedious trial-and-error solution. Since flash calculations
can be performed easily by computer, manual methods for
multicomponent flash calculations are not discussed here.
Instead, phase and equilibrium diagrams will be used in a binary
system to demonstrate and reinforce the concept of equilibrium.

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One-Stage Flash

Figure 7 shows an equilibrium separation. A propane and n-

butane vapor mixture from a distillation column is cooled to
120ºF at 100 psia . The vapor and liquid are separated and the
vapor is condensed and collected in a second drum.
Calculations are done to find the composition of the liquid in the
two drums and the minimum temperature required to condense
the vapor leaving the first drum. For simplicity, assume that the
entire system is at 100ºpsia. The system is a one-stage flash.
The second drum merely collects the condensed liquid. It is not
an equilibrium stage.

Figure 7. One-Stage Flash

The required compositions can be determined using a phase

diagram for the propane/ n-butane system at 100 psia. The
liquid in the flash drum is represented by a point (T, x1) =
(120ºF, 0.19) on the bubble point curve of the phase diagram at
120ºF (Figure 8). Similarly, the vapor is represented by a point
on the dew point curve at 120ºF (T, y1) = (120ºF, 0.4). Thus, the
propane mole fractions in the liquid and vapor phases are 0.19
and 0.4, respectively.
The minimum cooling required to condense the vapor leaving
the first drum corresponds to its bubble point temperature. This
is the maximum temperature at which all vapor from the first
drum can be condensed. Since the vapor from the first drum
and the liquid from the second drum have the same
composition, the bubble point can be located by drawing a
vertical line between the dew point and the bubble point curves.
The temperature obtained, 96ºF, is the minimum temperature
required to condense the vapor from the first drum.

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Figure 8. One-Stage Flash on a Phase Diagram

The equilibrium in the flash drum is represented in the

equilibrium diagram by a point, y1, x1 = 0.4, 0.19 (Figure 9).
Condensation in the second drum is represented by a horizontal
line from y1, x1, to the reference line, y1, x'1, where x'1 is the
mole fraction of the liquid of the first drum, which is equal to y1.
The equilibrium diagram does not provide temperature
information. Therefore, it cannot be used to determine the
equilibrium concentrations or the temperature of the second
drum. However, if the composition of one phase is known, it can
be used to determine the composition of the other phase.

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Figure 9. One-Stage Flash on an Equilibrium Diagram

Representation of Petroleum with Pseudo-Components

Computer simulations of distillation columns break the

hydrocarbon streams fractionated into their constituents.
Generally, hydrocarbons with five or six carbons are identified
as individual components. Hydrocarbons with more than five or
six components are represented by narrow fractions. The
narrow fractions are defined by their volume average boiling
point (VABP) and their average gravity. In other words,
components boiling within certain ranges are represented in the
simulation as one component. Such a component is called a
pseudo-component.
Figure 10 illustrates the division of a wide petroleum fraction into
11 pseudo-components. The fraction can be divided into
pseudo-components of equal volume or equal boiling range.
Alternatively, there can be an increased number of components
in the region where the distillation column will split the products.
In this case, the engineer has used few components in the
middle of the curve where the curve is flat. This provides an

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accurate description of the material. However, if this material is

to be divided into fractions, more components in the region of
the division should be included.
PROCESS/PRO II and HYSIM offer a variety of options for
representing petroleum fractions and determining their pseudo-
components.

Figure 10. Pseudo-Component Breakdown

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DETERMINATION OF KEY DISTILLATION PROCESS PARAMETERS

Conventional Distillation Column

Distillation is the separation of the constituents of a liquid
mixture by partial vaporization of the mixture, followed by
separate recovery of the vapor and liquid. The more volatile
(light) constituents of the mixture are obtained in increased
concentration in the vapor, while the less volatile (heavy)
constituents are components concentrated in the liquid residue,
also called the bottoms. The vapor is most frequently
condensed by cooling and is called the distillate or overhead
product. In most petroleum processing plants, the terms
distillation and fractionation are used interchangeably.
Figure 11 shows a conventional distillation column that has one
feed stream and two product streams. The section above the
feed is the rectifying or enriching section. In the rectifying
section, the concentrations of the light components increase
toward the top of the tower; that is, the light product is enriched.
The section of the column below the feed is the stripping
section. Here the light components are stripped out of the liquid
as it descends the column. In binary distillation, the feed
contains only two components.
Frequently in multi-component distillation, a light component
that must be recovered in the distillate is also present in the
residue in important amounts, while components lighter than
this component are present only in small amounts. This
component is called the light key. In the case of a depropanizer,
for example, the light key is the propane, with a concentration in
the residue of 0.7%.
Similarly, a heavy component present in the distillate in
important amounts is called the heavy key. If more than one of
the heavy components is present in the distillate in important
amounts, then the more volatile component is the heavy key. In
a depropanizer, where both the isobutane and the n-butane are
present in the distillate in important amounts, the heavy key is
the isobutane. Key components are used in shortcut distillation
calculations and in some tray and packing efficiency
calculations. Frequently product specifications are based on the
concentration of key components, for example, 0.7% maximum
propane in the depropanizer bottoms.

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Rate Mole Fraction Enthalpy Heat Duty

Bottoms B XB HB -

Distillate D XD HD -

Feed F XF HF -

Reflux R XD ND -

Condenser - - - QC

Reboiler - - - QR

Figure 11. Conventional Distillation Column

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Reflux

Part of the liquid condensed from the vapor leaving the column
top is returned to the column as reflux. The reflux provides the
liquid for the contact with the vapor in the rectifying section of
the column. The important role of reflux in distillation will be
discussed in later sections.
The proper definition of reflux ratio is the mole ratio of reflux to
total distillate, liquid plus vapor. Often, the ratio of reflux to liquid
distillate is also called reflux ratio. The mole ratio of liquid to
vapor between the column stages is called internal reflux ratio.

Major Equipment
The column or tower is the main piece of equipment in a
distillation unit. It contains vapor-liquid contacting devices, trays
in most cases, packing less frequently. Typical major auxiliary
equipment of a column includes the condenser, the reflux drum
(condenser separator), and the reboiler. The condenser
condenses the reflux and the part of the distillate that is
removed as liquid. When all the overhead vapor from the
column is condensed, it is called a total condenser. The reflux
drum separates any vapor distillate from the liquid and provides
surge capacity for the reflux and the liquid distillate.
The reboiler vaporizes part of the liquid leaving the bottom tray
of the column. This vapor, playing a role similar to that of the
reflux, contacts the descending liquid and strips the lighter
components. The heat source of the reboiler may be steam, a
process fluid, or a fired furnace.

McCabe-Thiele Diagram
The graphical representation of a binary system (two
components) distillation column in x,y mole fraction axes
through its operating and equilibrium lines is called a McCabe-
Thiele diagram. Figure 12 is a McCabe-Thiele diagram for an
eight-stage distillation column with one feed, a total condenser,
and a kettle type reboiler.
The slopes of the rectifying (top) and stripping section operating
lines are equal to the internal reflux ratios, L/V and L'/V', in
these sections. The intersections of the operating lines with the
x = y line indicate the distillate and bottoms compositions. The
column theoretical stages are indicated by steps between the

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equilibrium and operating lines. The points of the steps that are
on the equilibrium line indicate the vapor, y, and liquid, x,
compositions that are at equilibrium on each column stage. The
points that are on the operating lines indicate the vapor and
liquid compositions between stages.

The McCabe-Thiele technique can be used to determine the

required stages for a given separation or the expected product
qualities for a given column. Compared to available numerical
techniques, it is too slow and not sufficiently accurate. However,
it is an excellent tool for demonstrating some of the principles of
distillation. One of the simplifying assumptions, for example, is
that the vapor and liquid molar rates, V, L, in each section of the
column are constant. Additional information on the McCabe-
Thiele diagram can be found in ChE 205.03 and distillation
textbooks.

Figure 12. McCabe- Thiele Diagram

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Effect of Reflux on Required Stages

Figure 13 illustrates the effect of reflux on the required stages to

achieve a specified separation (xB,xD). We observe that as the
operating line slope is reduced (reflux decreased), the change in
composition from stage to stage is reduced and the required
stages increase.

Figure 13. Effect of Reflux on Required Stages

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Minimum Reflux

If the reflux is reduced to the point that the operating lines

intersect each other at the equilibrium line, the required number
of stages becomes infinite (Figure 14). This reflux is called the
minimum reflux.
The minimum reflux does not represent a practical operation.
However, it can be used to compare the difficulty of separation
for various product specifications. Also, the actual reflux is often
expressed in terms of minimum reflux; for example, a tower may
normally operate at 1.1 × minimum reflux.

Figure 14. Minimum Reflux

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Total Reflux - Minimum Stages

Total reflux represents an operation where the feed and product

streams of a column operating at steady state are
simultaneously blocked. The reflux is adjusted to maintain the
level in the accumulator, and the reboiler and condenser loads
are adjusted to maintain the enthalpy balance. When the
column regains the steady state, it is operating at total reflux. All
the vapor entering the condenser is condensed and returned to
the column as reflux, and all the liquid entering the reboiler is
vaporized and returned to the column (Figure 15).

Figure 15. Total Reflux - Minimum Stages

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There is one operating line for the entire column passing

through xB and xD, and coinciding with the y = x diagonal. The
operating line has the maximum slope possible, and the
corresponding stages are the minimum number of stages that
can provide the required separation. A column with the
minimum number of stages, similar to a column with minimum
reflux, does not represent a practical operation. However, it
represents the difficulty of separation, and it is used in shortcut
calculations.
The concept of total reflux can also be applied to a steady-state
operation with a feed and a bottoms product. In this case, the
bottoms composition is the feed composition.

Stages Versus Reflux

We have seen in the previous pages that there is a relationship
between reflux and the required number of stages in a column.
This relationship is illustrated in Figure 16. Operation of a
distillation column near the minimum reflux or minimum number
of stages is not stable or economical. When a column operates,
for example, near the minimum number of stages, a reduction in
tray efficiency, due to tray damage, can be compensated for
only by a very large reflux increase. Normal design is some
where in the knee of the curve.

Figure 16. Stages Versus Reflux

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Simulations with PROCESS/PRO II

PROCESS/PRO II is a flowsheet simulator. The main unit

operations used for the simulation of distillation columns are:
• Shortcut distillation.
• Rigorous distillation.
• Flash.
• Mixer/splitter.
• Exchanger.
The Multi-variable Controller and Flowsheet Optimizer options of
PROCESS/PRO II can be used to accomplish complex
objectives. The PROCESS/PRO II input manual provides further
information on simulations.
Figure 17 is a printout of a column profile from a Process
simulation of a depropanizer. We observe the following:
• The tray temperature is lowest at the tower top and highest
at the bottom of the tower. The gradual temperature change
reflects the change in tray composition and, to a lesser
extent, the change in pressure. The top tray represents the
condenser.
• The pressure changes very little from tray to tray. There is a
5 psi drop between the top tray and the condenser.
• The liquid loadings increase sharply below the feed. The
feed is 100% liquid as indicated by capital L to the right of
the feed rate.
• The vapor profile changes gradually as the composition and
thermal condition changes. There is a drop near the feed
indicating subcooled feed.

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Figure 17. Simulation Tower Profile

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Steps in Column Design

Below is a list of steps typically followed in designing a

distillation column.
• The basis is defined.
• The tower operating conditions, including pressure and
temperature are determined.
• The simulation components are defined.
• The distillation calculations are performed.
• The tray efficiencies are estimated.
• The hardware is designed.
The following section focuses on two important operating
conditions of a column: pressure and temperature. A section on
tray efficiencies follows. Hardware design is discussed in
ChE 104.02.

Column Operating Pressure, Temperature

Column pressure is normally selected so that the reflux and
distillate can be condensed by the available coolant. Allowances
must be made for the condenser approach temperature,
fluctuations in product rate and composition, and the need for
subcooling of reflux or liquid distillate. Further considerations
are:
• Use of a partial or a total overhead condenser.
• The limiting design pressure should correspond to the
highest design temperature level of the coolant, such as
summer conditions for ambient air. Below is a list of typical
approach temperatures of the cooling medium and the heat-
source for different reboiler and condenser services:
• Temperature approach to the cooling medium, ºF:
Refrigeration 5 - 20
Sea water 10 - 25
Air 15 - 30

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• Temperature approach to the heat source, ºF:

Process fluid 15 - 35
Steam 15 - 100
Hot oil 35 - 100
• Refrigeration of the overhead should be avoided if possible.
When refrigeration cannot be avoided, economics will
normally dictate that the refrigerant temperature be as high
as possible in order to minimize refrigerant compression
costs.
• Lower pressure increases the relative volatility and improves
the ease of separation. Therefore, the total stages or reflux
ratio can be reduced while still meeting a given design
specification.
• Lower pressures will give a somewhat larger diameter tower
(lower vapor density, higher actual volume) and possibly a
thinner tower shell. The exception may be towers
fractionating close-boiling components, where the
improvement in relative volatility and reduced reflux
requirements more than compensate for the reduced
density.
• With some heat-sensitive systems, severe fouling conditions
in the reboiler and lower tower stages can be avoided by
reducing the tower pressure to reduce bulk liquid
temperatures. Lower reboiler temperatures minimize
degradation of gas treating amines such as DGA.

Tray Efficiency
The overall efficiency, EO, is a measure of the effectiveness of
an entire column or section of a column. This efficiency is the
one most often used by designers in determining the number of
actual trays to provide. EO is simply the total number of
calculated theoretical trays required, divided by the total number
of actual trays required for the separation.
Approximate tray efficiencies can be predicted using the fluidity
method in Maxwell's Data Book on Hydrocarbons or the
viscosity-volatility method developed by F. J. Lockhart and
C. W. Leggett (E. J. Henley, J. D. Seader, Equilibrium-Stage
Separation Operations in Chemical Engineering, 1958, John
Wiley and Sons, Inc.). Efficiencies from these two sources do

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not take into account the tray geometry and the effects of
fouling. As a result, the approximate efficiencies may deviate
considerably from the actual efficiencies and should not be used
for design.

More accurate tray efficiencies can be obtained using unit

operating data. The unit can be simulated and the simulation-
actual heat and material balance and product quality matched
using the number of theoretical trays as a variable. Overall
efficiency is the ratio of theoretical to actual trays.
Vendors also use their own data and predictive methods to
estimate tray efficiencies. Typical tray efficiencies for some of
the Saudi Aramco units are listed in Table 1.

Table 1. Typical Overall Efficiencies, %

Above Feed Below Feed

Crude Stabilizer 25

Atmospheric Crude Unit

Top 70-80

Bottom (above Flash Zone) 50-60

NGL Fractionation 90 65-85

(DeC3…DeC5 (Low-end if
very heavy
components
are present)

Reformer Feed Stripper 85 75

Kero/Diesel Hydrotreater

Stripper 80-90 30-35

H2/Amine Absorber 10-25

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DETERMINATION OF REQUIRED OPERATIONAL CHANGES IN SAUDI

ARAMCO DISTILLATION PROCESSES

Crude Stabilization and Sweetening

Fluids flowing from oil and gas wells generally include a gas
phase and a liquid phase in equilibrium. The liquid phase
consists primarily of crude oil, water, and dissolved salts. Water
and salts are removed in desalters. The crude oil is sent to
stabilization plants. The objective of crude stabilization is to
remove light hydrocarbons and hydrogen sulfide (H2S) from the
crude in order to achieve an acceptable level of volatility for
storage and transportation and an acceptable concentration of
H2S. Generally, it is not desirable to remove from the crude any
hydrocarbons beyond the required minimum.
There are two obvious advantages to retaining the lighter
hydrocarbons in the stabilized crude. First, liquids can be stored
and transported to the user more economically than gas.
Second, retention of the lighter hydrocarbons in the liquid phase
is especially important when the separated gas has no market
value and is being flared or compressed and re-injected.
Conversely, retention of too many light ends can cause
problems. The quantity of light ends that must be removed from
the crude oil is controlled by the composition of the oil, the
ambient air temperature, the method of transportation, and
economic considerations. Refineries and tankers require crude
oil to meet maximum vapor pressure specifications. If the crude
contains H2S, the H2S concentration must be lowered to reduce
toxicity and corrosion and to meet sales specifications.
For Saudi Aramco units, the H2S content is generally the
controlling specification. The H2S content specification in the
stabilized crude leaving the units generally varies between 1
and 60 parts per million (ppm), and it is typically 10 ppm. At
such low H2S contents, volatility specifications, for example, 13
psia Reid vapor pressure (RVP), are normally met. The
maximum allowable H2S content in the stabilized crude is 70
ppm, however, the Ras Tanura refinery requires a lower H2S
content when maximizing NGL.

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Figure 18 illustrates a typical stabilization unit, the Abqaiq Crude

Stabilizer No. 17. Desalted crude at ambient temperature is fed
to the top tray. A set of reboilers provides the driving force for
stripping light material and H2S. Live steam is also injected in
the reboiler return line. The bottoms temperature is a function of
the quality of stripping and is a good indication of the H2S
content in the bottoms.
The gas from the column overhead is sent to compression and
recovery of light hydrocarbons. The bottoms, the stabilized
crude, is cooled and stored or shipped.

Figure 18. Abqaiq Crude Stabilizer No. 17

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Condensate Stripping

The objective of condensate stripping is to remove ethane from

condensed stabilizer overhead vapor. Condensate stripping is a
relatively simple fractionation process (Figure 19). The feed,
which is stabilizer gas that has been compressed and
condensed, is sent to the top tray of the tower. The tower is
reboiled; however, there is no overhead condensing system or
external reflux. The feed condenses, providing the internal reflux
required for fractionation.
The key product specification is the bottoms C2 content, which
is controlled by the reboiler duty. Due to the lack of overhead
condenser and external reflux, there is no control on the content
of C3+ material in the tower overhead product. The tower
pressure is typically determined by downstream transportation
and processing requirements for the overhead gas.
The overhead product is sent to gas plants for treating and
separation of the light hydrocarbons. The bottoms product is
sent to NGL recovery units.

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Figure 19. Abqaiq Condensate Stripper

Crude Distillation
Crude petroleum (oil) is a complex mixture of an extremely large
number of hydrocarbons, from gases like methane and ethane,
to molecules having more than 70 carbon atoms and boiling
points above 1000ºF. Because of this wide range of
components, it is rarely used directly as it is produced from the
field. However, it can be refined and further processed into any
number of products according to boiling range. The objective of
crude distillation is to separate the crude into fractions according
to boiling range. These fractions may be products for sale or
may be feedstocks for other refining or processing units.

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In most refineries, the distillation of crude is carried out in two

stages (Figure 20). The crude oil is first heated and then fed to a
fractionating tower that operates at a pressure slightly above
atmospheric pressure. This tower is usually called the
atmospheric tower or column. It yields several distillate products
and a bottoms product, which is the residual liquid material that
could not be vaporized under the conditions of temperature and
pressure existing in the atmospheric tower. This bottoms liquid
is then reheated to the maximum allowable temperature, usually
higher than the maximum temperature allowed for the feed to
the atmospheric tower, and fed to a fractionating tower
operating at subatmospheric pressure. This tower is usually
called the vacuum tower or vacuum column.

Figure 20. Plant 15 Crude Unit - Ras Tanura Refinery

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Atmospheric Unit

Figure 21 is a simplified process flow diagram of the Ras

Tanura Plant 15 Crude Unit. Crude from storage is mixed with
water for desalting and heated in a series of exchangers. The
exchangers typically utilize heat available in the crude unit. After
the crude reaches about 250-290ºF, it enters the desalter.
Downstream of the desalter, the crude is flashed and the vapors
fed to the flash zone of the atmospheric unit. The purpose of the
flash drum is to remove light components that may vaporize
upstream of the furnace and cause high pressure drop and
maldistribution between the furnace passes. The liquid from the
flash drum is pumped through a series of exchangers and fed to
the furnace.
Following is a discussion of the main features of an atmospheric
crude unit. Figure 21 shows the Ras Tanura Plant 15
Atmospheric Crude Unit in greater detail.
Furnace - Generally, the furnace heats the crude to the highest
temperature allowable for the crude being processed. The
furnace transfer temperature, also known as the coil outlet
temperature (COT), is the most significant parameter for
achieving the desired total yield of distillates. The COT is limited
by the effect of temperature on cracking and the effects of the
cracked products on product quality. Other limitations on COT
are coking of the furnace tubes or tower internals and the
maximum permissible furnace tube metal temperature (TMT).
The crude leaves the furnace partially vaporized and enters the
atmospheric column at the flash zone.
Bottoms and Sidestream Stripping – Liquid from the flashed
feed plus the overflash enters the section of the column below
the flash zone, the stripping section. Superheated steam
injected below the bottom tray removes light components from
the bottoms (vacuum unit feed). Excessive amounts of light
components in the feed to the vacuum unit increase the furnace
duty and pressure drop and increase the difficulty of maintaining
vacuum.
Crude unit sidestream stripper towers are used primarily to
achieve certain product specification targets such as flash point
and front-end distillation requirements. These specifications are
achieved through a combination of stripping steam and the
number of theoretical stages present in the stripper.

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Figure 21. Plant 15 Atmospheric Crude Unit - Ras Tanura Refinery

Pumparounds – Liquid from fractionator trays is pumped

through exchangers and returned to a higher tray. The cold
liquid contacts and condenses the rising vapor. The purpose of
the pumparounds (PA) is to:
• Reduce the rate of vapor rising up the tower and thus reduce
the required tower diameter.
• Recover heat at higher temperature levels relative to the
overhead condenser.
The trays between the pumparound draw and return are heat
transfer trays, and their contribution to the separation of
components is minimal due to the backmixing of liquid.

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Vacuum Unit

Reduced crude from the atmospheric unit bottom is fed to the

vacuum unit, operated under vacuum in order to recover
additional distillate liquids (Figure 22). The reduced crude is
heated in the vacuum furnace, which provides the heat required
to partially vaporize the feed. Steam is added in the furnace to
limit process stream temperatures and to reduce the residence
time, thereby minimizing coking and cracking. Typically two
sidestream products are taken: light (LVGO) and heavy (HVGO)
vacuum gas oils (in addition to a small overhead distillate or slop
stream). The sidestreams are used for fuel oil or as cracker feed
and are condensed via pumparounds, which recover heat for
reuse either for feed to cat crackers or other conversion units.
Ras Tanura Vacuum Unit 15 has three sidestreams. The top
(S/C 6) is a diesel blending component (recovery of light
components, such as diesel, in the vacuum units is not typical),
while the heavier sidestreams and bottoms go to fuel oil. The
LVGO (S/C 7) and HVGO (S/C 8) are typically blended with
vacuum resid to produce a variety of fuel oil grades. In many
refineries, the LVGO and HVGO are fed to fluid catalytic
cracking units and converted to lighter products.
Because fractionation between the sidestreams is usually not
critical, fractionation sections are not usually provided; only the
amount of trays or packing required for heat transfer is provided
in the pumparounds. Packing is often used instead of trays to
minimize the tower pressure drop and maximize the sidestream
or gas oil recovery. In the Ras Tanura Vacuum Unit 15, there
are four trays between the top (S/C 6) and the second (S/C 7)
sidestreams to provide the fractionation needed for the diesel oil
(S/C 6) blending component. A wash section, consisting of a
Glitsch grid or a similar structured packing, is located above the
flash zone to reduce resid entrainment from the flash zone. A
bottom steam stripping section recovers distillate product from
the feed flash liquid and the wash section liquid or overflash.
The tower vacuum is maintained by steam jet ejectors either
alone or in combination with mechanical vacuum pumps.

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Figure 22. Plant 15 Vacuum Crude Unit - Ras Tanura Refinery

Following are highlights of features particular to vacuum units.

Furnace Coil Outlet Temperature (COT) and Coil Steam –
The coil outlet temperature is the most significant parameter for
achieving the desired vacuum pipestill cutpoint. A COT as high
as 780°F has been achieved in practice, but 740 to 750°F is
more the norm in actual operations. The maximum permissible
COT is limited by three factors:
• Cracking, which yields noncondensible gases that potentially
could overload the overhead vacuum system.
• Coking of the furnace tubes or the tower internals, which
limits the run length.

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• Maximum permissible tube metal temperature.

Steam is added to the furnace radiant coil to suppress coking by
increasing velocity.
Bottoms Quenching – The bottoms is quenched to reduce the
noncondensible gas produced through cracking, as well as to
reduce coking of the resid. If not quenched, the bottoms would
produce considerably more gas due to its high temperature and
long residence time.
Wash Zone – The primary purpose of the fresh wash and
overflash is to wash out the entrained pitch that is removed from
the feed flash vapor in the wash section. Fractionation
requirements in the wash section are minimal. Thus, the
minimum fresh wash and overflash rates are set by the
minimum liquid rate required to wet the wash zone internals.
Steam Jet Ejectors – Due to their principle of operation, steam
jet ejectors are virtually intolerant of increases in vapor load and
decreases in motive steam pressure and rate compared to their
design values. Common causes of difficulties in obtaining the
desired vacuum are associated with:
• Insufficient heat removal in the condensers.
• Low motive steam pressure or rate.
• Excessive steam or air leakage into the system.
• Ejector cycling occurring when the vapor load to the ejector
is reduced significantly below design.

NGL Fractionation

Gas plants such as Uthmaniyah and Shedgum receive GOSP

gas (for example, Uthmaniyah GOSP), offgas from crude
stabilization units (for example, Abqaiq), and nonassociated gas
KHUFF. The objective of NGL fractionation is to separate the
raw hydrocarbon mixture into individual products suitable for
sale, injection into crude, or downstream processing. Refineries
such as Ras Tanura also produce gas and light hydrocarbons,
often referred to as light-ends, by removing the light fractions of
the crude or by conversion (chemical reaction) in units such as
reformers and fluid catalytic crackers. NGL fractionation
separates these light-ends components.

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The raw stream to be fractionated normally contains the

hydrocarbon components shown below in order of decreasing
volatility (ease of vaporization) as denoted by increased boiling
points in the following table:

Table 2. Hydrocarbon Components

Light Hydrocarbon Atmospheric Pressure Vapor Pressure at

Boiling Point, ºF 120ºF, psia

Methane - 259 Critical temp. 116ºF

Ethane - 127 Critical temp. 90ºF

Propane - 44 225

iso-Butane 11 96

n-Butane 31 70

iso-Pentane 82 29

n-Pentane 97 22

Hexanes and Heavier 122-250 < 15

If fractionation is to be feasible, the top product and bottom

product from a fractionator must have different boiling points.
The difference in boiling points between the top and bottom
products indicates the degree of difficulty in separating these
products in a single fractionator. The smaller the boiling point
difference, the more difficult the separation; that is, the
separation will require more trays (tower size), more reflux
(pump size, tower size), and additional reboiler heat (reboiler
size) to perform the job. This means that higher investment will
be required and greater operating costs will be incurred to distill
the products.
The column pressures are generally set such that the
condensing temperature of the overhead products are the
same. This permits the use of a single cooled utility system.
This does not apply to demethanizers and deethanizers, which

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normally require refrigeration. The last column of the table on

the previous page lists the vapor pressures of light
hydrocarbons at 120ºF. These are the pressures required to
condense the hydrocarbons in a tower overhead condenser and
separator if we are to use a cooling medium, such as air in air
fin coolers, that can cool to 120ºF. For instance, propane
product separation may require a 225-psia operating pressure
using air coolers, while isobutane product separation may need
only a 100-psia pressure. The actual pressure of the tower may
be somewhat different since the overhead product is not pure.
It generally contains small amounts of lighter and heavier
components that result in a different vapor pressure.
A single fractionator tower generally removes only one single
finished product stream from the raw feed mixture. A common
order of fractionation in gas plants is to remove the ethane, the
propane, the butanes as a mix, and then the pentanes.
Separation of the butane mix into finished normal and isobutane
products is then done last in another tower, the deisobutanizer.
Pentanes may be separated similarly. Other schemes remove
more components in one tower (for example, propane/butanes)
and fractionate them in a separate tower (for example, a C3/C4
splitter).

NGL Fractionators

The degree of fractionation required and the number of products

vary from plant to plant and are determined by factors such as
product outlets, customer needs, and plant location. As a result,
there are several types of NGL fractionation towers, and many
fractionation sequences are found in refineries and gas plants.
Table 3 summarizes the main types of NGL fractionators.

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Table 3. NGL (Light-Ends) Fractionators

Fractionator Products
Type of Fractionator Feed Stream Overhead Bottom
Product Product

Demethanizer (DeC1) GOSP, Methane C2+

Condensed Stripout
Deethanizer (DeC2) DeC1 Bottoms, Ethane, C2- C3+
Condensed Stripout
Depropanizer (DeC3) DeC2 Bottoms, Propane C4+
Condensed Stripper
Bottoms
Debutanizer (DeC4) DeC3 Bottoms Butane C5+
Debutanizer (DeC4) DeC2 Bottoms Propane + Butane C5+
C3/C4 Splitter DeC4 Overhead Propane Butane
Deisobutanizer (DeiC4) DeC4 Overhead Iso-Butane n-Butane
Depentanizer (DeC5) DeC4 Bottoms Pentane C6+

Depentanizer (DeC5) DeC3 Bottoms Pentane + Butane C6+

C4/C5 Splitter DeC5 Overhead Butane Pentane

Figure 23 is a simplified process flow diagram of Yanbu

deethanizer gas plant. Figure 26 and 27 are simplified process
flow diagrams for the depropanizer, debutanizer, rerun, and
butane/pentane splitter of the Ras Tanura Plant 10 NGL
fractionation. The figures illustrate the flow sequence, operating
conditions, and key product qualities.

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Figure 23. YANBU Fractionation and Treating: Deethanizer

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Figure 24. Ras Tanura NGL Fractionation Plant 10

Figure 25. Ras Tanura Plant 10 - NGL Fractionation

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Hydrocracker Fractionator

Figure 26 is a Process Flow Schematic of the Ras Tanura

Hydrocracker fractionation. The Ras Tanura Hydrocracker has a
hot and cold flash separators to remove hydrogen and C2- from
the reactor effluent. The liquid reactor effluent from the cold
flash separator is fed to a Debutanizer to remove LPG. The LPG
is Amine treated to remove traces of H2S and water washed to
remove traces of amine. The Debutanizer is reboiled by a fired
heater. The bottoms form the Debutanizer is fed through a fired
heater to the main fractionator where the products of naphtha
(C5- 400), kerosene (300-500), and diesel (400 - 700) are
separated from the bottoms. The bottoms are recycled to the
second stage Hydrocracker reactor or to storage. The naphtha
stream is split into light naphtha (C5-C6) and heavy naphtha (C6-
400). The light naphtha is caustic washed to remove traces of
sulfur compounds. The heavy naphtha which will be feed to a
Reformer is treated in a sulfur guard reactor to remove traces of
sulfur compounds. The kerosene and diesel streams have
steam strippers to control light ends in these products. Low
pressure (60 psi) superheated steam is used in the strippers.
The kerosene can be added to the diesel steam to maximize
diesel production. The kerosene cut can be recycled to the
kerosene stripper to maintain stable stripper operation when the
Hydrocracker is operating in maximum naphtha mode.

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Figure 26. Hydrocracker Fractionation

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The Hydrocracker fractionation system is designed to operate

over a wide range of operation. The product yields change over
the run as the catalyst ages as shown in Table 4. The
Hydrocracker can also be operated to maximize either naphtha
or diesel which allows for a large shift in the refinery production
as shown in Figure 3. Designing a fractionation system with this
much flexibility is a challenge.

Table 4. Hydrocracker Production

Maximum Diesel Maximum Naphtha

Start Run End Run Start Run End Run

BPSD BPSD BPSD BPSD
LPG 36 2,020 Aramco to Supply
Lt Naphtha (???) 4,554 6,506 “
Hvy Naphtha 10,493 10,841 “
Diesel 34,013 30,294 “
Bolts 32,820 33,021 “

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Summary
• Raoult's law is a simple VLE relationship that can be applied
to low-pressure systems with chemically similar, close-boiling
components. The use of fugacities in place of vapor pressure in
Raoult's law extends its applicability to higher pressures.

• Equilibrium diagrams and phase diagrams demonstrate

differences in the composition of the vapor and liquid phases
and the bubble point and dew point concepts. They can be
used in equilibrium flash separation calculations.

• Distillation concepts and process characteristics, which for

ease of comprehension were introduced under binary
distillation, apply to multicomponent distillation. These concepts
include minimum reflux, minimum stages, and the tradeoff
between reflux and number of stages.

• The key steps in column design are:

- Definition of the basis, including feed, specs, and
composition.
- Determination of key operating conditions, such as pressure.
- Definition of simulation components and performance of
shortcut and rigorous simulations.
- Estimation of tray efficiencies and actual column stages.
- Design of hardware, which is covered in later modules.

• The objectives and operation of the main Saudi Aramco

distillation processes were reviewed.

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GLOSSARY

Blowdown Water removed from a cooling system used to control the

concentration of impurities in the recirculated water.
Caustic Embrittlement Intercrystalline cracking of metal due to the presence of
caustic in hot water when the metal is under stress.
Condensers Mechanical devices used to convert a substance from the
vapor state to the liquid state.
Cooling Towers Devices, usually exterior mounted, that remove heat from
a liquid by passing it through a series of coils that are
exposed to a rapidly moving air current.
Electrochemical Pertaining to the relationship of electricity to chemical
changes, and the mutual conversion of chemical and
electrical energy.
Evaporators Equipment that supplies heat to a liquid to convert it to the
vapor state.
Filtration The removal of suspended solids from water.
Ions Electrically charged particles formed in water when a salt
dissociates into positive and negative particles.
Macrobiological Organisms visible without magnification.
Makeup Water Water added to a cooling water system or boiler system to
replace water losses caused by evaporation, drift, and
blowdown.
Microbiological Organisms visible under magnification.
Organism Any living thing or being.
Precipitation The formation of insoluble salts from soluble salts due to a
chemical change.
Raw Water Water that has not been treated.
Slime Insoluble organic matter of a viscous gelatinous
composition normally resulting from microbiological growth
that traps other suspended impurities in the cooling water.

Suspended Solids Insoluble matter in water larger than 0.45 microns in size.

Turbidity A measurement of the clarity of water using light

transmission. It is normally expressed as nephlometric
turbidity units (NTU).

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ADDENDUM

References
1. DeWiest, Roger J. M., Geohydrology, Figure 2-1, John Wiley & Sons, Inc. 1965;
Frank N. Kemmer, Editor, McGraw Hill Book Company, 1988

2. Global Ecology Handbook, The, Figure 9-2, Walter H. Corson, Editor, Beacon
Press, 1990

3. The Nalco Water Handbook, Table 2.4, Second Edition, Nalco Chemical Company,
Frank N. Kemmer, Editor, McGraw Hill Book Company, 1988

4. The Nalco Water Handbook, Figure 1.4, Second Edition, Nalco Chemical
Company, Frank N. Kemmer, Editor, McGraw Hill Book Company, 1988

5. Perry’s Handbook, Figure 23-2, Sixth Edition; Perry, Robert H. and Green, Don,
Editors; McGraw Hill Book Company, 1984

6. The Nalco Water Handbook, Table 35.1B & 35.2, Second Edition, Nalco Chemical
Company, Frank N. Kemmer, Editor, McGraw Hill Book Company, 1988newidth
SD stroke}

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