1 Hydrocarbon Dew Point Control

Hydrocarbon Dew Point Control

Introduction

Dew point is defined as the temperature at which vapor begins to condense. We see it in action
every foggy morning. Air is cooled to its water dew point and the water starts condensing and
collects into small droplets. We also see it demonstrated by a cold glass "sweating" on a humid
day. The cold glass lowers the air temperature below the water dew point temperature and the
water condenses on the sides of the cold glass. Water dew point is relatively simple and easy to
predict since it is a single component system.

Hydrocarbon dew point (HCDP) is similar to the water dew point issue, except that we have a
multi-component system. Natural gas typically contains many liquid hydrocarbon components
with the heavier components found in smaller amounts than the lighter gaseous ends. It is the
heaviest weight components that first condense and define the hydrocarbon dew point
temperature of the gas. The dew point temperature also moves in relation to pressure.

Why Control Hydrocarbon Dew Point?

To answer this question lets take a general example. A production separator separating oil from
gas operates at vapor-liquid equilibrium. Therefore, the gas leaving the separator is in
equilibrium with the oil. In other words, the gas leaving the separator is at its hydrocarbon dew
point which equals the separator operating temperature. If the separator is operating at 100 F,
then the gas has a 100 F dew point at separator pressure. As the gas leaves the separator and
cools flowing through the piping system, liquids condense and the dew point decreases as the
heavy ends condense.

The gas transportation companies have come to the realization that managing hydrocarbon dew
point reduces system liabilities, opens up new gas markets and generates operating revenue. By
managing hydrocarbon dew point, hydrocarbon condensation can be prevented in cold spots
under rivers and lakes where the liquids collect in the low areas and then often move as a slug
through the system, over pressuring the pipe, and overpowering liquid handling facilities,
flowing into compressors and end user sales points. Liquids in burners and pilots can cause fire
and explosion hazards. Removing pipeline liquids helps prevent pipe corrosion in the low areas
where water is trapped under the hydrocarbon liquid layer and slowly destroys the pipe integrity.
Proper managing of gas dew point can also prevent liquids from forming as the gas cools while
1 flowing through pressure reduction stations that feed end user supply systems. Controlling dew
point can also qualify the pipeline to market gas to high efficiency gas turbine end users that
require a dry and consistent quality fuel.

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 2 Hydrocarbon Dew Point Control

Specifications for HCDP

Pipelines use two main methods to specify contractual natural gas hydrocarbon dew points.

1. Limit on C5+ or C6+ components by analyzing for:

• GPM (gallons of liquid per thousand SCF)

• Mole %

2. Specifying an actual HCDP by:

• Setting a hydrocarbon dew point temperature maximum at any operating pressure

• Setting a maximum cricondentherm* hydrocarbon dew point

* The cricondentherm temperature is the highest dew point temperature seen on a liquid-vapor
curve for a specific gas composition over a range of pressure

Techniques Available to Control HCDP

HCDP is most commonly controlled by following techniques

a) J-T Expansion valve

b) Mechanical Refrigeration
c) Turbo Expander
d) Solid Membrane

A) J-T Expansion Valve

Processes which use the cooling effect of the expansion of a gas across a valve or choke are
sometimes called LTS (Low Temperature Separation) or LTX (Low Temperature Extraction)
units. A more common name for this type of process is a J-T plant, where the J-T refers to Joule-
Thomson. This is because the thermodynamic principle that explains the expansion of gas across
a valve is called a "Joule-Thomson expansion," and is named for the scientists who first
explained it.

Process Introduction

2 Figure below presents a process flow schematic for a typical Low Temperature Separation. The
full well stream fluid enters a Free Water knockout vessel (separator) where any condensed
liquids are removed. The gas then flows to a heat exchanger where the incoming gas is cooled by

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 3 Hydrocarbon Dew Point Control

the processed gas stream. A pressure drop is then taken across an expansion valve and the gas
temperature is further reduced. The condensed liquids are then separated from the gas stream in a
low temperature separator. The cold gas is routed through the inlet gas-gas heat exchanger to
cool the incoming gas, and then sent on to sales. Since the gas is cooled in the process to an
extent that it passes into the hydrate formation region, Ethylene Glycol is injected upstream of
the low temperature separator (or upstream of the gas-gas heat exchanger, depending on
temperature and pressure levels) in sufficient quantities to depress the hydrate formation
temperature below that of the low temperature separator's temperature. As Ethylene Glycol is a
valuable content, we place a glycol regeneration system. Ethylene Glycol is collected in the boot
of Low Temperature Separator and is sent to EG Regeneration skid as a rich glycol steam. From
here, the lean glycol is pumped to the gas-gas exchanger by the help of lean glycol injection
pump.

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 4 Hydrocarbon Dew Point Control

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 5 Hydrocarbon Dew Point Control

Major Equipments used in J-T Expansion Valve Technique

1. Water Knockout Vessel

With respect to this process, the Water Knockout Vessel is a simple two phase Vertical Separator
which function is to remove the free water entering the vessel with the gas. The removal of water
from gas is required to avoid hydrate formation in any further step of the process. In Low
Temperature Separation, the temperature of the gas drops below the freezing point of water
which can lead to ice formation. To avoid this problem in the process water is removed from gas
before gas is cooled down.

Working and Design

Water Knockout Vessels are usually Vertical

separators and are selected when the gas-
liquid ratio is high or total gas volumes are
low. In Knockout vessels, the fluids enter the
vessel through an inlet diverter whose
primary objectives are to achieve efficient
bulk separation of liquid from the gas and to
improve flow distribution of both phases
through the separator. Liquid removed by the
inlet diverter is directed to the bottom of the
vessel.
The gas moves upward, usually passing through
a mist extractor to remove any small entrained
liquid droplets, and then the vapor phase flows
out of the vessel. Liquid removed by the mist
extractor is coalesced into larger droplets that
then fall through the gas to the liquid reservoir in
the bottom. The ability to handle liquid slugs is
typically obtained by increasing vessel height to
accommodate additional surge volume. Level
control is normally not highly critical and liquid
level can fluctuate several inches without
affecting the separation performance or capacity
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of the vessel.

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 6 Hydrocarbon Dew Point Control

Talking about the vessel sizing, typical knockout vessel L/D ratio is in the range of 2–4.

The gas handling capacity of conventional knockout vessel that employs mist extractors has
normally been calculated from the Souder and Brown equation, using “experience-based” K
factors.

And the required Area of the mist extractor is calculated by

Typical K values for vertical separators from API 12J are presented as;

In qualitative terms, the ranges of K given above may be taken to reflect difficulty of thee
separation conditions, i.e. from non-ideal/difficult to ideal/easy. As indicated in above table, K is
also a function of vessel height. This reflects the fact that a certain minimum distance is required
to establish a relatively uniform velocity profile before the gas reaches the mist extractor.
Theoretically, it is not simply the vessel height that is important with respect to velocity profile,
but the vertical height between the inlet device and the mist extractor. As gas handling capacity
is based on an allowable limit for liquid carryover into the separated gas stream, and the final
liquid removal element is the mist extractor, the mist extractor has a significant influence on the
K value used for separator sizing.

A design that optimizes the inlet feed flow condition and utilizes an efficient inlet device, may
provide enough feed gas pre-conditioning to allow the vessel diameter to be sized equivalent to
the mist extractor. However, traditionally the method typically used has been to “oversize” the
6 vessel diameter, relative to the mist extractor. This is generally done by selecting a separable
droplet size (A droplet size of 150 microns has been typically specified) and size for the vessel
diameter using below Equation.

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 7 Hydrocarbon Dew Point Control

Where,
Dp = droplet size
g = force of gravity
ρl = density of liquid
ρg = density of gas
C` = drag coefficient

C` is determined by following graph

Here,

And
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 8 Hydrocarbon Dew Point Control

2. Gas /Gas Heat Exchanger

The dry gas from separator enters a heat exchanger. Heat exchanger is device used to transfer
heat between two streams which are relatively at different temperatures. In this process, we
usually use shell and tube heat exchangers. As the heat is exchanged between two gas streams so
we call it Gas/Gas heat exchanger.

Working and Design

The gas entering the exchanger is needed to be cooled down; its heat is exchanged with the lean
gas stream coming from the Low Temperature Separated. The gas from KOD enters the one side
exchanger shell and leaves from other side of the exchanger shell with comparatively lower
temperature. The lean gas from separator enters the tube side and leaves with comparatively
higher temperature. As the temperature of the shell side gas is lowered enough to facilitate
hydrate formation, so to avoid this ethylene glycol is injected in the shell. This EG also travels to
the Low Temperature Separator along with the gas and is separated from gas there.

When designing the shell and tube heat exchanger, we usually follow these steps

1. Define the duty: heat-transfer rate, fluid flow-rates, and temperatures.

8 2. Collect together the fluid physical properties required: density, viscosity, thermal
conductivity.

3. Decide on the type of exchanger to be used.

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 9 Hydrocarbon Dew Point Control

4. Select a trial value for the overall coefficient, U.

5. Calculate the mean temperature difference, Tm.

6. Calculate the area required.

7. Decide the exchanger layout.

8. Calculate the individual coefficients.

9. Calculate the overall coefficient and compare with the trial value. If the calculated
value differs significantly from the estimated value, substitute the calculated for the
estimated value and return to step 6.

10. Calculate the exchanger pressure drop; if unsatisfactory return to steps 7 or 4 or 3, in

that order of preference.

11. Optimize the design: repeat steps 4 to 10, as necessary, to determine the cheapest
exchanger that will satisfy the duty. Usually this will be the one with the smallest
area.

The prime objective in the design of a heat exchanger is to determine the surface area required
for the specified duty (rate of heat transfer) using the temperature differences available. This can
be done by using the general equation for heat transfer across a surface is:

Where,
Q= heat transferred per unit time, W
U= the overall heat transfer coefficient, W/m2 0C,
A= heat-transfer area, m2,
Tm = the mean temperature difference, the temperature driving force, 0C.

9
The overall heat transfer coefficient is the reciprocal of the overall resistance to heat transfer,
which is the sum of several individual resistances. For heat exchange across a typical heat

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 10 Hydrocarbon Dew Point Control

exchanger tube the relationship between the overall coefficient and the individual coefficients,
which are the reciprocals of the individual resistances, is given by:

Where,
Uo = the overall coefficient based on the outside area of the tube, W/m2 0C,
ho = outside fluid film coefficient, W/m2 0C,
hi = inside fluid film coefficient, W/m2 0C,
h od = outside dirt coefficient (fouling factor), W/m2 0C,
h id = inside dirt coefficient, W/m2 0C,
kw = thermal conductivity of the tube wall material, W/m0C,
di = tube inside diameter, m,
do = tube outside diameter, m.

Typical values of the overall heat-transfer coefficient for various types of fluids that can be used
in shell and tube heat exchanger are given in Table below. And in our current case, we will take
U between 10 - 50 w/m2.0C.

The mean temperature difference, the well-known “logarithmic mean” temperature difference
is only applicable to sensible heat transfer in true co-current or counter-current flow. This will
normally be calculated from the terminal temperature differences: the difference in the fluid
temperatures at the inlet and outlet of the exchanger. For counter-current flow, the logarithmic
10 mean temperature is given by:

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 11 Hydrocarbon Dew Point Control

Where,
T lm = log mean temperature difference,
T1 = hot fluid temperature, inlet,
T2 = hot fluid temperature, outlet,
t1 = cold fluid temperature, inlet,
t2 = cold fluid temperature, outlet.

The equation is the same for co-current flow, but the terminal temperature differences will be
(T 1 - t 1 ) and (T 2 -t 2 ).

3. J -T Valve

In thermodynamics, the Joule–Thomson effect describes the

temperature change of a gas or liquid when it is forced through a valve or
porous plug while kept insulated so that no heat is exchanged with the
environment. This procedure is called a throttling process or Joule–
Thomson process. So the valve which has capability to produce this effect
is called J – T valve.

The after passing through the heat exchanger passes through the J-T valve
before entering the separator. Here, the refrigeration or cooling effect takes
place which lower the temperature of the gas to extent that most of it
liquefies. J-T valve is considered as the essential part of this system.

Pressure Drop across the valve

Pressure drop across a valve (choke) is an isenthalpic process, as noted previously. If no liquid
forms, the following equation applies

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 12 Hydrocarbon Dew Point Control

The symbol "p" is known as the Joule-Thomson coefficient. It is positive or negative, depending
on the relative size of the two terms in the numerator.

Curve A below shows a case where the instantaneous slope is greater than the average slope.
Therefore, the gas will cool on expansion. The curve C gas is just the opposite and will heat on
expansion. Curve B is for an ideal gas, which will not change temperature on expansion.

Many gases exhibit a characteristic wherein the slope of the V-T curve changes sign. The
temperature at which the slope changes sign (μ = O) is known as the inversion temperature. The
right-hand plot above shows inversion temperature versus pressure. The shape shown is general
for all actual gases. Outside the curve, the gas represented would heat upon expansion. Inside, it
cools on expansion.

Because of the location of the curve, hydrogen heats on expansion at normal pressures, whereas
most light hydrocarbons cool. At very high pressures, of the order of 60MPa [8700psia], many
naturally occurring hydrocarbon gases heat on expansion.

Curves which show the temperature drop expected for a given pressure drop across a choke are
only applicable if no liquid forms on such expansion.

Liquid Condensation across the Valve

Across a valve the First Law of Thermodynamics reduces to h l = h 2 . No work is possible and the
process is almost adiabatic. The amount of heat transfer across a valve body is poor, and the gas
12 is in it for only a short time. The calculation is inherently trial-and-error

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1. Calculate the total enthalpy of the feed stream at P I and T I . If it is a two-phase

stream, the total enthalpy is found by adding that of the liquid and vapor
phases.
2. Assume the unknown temperature T 2 .
3. Make a flash calculation at P 2 and T 2 to find relative amount and analysis of
each phase.
4. Find the total enthalpy at Point Two from the above flash and the assumed T 2 .
5. If h 2 = h l , you assumed the right temperature. If not, repeats Steps 2-5 until the
h are equal, within the desired limits of accuracy.

Since the above is very tedious, a reasonably good answer usually can be obtained by assuming
two different temperatures and plotting them on the
following type of figure.

A straight line between the h 2 - h l , found for two

assumed temperatures, is connected by a straight line.
The intersection at h 2 - h l = 0 gives approximate true
T2.

For the systems shown in Figure aside, the inlet gas

stream is cooled by the exit separated gas before going
to the choke. All one knows are the inlet conditions
and composition and the sales gas limitations. These fix the pressure drop across the heat
exchanger and choke in series. The drop across the former should not exceed 70kPa [10 psi].

The following is the general procedure to find the change in enthalpy

1. Assume temperature of gas downstream from gas/gas exchanger.

2. Run flash calculation* at this temperature and inlet pressure, minus 70kPa
[10 psia].
3. Determine enthalpy of total stream at this point from composition in Step 2.
4. Use the previous procedure for a choke to find the temperature in the low
13 temperature separator.
5. Run flash at separator conditions.
6. Find the enthalpy of the vapor leaving the separator.

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 14 Hydrocarbon Dew Point Control

7. The ∆H of the sales gas must equal the H of the inlet gas across the gas/gas
exchanger. If this is not found to be true, Steps 1-6 must be repeated.
8. Once Step 7 is satisfied, the heat exchanger may be found by conventional heat
transfer principles.

*A flash calculation is needed on the inlet gas to the gas/gas exchanger if it is two-phase and composition
and relative quantity of each phase is not known.

The above procedure illustrates the general conditions that must be satisfied for all
systems where expansion across a valve is involved.

Expansion across a valve may be the proper choice over an expander but the temperature drop is
less and no useful work is produced.

4. Low Temperature Separator

The gas after getting condensed enters the

Low Temperature Separator. This is basically
a three phase horizontal separator whose
basic function is separate the uncondensed
gas coming from J-T Valve from liquid gas
as well as to separator ethylene glycol from
the condense gas. From here, the separated
gases are removed as dry lean gas and used to
lower the inlet temperature of the incoming
gas by exchanger. On the other the NGL hydrocarbons after separation are sent to stabilization
phase and separated glycol is moved back to EG regeneration skid for further processing.

Working and Design

In a horizontal separator, the liquid that has been separated from the gas moves along the bottom
of the vessel to the liquid outlet. The gas and liquid occupy their proportionate shares of the shell
cross-section. Increased slug capacity is obtained through shortened retention time and increased
liquid level.
14 Figure below also illustrates the separation of two liquid phases (glycol and hydrocarbon). The
denser glycol settles to the bottom and is withdrawn through the boot. The glycol level is
controlled by an interface level control instrument.

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 15 Hydrocarbon Dew Point Control

This separator has certain advantages with respect to gravity separation performance in that the
liquid droplets or gas bubbles are moving perpendicular to the bulk phase velocity, rather than
directly against it as in vertical flow, which makes separation easier.

Typical L/D ratios for horizontal separators normally fall in the range of 3–9.

While designing this separator, we have to take care of the gas and liquid capacity of the
separator.

Gas capacity of the separator may be determined by modification of Stroke`s law. This law is
based on the principal of the minimum droplet size that will settle out of a moving gas stream at
a given velocity. The maximum superficial velocity of the gas at operating conditions is
calculated by

15
Where,
Vt = Max allowable superficial velocity
ρL = Density of the liquid
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 16 Hydrocarbon Dew Point Control

ρg = Density of the gas

K = Constant (Depending upon the design and operating conditions)

The API 12J recommended K values are shown in the below table. Many separators are greater
than 10 feet in length, with some reaching 50 feet or more. The relationship shown in the table
for adjusting for length will give K factors greater than 1 ft/sec for large separators. These higher
values of K for large (long) horizontal separators are generally considered to be overly
optimistic. In practice, K = 0.5 ft/sec is normally used as an upper limit for horizontal separators
equipped with wire-mesh mist extractors.

The required Area of the mist extractor is calculated by

The Liquid Capacity of the separator is typically specified in terms of residence time, which
must be translated into vessel layout requirements for dimensioning purposes. Residence time
establishes the separator volume required for the liquid as shown in the below equation

Where,

U = Liquid Capacity
W = Liquid settling volume
16
t = Retention Time

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 17 Hydrocarbon Dew Point Control

Table below provides suggested residence times for various liquid- liquid separation
applications. These figures generally assume equal residence times for both the light and heavy
liquid phases.

5. Ethylene Glycol Injection & Regeneration System

The saturated water content of a gas depends on pressure, temperature, and composition. The
effect of composition increases with pressure and is particularly important if the gas contains
CO 2 and/or H 2 S. For lean, sweet natural gases containing over 70% methane and small amounts
of heavy hydrocarbons, generalized pressure-temperature correlations are suitable for many
applications.

Graph below is an example of one such correlation which has been widely used for many years
in the design of “sweet” natural gas dehydrators. The gas gravity correlation should never be
used to account for the presence of H 2 S and CO 2 and may not always be adequate for certain
hydrocarbon effects, especially for the prediction of water content at pressures above 1500 psia.
17 The hydrate formation line is approximate and should not be used to predict hydrate formation
conditions.

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Hydrate formation is a time dependent process. The rate at which hydrate crystals form depends
upon several factors including gas composition, presence of crystal nucleation sites in the liquid
phase, degree of agitation, etc. During this transient “hydrate formation period” the liquid water
present is termed “Metastable Liquid”. Metastable water is liquid water which, at equilibrium,
will exist as a hydrate.

When designing dehydration systems meet extremely low water dewpoint specifications, it is
necessary to determine the water content of the gas in equilibrium with a hydrate using a
correlation like that presented in Figure. If a Metastable correlation is used, one will overestimate
the saturated water content of the gas at the dewpoint specification. This, in turn, may result in a
dehydration design which is unable to meet the required water removal.

Where experimental data is unavailable, utilization of a sound thermodynamic-based correlation

can provide an estimate of water content in equilibrium with hydrates.

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Hydrate Prediction Based on Composition for Gas

Several correlations have proven useful for predicting hydrate formation of sweet gases and
gases containing minimal amounts of CO 2 and/or H 2 S. The most reliable ones require a gas
analysis. The Katz method utilizes vapor solid equilibrium constants defined by the Equation

The applicable K-value correlations for the hydrate forming molecules (methane, ethane,
propane, isobutane16, normal butane17, carbon dioxide, and hydrogen sulfide) are shown in
Figures below. Normal butane cannot form a hydrate by itself but can contribute to hydrate
formation in a mixture.

For calculation purposes, all molecules too large to form hydrates have a K-value of infinity.
These include all normal paraffin hydrocarbon molecules larger than normal butane.
Nitrogen is assumed to be a non-hydrate former and is also assigned a K-value of infinity.

The K vs values are used in a “dewpoint” equation to determine the hydrate temperature or
pressure. The calculation is iterative and convergence is achieved when the following objective
function is satisfied.

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Process and working

Inhibition utilizes injection of ethylene glycols into a process stream where it can combine with
the condensed aqueous phase to lower the hydrate formation temperature at a given pressure.

Ethylene Glycol can be recovered with the aqueous phase, regenerated and re-injected. At
cryogenic conditions (below –40°F) methanol usually is preferred because glycol’s viscosity
makes effective separation difficult.

Ethylene glycol (EG), di-ethylene glycol (DEG), and tri-ethylene glycol (TEG) glycols have
been used for hydrate inhibition. The most popular has been ethylene glycol because of its lower
cost, lower viscosity, and lower solubility in liquid hydrocarbons.

Physical properties of the most ethylene glycol and ethylene glycol-water mixtures are given in
Figures below

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Tabular information for the Ethylene glycol and methanol is provided in below

In this system, ethylene glycol is typically sprayed on the tube-sheet faces of the gas/gas
exchanger so that it can flow with the gas. As water condenses, the inhibitor is present to mix
with the water and prevent hydrates. Injection must be in a manner to allow good distribution
throughout the heat exchanger operating below the gas hydrate temperature.

The viscosities of ethylene glycol and its aqueous solutions increase significantly as temperature
decreases, and this must be allowed for in the rating of plant exchanger.

The inhibitor and condensed water mixture is separated from the gas stream along with a
separate liquid hydrocarbon stream. At this point, the water dewpoint of the gas stream is
29 essentially equal to the separation temperature. Ethylene Glycol-water solutions and liquid
hydrocarbons can emulsify when expanded from a high pressure to a lower pressure by J-T
expansion valve.

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Careful separator design will allow nearly complete recovery of the diluted Ethylene glycol for
regeneration and reinjection. The regenerator in an ethylene glycol injection system should be
operated to produce a regenerated ethylene glycol solution that will have a freezing point below
the minimum temperature encountered in the system. This is typically 75-80 wt%. Figure shows
the freezing point of various concentrations of Ethylene glycol water solutions.

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The minimum inhibitor concentration in the free water phase may be approximated by
Hammerschmidt’s equation
.

Here, K H is taken from 2335 to 4000 and this equation is used for 40 – 70% wt of glycol

Once the required inhibitor concentration has been calculated, the mass of inhibitor required in
the water phase may be calculated from Equation

The amount of inhibitor to be injected not only must be sufficient to prevent freezing of the
inhibitor water phase, but also must be sufficient to provide for the equilibrium vapor phase
content of the inhibitor and the solubility of the inhibitor in any liquid hydrocarbon.

Solubility of EG in the liquid hydrocarbon phase is extremely small. Solubility of 0.3 lb per 1000
gallon (U.S.) of NGL is often used for design purposes. However, entrainment and other physical
losses may result in total losses significantly higher than this.

EG Regeneration Equipments

Conventional Ethylene glycol type dehydration unit furnished to this specification is to be a skid
mounted assembly. The inlet scrubber and contactor may be skid or foundation mounted separate
from the re-concentrator skid upon agreement between the purchaser and the manufacturer.

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a. Inlet Scrubber

The inlet scrubber furnished to this specification

is to be vertical and is normally available in
sizes and maximum allowable working pressure
ratings shown in Table. Table is for nominal
industry standards. Available sizes and working
pressures may vary from the stated ratings.
Other types, sizes, pressures, and temperature
ratings may be furnished by agreement between
the purchaser and manufacturer provided they
conform to API Specification 125.

The inlet scrubber may be separate or integral with the contactor as specified by the purchaser.
The inlet scrubber requires a mist extractor and an integral scrubber also requires a chimney tray
with a sufficient volume to prevent glycol overflow into the scrubber during shutdown.

b. Reboiler

A reboiler furnished to this specification is to be horizontal. The fire tube shall be field
removable for inspection. The heat duty
requirement of the reboiler shall include
the benefit of a heat exchanger used for
heat recovery in the re-concentrator
system.

The reboiler should be designed for a

minimum of 1% psi of internal pressure or
full of water, whichever is greater. The
deflection of fiat end closures should be
limited to the diameter divided by 500 with
1% psi internal pressure or full of water,
whichever is greater. Typical reboiler
nominal duties are given in Table.

c. Still Column
32

A still column furnished to this specification is to be integral with the reboiler. The column is to
be flanged such that it is removable and is to be provided with one or more lugs for removal.

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The glycol reboiler should be equipped with a still column complete with packing in order to
minimize glycol vaporization losses. On larger systems it may be economical to include a reflux
system which utilizes the incoming rich glycol in an internal coil to cool the outlet vapor stream.
Outlet vapor piping should be sized for minimum pressure loss. Vapor piping should not be
restricted

d. Surge Tank

The surge tank is used to provide glycol for operating purposes and is not considered as storage.
Additional holding capacity for glycol from the other equipment during shutdown must be
specified by the purchaser. A surge tank furnished to this specification is to be horizontal and
may be integral with the reboiler. Holdup time is 15min to 30min.

e. Flash Vessel

A two or three phase separator is used in the rich glycol stream, to remove entrained gas and
hydrocarbon liquids.

A frequently used option in regeneration systems is a gas-condensate glycol separator, and

should be included when the inlet gas contains condensate. It may be located upstream or
downstream of the glycol/glycol heat exchanger and usually operates at a pressure of 25-75 psig.
It removes condensate from the glycol prior to the reboiler, which minimizes coking and
foaming problems. The separator also captures flash gas that is liberated from the glycol and
exhaust gas from the glycol-gas powered pumps, so that the gas may be used as fuel. Glycol is
regulated from the separator to the reboiler by means of a level controller and dump valve.
Condensate removal may be controlled automatically or manually.

f. Glycol / Glycol Exchanger

A shell-and-tube, plate type, double pipe, internal coil within the surge tank or other type heat
exchanger employed to recover heat from the outgoing hot lean glycol from the reboiler and
preheating the incoming cool rich glycol from the contactor.

A glycol to glycol heat exchanger furnished to this specification is to be either external or

integral with the surge tank. The exchanger shall be capable of cooling the lean glycol to the
33 maximum temperature.

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g. Gas/Glycol heat exchanger

A shell-and-tube, pipe-in-pipe, or other type heat exchanger employed to cool the lean glycol
with the gas leaving the contactor before the glycol enters the contactor.

It is important that the glycol entering the contactor be cooled to a 10" to 30°F above the
temperature of the gas stream. This is necessary because the equilibrium conditions between the
glycol and the water vapor in the gas are affected by temperature. At higher temperatures, more
water vapor will remain in the gas stream. A cooler glycol temperature will decrease the glycol
vaporization losses but hydrocarbons may condense in the contactor.

The Gas/Glycol heat exchanger may be either external or internal to the contactor

Absorption is improved with lower temperature glycol. A gas/glycol heat exchanger is required
which uses dehydrated gas to cool the lean (dry) glycol before it enters the top of the contactor.

h. Skid

Some skid mounted items may be shipped separately from the skid by agreement of the
purchaser and manufacturer. The skid provided to this specification is to have a pull bar or lift
lugs for loading and unloading for shipment. The skid is to be capable of a single end lift as
assembled for shipment.

i. Filters and Strainers

Regeneration systems contain various types of filters and strainers. A particle filter or fine mesh
strainer is required to protect the pump. To reduce foaming, an activated carbon filter may be
installed to remove heavy hydrocarbons from the glycol. There is no standard arrangement for
these items in the system.

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 35 Hydrocarbon Dew Point Control

Following diagram shows the ethylene glycol regeneration system

Process Limitation

J-T plants are simple and easily operated facilities. However, they have the limitation that the
flowing wellhead pressure must be at least 2000-3000 kPa [300-500 psia] above the sales
pressure for the system to reach low enough temperatures to meet normal dewpoint
requirements. When the reservoir depletes to the point that this excess pressure is not available,
the process ceases to function as a dew-point control method unless front-end compression (or
35 residue compression) is installed to maintain the inlet pressure or mechanical refrigeration is
added to assist in cooling the gas. These units are most commonly used to process high pressure,
non-associated gas with low flows (less than 10 MMscfd). The mechanical refrigeration is
required for rich gas streams and to obtain high recoveries.
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 36 Hydrocarbon Dew Point Control

J-T plant had a free pressure drop and maximized propane plus recovery as shown below:

B) MECHANICAL REFRIGERATION

The process involved in Mechanical Refrigeration in same as used in J-T Expansion Valve. The
only difference between two processes is the use of chiller instead of J-T Valve. In the
mechanical refrigeration, the chiller duty is fulfilled by the process of refrigeration due to which ,
more less temperature is achievable.

Process and Working

Figure shows the simplest compression refrigeration system. Saturated liquid at Point A expands
across a valve (isenthalpically). On expansion some vaporization occurs. The mixture of
36
refrigerant vapor and liquid enters the chiller at 3-6°C [5-10°F] lower than temperature to which
the process stream is to be cooled. The liquid vaporizes. Leaving at Point C is a saturated vapor

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 37 Hydrocarbon Dew Point Control

at the P and T of the chiller. This vapor is compressed and then enters the condenser as a
superheated vapor.

The refrigerant must leave the condenser as a saturated liquid or slightly sub-cooled. Nothing
happens in the accumulator. It merely serves as a reservoir for refrigerant as levels vary in the
chiller(s) and condenser.

Calculation of a Simple System

There are several discrete steps in the sizing of the system shown in Figure. These are
summarized below:

Determination of refrigerant circulation rate

The balance at right is around the chiller and expansion valve. At Point A the refrigerant is a
saturated liquid (or very close to it). At Point C it is a saturated vapor. Q chiller is the heat load
determined by specifications on the stream being
cooled.

If one writes an energy balance around the system,

Q chiller + mA h A = mc h c
But m A = m c = m, so

Where,

Q chiller = Chiller Heat Load, btu/hr

hc = Saturated Vapor Enthaphly, btu/lbm

hA = Saturated Liquid Enthaphly, btu/lbm

m = Circulation Rate, lbm/hr

37 Determination o f Condenser Heat Load

There are two ways to do this. Knowing Q chiller and W, you can write an overall balance as
shown in Figure to find Q cond .

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 38 Hydrocarbon Dew Point Control

If we are performing the calculation manually and wish an

independent check of the previous work, write the balance
shown below

The Q cond found from above equation will be negative.

This merely signifies that heat is leaving the system. The value found from the overall balance
will not check the condenser balance exactly because only part of the compressor inefficiency
shows up in the refrigeration system. For practical purposes, the difference is trivial.

Calculation of Chiller Load

The refrigeration load must be calculated from the specifications of the system in which chiller is
placed. Figure below shows a simple system using glycol injection to inhibit hydrate formation.

The first step in the design is to fix the temperature (T 3 ) in the low temperature separator (LTS).
The pressure in the LTS must be high enough above the specified sales pressure to allow for
pressure drop in the gas-gas exchanger and lines.

The minimum temperature coming to the gas-gas exchanger is fixed by the economics of pre-
cooling the feed stream. The maximum sales gas temperature is usually fixed by contract and is
seldom allowed to exceed 50°C [122°F]. Consequently, the heat load between P 1 , T 1 and P 3 , T 3
is fixed by these considerations. The problem revolves around the distribution of this load
between the gas-gas exchanger and the refrigerated chiller. For this calculation, it is convenient
to assume a 34 kPa [5 psi] drop in each heat exchanger.

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 39 Hydrocarbon Dew Point Control

The following general procedure is suggested if the gas-gas exchanger and chiller are to be sized
as part of the exercise.

• Determine the amount and composition of vapor and liquid at T 1 , P 1 and T 3 , P 3 .

(Don’t forget the water and glycol.)

• Calculate the total ∆H between points 1 and 3 (H I and H 3 ).

• Calculate the cooling capacity for the sales gas in the gas-to-gas exchanger (H 6 - H 4 ), T 6
should be fixed at (T I - 5°C) or the contractual maximum temperature, whichever is
lower.

• Calculate the chiller duty, Q chi1ler , by difference

Q chiller = (H I - H 3 ) - (H 4 - H 6 )

• Perform a series of flashes on the feed stream at temperatures and pressures between
point l and 3.

The cold liquid from the LTS is also available for cooling service since usually it must be heated
before entering the fractionation system. Although not shown in Figure above, it may be used for
cooling the feed or for any other cooling function within the system. If this liquid is not heated, a
cold-feed stabilizer might be specified

Refrigeration
By utilizing the Pressure-Enthalpy (P-H) diagram, the mechanical refrigeration can be broken
down into four distinct steps:

1. Expansion
2. Evaporation
3. Compression
4. Condensation

The vapor-compression refrigeration cycle can be represented by the process flow and P-H
diagram is shown

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 40 Hydrocarbon Dew Point Control

1. Expansion Step

The starting point in mechanical refrigeration is the availability of liquid refrigerant. Point A in
above figure represents a bubble point liquid at its saturation pressure, P A , and enthalpy, h LA . In
the expansion step, the pressure and temperature are reduced by flashing the liquid through a
control valve to pressure P B . The lower pressure, P B , is determined by the desired refrigerant
temperature, T B (point B).

At point B, the enthalpy of the saturated liquid is h LB , while the corresponding saturated vapor
enthalpy is h VB . Since, the expansion step (A – B) occurs across an expansion valve and no
energy has been exchanged, the process is considered to be isenthalpic. Thus the total stream
enthalpy at the outlet of the valve is the same as the inlet, h LA .

Since point B is inside the envelope, vapor and liquid coexist. In order to determine the amount
40 of vapor formed in the expansion process, let X be the fraction of liquid at pressure P B with an
enthalpy h LB . The fraction of vapor formed during the expansion process with an enthalpy h VB is
(1-X). Equations for the heat balance and the fraction of liquid formed are:

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 41 Hydrocarbon Dew Point Control

2. Evaporation Step

The vapor formed in the expansion process (A-B) does not provide any refrigeration to the
process Heat is absorbed from the process by the evaporation of the liquid portion of the
refrigerant. As shown in previous figure, this is a constant temperature, constant pressure step
(B-C). The enthalpy of the vapor at point C is h VB .

Physically, the evaporation takes place in a heat exchanger referred to as an evaporator or a

chiller. The process refrigeration is provided by the cold liquid, X, and its refrigerant effect can
be defined as X (h VB – h LB ) and substituting from Equations above, the effect becomes:

Effect = h VB − h LA
The refrigeration duty (or refrigeration capacity) refers to the total amount of heat absorbed in
the chiller by the process, generally expressed as “tons of refrigeration,” or Btu/unit time.

The refrigerant flow rate is given by:

m= Q ref ____
(h VB − h LA )

3. Compression Step

The refrigerant vapors leave the chiller at the saturation pressure P C . The corresponding
temperature equals T C at an enthalpy of h VB . The entropy at this point is SC. These vapors are
compressed isentropically to pressure PA along line C – D′ as shown in figure.

The isentropic (ideal) work, W i , for compressing the refrigerant from P B to P A is given by:

W i = m (h′ VD − h VB )
41
The quantity h′ VD is determined from refrigerant properties at P A and entropy of S C . Since the
refrigerant is not an ideal fluid and since the compressors for such services do not operate

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 42 Hydrocarbon Dew Point Control

ideally, isentropic efficiency, η i , has been defined to compensate for the inefficiencies of the
compression process. The actual work of compression, W, can be calculated from

The enthalpy at discharge is given by:

The work of compression can also be expressed as:

Where, 2544.4 Btu/hr = hp.

4. Condensation Step

The superheated refrigerant leaving the compressor at P A and T D (Point D in Figure) is cooled at
nearly constant pressure to the dew point temperature, T A , and refrigerant vapors begin to
condense at constant temperature.

During the de-superheating and condensation process, all heat and work added to the refrigerant
during the evaporation and compression processes must be removed so that the cycle can be
completed by reaching Point A (the starting point) on the P-H diagram, as shown in Figure.

By adding the refrigeration duty to the heat of compression,

We calculate the condensing duty, Q cd , from:

The condensing pressure of the refrigerant is a function of the cooling medium available (air,
cooling water, or another refrigerant). The cooling medium is the heat sink for the refrigeration
cycle.
42 Because the compressor discharge vapor is superheated, the refrigerant condensing curve is not a
straight line. It is a combination of de-superheating and constant temperature condensing. This
fact must be considered for proper design of the condenser.

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 43 Hydrocarbon Dew Point Control

System Pressure Drop

Some typical values for pressure drops that must be considered are:

Condenser pressure drop 3.0 to 7.0 psi

Line hydraulic losses

Evaporator to Compressor* 0.1 to 1.5 psi

Compressor to Condenser 1.0 to 2.0 psi

Condenser to Receiver 0.5 to 1.0 psi

* This is an important consideration in refrigeration services with low suction pressure to

compressor.

Facts relating to refrigeration

Figure below illustrates the ethane recovery efficiency which can be expected at certain.
Knowing the pressure and achievable temperature by different technique we can easily conclude
to which extent we can be successful in ethane recovery. As with propane recovery, for a given
temperature level, higher extraction efficiency can be achieved with richer gas. However, ethane
recovery of over 30% can be achieved from a gas as lean as 3 GPM (C3+).

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 44 Hydrocarbon Dew Point Control

Figure below illustrates the effect of gas pressure on plant performance in propane plus recovery
operation.

Choice of Refrigerant
The ideal refrigerant is nontoxic, noncorrosive, has PVT and physical properties compatible with
the system needs, and has a high latent heat of vaporization. Any material could be used as a
refrigerant. The practical choice reduces to one which has desirable physical properties and will
vaporize and condense at reasonable pressures, at the temperature levels desired. The usual
choice is propane, ammonia, R-12 or R-22 at chiller temperatures above about 40°C. At
cryogenic conditions, ethylene and methane might be used. In general, the lower practical limit
44 of any refrigerant is its atmospheric pressure boiling point. It is desirable to carry some positive
pressure on the chiller to obtain better efficiency in the compressor, reduce equipment size and
avoid air induction into the system.

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 45 Hydrocarbon Dew Point Control

Ammonia is seldom chosen because of emotional reactions to its odor. However, it is easy to
handle in ordinary steel equipment containing no copper and brass and is really less dangerous
than propane because of its pungent odor. No dangerous accumulation can build up unnoticed.

Propane is by far the most popular refrigerant in the gas processing applications. It is readily
available (often manufactured on-site), inexpensive and has a "good" vapor pressure curve. It is
flammable but this is not a significant problem if proper consideration is given to the design and
operation of the facility.

Temperature and pressure relationship of different refrigerants is given in following graph.

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 46 Hydrocarbon Dew Point Control

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 47 Hydrocarbon Dew Point Control

Determination of the Enthalpies

The calculation requires one to find the enthalpy per unit mass at Points A, B and C of the last
shown figure. This can be done from a computer routine. Convenient tables and figures are also
available for all of the common commercial refrigerants. Pressure-Enthalpy (P-H) figures for
propane as well as vapor pressure and physical property information on all common refrigerants
can be easily found

The P-H diagram is very convenient for solving the energy balance for a simple system.

The left-hand figure is a representation of the P-H diagram. The refrigerant is all liquid to the left
of the saturated liquid curve; it is two-phase inside the saturation curve and all vapor to the right
of the saturated vapor curve. The lines of constant temperature are horizontal between the
saturated vapor and liquid curves and then rise almost vertically in the liquid section.

The calculation process starts by choosing the temperature of Point A. Considering, water, air or
some other stream which is used for condensation of the refrigerant. And which temperature is
realistically achievable in the condenser. That is Point A. It is on the liquid saturation curve,
since it leaves the condenser as a liquid.

What is the temperature at Points B and C? Normally, it will be 3-6°C less than the minimum
desired temperature for the fluid being cooled. This approach fixes the location of Point C. It is
on the saturated vapor curve, since it is in equilibrium with the liquid in the chiller (evaporator).
47 The expansion across the choke from Point A is an isenthalpic process; a vertical line on a P-H
diagram. Draw a vertical line from A to B, the pressure of Point C, and then go horizontally to C.

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 48 Hydrocarbon Dew Point Control

We can also read the ∆H required for Equation mentioned in determination of refrigeration
circulation rate.

The theoretical compression is isentropic. Starting at Point C, draw parallel entropy lines until it
intersects the pressure line of Point A. This is theoretical Point D.

Equation used for condenser heat load is found from the ∆h between Points D and A.

For a commercially pure refrigerant, use of a P-H diagram is as reliable as any method.
Following are the P-H diagrams for propane.

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 49 Hydrocarbon Dew Point Control

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 51 Hydrocarbon Dew Point Control

Process Limitations
The drawback is that the refrigeration must be in operation to accomplish the dehydration. If it is
desired to operate the dehydration at times independent of the refrigeration, then separate units
are used.

Another problem in using dew point control units of mechanical refrigeration systems is the
disposition of the liquids removed. The liquids must be stabilized by flashing to lower pressure
or by the use of a stabilization column. When the condensate is flashed to a lower pressure, light
hydrocarbons are liberated which may be disposed of in a fuel gas system. The stabilization
column can produce a higher quality and better controlled product.

C) TURBO EXPANDER
The gas expander had its beginnings as a modified form of steam turbine, which is a common
machine used to drive pumps, generators, and other rotating equipment. However, development
of a high efficiency gas expander has occurred over the past
years. The expanders available today can recover up to 85%
of the energy given up by gas as its pressure is lowered.
This energy is transmitted to the rotating device such as a
pump, compressor, generator, etc. In typical cryogenic
plants, a compressor is attached to the expander shaft. The
compressor also has a single impeller or wheel. It rotates at
the same speed as the expander.

High pressure gas enters the expander and is directed at the

outer tip of the expander impeller blades, causing it to
rotate. The gas flows to the center of the impeller and exits the expander at a lower pressure. The
rotating speed of the expander can be in excess of 50,000 rpm, depending upon the volume of
gas entering the unit, and the pressure drop the gas takes in flowing through the unit (inlet
pressure minus discharge pressure). A high flow rate and high pressure drop result in a high
expander speed and obviously a high power output.

The power developed at the rotating shaft of the expander is used to drive a single impeller
compressor attached to the other end of the shaft. Low pressure gas enters the center of the
impeller and discharge gas is withdrawn from the tip of the impeller blades at a pressure about
1.1 to 1.4 times that of suction pressure.
51
In the event the where flow of gas to the expander is more than the expander can handle, a
bypass valve will open and allow the excess gas through it. The bypass valve is often referred to
as a J-T (Joule Thomson) valve. This is typically done by use of a split range flow controller.
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 52 Hydrocarbon Dew Point Control

Most turbo expander plants are

designed to operate at very low
temperatures, -101°C [-15O °F].
These temperatures are required
to achieve high NGL extraction
levels, especially ethane. In
these plants the expander outlet
stream flows to a low
temperature separator which is
usually installed at the top of a
demethanizer column. The condensed liquids drop down on the top tray of the demethanizer
where they are stabilized to meet an NGL product specification. The demethanizer bottom
product is a mixture of C2, C3, C4's and C5+.

In hydrocarbon dewpoint control plants, these extremely low temperatures are not required.
The expander outlet temperature is usually O to -20°C [32° to -4°F]. Expansion ratios are
typically about 1.5. A typical hydrocarbon dewpoint control plant using a turbo expander is
shown in figure below

52 As gas flows through the expander, its temperature is lowered and some of the stream condenses.
The liquid which forms has no detrimental effect on the expander. If the gas stream entering the
expander contains solid particles of dirt or debris or contains moisture or carbon dioxide which
will freeze at the low temperatures in the unit, serious damage to the machine may result. At the
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 53 Hydrocarbon Dew Point Control

high operating speed of the machine, the presence of solid materials, like debris or ice, will
quickly sandblast the wheel and casing. The expander inlet separator is a very important piece of
equipment. It must be sized properly to remove these components from the gas.

A screen is normally installed on the inlet gas line to the expander to remove solid particles from
the gas stream. Moisture is removed from the gas in the dehydrators in the front end of the plant.

Process Limitations
Turbo Expander is used only when “Free” pressure drop in the lean gas stream. Mostly it is used
when high ethane recovery (i.e., over 30% ethane recovery) our requirement

It has high utility costs but it operation is quite flexible that is it can be easily adapted to wide
variation in pressure and products.

D) Solid Membrane
Molecular Sieves are used for separation of heavier
hydrocarbons. Molecular sieves are manufactured in
two crystal types, a simple cubic or Type A crystal and
a body-centered cubic or Type X crystal. The Type A
sieve is available in sodium, calcium, and potassium
forms. The Type X sieve is available in sodium and
calcium forms. The sodium forms of the sieves are the
most common and are shown below in oxide formulas.

Type 4A: Na20-A1203.2 SiQ.YH20

Type 13X: Na20.Al203.2.5 Si02.YH20

The value of Y depends on the extent of activation.

A sieve is manufactured by crystallizing the proper crystal type in sodium form from a solution
of sodium silicate, aluminum tri-hydrate, and sodium hydroxide. If the sodium formed, is not the
desired product, then either calcium or potassium ions are substituted for the sodium ions by
soaking the crystals in a solution of the appropriate chloride salt. The ion exchange is never
complete. The small crystals (1-4 microns in size) are then mixed with a clay binder and either
extruded into cylindrical pellets or rolled into spherical beads. The sieve is finally activated
53 (calcined) by heat at 650°C [1200° F] to remove most of the hydrated water.

Table below contains data on the standard molecular sieves. Openings or ports into the crystal
cavities are 3, 4, 5, and 10 Angstroms (1 A = 10-lo meter). These dimensions are of the same
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 54 Hydrocarbon Dew Point Control

order of magnitude as the diameters of small molecules. Type 3A is not as stable as Type 4A. It
is used in cracked gas dehydrators to dry light olefin streams, e.g., ethylene, propylene, etc.
Rarely is a sieve other than Type 4A, 5A or Type 13X needed in a natural gas plant. Type 4A is
used for dehydration of gases and liquids. Type 5A is frequently used for simultaneous H 2 0 and
H 2 S removal from natural gas. Type 13X is used for NGL product treating.

Line Sizing Criteria as per API 14-E

In determining the diameter of pipe to be used in platform piping systems, both the flow velocity
and pressure drop should be considered. When determining line sizes, the maximum flow rate
expected during the life of the facility should be considered rather than the initial flow rate. It is
also usually advisable to add a surge factor of 20 to 50 percent to the anticipated normal flow
rate, unless surge expectations have been more precisely determined by pulse pressure
measurements in similar systems or by specific fluid hammer calculation.

Determination of pressure drop in a line should include the effect of valves and fittings.
Calculated line sizes may need to be adjusted in accordance with good engineering judgment.
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 55 Hydrocarbon Dew Point Control

A) Sizing Criteria For Liquid Lines

Single-phase liquid lines should be sized primarily on the basis of flow velocity. For lines
transporting liquids in single phase from one pressure vessel to another by pressure differential,
the flow velocity should not exceed 15 feet/second at maximum flow rates to minimize flashing
ahead of the control valve. If practical, flow velocity should not be less than 3 feet/second to
minimize deposition of sand and other solids. At these flow velocities, the overall pressure drop
in the piping will usually be small. Most of the pressure drop in liquid lines between two
pressure vessels will occur in the liquid dump valve and/or choke.

Flow velocities in liquid lines may be calculated using the following derived equation

Where,

Vi = average liquid flow velocity, feet/second

QI = liquid flow rate, barrel/day
di = pipe inside diameter, inches.

Pressure drop (psi per 100 feet of flow length) for single phase liquid lines may be
calculated using the following (Fanning) equation:

Where,

∆P = pressure drop, psi/100 feet.

f = Moody friction factor, dimensionless.
QI = liquid flow rate, barrels/day.
Si = liquid specific gravity (water = 1).
di = pipe inside diameter, inches.
55

The Moody friction factor, f, is a function of the Reynolds number and the
surface roughness of the pipe. The modified Moody diagram, Figure below, may be used to
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 56 Hydrocarbon Dew Point Control

determine the friction factor once the Reynolds number is known. The Reynolds number may be
determined by the following equation:

Where,

Re = Reynolds number, dimensionless.

ρi = liquid density, lb/ft3.
df = pipe inside diameter, ft.
Vi = liquid flow velocity, ft/sec.
μ1 = liquid viscosity, lb/ft-sec,

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 58 Hydrocarbon Dew Point Control

B) Sizing Criteria for Single-phase Gas Lines

Single phase gas lines should be sized so that the resulting end pressure is high enough to satisfy
the requirements of the next piece of equipment. Also velocity may be a noise problem if it
exceeds 60 feet/second. However, the velocity of 60 feet/second should not be interpreted as
absolute criteria. Higher velocities are acceptable when pipe routing, valve choice and placement
are done to minimize or isolate noise.

The design of any piping system where corrosion inhibition is expected to be utilized should
consider the installation of additional wail thickness in piping design and/or reduction of velocity
to reduce the effect of stripping inhibitor film from the pipe wall. In such systems it is suggested
that a wall thickness monitoring method be instituted.

General Pressure Drop Equation is

Other equations to find pressure drop are

Weymouth Equation

Panhandle Equation

Spitzglass Equation

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 59 Hydrocarbon Dew Point Control

Gas Velocity Equation

C) Sizing Criteria for Gas/Liquid Two-Phase lines

Erosional Velocity

Flow lines, production manifolds, process headers and other lines transporting gas and
liquid in two-phase flow should be sized primarily on the basis of flow velocity. Experience has
shown the loss of wall thickness occurs by a process of erosion/corrosion. This process is
accelerated by high fluid velocities, presence of sand, corrosive contaminants such as CO 2 and
H 2 S, and fittings which disturb the flow path such as elbows.

The following procedure for establishing an “Erosional velocity’, can be used where no
specific information as to the erosive/corrosive properties of the fluid is available.

(1) The velocity above which erosion may occur can be determined by the following
empirical equation:

Where,
Ve = Fluid Erosional Velocity, feet/second
C = empirical constant
ρm = gas/liquid mixture density at flowing pressure and temperature, lbs/ft

Industry experience to date indicates that for solids-free fluids values of c = 100 for
continuous service and c = 125 for intermittent service are conservative. For solids-free fluids
where corrosion is not anticipated or when corrosion is controlled by inhibition or by employing
corrosion resistant alloys, values of c = 150 to 200 may be used for continuous service; values up
to 250 have been used successfully for intermittent service. If solids production is anticipated,
59 fluid velocities should be significantly reduced, Different values of “c” may be used where
specific application studies have shown them to be appropriate.

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 60 Hydrocarbon Dew Point Control

Where solids and/or corrosive contaminants are present or where “c” values higher than
100 for continuous service are used, periodic surveys to assess pipe wall thickness should be
considered. The design of any piping system where solids are anticipated should consider the
installation of sand probes, cushion flow tees, and a minimum of three feet of straight piping
downstream of choke outlets.

(2) The density of the gas/liquid mixture may be calculated using the following derived
equation:

Where,

P = operating pressure, psia.

Si = liquid specific gravity (water = 1, use average gravity for
hydrocarbon water mixtures) at standard conditions.
R = gas/liquid ratio, fts/barrel at standard conditions.
T = operating temperature, “R.
Sg = gas specific gravity (air = 1) at standard
B = gas compressibility factor, dimensionless.

(3) Once Ve is known, the minimum cross-sectional area required to avoid fluid erosion
may be determined from the following derived equation:

Where,

A = minimum pipe cross-sectional flow area required, in2/1000 barrels

liquid per day.

Minimum Velocity

60 If possible, the minimum velocity in two-phase lines should be about 10 feet per second to
minimize slugging of separation equipment. This is particularly important in long lines with
elevation changes.

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 61 Hydrocarbon Dew Point Control

Pressure Drop
The pressure drop in a two-phase steel piping system may be estimated using a simplified
Darcy equation from the GPSA Engineering Data Book (1981 Revision)

Where,

∆P = pressure drop, psi/100 feet.

di = pipe inside diameter, inches
f = Moody friction factor, dimensionless.
ρm = gas/liquid density at flowing pressure and temperature, lbs/ft3
W = total liquid plus vapor rate, lbs/hr.

W may be calculated using the following derived equation:

Where,

W = 3180 Q g S g + 14.6 Q I S i

Qg = gas flow rate, million cubic feet/day (14.7 psia and 60°F).
Sg = gas specific gravity (air = 1).
QI = liquid flow rate, barrels/day.
Si = liquid specific gravity (water = 1).

*It should be noted this pressure drop calculation is an estimate only.

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 62 Hydrocarbon Dew Point Control

Given Assignment
Inputs of the Assignment
Gas Flow rate = 25 MMSCFD

Gas Molecular weight = 24

Inlet Temperature = 100 °F

Outlet Temperature = 120 °F

Operating Pressure = 1200 psia

Composition
Sales Gas Composition after HCDP
Component unit inclusion (Mole Fraction)
(with both flowing wellheads)
Nitrogen 0.127290
CO2 0.028573
Methane 0.833896
Ethane 0.005841
Propane 0.001312
i-butane 0.000473
n-butane 0.000263
i-pentane 0.000227
n-pentane 0.000107
n-hexane 0.000282
n-heptane 0.000278
n-octane 0.000172
n-nonane 0.000034
n-decane 0.000000
H2O 0.000039
Oxygen 0.000000
Helium 0.001212
TOTAL 1.000000

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1. J-T Valve based on HCDP

(Feed gas compression)

Inlet gas flow = 25 MMscfd

Inlet Pressure = 1200 psi
Outlet pressure required = 4000 psi

Compressor power can be calculated by Eq.. 13-4 of GPSA

Where,
No. of stages = 1
Compression ratio = 3.34
F (for single stage) = 1

Thus,
Compressor power = 1837 hp

• Size of Free water Knockout Vessel

(Vertical two phase separator)

Inlet gas flow rate = 25 MMscfd

Gas density = 11.85 lb/ft3
Liquid density = 53.67 lb/ft3
Inlet Temperature = 100 0F
Inlet Pressure = 3995 psi
Compressibility factor = 0.9
Viscosity = 0.0226 cP
Particle size = 100 μm

(From API – 12 J)
Interpolating Pressure adjustment for 3995 psi = 23.5

63
K value = 0.042 ft/sec

July 2011 | By Muhammad Mansoor Anwar

 64 Hydrocarbon Dew Point Control

From GPSA Eq. 7-3

C`(Re)2 = 6144

From GPSA Fig 7-4, given below

C` = 1.2

64

July 2011 | By Muhammad Mansoor Anwar

 65 Hydrocarbon Dew Point Control

From Eq. 7-1 of GPSA

Terminal Velocity = 0.156 ft/sec

Gas flow rate can be calculated from the following Equation

QA = 1.34 ft3/sec

Vertical Separator Area is calculated by

65

Area of Separator = 8.60 ft2

July 2011 | By Muhammad Mansoor Anwar

 66 Hydrocarbon Dew Point Control

Dia of Separator = 3.0 ft

Taking L/d = 3

Length of separator = 9.0 ft

Mist Extractor AREA

From Eq. 7-8 of GPSA

Vt = 0.258 ft/sec

From Eq. 7-9 of GPSA

Area of Extractor = 5.17 ft2

66

July 2011 | By Muhammad Mansoor Anwar

 67 Hydrocarbon Dew Point Control

• Thermal Design of the Gas/Gas Exchanger

(Shell and Tube Heat Exchanger)

1) Thermal Duty = 1.0 MMbtu/hr (according to 1st assumption)

From Richardson & Coulson Volume 6

Log mean Temperature Difference

Inlet hot fluid temperature, T 1 = 100 0F

Outlet hot fluid temperature, T 2 = 50 0F
Inlet cold fluid temperature, t 1 = -21.43 0F
Outlet cold fluid temperature, t 2 = 30.30 0F
Log mean Temperature Difference = 70.56 0F

Overall Heat Transfer Coefficient = 0.1476 Btu/hr.ft2. 0F

(Trial value From Richardson & Coulson Vol 6 below table)

Area required can be calculated by Equation

67
Area 1 = 96018.6 ft2 (according to 1st assumption)

July 2011 | By Muhammad Mansoor Anwar

 68 Hydrocarbon Dew Point Control

2) Thermal Duty = 1.5 MMbtu/hr (according to 2nd assumption)

Overall Heat Transfer Coefficient = 0.1476 Btu/hr.ft2. 0F

(Trial value From R&C Volume 6 above table)

Area 2 = 144027.8 ft2 (according to 2nd assumption)

3) Thermal Duty = 2.0 MMbtu/hr (according to 3rd assumption)

Overall Heat Transfer Coefficient = 0.1476 Btu/hr.ft2. 0F

(Trial value From R&C Volume 6 above table)

Area 3 = 192037 ft2 (according to 3rd assumption)

• I select 2.0 MMbtu/hr Thermal Duty because it is enough for current process conditions along
with a reasonable area of gas/gas heat exchanger.

68

July 2011 | By Muhammad Mansoor Anwar

 69 Hydrocarbon Dew Point Control

• Size of Low Temperature Separator

(Horizontal three phase separator)

Inlet Gas flow = 25 MMscfd

Inlet Temperature = -21.43 0F
Inlet Pressure = 1200 psi
Liquid retention time = 20 min (From GPSA table as below)

K value = 0.30 ft/sec (from API – 12 J)

69

July 2011 | By Muhammad Mansoor Anwar

 70 Hydrocarbon Dew Point Control

Design parameters can be as followings

L/d 3 4 5

Diameter of
9 ft 8.5 ft 8 ft
Separator

L SS of Separator 31.5 ft 35.3 ft 39.8 ft

70

July 2011 | By Muhammad Mansoor Anwar

 71 Hydrocarbon Dew Point Control

• EG concentration

From GPSA Eq. 20-6

Where,
Mw = 62.1
KH = 2335
d = 10.2 0F
Thus,
XR = 0.37

Solving Eq. 20-8

Where,

XR = 0.37

XL = 0.8

m H2O = 100 * (W in – W out )

= 100 * (1.9292 – 1.8478)
= 8.13 lb/hr

Thus,

mI = 7.0 lb/hr

71

July 2011 | By Muhammad Mansoor Anwar

 72 Hydrocarbon Dew Point Control

2. J-T Valve based on HCDP

(With Sales gas compression)

• Size of Free water Knockout Vessel

(Vertical two phase separator)

Inlet gas flow rate = 25 MMSCFD

Gas density = 4.17 lb/ft3
Liquid density = 41.63 lb/ft3
Inlet Temperature = 100 °F
Inlet Pressure = 1200 psia
Compressibility factor = 0.8
Viscosity = 0.014 cP
Particle size = 100 μm
K value = 0.135 ft/sec (from API – 12 J)

72

July 2011 | By Muhammad Mansoor Anwar

 73 Hydrocarbon Dew Point Control

From GPSA Eq. 7-3

C`(Re)2 = 2598

From GPSA Fig 7-4, given below

C` = 1.7

From Eq. 7-1 of GPSA

73

Terminal Velocity = 0.314 ft/sec

July 2011 | By Muhammad Mansoor Anwar

 74 Hydrocarbon Dew Point Control

Gas flow rate can be calculated from the following Equation

QA = 4.40 ft3/sec

Vertical Separator Area is calculated by

Area of Separator = 14.01 ft2

Dia of Separator = 4.5 ft

Taking L/D = 3

Length of separator = 13.5 ft

Mist Extractor AREA

From Eq. 7-8 of GPSA

Vt = 0.465 ft/sec

From Eq. 7-9 of GPSA

Area of Extractor = 9.44 ft2

74

July 2011 | By Muhammad Mansoor Anwar

 75 Hydrocarbon Dew Point Control

• Thermal Design of the Gas/Gas Exchanger

(Shell and Tube Heat Exchanger)

1) Thermal Duty = 1.0 MMbtu/hr (Assumed for 1st iteration)

From Richardson & Coulson, Volume 6

Log mean Temperature Difference

Inlet hot fluid temperature, T 1 = 100 0F

Outlet hot fluid temperature, T 2 = 50 0F
Inlet cold fluid temperature, t 1 = 20 0F
Outlet cold fluid temperature, t 2 = 75 0F
Log mean Temperature Difference = 27.71 0F

Overall Heat Transfer Coefficient = 0.1476 Btu/hr.ft2. 0F

(Trial value From Richardson & Coulson Vol 6 below table)

Area required can be calculated by Equation

75

Area 1 = 244499 ft2 (according to 1st assumption)

July 2011 | By Muhammad Mansoor Anwar

 76 Hydrocarbon Dew Point Control

2) Thermal Duty = 1.5 MMbtu/hr (according to 2nd assumption)

Overall Heat Transfer Coefficient = 0.0592 Btu/hr.ft2. 0F

(Trial value From R&C Volume 6 above table)

Area 2 = 366748.5 ft2 (according to 2nd assumption)

3) Thermal Duty = 2.0 MMbtu/hr (according to 3rd assumption)

Overall Heat Transfer Coefficient = 0.0888 Btu/hr.ft2. 0F

(Trial value From R&C Volume 6 above table)

Area 3 = 488998 ft2 (according to 3rd assumption)

• I select 1.5 MMbtu/hr Thermal Duty because it is enough for current process conditions along
with a reasonable area of gas/gas heat exchanger

76

July 2011 | By Muhammad Mansoor Anwar

 77 Hydrocarbon Dew Point Control

• Size of Low Temperature Separator

(Horizontal three phase separator)

Inlet Gas flow = 25 MMSCFD

Inlet Temperature = 20 0F
Inlet Pressure = 640 psi
Liquid retention time = 20 min (From GPSA table as given below)

K value = 0.32 ft/sec (from API – 12 J)

77

July 2011 | By Muhammad Mansoor Anwar

 78 Hydrocarbon Dew Point Control

Design parameters can be as following;

L/d 3 4 5

Diameter of
9 ft 8.5 ft 8 ft
Separator

L SS of Separator 31.5 ft 35.3 ft 39.8 ft

• Compressor Horsepower

Inlet gas flow = 25 MMscfd

Inlet Pressure = 630 psi
Outlet pressure required = 1205 psi

Compressor power can be calculated by Eq. 13-4 of GPSA

Where,
No. of stages = 1
Compression ratio = 1.9
F (for single stage) = 1

78
Thus,
Compressor power = 1045 hp

July 2011 | By Muhammad Mansoor Anwar

 79 Hydrocarbon Dew Point Control

3. Mechanical Refrigeration Based Separation

The designing of the knockout Vessel and gas/gas exchanger is same as in last case due to
similar conditions

• Size of Low Temperature Separator

(Horizontal three phase separator)

Inlet Gas flow = 25 MMSCFD

Inlet Temperature = -10 0F
Inlet Pressure = 1185 psi
Liquid retention time = 20 min (From GPSA table as below)

K value = 0.30 ft/sec (from API – 12 J)

79

July 2011 | By Muhammad Mansoor Anwar

 80 Hydrocarbon Dew Point Control

Design parameters can be as followings

L/d 3 4 5

Diameter of
9 ft 8.5 ft 8 ft
Separator

L SS of Separator 31.4 ft 35.2 ft 39.7 ft

• Compressor Horsepower

Inlet gas flow = 25 MMscfd

Inlet Pressure = 1180 psi
Outlet pressure required = 1205 psi

Compressor power can be calculated by Eq. 13-4 of GPSA

Where,
No. of stages = 1
Compression ratio = 1.02
F (for single stage) = 1

80
Thus,
Compressor power = 561 hp

July 2011 | By Muhammad Mansoor Anwar

 81 Hydrocarbon Dew Point Control

• EG concentration

From GPSA Eq 20-6

Where,
Mw = 62.1
KH = 2335
d = 21.12 0F
Thus,
XR = 0.35

Solving Eq 20-8

Where,
XR = 0.35
XL = 0.8

m H2O = 100 * (W in – W out )

= 100 * (1.9292 – 1.8778)
= 5.14 lb/hr
Thus,

mI = 4.0 lb/hr

81

July 2011 | By Muhammad Mansoor Anwar

 82 Hydrocarbon Dew Point Control

Comparison of Actual Results

Process J-T Valve Based on J-T Valve Based on Mechanical

Name HCDP UNIT - 1 HCDP UNIT - 2 Refrigeration

Gas/Gas Exchanger
Duty (MMbtu/hr) 1.5 2.0 1.5

Diameter of Free Water

Knockout Drum 4.5 3.5 4.5
(ft)
Length of Free Water
Knockout Drum 13.5 10.5 13.5
(ft)
Diameter of Low
Temperature Separator 9.0 8.5 8.0
(ft)
Length of Low
Temperature Separator 31.5 40 35
(ft)
Compressor Power
(hp) 1045 1837 561

Log Mean Temperature

of gas/gas heat 27.71 70.56 27.71
exchanger (0F)
Concentration Ethylene
Glycol added 2.0 7.0 4.0
(lb/hr)
HCDP of Sales gas
(0F) 24 24 32

Recommended Option
I recommend that in these processes, Mechanical Refrigeration will be a better option due to its
lower gas/gas exchanger duty, lesser compressor horsepower and less separators length and
diameter.

82

July 2011 | By Muhammad Mansoor Anwar

 83 Hydrocarbon Dew Point Control

Comparison of HYSYS Results

Process J-T Valve Based on J-T Valve Based on Mechanical

Name HCDP UNIT - 1 HCDP UNIT - 2 Refrigeration

Gas/Gas Exchanger
Duty (MMbtu/hr) 1.457 1.77 1.464

Gas/Gas Exchanger
Approach (0F) 26.05 69.59 59.95

Pressure Drop across JT

valve / Chiller (psi) 545 2780 5

Temp of Gas coming

out of LTS (0F) 19.47 -21.43 -10

Sales Gas Out Temp 120

(0F) 120 30.30

Flow rate of Liquid out

of LTS (bbl/day) 2.806 3.261 0.7688

Flow rate of Ethylene

Glycol out of LTS (GPM) 0.0023 0.0097 0.0023

HCDP of Sales gas

(0F) 23.58 24.52 30.13

Hydrate Formation
Temp upstream of
4.4 11.23 11.16
gas/gas heat exchanger
(0F)

 The above results are concluded from ASPEN HYSYS simulations which are given below

83

July 2011 | By Muhammad Mansoor Anwar

 JT Valve
based on
HCDP unit 1

LTS
Feed VAP
Std Gas Flow 25.00 MMSCFD
Temperature 100.0 F
Pressure 1200 psia

NGL
Sep
Vap mixer cold LTS Liq Vol Flow @Std Cond 2.758 barrel/day
out gas/gas gas JT Feed
Feed MIX-100 VALVE
HEX
EG NGL
Free IN LTS
Water gas/gas HEX
knockout Duty -1.457e+006 Btu/hr
Vessel

Sep Gas
Liq out
shaft
Compressor work
RICH
Glycol
Gas out shaft work
73.72 F 989.3 hp Sales gas

630.0 psia Air Temperature 120.0 F

cooler
Pressure 1200 psia
HCDP
compresed Sales Molar Flow 25.00 MMSCFD
gas out gas HCDP A1: 23.56 F

Mon Jul 25 12:41:04 2011 Case: D:\Assignments\HCDP\JT vlave Based 1.hsc Flowsheet: Case (Main)
1
Case Name: D:\Assignments\HCDP\JT vlave Based 1.hsc
2 LEGENDS
3 Calgary, Alberta Unit Set: Amine
4 CANADA
Date/Time: Mon Jul 25 12:42:44 2011
5
6
7
Workbook: Case (Main)
8
9
10
Material Streams Fluid Pkg: All

11 Name Feed Sep Vap cold gas LTS VAP Gas out
12 Vapour Fraction 1.0000 1.0000 1.0000 1.0000 1.0000
13 Temperature (F) 100.0 * 99.80 50.00 * 19.48 73.72
14 Pressure (psia) 1200 * 1195 1185 640.0 630.0
15 Molar Flow (MMSCFD) 25.00 * 25.00 25.00 25.00 25.00
16 Mass Flow (lb/hr) 5.099e+004 5.099e+004 5.099e+004 5.096e+004 5.096e+004
17 Liquid Volume Flow (USGPM) 282.3 282.3 282.3 282.2 282.2
18 Heat Flow (Btu/hr) -8.904e+007 -8.904e+007 -9.051e+007 -9.047e+007 -8.902e+007
19 Master Comp Mole Frac (Nitrogen) 0.1273 * 0.1273 0.1273 0.1273 0.1273
20 Master Comp Mole Frac (CO2) 0.0286 * 0.0286 0.0286 0.0286 0.0286
21 Master Comp Mole Frac (Methane) 0.8341 * 0.8341 0.8341 0.8342 0.8342
22 Master Comp Mole Frac (Ethane) 0.0058 * 0.0058 0.0058 0.0058 0.0058
23 Master Comp Mole Frac (Propane) 0.0013 * 0.0013 0.0013 0.0013 0.0013
24 Master Comp Mole Frac (i-Butane) 0.0005 * 0.0005 0.0005 0.0005 0.0005
25 Master Comp Mole Frac (i-Pentane) 0.0002 * 0.0002 0.0002 0.0002 0.0002
26 Master Comp Mole Frac (n-Pentane) 0.0001 * 0.0001 0.0001 0.0001 0.0001
27 Master Comp Mole Frac (n-Hexane) 0.0003 * 0.0003 0.0003 0.0003 0.0003
28 Master Comp Mole Frac (n-Octane) 0.0002 * 0.0002 0.0002 0.0001 0.0001
29 Master Comp Mole Frac (n-Nonane) 0.0000 * 0.0000 0.0000 0.0000 0.0000
30 Master Comp Mole Frac (EGlycol) 0.0000 * 0.0000 0.0000 0.0000 0.0000
31 Master Comp Mole Frac (TEGlycol) 0.0000 * 0.0000 0.0000 0.0000 0.0000
32 Master Comp Mole Frac (H2O) 0.0000 * 0.0000 0.0000 0.0000 0.0000
33 Master Comp Mole Frac (Helium) 0.0012 * 0.0012 0.0012 0.0012 0.0012
34 Master Comp Mole Frac (n-Hex-CC5) *** *** *** *** ***
35 Master Comp Mole Frac (n-Heptane) 0.0003 * 0.0003 0.0003 0.0003 0.0003
36 Name Sep Liq LTS Feed NGL RICH Glycol compresed gas out
37 Vapour Fraction 0.0000 0.9999 0.0000 0.0000 1.0000
38 Temperature (F) 99.80 19.48 19.48 19.48 188.8
39 Pressure (psia) 1195 640.0 * 640.0 640.0 1205 *
40 Molar Flow (MMSCFD) 0.0000 25.00 2.852e-003 3.558e-004 25.00
41 Mass Flow (lb/hr) 0.0000 5.099e+004 27.20 1.818 5.096e+004
42 Liquid Volume Flow (USGPM) 0.0000 282.3 8.184e-002 3.331e-003 282.2
43 Heat Flow (Btu/hr) 0.0000 -9.051e+007 -2.851e+004 -6677 -8.650e+007
44 Master Comp Mole Frac (Nitrogen) 0.0029 0.1273 0.0091 0.0004 0.1273
45 Master Comp Mole Frac (CO2) 0.0227 0.0286 0.0210 0.0113 0.0286
46 Master Comp Mole Frac (Methane) 0.0000 0.8341 0.2036 0.0001 0.8342
47 Master Comp Mole Frac (Ethane) 0.0000 0.0058 0.0086 0.0000 0.0058
48 Master Comp Mole Frac (Propane) 0.0000 0.0013 0.0072 0.0000 0.0013
49 Master Comp Mole Frac (i-Butane) 0.0000 0.0005 0.0067 0.0000 0.0005
50 Master Comp Mole Frac (i-Pentane) 0.0000 0.0002 0.0117 0.0000 0.0002
51 Master Comp Mole Frac (n-Pentane) 0.0000 0.0001 0.0076 0.0000 0.0001
52 Master Comp Mole Frac (n-Hexane) 0.0000 0.0003 0.0661 0.0000 0.0003
53 Master Comp Mole Frac (n-Octane) 0.0000 0.0002 0.3268 0.0000 0.0001
54 Master Comp Mole Frac (n-Nonane) 0.0000 0.0000 0.1353 0.0000 0.0000
55 Master Comp Mole Frac (EGlycol) 0.0000 0.0000 0.0000 0.6406 0.0000
56 Master Comp Mole Frac (TEGlycol) 0.0000 0.0000 0.0000 0.0000 0.0000
57 Master Comp Mole Frac (H2O) 0.9743 0.0000 0.0000 0.3476 0.0000
58 Master Comp Mole Frac (Helium) 0.0000 0.0012 0.0001 0.0000 0.0012
59 Master Comp Mole Frac (n-Hex-CC5) *** *** *** *** ***
60 Master Comp Mole Frac (n-Heptane) 0.0000 0.0003 0.1963 0.0000 0.0003
61
62
63
64
65
66
67
68
69 Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728) Page 1 of 2
Licensed to: LEGENDS * Specified by user.
1
Case Name: D:\Assignments\HCDP\JT vlave Based 1.hsc
2 LEGENDS
3 Calgary, Alberta Unit Set: Amine
4 CANADA
Date/Time: Mon Jul 25 12:42:44 2011
5
6
7
Workbook: Case (Main) (continued)
8
9
10
Material Streams (continued) Fluid Pkg: All

11 Name mixer out EG IN Sales gas

12 Vapour Fraction 1.0000 0.0000 1.0000
13 Temperature (F) 99.76 100.0 * 120.0 *
14 Pressure (psia) 1195 1250 * 1200
15 Molar Flow (MMSCFD) 25.00 4.370e-004 25.00
16 Mass Flow (lb/hr) 5.099e+004 2.000 * 5.096e+004
17 Liquid Volume Flow (USGPM) 282.3 3.677e-003 282.2
18 Heat Flow (Btu/hr) -8.905e+007 -7705 -8.844e+007
19 Master Comp Mole Frac (Nitrogen) 0.1273 0.0000 * 0.1273
20 Master Comp Mole Frac (CO2) 0.0286 0.0000 * 0.0286
21 Master Comp Mole Frac (Methane) 0.8341 0.0000 * 0.8342
22 Master Comp Mole Frac (Ethane) 0.0058 0.0000 * 0.0058
23 Master Comp Mole Frac (Propane) 0.0013 0.0000 * 0.0013
24 Master Comp Mole Frac (i-Butane) 0.0005 0.0000 * 0.0005
25 Master Comp Mole Frac (i-Pentane) 0.0002 0.0000 * 0.0002
26 Master Comp Mole Frac (n-Pentane) 0.0001 0.0000 * 0.0001
27 Master Comp Mole Frac (n-Hexane) 0.0003 0.0000 * 0.0003
28 Master Comp Mole Frac (n-Octane) 0.0002 0.0000 * 0.0001
29 Master Comp Mole Frac (n-Nonane) 0.0000 0.0000 * 0.0000
30 Master Comp Mole Frac (EGlycol) 0.0000 0.5372 * 0.0000
31 Master Comp Mole Frac (TEGlycol) 0.0000 0.0000 * 0.0000
32 Master Comp Mole Frac (H2O) 0.0000 0.4628 * 0.0000
33 Master Comp Mole Frac (Helium) 0.0012 0.0000 * 0.0012
34 Master Comp Mole Frac (n-Hex-CC5) *** *** ***
35 Master Comp Mole Frac (n-Heptane) 0.0003 0.0000 * 0.0003
36
37
38
39
40
41
42
43
44
45
46
47
48
49
50
51
52
53
54
55
56
57
58
59
60
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69 Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728) Page 2 of 2
Licensed to: LEGENDS * Specified by user.
 LTS
Feed VAP
Std Gas Flow 25.00 MMSCFD
Temperature 100.0 F
gas/gas HEX
Pressure 1200 psia
Duty -1.776e+006 Btu/hr
AIR NGL
Cooler Sep
Vap mixer cold LTS Liq Vol Flow @Std Cond 3.247 barrel/day
Feed SEP out gas/gas gas JT Feed
compresed MIX-100 VALVE
gas out IN HEX
EG NGL
Free IN
Compressor Water LTS
knockout
Vessel
shaft
work
Sep
shaft work Liq
2092 hp

RICH
Glycol

HCDP
Sales
gas A1: 24.52 F
HCDP
Sales gas
Temperature 30.30 F
Pressure 1190 psia
Molar Flow 25.00 MMSCFD

Mon Jul 25 12:46:03 2011 Case: D:\Assignments\HCDP\JT vlave Based 2.hsc Flowsheet: Case (Main)
1
Case Name: D:\Assignments\HCDP\JT vlave Based 2.hsc
2 LEGENDS
3 Calgary, Alberta Unit Set: Amine
4 CANADA
Date/Time: Mon Jul 25 12:46:31 2011
5
6
7
Workbook: Case (Main)
8
9
10
Material Streams Fluid Pkg: All

11 Name Feed Sep Vap cold gas LTS VAP Sep Liq
12 Vapour Fraction 1.0000 1.0000 1.0000 1.0000 0.0000
13 Temperature (F) 100.0 * 99.95 50.00 * -21.43 99.95
14 Pressure (psia) 1200 * 3990 3980 1200 3990
15 Molar Flow (MMSCFD) 25.00 * 25.00 25.00 25.00 0.0000
16 Mass Flow (lb/hr) 5.099e+004 5.099e+004 5.099e+004 5.095e+004 0.0000
17 Liquid Volume Flow (USGPM) 282.3 282.3 282.3 282.2 0.0000
18 Heat Flow (Btu/hr) -8.904e+007 -9.112e+007 -9.292e+007 -9.285e+007 0.0000
19 Master Comp Mole Frac (Nitrogen) 0.1273 * 0.1273 0.1273 0.1273 0.0086
20 Master Comp Mole Frac (CO2) 0.0286 * 0.0286 0.0286 0.0286 0.0278
21 Master Comp Mole Frac (Methane) 0.8341 * 0.8341 0.8341 0.8342 0.0001
22 Master Comp Mole Frac (Ethane) 0.0058 * 0.0058 0.0058 0.0058 0.0000
23 Master Comp Mole Frac (Propane) 0.0013 * 0.0013 0.0013 0.0013 0.0000
24 Master Comp Mole Frac (i-Butane) 0.0005 * 0.0005 0.0005 0.0005 0.0000
25 Master Comp Mole Frac (i-Pentane) 0.0002 * 0.0002 0.0002 0.0002 0.0000
26 Master Comp Mole Frac (n-Pentane) 0.0001 * 0.0001 0.0001 0.0001 0.0000
27 Master Comp Mole Frac (n-Hexane) 0.0003 * 0.0003 0.0003 0.0003 0.0000
28 Master Comp Mole Frac (n-Octane) 0.0002 * 0.0002 0.0002 0.0001 0.0000
29 Master Comp Mole Frac (n-Nonane) 0.0000 * 0.0000 0.0000 0.0000 0.0000
30 Master Comp Mole Frac (EGlycol) 0.0000 * 0.0000 0.0000 0.0000 0.0000
31 Master Comp Mole Frac (TEGlycol) *** *** *** *** ***
32 Master Comp Mole Frac (H2O) 0.0000 * 0.0000 0.0000 0.0000 0.9633
33 Master Comp Mole Frac (Helium) 0.0012 * 0.0012 0.0012 0.0012 0.0002
34 Master Comp Mole Frac (n-Hex-CC5) *** *** *** *** ***
35 Master Comp Mole Frac (n-Heptane) 0.0003 * 0.0003 0.0003 0.0003 0.0000
36 Name LTS Feed NGL RICH Glycol compresed gas out mixer out
37 Vapour Fraction 0.9998 0.0000 0.0000 1.0000 1.0000
38 Temperature (F) -21.43 -21.43 -21.43 325.3 99.90
39 Pressure (psia) 1200 * 1200 1200 4000 * 3990
40 Molar Flow (MMSCFD) 25.00 4.085e-003 2.002e-003 25.00 25.00
41 Mass Flow (lb/hr) 5.099e+004 29.25 8.712 5.099e+004 5.099e+004
42 Liquid Volume Flow (USGPM) 282.3 9.511e-002 1.620e-002 282.3 282.3
43 Heat Flow (Btu/hr) -9.292e+007 -3.470e+004 -3.529e+004 -8.372e+007 -9.114e+007
44 Master Comp Mole Frac (Nitrogen) 0.1273 0.0193 0.0007 0.1273 0.1273
45 Master Comp Mole Frac (CO2) 0.0286 0.0394 0.0213 0.0286 0.0286
46 Master Comp Mole Frac (Methane) 0.8341 0.3962 0.0000 0.8341 0.8341
47 Master Comp Mole Frac (Ethane) 0.0058 0.0140 0.0000 0.0058 0.0058
48 Master Comp Mole Frac (Propane) 0.0013 0.0101 0.0000 0.0013 0.0013
49 Master Comp Mole Frac (i-Butane) 0.0005 0.0081 0.0000 0.0005 0.0005
50 Master Comp Mole Frac (i-Pentane) 0.0002 0.0119 0.0000 0.0002 0.0002
51 Master Comp Mole Frac (n-Pentane) 0.0001 0.0077 0.0000 0.0001 0.0001
52 Master Comp Mole Frac (n-Hexane) 0.0003 0.0572 0.0000 0.0003 0.0003
53 Master Comp Mole Frac (n-Octane) 0.0002 0.2098 0.0000 0.0002 0.0002
54 Master Comp Mole Frac (n-Nonane) 0.0000 0.0811 0.0000 0.0000 0.0000
55 Master Comp Mole Frac (EGlycol) 0.0000 0.0000 0.4781 0.0000 0.0000
56 Master Comp Mole Frac (TEGlycol) *** *** *** *** ***
57 Master Comp Mole Frac (H2O) 0.0000 0.0000 0.4998 0.0000 0.0000
58 Master Comp Mole Frac (Helium) 0.0012 0.0002 0.0000 0.0012 0.0012
59 Master Comp Mole Frac (n-Hex-CC5) *** *** *** *** ***
60 Master Comp Mole Frac (n-Heptane) 0.0003 0.1451 0.0000 0.0003 0.0003
61
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69 Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728) Page 1 of 2
Licensed to: LEGENDS * Specified by user.
1
Case Name: D:\Assignments\HCDP\JT vlave Based 2.hsc
2 LEGENDS
3 Calgary, Alberta Unit Set: Amine
4 CANADA
Date/Time: Mon Jul 25 12:46:31 2011
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Workbook: Case (Main) (continued)
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Material Streams (continued) Fluid Pkg: All

11 Name EG IN Sales gas SEP IN

12 Vapour Fraction 0.0000 1.0000 1.0000
13 Temperature (F) 100.0 * 30.30 100.0 *
14 Pressure (psia) 4100 * 1190 3995
15 Molar Flow (MMSCFD) 1.197e-003 25.00 25.00
16 Mass Flow (lb/hr) 7.000 * 5.095e+004 5.099e+004
17 Liquid Volume Flow (USGPM) 1.268e-002 282.2 282.3
18 Heat Flow (Btu/hr) -2.351e+004 -9.107e+007 -9.112e+007
19 Master Comp Mole Frac (Nitrogen) 0.0000 * 0.1273 0.1273
20 Master Comp Mole Frac (CO2) 0.0000 * 0.0286 0.0286
21 Master Comp Mole Frac (Methane) 0.0000 * 0.8342 0.8341
22 Master Comp Mole Frac (Ethane) 0.0000 * 0.0058 0.0058
23 Master Comp Mole Frac (Propane) 0.0000 * 0.0013 0.0013
24 Master Comp Mole Frac (i-Butane) 0.0000 * 0.0005 0.0005
25 Master Comp Mole Frac (i-Pentane) 0.0000 * 0.0002 0.0002
26 Master Comp Mole Frac (n-Pentane) 0.0000 * 0.0001 0.0001
27 Master Comp Mole Frac (n-Hexane) 0.0000 * 0.0003 0.0003
28 Master Comp Mole Frac (n-Octane) 0.0000 * 0.0001 0.0002
29 Master Comp Mole Frac (n-Nonane) 0.0000 * 0.0000 0.0000
30 Master Comp Mole Frac (EGlycol) 0.8000 * 0.0000 0.0000
31 Master Comp Mole Frac (TEGlycol) *** *** ***
32 Master Comp Mole Frac (H2O) 0.2000 * 0.0000 0.0000
33 Master Comp Mole Frac (Helium) 0.0000 * 0.0012 0.0012
34 Master Comp Mole Frac (n-Hex-CC5) *** *** ***
35 Master Comp Mole Frac (n-Heptane) 0.0000 * 0.0003 0.0003
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 1.58e06
kj/hr

Mechical
4 1
Refrigeration Condensor
BASED
HCDP UNIT comp J-T
Comp HP
Ref

3 2
LTS
VAP E-102

Feed
Std Gas Flow 25.00 MMSCFD
Q-100
Temperature 100.0 F
Pressure 1200 psia
LTS NGL
Sep
Vap HEX HEX Feed Liq Vol Flow @Std Cond 0.7620 barrel/day
IN gas/gas OUT cold
EG HEX gas
mixer E-100
EG EG
Feed 1 EG mixer NGL
in IN 2 LTS
2
Free
Water gas/gas HEX
knockout
Vessel Duty -1.466e+006 Btu/hr
shaft work
Sep
Liq 24.27 hp

shaft
work RICH
Glycol
Gas
out

Gas out Compressor

34.20 F HCDP
1180 psia A1: 30.13 F
energy HCDP
in

Compressed sales
gas E-101 gas

sales gas
Temperature 120.0 F
Pressure 1200 psia
Molar Flow 25.00 MMSCFD

Mon Jul 25 12:48:49 2011 Case: D:\Assignments\HCDP\MECH REFRIGERATION.hsc Flowsheet: Case (Main)
1
Case Name: D:\Assignments\HCDP\MECH REFRIGERATION.hsc
2 LEGENDS
3 Calgary, Alberta Unit Set: Amine
4 CANADA
Date/Time: Mon Jul 25 12:49:06 2011
5
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Workbook: Case (Main)
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Material Streams Fluid Pkg: All

11 Name Feed Sep Vap cold gas LTS VAP Gas out
12 Vapour Fraction 1.0000 1.0000 1.0000 1.0000 1.0000
13 Temperature (F) 100.0 * 99.80 50.00 * -10.00 34.20
14 Pressure (psia) 1200 * 1195 1190 1185 1180
15 Molar Flow (MMSCFD) 25.00 * 25.00 25.00 25.00 25.00
16 Mass Flow (lb/hr) 5.099e+004 5.099e+004 5.099e+004 5.098e+004 5.098e+004
17 Liquid Volume Flow (USGPM) 282.3 282.3 282.3 282.3 282.3
18 Heat Flow (Btu/hr) -8.904e+007 -8.904e+007 -9.052e+007 -9.243e+007 -9.096e+007
19 Master Comp Mole Frac (Nitrogen) 0.1273 * 0.1273 0.1273 0.1273 0.1273
20 Master Comp Mole Frac (CO2) 0.0286 * 0.0286 0.0286 0.0286 0.0286
21 Master Comp Mole Frac (Methane) 0.8341 * 0.8341 0.8341 0.8342 0.8342
22 Master Comp Mole Frac (Ethane) 0.0058 * 0.0058 0.0058 0.0058 0.0058
23 Master Comp Mole Frac (Propane) 0.0013 * 0.0013 0.0013 0.0013 0.0013
24 Master Comp Mole Frac (i-Butane) 0.0005 * 0.0005 0.0005 0.0005 0.0005
25 Master Comp Mole Frac (i-Pentane) 0.0002 * 0.0002 0.0002 0.0002 0.0002
26 Master Comp Mole Frac (n-Pentane) 0.0001 * 0.0001 0.0001 0.0001 0.0001
27 Master Comp Mole Frac (n-Hexane) 0.0003 * 0.0003 0.0003 0.0003 0.0003
28 Master Comp Mole Frac (n-Octane) 0.0002 * 0.0002 0.0002 0.0002 0.0002
29 Master Comp Mole Frac (n-Nonane) 0.0000 * 0.0000 0.0000 0.0000 0.0000
30 Master Comp Mole Frac (TEGlycol) *** *** *** *** ***
31 Master Comp Mole Frac (H2O) 0.0000 * 0.0000 0.0000 0.0000 0.0000
32 Master Comp Mole Frac (Helium) 0.0012 * 0.0012 0.0012 0.0012 0.0012
33 Master Comp Mole Frac (n-Hex-CC5) *** *** *** *** ***
34 Master Comp Mole Frac (n-Heptane) 0.0003 * 0.0003 0.0003 0.0003 0.0003
35 Master Comp Mole Frac (EGlycol) 0.0000 * 0.0000 0.0000 0.0000 0.0000
36 Name Sep Liq LTS Feed NGL RICH Glycol sales gas
37 Vapour Fraction 0.0000 0.9999 0.0000 0.0000 1.0000
38 Temperature (F) 99.80 -10.00 * -10.00 -10.00 120.0 *
39 Pressure (psia) 1195 1185 1185 1185 1200
40 Molar Flow (MMSCFD) 0.0000 25.00 9.295e-004 1.288e-003 25.00
41 Mass Flow (lb/hr) 0.0000 5.099e+004 7.003 5.252 5.098e+004
42 Liquid Volume Flow (USGPM) 0.0000 282.3 2.243e-002 9.800e-003 282.3
43 Heat Flow (Btu/hr) 0.0000 -9.246e+007 -8099 -2.208e+004 -8.845e+007
44 Master Comp Mole Frac (Nitrogen) 0.0029 0.1273 0.0184 0.0007 0.1273
45 Master Comp Mole Frac (CO2) 0.0227 0.0286 0.0364 0.0166 0.0286
46 Master Comp Mole Frac (Methane) 0.0000 0.8341 0.3723 0.0000 0.8342
47 Master Comp Mole Frac (Ethane) 0.0000 0.0058 0.0130 0.0000 0.0058
48 Master Comp Mole Frac (Propane) 0.0000 0.0013 0.0093 0.0000 0.0013
49 Master Comp Mole Frac (i-Butane) 0.0000 0.0005 0.0074 0.0000 0.0005
50 Master Comp Mole Frac (i-Pentane) 0.0000 0.0002 0.0110 0.0000 0.0002
51 Master Comp Mole Frac (n-Pentane) 0.0000 0.0001 0.0071 0.0000 0.0001
52 Master Comp Mole Frac (n-Hexane) 0.0000 0.0003 0.0533 0.0000 0.0003
53 Master Comp Mole Frac (n-Octane) 0.0000 0.0002 0.2251 0.0000 0.0002
54 Master Comp Mole Frac (n-Nonane) 0.0000 0.0000 0.1059 0.0000 0.0000
55 Master Comp Mole Frac (TEGlycol) *** *** *** *** ***
56 Master Comp Mole Frac (H2O) 0.9743 0.0000 0.0000 0.5587 0.0000
57 Master Comp Mole Frac (Helium) 0.0000 0.0012 0.0002 0.0000 0.0012
58 Master Comp Mole Frac (n-Hex-CC5) *** *** *** *** ***
59 Master Comp Mole Frac (n-Heptane) 0.0000 0.0003 0.1407 0.0000 0.0003
60 Master Comp Mole Frac (EGlycol) 0.0000 0.0000 0.0000 0.4240 0.0000
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69 Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728) Page 1 of 2
Licensed to: LEGENDS * Specified by user.
1
Case Name: D:\Assignments\HCDP\MECH REFRIGERATION.hsc
2 LEGENDS
3 Calgary, Alberta Unit Set: Amine
4 CANADA
Date/Time: Mon Jul 25 12:49:06 2011
5
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Workbook: Case (Main) (continued)
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Material Streams (continued) Fluid Pkg: All

11 Name 1 3 2 4 Compressed gas

12 Vapour Fraction 0.0000 * 1.0000 * 0.4643 1.0000 1.0000
13 Temperature (F) 120.0 * -4.000 * -2.580 160.3 37.43
14 Pressure (psia) 243.5 35.41 36.43 248.6 1205 *
15 Molar Flow (MMSCFD) 4.356 4.356 4.356 4.356 25.00
16 Mass Flow (lb/hr) 2.109e+004 2.109e+004 2.109e+004 2.109e+004 5.098e+004
17 Liquid Volume Flow (USGPM) 83.14 83.14 83.14 83.14 282.3
18 Heat Flow (Btu/hr) -2.403e+007 -2.209e+007 -2.403e+007 -2.099e+007 -9.090e+007
19 Master Comp Mole Frac (Nitrogen) 0.0000 * 0.0000 0.0000 0.0000 0.1273
20 Master Comp Mole Frac (CO2) 0.0000 * 0.0000 0.0000 0.0000 0.0286
21 Master Comp Mole Frac (Methane) 0.0000 * 0.0000 0.0000 0.0000 0.8342
22 Master Comp Mole Frac (Ethane) 0.0000 * 0.0000 0.0000 0.0000 0.0058
23 Master Comp Mole Frac (Propane) 1.0000 * 1.0000 1.0000 1.0000 0.0013
24 Master Comp Mole Frac (i-Butane) 0.0000 * 0.0000 0.0000 0.0000 0.0005
25 Master Comp Mole Frac (i-Pentane) 0.0000 * 0.0000 0.0000 0.0000 0.0002
26 Master Comp Mole Frac (n-Pentane) 0.0000 * 0.0000 0.0000 0.0000 0.0001
27 Master Comp Mole Frac (n-Hexane) 0.0000 * 0.0000 0.0000 0.0000 0.0003
28 Master Comp Mole Frac (n-Octane) 0.0000 * 0.0000 0.0000 0.0000 0.0002
29 Master Comp Mole Frac (n-Nonane) 0.0000 * 0.0000 0.0000 0.0000 0.0000
30 Master Comp Mole Frac (TEGlycol) *** *** *** *** ***
31 Master Comp Mole Frac (H2O) 0.0000 * 0.0000 0.0000 0.0000 0.0000
32 Master Comp Mole Frac (Helium) 0.0000 * 0.0000 0.0000 0.0000 0.0012
33 Master Comp Mole Frac (n-Hex-CC5) *** *** *** *** ***
34 Master Comp Mole Frac (n-Heptane) 0.0000 * 0.0000 0.0000 0.0000 0.0003
35 Master Comp Mole Frac (EGlycol) 0.0000 * 0.0000 0.0000 0.0000 0.0000
36 Name EG in HEX IN HEX OUT EG IN 2
37 Vapour Fraction 0.0000 * 1.0000 1.0000 0.0000 *
38 Temperature (F) 100.0 * 99.77 49.95 792.6
39 Pressure (psia) 1250 * 1195 1190 1200 *
40 Molar Flow (MMSCFD) 3.420e-004 25.00 25.00 3.420e-004
41 Mass Flow (lb/hr) 2.000 * 5.099e+004 5.099e+004 2.000 *
42 Liquid Volume Flow (USGPM) 3.624e-003 282.3 282.3 3.624e-003
43 Heat Flow (Btu/hr) -6735 -8.905e+007 -9.052e+007 -5329
44 Master Comp Mole Frac (Nitrogen) 0.0000 * 0.1273 0.1273 0.0000 *
45 Master Comp Mole Frac (CO2) 0.0000 * 0.0286 0.0286 0.0000 *
46 Master Comp Mole Frac (Methane) 0.0000 * 0.8341 0.8341 0.0000 *
47 Master Comp Mole Frac (Ethane) 0.0000 * 0.0058 0.0058 0.0000 *
48 Master Comp Mole Frac (Propane) 0.0000 * 0.0013 0.0013 0.0000 *
49 Master Comp Mole Frac (i-Butane) 0.0000 * 0.0005 0.0005 0.0000 *
50 Master Comp Mole Frac (i-Pentane) 0.0000 * 0.0002 0.0002 0.0000 *
51 Master Comp Mole Frac (n-Pentane) 0.0000 * 0.0001 0.0001 0.0000 *
52 Master Comp Mole Frac (n-Hexane) 0.0000 * 0.0003 0.0003 0.0000 *
53 Master Comp Mole Frac (n-Octane) 0.0000 * 0.0002 0.0002 0.0000 *
54 Master Comp Mole Frac (n-Nonane) 0.0000 * 0.0000 0.0000 0.0000 *
55 Master Comp Mole Frac (TEGlycol) *** *** *** ***
56 Master Comp Mole Frac (H2O) 0.2000 * 0.0000 0.0000 0.2000 *
57 Master Comp Mole Frac (Helium) 0.0000 * 0.0012 0.0012 0.0000 *
58 Master Comp Mole Frac (n-Hex-CC5) *** *** *** ***
59 Master Comp Mole Frac (n-Heptane) 0.0000 * 0.0003 0.0003 0.0000 *
60 Master Comp Mole Frac (EGlycol) 0.8000 * 0.0000 0.0000 0.8000 *
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69 Hyprotech Ltd. Aspen HYSYS Version 2006 (20.0.0.6728) Page 2 of 2
Licensed to: LEGENDS * Specified by user.