A COMPUTER MODEL OF THE RECTISOLPROCESS USING THE ASPEN SIMULATOR

by

Rosalyn A. preston

SUBMITTED IN PARTIAL FULFILLMENTOF THE REQUIREMENTS FOR THE

DEGREE OF

MASTER OF SCIENCE INCHEMICAL ENGINEERING

at the

MASSACHUSETTS INSTITUTE OF TECHNOLOGY

December, 1981

Q Massachusetts Institute of Technology 1981

Signature of AuthorSignature of Author Department of Chemical Engineering

December 7, 1981

Certified byLawrence B. EvansThesis Supervisor

Accepted byGlenn C Williams

Chairman, Departmental Graduate Committee

ArchivesMASSACHUSETTS INSTiTUTE

OF TECHMflLOGY

JUN 1 13982

i IRRA!ES

A COMPUTER MODEL OF THE RECTISOLPROCESS USING THE ASPEN SIMULATOR

by

Rosalyn A. Preston

Submitted to the Department of Chemical Engineering

on December 7, 1981 in partial fulfillment of the

requirements for the Degree of Master of Science

in Chemical Engineering

ABSTRACT

A computer model of the Rectisol process for acidgas absorption has been developed with the ASPENprocess simulator. The model uses a modified ver-sion of the Redlich-Kwong-Soave equation-of-stateto represent the phase equilibria for the methanol-water-aromatics-acid gas system. This report de-scribes the development of the physical propertiesmodel and the flowsheet simulation. The work isbased on the CONOCO design for a commercial scalecoal gasification facility. The computer modellingeffort was carried out under contract to the Morgan-town Energy Technical Center, U. S. Department ofEnergy.

TABLE OF CONTENTS

ACKNOWLEDGEMENTS . . . . . . . . . . . . . .

I. INTRODUCTION . . . . . . . . . . . . . . . .

II. AC:ID GAS ABSORPTION PROCESSES . . . . . . .

General Features of Acid Gas AbsorptionProcesses . . . . . . . . . . . . . . .

Absorption . . . . . . . . . . . . .

Regeneration . . . . . .. . . . .

Solvent Selection . . . . . . . .

Heat Integration . . . . . . . . . .

Solvent Recovery . . . . . . . .a

Advantages of the Rectisol Process . . .

Rectisol Process Description . . . . . .

III. DEVELOPMENT OF PHYSICAL PROPERTIES MODEL ..

Selection of the Thermodynamic Model . .

vii.. . 6. a . 6

· · · 6

a1 a a 7

a a a 7* a a 9

* a a 10

. . . 12

. a a 12

* a a 13

. a a 19

a . . 19

Modified Redlich-Kwong-Soave Equation-of-State

Experimental Phase Equilibrium Data . . . . .

Development of Modified Redlich-Kwong-SoaveModel Parameters . . . . . . . . . . . . . . .

Evaluation of the Preliminary ThermodynamicModel . . . . . . . . .. . .. . .. . . .

Enthalpy Calculations . . . . . . . . . . .

IV. SIMULATION OF THE RECTISOL PROCESS . . . . . . . .

Basis for the Process Simulation . . . . . . .

Absorption Section . . . . . . . . . . . .

Flash Regeneration Section . . . . . . . .

Naphtha Extractor . . . . . . ... . . .

Azeotrope Column .... . . .a a .a

Methanol-Water Column . . . . . . . . .

iii

21

23

25

27

32

36

36

38

43

43

44

49

. . .

Hot Regeneration Column . .

Water Wash Column . . ..

Heat Exchangers . . . . . .

Methanol Make-up . . . . .

Flowsheet Convergence . ....V. RECTISOL PROCESS SIMULATION RESULTSVI. CONCLUSION. ........ ...

APPENDIX A OVERALL RECTISOL MODEL

APPENDIX B ABSORPTION SECTION SIM

APPENDIX C AZEOTROPE COLUMN SIMUIL

APPENDIX D METHANOL-WATER COLUMN

APPENDIX E HOT REGENERATION COLUMI

APPENDIX F WATER WASH COLUMN SIMU

LITERATURE CITED

* O . . . . ·

a 0 0 · a a a

* a a a a e a a a a a a ·· · ··

U a a a a a

[LATION

MTION

SIMULATION

[ SIMULATION

LATION

BIBLI OGRAPHY

iv

50

52

52

53

53

55

59a

LIST OF FIGURES

1. The Pipeline Gas Demonstration Plant . . . . . . . 2

2. The Rectisol Process . . . . . . . . . . . . . . 4

3. Comparison of Equilibrium Loading of Chemical andPhysical Solvents . . . .. . . . . . . . . . 11

4. Block Flow Diagram of the Rectisol Process . . . . 14

5. Process Flow Diagram of the Rectisol Process . . . 15

6. Process Flow Diagram of the Rectisol Process . . . 16

7. Henry's Law Constants for Nitrogen in a Methanoland Carbon Dioxide Mixture . . . . . . . . . . . . 30

8. Block Flow Diagram of the Rectisol Computer Model 39

9. Effect of Solvent Rate and Number of Trays onCarbon Dioxide Absorption . . . . . . . . . . . 42

10. Ternary Diagram for Methanol-Benzene-Water . . . . 45

11. McCabe-Thiele Plot for Azeotrope Column . . . . . 48

v

LIST OF TABLES

1. Comparison of Modified RKS Model Predictions withExperimental Data for Methanol-Hydrogen Sulfide-Carbon Dioxide- Nitrogen . . . . . . . . . . . . . 33

2. Binary Interaction Parameters for the ModifiedRedlich-Kwong-Soave Equation-of-State ......34

3. Composition of Feed Gas to Rectisol Process . . . 37

4. Desulfurized Gas Composition . o . . .....56

5. Acid Gas Composition . . . . ... . . . .. . . 56

6. Regenerated Methanol Composition . ... . . . . . 56

7. Cooling and Heating Duties for the Rectisol Process 58

vi

ACKNOWLEDGMENTS

I would like to express my appreciation to all the ASPEN

staff who gave me help and support on this project, especially

to Herb Britt and Chau-Chyun Chen for assisting me in almost

every aspect, to Joe Boston and Paul Mathias for their help

with modelling the physical properties, and to Willie Chan

and Fred Ziegler for their debugging efforts. I would also

like to thank my advisor, Professor L. B. Evans, forhis gui-

dance and optimism throughout.

vii

I. INTRODUCTION

The objective of this work was to develop a computer

model of the Rectisol process (Hochgesand, 1970; Scholz, 1969)

using the ASPEN process simulator. This simulation was pre-

pared as part of a program to transfer ASPEN technology to the

Morgantown Energy Technology Center (METC) of the United States

Department of Energy. The overall program involved develop-

ment of computer models for the major process units in the

CONOCO Slagging Lurgi Coal Gasification process.

The computer modelling effort took CONOCO's design for

a commercial scale coal gasification facility as the basis

(CONOCO, 1980). This facility, shown schematically in Figure

1, is designed to generate pipeline quality substitute natural

gas from coal. Briefly, the process gasifies coal with steam

and oxygen, generated by cryogenic air separation. The gas-

ification takes place at high pressure in a moving bed slagging

gasifier using Lurgi technology. Crude synthesis gas is re-

covered from the top of the gasifier, while the ash is with-

drawn in a molten slag form from the bottom. The synthesis

gas is subsequently directed to the Rectisol unit for purifi-

cation. Here, light oils and sulfur compounds are removed

from the gas and recovered. Following purification, the gases

are allowed to undergo water-gas shift and methanation reac-

tions simultaneously. Carbon dioxide produced in the shift

reaction must then be washed out with hot carbonate solution

to improve the BTU quality of the substitute natural gas. The

final gas treatment involves compression and drying.

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The Rectisol process itself features refrigerated meth-

anol as the solvent into which the acid gases in the synthesis

gas are absorbed. The process is comprised of an absorption

section followed by a series of regeneration processes in which

the rich solvent is purified and the absorbed compounds recov-

ered. The diagram in Figure 2 illustrates the particular

Rectisol design proposed for the CONOCO coal gasification

facility.

The Rectisol computer model described in this report

provides a stream-by-stream heat and material balance for the

process. In addition, the model predicts the composition of

the outlet streams for a given set of operating conditions.

Therefore, the model can be used to assess the effect of var-

ious process variables on the composition of the desulfurized

gas and acid gas products. The model also predicts the purity

of the regenerated methanol solvent for specific regeneration

operating conditions. Separate models of the absorbers and

distillation columns yield a rigorous tray-to-tray analysis

of each column. Overall, the Rectisol simulation represents

a flexible tool which may be applied in future design and sen-

sitivity studies and in trouble-shooting after process start-

up.

The input to the simulation is based on the process flow-

sheet and feed streams from the CONOCO commercial design.

Since Rectisol is a proprietary process licensed by Lurgi Min-

eraloltechnik GmbH, the CONOCO design did not provide any in-

formation on the actual process operating conditions. For

this reason, development of the computer model entailed a

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fair amount of design work to establish reasonable process

operating conditions.

The following chapters describe the development of the

Rectisol model. The work initially involved a literature

search to locate published process data regarding Rectisol.

Subsequently, considerable effort was devoted to developing

a good thermodynamic model of the system to provide adequate

physical properties predictions. The actual computer model-

ling work required testing of several ASPEN models which were

still in the development stage. Further, it was necessary to

perform sensitivity studies on the design variables in order

to establish suitable operating conditions. It was orig-

inally intended to use the final Rectisol model to examine

the sensitivity of the process to various design variables.

However, lack of computer funds curtailed the project once

the model was developed.

6

II. ACID GAS ABSORPTION PROCESSES

This section discusses the general features of acid gas

absorption processes in order to provide some background for

the detailed description of the Rectisol process. In addition,

the discussion serves to present the basic design variables

which must be considered in designing a process to remove acid

gases.

General Features of Acid Gas Absorption Processes

Removal of acid gases, primarily carbon dioxide and hydro-

gen sulfide, is important for many industrial plants producing

a gaseous intermediate or final product. In the production

of substitute natural gas by coal gasification, hydrogen is

generated together with a substantial quantity of carbon dioxide.

Since carbon dioxide has no heating value, it must be separated

out of the final pipeline gas to raise the overall BTU value.

The separation may occur at several points in the process.

Removal prior to shift conversion of the carbon monoxide and

water to hydrogen and carbon dioxide favors the equilibrium

hydrogen concentration. However, since carbon dioxide is pro-

duced in this reaction, further treating to remove carbon di-

oxide may be required. All sulfur bearing compounds, on the

other hand, must be removed before methanation to avoid poison-

ing the methanation catalyst.

Rectisol is just one example of the many acid gas absorp-

tion processes. In general, these processes feature an ab-

sorption section in which the acid gases are absorbed into a

liquid solvent, and a regeneration section where the acid gases

7

are recovered to yield a lean solvent of sufficient purity

for recycle. In addition, the processes usually include a

final wash to recover solvent carried over with the acid gas

stream. The details of each section are discussed below,

highlighting the major factors affecting each section.

Absorption

Acid gases are removed from the sour gas by absorption

into a liquid solvent. The gas stream is contacted with the

solvent in an absorber column where it flows counter-current

to the downflowing liquid. The column itself may contain pack-

ing or trays, as dictated by the economics of the specific

design. Plate columns generally operate at lower efficiency.

Thus, a packed column design, which requires less height, may

be preferable when a large concentration of acid gas is to be

absorbed (Kohl and Riesenfeld, 1979).

Most absorbers operate at low temperature where equilib-

rium favors absorption. However, the fact that the absorption

rate increases with an increase in temperature justifies oper-

ating some absorption processes at elevated temperature.

Regeneration

The solvent regeneration section serves two functions:

namely, to recover the absorbed acid gases from the solvent,

and to reduce the gas concentration in the solvent to a low

level prior to its recycle to the absorber. A recycled sol-

vent of higher purity will absorb a greater quantity of gas

per pass, thereby reducing the solvent recirculation rate.

The tradeoff must therefore be considered between regeneration

costs andthe costs associated with a higher recycle rate, such

8

as larger equipment and higher pumping requirements. Further,

the purity of the treated gas is limited by the degree of re-

generation. The gas will approach equilibrium with the fresh-

ly regenerated solvent which it contacts in the final stage

of the absorber (Hochgesand, 1970).

There are three primary methods of regeneration. The

choice is determined by the nature of the solvent and the de-

sired purity. Flash desorption represents the cheapest method

where the pressure of the solvent is let down in several stages.

The final stage is usually atmospheric. However, a vacuum

flash may be employed to remove the dissolved gases further.

If flash desorption is not adequate, the gases may be stripped

from the solvent following a single flash. Heat must be sup-

plied to the stripper to provide the heat of desorption and

to vaporize the stripping medium (often steam). This heat

is frequently supplied in the form of steam through the strip-

per reboiler. As such, the regeneration step can be a sig-

nificant energy consumer.

Hot regeneration provides the most complete regeneration

and consequently leads to the highest purity for the feed gas.

Gas solubility tends to decrease with increased temperature.

This effect is augmented by heating the solvent to its boil-

ing point so that solvent vapors provide some stripping action.

In this case, heat is required to raise the system to the sol-

vent boiling point, to supply the heat of desorption and fur-

ther to supply the heat of vaporization for the solvent vapors.

Although some of this heat may be obtained through heat ex-

change of the hot overhead gases and hot regenerated solvent

9

with the incoming solvent, heat exchanger losses make external

heating necessary (Hochgesand, 1970).

Solvent Selection

A variety of solvents are employed to absorb acid gases.

Despite the varied chemical nature of these solvents, all

should exhibit certain characteristics to render them suit-

able as solvents. Foremost, the solvent must have a high sol-

vency for the acid gas constituents at reasonable operating

conditions. In addition, the absorption process must be rever-

sible so that the solvent will release the gas components with

relative ease to yield both gas and solvent of high purity.

It is highly undesirable for the solvent to form an irreversi-

ble compound with any of the feed gas constituents since this

may lead to appreciable solvent losses. The solvent should

also have low volatility to prevent excessive solvent losses

in the gas stream by vaporization. Other desirable character-

istics include low viscosity for minimal pumping costs, good

heat transfer properties and low heat capacity to minimize

heating requirements, stability, and low corrosivity. Finally,

low cost and availability are obviously desirable (Nonhebel,

1972).

The acid gas components may be absorbed into the solvent

by physically dissolving or by reacting chemically with an

active solvent component. The absorption process involves

transfer of the acid gases from the gas phase, through the

gas-liquid interface, into the liquid phase. In the case of

a chemical solvent, the absorbed gases then react reversibly

to form compounds. Chemical and physical solvents have

10

several distinguishing characteristics which account for the

different absorption mechanisms. As the diagram in Figure 3

shows, the equilibrium loading of a physical solvent increases

in direct proportion to partial pressure, provided Henry's

Law applies. In contrast, the loading of a chemical solvent

rises rapidly at low gas partial pressure, but then reaches

a point of saturation beyond which little absorption occurs.

This distinction makes chemical solvents the best choice for

purifying gas streams at low gas partial pressures, and phys-

ical solvents the choice for purifying high pressure concen-

trated streams. The same behaviour dictates the method of

regeneration. Physical solvents are easily regenerated by

flash desorption, although stripping or hot regeneration may

be used for greater purity. However, a saturated chemical

solvent subjected to a pressure reduction will release only

a fraction of the total absorbed gas. Some form of reboiling

is required to effect adequate regeneration (Hochgesand, 1970).

In a chemical absorption system, chemical reactions in-

crease the liquid film coefficient, thereby raising the over-

all absorption rate (Kohl, Riesenfeld, 1970). Since chemical

absorption is more rapid than physical absorption, chemical

systems require fewer trays and thus smaller towers.

Heat Integration

Solvent regeneration usually takes place at higher temp-

erature than absorption. As a result, most acid gas absorption

processes are highly heat integrated. Often the rich solvent

may be preheated and the lean solvent cooled by heat exchange.

The absorption process is exothermic so that the solvent

11

Figure 3

COMPARISON OF EQUILIBRIUM LOADINGnp' rT4.-.MTTAT. ATTh T.TVTAT. (NT.'T/1XTr1Z'

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12

temperature is raised within the absorber. This means that

additional cooling beyond heat exchange is necessary.

Solvent Recovery

To minimize the solvent losses from the system, the gas

streams may be washed with water or other suitable wash sol-

utions to recover the vaporized solvent leaving the absorption

or regeneration sections.

Advantages of the Rectisol Process

The Rectisol process employs refrigerated methanol as a

physical solvent to absorb both hydrogen sulfide and carbon

dioxide. According to the literature, the absorption process

is carried out at low temperatures, generally between O°F and

-75°F. At this temperature level, both hydrogen sulfide and

carbon dioxide are highly soluble in the methanol, although

the solubility of hydrogen sulfide exceeds that of carbon di-

oxide by a factor of 5-6 (Hochgesand, 1970). Methanol has a

high vapor pressure compared to most solvents. Thus, the low

temperature is also desirable to reduce the methanol losses

due to vaporization.

Rectisol has several attractions supporting its applic-

ation in the CONOCO design. Firstly, Rectisol is ideal for

high pressure processes since the equilibrium loading of a

physical solvent is directly proportional to pressure. In the

CONOCO design, the synthesis gas comes directly from the gas-

ifier and is therefore at high pressure. Further, since Rect-

isol favors absorption of hydrogen sulfide over carbon dioxide,

the hydrogen sulfide may be totally removed with only a frac-

tion of the carbon dioxide. Such selective removal is important

13

for two reasons. Firstly, if the sulfur compounds are recov-

ered with a low concentration of carbon dioxide, the gas

stream may be handled directly in a sulfur recovery unit,

such as a Claus Plant, without overloading the sulfur plant

or without recourse to further separation. Secondly, since

carbon dioxide removal prior to shift conversion and methan-

ation is not necessary, utilities would be needlessly consumed

in removing a high proportion of the carbon dioxide at this

stage (Hochgesand, 1970). However, as mentioned earlier, it

may be attractive to remove a fraction of the carbon dioxide

on account of the equilibrium considerations associated with

the subsequent water-gas shift reaction.

Rectisol Process Description

This section discusses the specific design of the Rect-

isol process developed as part of the CONOCO Slagging Lurgi

Coal Gasification commercial design. Figure 4 shows a simpli-

fied block flow diagram of the process, indicating the major

process sections. A copy of the Rectisol process flowsheet,

taken from the CONOCO commercial design report,is displayed

in Figures 5 and 6. The details of each section are outlined

briefly below.

In the Rectisol unit, the synthesis gas produced in the

gasification section is treated to remove the undesirable

sulfur compounds and to recover any light oils in the gas.

Purified gas is then directed to the joint shift/methanation

section. The removed sulfur compounds proceed to the sulfur

recovery unit.

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The feed gas is first cooled to absorption conditions by

heat exchange with the purified gas and by supplemental refrig-

eration. The cooled gas then proceeds to the absorption sec-

tion where the sulfur compounds, naphtha, hydrogen cyanide and

water are physically absorbed with refrigerated methanol. The

absorption process occurs at high pressure and low temperature

in two stages; in the prewash section, all the naphtha, water

and hydrogen cyanide are removed with some of the acid gas,

while in the main wash section, the remaining sulfur compounds

and a portion of the carbon dioxide are absorbed.

The ensuing process steps are devoted to solvent regen-

eration in which acid gases are recovered from the methanol

to yield a lean solvent of sufficient purity for recirculation.

In the first step, the rich solvent streams from the prewash

and main wash absoprtion sections are separately flash regen-

erated. The pressure of the solvent is reduced causing a frac-

tion of the acid gases to desorb.

The liquid from the prewash flash regenerator is sent to

the naphtha extractor to recover the absorbed naphtha. Here

a liquid-liquid separation is attained on adding the water

rich bottoms from the water wash column. The high purity

naphtha layer is decanted and sent to storage. The remaining

methanol-water-naphtha liquid is fed to the azeotrope column.

This column produces an overhead composed of naphtha, carbon

dioxide, hydrogen sulfide and methanol. A mixture of methanol

and water is recovered in the bottoms which is subsequently

separated in the methanol-water column.

18

The hot regenerator serves to strip the main wash solvent

of the gases that were not desorbed during flash regeneration.

The methanol overhead from the methanol-water column is also

fed to this unit where the methanol vapors, in conjunction

with those generated in the reboiler, act as the stripping

medium. The acid gases are recovered in the overhead and com-

bined with the gases from flash regeneration, the azeotrope

column and the naphtha extractor. This total gas stream is

finally water washed to recover any vaporized methanol prior

to being sent to the sulfur plant.

A portion of the regenerated methanol from the hot regen-

eration column bottoms is used as reflux for the methanol-

water column. The remaining methanol is cooled by heat ex-

change with the naphtha extractor feed and the hot regeneration

column feed. After further refrigeration, the regenerated

methanol is recirculated to the absorption section.

19

III. DEVELOPMENT OF PHYSICAL PROPERTIES MODEL

Accurate representation of the phase equilibria was es-

sential to the simulation work in view of the lack of process

data. Moreover, the physical properties estimation was esp-

ecially important since the hydrogen sulfide-carbon dioxide-

nitrogen-hydrogen-methanol-water-aromatics system is highly

non-ideal. The presence of methanol and water, which are

significantly polar, and hydrogen, nitrogen, carbon monoxide

and methane, which are supercritical, render this system hard

to describe. Therefore, considerable effort was devoted to

developing a good thermodynamic model.

In order to simplify the system, all the sulfur compounds

are treated as hydrogen sulfide, and the C2, C3, C4 compounds

are lumped into the methane component throughout this work.

The naphtha in the feed is represented as benzene since coal-

derived liquids are generally aromatic. These simplifications

reduce the system to the following nine components: H2, CO, N2,

C02, CH4, H2S, H20, C6H6, and CH30H.

Selection of the Thermodynamic Model

There are two approaches to modelling the phase behaviour

of a mixture. The conventional method invokes a liquid act-

ivity coefficient model to represent the liquid and an equation-

of-state for the vapor. Thus, the system is described according

to:

iYiPT xfwherei = vapor fugacity coefficient

PT = total pressure

20

Xi = liquid activity coefficient

fol= liquid reference fugacity1

yixi = vapor, liquid mole fractions

The alternative approach uses a single equation-of-state to

describe both the vapor and liquid phases. Noncondensable

components in the mixture present a problem with regard to

establishing the reference fugacity for the activity coeffi-

cient approach. Generally, the standard state fugacity is

that of the pure liquid at the system temperature and pressure.

It is meaningless, however, to consider a pure liquid state

for a supercritical component and, thus, it is necessary to

follow the unsymmetric convention. In this case, the activity

coefficient is normalized to 1.0 at infinite dilution and the

Henry's Law constant becomes the standard state fugacity. As

O'Connell (1977) points out, there are difficulties associated

with referencing to the Henry's Law constant for multicompon-

ent solvents since the constant is dependent on solvent comp-

osition. As such, the composition dependence of the standard

state fugacity should be built into the thermodynamic model.

On the other hand, equation-of-state models have generally

proved to be inadequate for predicting the non-idealities of

mixtures containing polar substances (Mathias, 1981). How-

ever, ASPEN has recently adapted the Redlich-Kwong-Soave

(RKS) equation-of-state to improve its ability to represent

systems containing highly polar compounds. Mathias (1981)

has demonstrated that this model is effective in describing

the coexisting vapor and liquid phases for systems such as

methanol-water. An alternative model applicable to polar

21

systems is the Perturbed Hard Chain equation-of-state (Gmeh-

ling et al., 1979). In this case, the model handles polar

components by assuming that they dimerize. Thus, the model

must evaluate the chemical equilibrium for the dimerization

process. Such treatment proves to be a major drawback for

this equation-of-state since it augments the computational

effort substantially. In addition, as Mathias (1981) notes,

dimers are rarely observed experimentally, so that the equation

rests on a weak theoretical foundation.

Consideration of the difficulties and inaccuracies assoc-

iated with the activity coefficient model approach led to the

selection of an equation-of-state to correlate the Rectisol

system for this work. The particular equation-of-state that

was chosen was the modified version of the Redlich-Kwong-

Soave model which is described in the next section.

Modified Redlich-Kwong-Soave Equation-of-State

The basic Redlich-Kwong-Soave equation is given by

RT aV - b V( + b)

where 'a' and 'b' are functions of critical temperature and

pressure as follows

a = (T)a

2 2ac= 0.42748 R TciciP

b = 0.08664 RTci

P .ci

To some extent, the 'a' parameter may be regarded as an attrac-

tive parameter, accounting for the intermolecular attractive

22

forces. On the other hand, the 'b' parameter is a repulsive

term, representing the finite volume occupied by the component.

c(T) is used to correlate pure component vapor pressure data

and has been modified by ASPEN to the following form

05 = 1 + m(1 - Tr05) - p( - Tr)(0a7 - Tr)

Tr = reduced temperature

m = function of acentric factor

p = polar parameter

The correlation for m, taken from the work of Graboski and

Daubert (1978), is

2m = .48508 - 1.55191 - 0.15613ao

where = acentric factor.

The ASPEN work has introduced the polar parameter, p, in the

expression for to improve the vapor pressure predictions.

This parameter is empirical and must be determined by fitting

experimental pure component vapor pressure data.

For supercritical components, Boston and Mathias (1980)

recommend a different expression for to extrapolate into

the supercritical region

= exp(c (1 - T))

with c = 1 +0.5m +0.3p

d= c- 1

c

Mixtures are represented by the above equations using

the one-fluid theory. This means that the parameters 'a' and

'b' become quadratic mole fraction averages with the following

form

a - zzx x.a.1 1 3a = '.: f i

b i xix .b..1 jj

23

where

aij = (a.a.)0 5(1 - K0 - K1 T )l13ai a i j 1000

b. (b + b )(1 - K ° - K1 T )13 i 2 i bij bi 10002 1

As these equations show, the binary interaction parameters,

Kai and Kb , are expressed as linear functions of the absol-

ute temperature in Kelvin. These parameters account for the

deviation of the behaviour of a binary pair from that predicted

according to the basic mixing rules. Any combination of these

four parameters, K K K0 K may be used to fit binary'a' a' Kb' Kb'

vapor-liquid equilibrium (VLE) data to the modified RKS equa-

tion-of-state.

Experimental Phase Equilibrium Data

The physical properties model requires only binary inter-

action parameters in order to describe a multicomponent system.

In the absence of multicomponent vapor-liquid equilibrium data,

values for these parameters may be established by fitting bi-

nary phase equilibrium data alone. Prausnitz (1980) indicates

that his experience has been good in employing only binary

experimental data to predict multicomponent equilibria, pro-

vided that the system is well removed from the critical for

each component. He cautions that, in cases where the critical

conditions are exceeded, multicomponent predictions may be

unreliable even if the binaries are fit well. Unfortunately,

published multicomponent data are not available for the spe-

cific Rectisol system. This deficiency made it necessary to

rely on experimental binary equilbrium data to a great extent.

As the work proceeded towards determining the parameters for

24

the modified RKS model, the results of a recent study at North

Carolina State University became available, presenting quater-

nary data for the hydrogen sulfide-carbon dioxide-nitrogen-

methanol system (Rousseau et al., 1981). These data initially

served as a comparison for the multicomponent equilibria pre-

dicted by the preliminary physical properties model. Subse-

quently, the quaternary data were used to improve the binary

parameter estimates, leading ultimately to an improvement in

the model predictions.

In order to maintain the data fitting effort at a manage-

able level, it was initially decided to ignore the interactions

between two gas components. Justification for this simpli-

fication lay in the fact that the non-idealities of the liquid

phase were expected to dominate over the vapor phase non-ideal-

ities, as Prausnitz notes (1980). This decision reduced the

data fitting to the 21 possible interactions of the gas com-

ponents with methanol, benzene and water and the interactions

between methanol, benzene and water themselves.

The temperatures in the Rectisol process span a broad

range, varying from -75°F in the absorption section to as high

as 250°F during solvent regeneration. To minimize the degree

of extrapolation, the parameter estimation work called for

data covering as wide a temperature range as possible. A

literature search revealed experimental phase equilibria for

each of the binary pairs. Although this information did not

provide complete temperature coverage, the data adequately

spanned the temperatures where a given interaction was most

relevant. For instance, gas-liquid non-idealities are most

25

important at the lower end of the temperature range where ab-

sorption takes place. On the other hand, the high temperature

processes mostly involve the separation of benzene, methanol

and water.

In cases where more than one data set were located for

a binary pair, only the most recent was selected for the model

development unless these data appeared to be generally incon-

sistent with previous data.

Development of Modified Redlich-Kwong-Soave Model Parameters

Fitting the modified RKS model to binary experimental

data involved adjustment of a combination of the four inter-

action parameters, K, K, K, K1, as discussed previously.

The best estimates for the combination of parameters were

generated for each set of binary data using the ASPEN Data

Regression System (DRS). DRS employs the maximum-likelihood

principle in reducing experimental data. This analysis recog-

nizes that experimental data contain random measurement errors

which it takes into consideration, assuming that the errors

are randomly distributed about zero with a normal distribution.

The standard deviation associated with a particular measure-

ment is supplied by the user. The system creates an objective

function comprised of sum-of-squares terms for each observed

variable, for example:

N e 2 o 2 o. e 2(Pi-Pi) -T i ) (Xlil)

T. + . +1 a,.2 x2

P, Cr1 1 .1

in which the superscript 'o' stands for the value adjusted

according to the thermodynamic model, while 'e' refers to the

26

experimental value. The terms are summed over all data points.

In DRS, the Britt-Luecke (1973) generalized least-squares al-

gorithm computes estimates for the model parameters such that

the objective function is minimized. At the same time, the

parameters must satisfy the equilibrium constraint that the

fugacities of all equilibrium phases are equal.

Prausnitz (1980) presents an informative discussion of

the methods of estimating parameters. Many of his recommend-

ations were adopted for this work. For instance, the follow-

ing values were used for the standard deviation of the measured

variables, as Prausnitz suggests: temperature - 0.050°C, pres-

sure - 1 mmHg, liquid mole fraction - 0.001, vapor mole frac-

tion - 0.003.

The data regression work demonstrated that, in most cases,

two parameters adequately correlate the equilibrium data. How-

ever, in the final set of interaction parameters, additional

parameters are included for some pairs if these improve the

fit significantly. One notable example is the benzene-water

system for which all four parameters are essential to modelling

the observed liquid immiscibility. It should be noted that,

in order for the 'a' and 'b' parameters in the fundamental RKS

equation to retain their theoretical significance, the term

(1 - K0 - K1 T ) and the analogous term for the 'b' para-aij aij 100-0

meter must have a positive value. Unfortunately, in the case

of benzene and water, large parameter values are necessary to

represent the equilibrium, rendering the above terms negative.

While this anomaly causes no computational difficulty, it does

mean that the model must be considered to be purely empirical

27

as far as benzene and water are concerned. As such, there

may be some concern about using the equation-of state to ex-

trapolate information. However, the benzene-water experimental

data adequately cover the temperature range of importance in

the Rectisol process. Further, a comparison with experimental

data demonstrates that the model provides a good prediction

of the liquid-liquid equilibrium for the methanol-benzene-

water system. This comparison is illustrated in the follow-

ing chapter.

It has been suggested that highly non-ideal mixtures,

such as the benzene-water mixture, may be better represented

using the recently proposed mixing rules based on the local-

composition concept (Boston, 1981; Whiting, Prausnitz, 1981).

Further work in this area lay beyond the scope of this project.

In general, good agreement was obtained between the ex-

perimental binary data and the thermodynamic model predictions.

The maximum error in the liquid composition for a set of bi-

nary data was, on the whole, less than 6 percent. For a given

binary pair, the binary parameters exhibit a high degree of

correlation which indicates that the parameters are not uniquely

determined. In other words, the experimental data may be rep-

resented by several alternate parameter combinations with

little loss of accuracy.

Evaluation of the Preliminary Thermodynamic Model

The preliminary thermodynamic model of the Rectisol

system was comprised of the modified RKS equation-of-state

with interaction parameters generated by regression of binary

28

data, as just described. This model neglected any gas-gas

interactions, considering them to obey the simple mixing rules

without correction. To assess the validity of the model, the

predicted equilibria for the hydrogen sulfide-carbon dioxide-

nitrogen-methanol subsystem were checked against the data pre-

sented by the North Carolina State University team (Rousseau

et al., 1981). Bubble point calculations using the experimental

liquid composition and temperature permitted a comparison of

the predicted bubble point pressure and vapor composition with

the actual data. Unfortunately, the model did not prove to

represent the data adequately, overpredicting the bubble point

pressure by as much as 30 percent and overestimating the nitro-

gen in the vapor at the expense of the carbon dioxide. In an

effort to ameliorate the situation, parameters modelling the

gas-gas interactions (carbon dioxide-nitrogen, hydrogen sul-

fide-carbon dioxide, hydrogen sulfide-nitrogen) were regressed

from binary data and included in the model. The resulting

improvement in the model predictions was not considered suf-

ficient. Further, the predicted Henry's Law constants for

nitrogen in mixed methanol-carbon dioxide and methanol-hydrogen

sulfide solvents exhibited unexpected behaviour. The experi-

mental data appeared to be consistent with the approximate

relationship that

~~,ln HN mix =Zixi ln HN, i(Prausnitz, 1980)

with x. = solvent mole fraction on solute free basis21

However, rather than displaying linearity on a logarithmic basis

according to the above expression, the predicted Henry's Law

constants increased initially with a small increase in the

29

amount of dissolved carbon dioxide or hydrogen sulfide. The

constants appeared to reach a maximum before decreasing to the

Henry's Law constant for nitrogen in pure carbon dioxide or

hydrogen sulfide. Figure 7 illustrates the variation of the

Henry's Law constant with solvent composition. The plot in-

cludes three curves. The upper two demonstrate the improve-

ment in the behaviour of the constant on inclusion of the gas-

gas interactions. The lower results from utilizing the quaternary

data in estimating the parameters for the physical proper-

ties model, which is discussed in detail in the following

section. The abnormal solubility behaviour is in part re-

sponsible for the deviations between the predicted and experi-

mental compositions. In the model, a small amount of dissolved

gas in the methanol results in an unexpectedly high Henry's

constant and consequently a lower solubility for nitrogen.

Therefore, the vapor mole fraction of nitrogen is overpredicted.

From an analysis of the experimental data and the model

predictions, it was concluded that the vapor interactions

between carbon dioxide and nitrogen are strong and conceal

any non-idealities in the liquid. This fact makes the system

hard to analyze. It did not appear possible to isolate the

vapor and liquid behaviour from binary data using a single

equation-of-state. The apparant success of Rousseau et al.

(1981) in modelling the quaternary system may be attributed

to the fact that they adopted the activity coefficient ap-

proach. -This means that they employed a separate equation

for each phase. In their model, the vapor is represented by

the RKS equation-of-state with interaction coefficients set

30

0

omq'e 'ueosjqTN oJ us!suoo M1j sRJuaH

R~

0 ~:E-q H

Z X00 Hi--.tH

zO P4Z

EOZCl . m

z .H Hr- CD Q¢ a4

> E1Z~ vN <

--I

0C0

0 ,

o*H .,i

0000Oi *H

o .

a) *I-- 0H 0

()0

00 O\ 0 0 0 0 00C) C- 0 0 0 0 V0oo 0 0 V 0 C1,2_) n cn CM -A-

00

;

31

to the standard vapor values available from separate vapor

phase studies. Having established the vapor behaviour in

this manner, the liquid behaviour could be considered indep-

endently.

In working with the modified RKS equation, it appeared

that the 'a' parameter, the attractive term in the basic

equation, could be treated as the vapor term, while the re-

pulsive 'b' term could be considered as representing the liq-

uid. This observation suggested that the vapor interactions

could be isolated from those in the liquid by fixing the in-

teraction parameter for the 'a' term to the standard RKS vapor

value reported in the literature. The remaining non-idealities

due to the liquid phase could then be reflected in the 'b'

parameter. It should be noted that this approach is only

suitable .for subcritical components since the modified RKS

equation only reduces to the original RKS equation at condi-

tions below the critical.

For the final model. it was decided to use the quater-

nary experimental data in the data regression runs to gener-

ate improved parameter estimates for the interactions between

carbon dioxide, hydrogen sulfide and nitrogen. So that the

DRS runs would converge, several parameters were fixed before

the final DRS run rather than fitting all the gas-gas inter-

actions with the quaternary data. Hydrogen sulfide-nitrogen

gas interactions appeared less significant, so the original

parameters developed from binary data were used. The effect

of the vapor interactions for the hydrogen sulfide-carbon di-

oxide pair, which are both subcritical, was modelled with the

32

standard RKS value for K (Graboski and Daubert, 1978). Dataa

regression of the quaternary data then established parameter

estimates for aKtrge and thestimates for nd Kb for carbon dioxide-nitrogen and the

liquid K for hydrogen sulfide-carbon dioxide. The resultant

model predictions agree substantially better with the experi-

mental data as the comparison in Table 1 shows. The bubble

point pressure predictions still appear to deviate from the

data, especially at the higher pressure levels, however.

Table 2 displays the final parameter estimates, indicating

the literature source of the data. This table shows that the

final set includes values for several additional gas pairs

which were readily available from previous work (Mathias, 1981).

The parameter set is stored in a user library under RKAKIJ,

to be called directly by the simulation program.

Enthalpy Calculations

In order to guarentee reliable enthalpy predictions from

the modified RKS model, enthalpy data should be included in

the data regression runs to determine interactions parameters.

However, such data were not available for the Rectisol system.

Under these circumstances, the Lee-Kesler equation-of-state

was favored over the modified RKS equation-of-state for enthal-

py calculations. The former correlation has a good reputation

for predicting reasonable mixture enthalpies based on the pure

component properties.

33

Nc O N N0 ~ 0 O\ V

o O cox 0 -0i- co 0i

0 o 0 0U O C' O Co

0 0 0 0000o o o

q C0 0 0 0

0 n C O

o 0 N nx * * *

O~EV2 ~ O O O O

M ~ O ad) E-4 c o # u-4 i2C 0 0 0 0H o 0 No c N

:HE0 0 0 0s a 72 N 0 0~C \0

O¢ C TO CO CO N

o fail 0H M Co 0 \tz H N iH N H

UE S\ 0 0 0 00 r z0 0 0 0H \PO O O OC'2P X ¢ O0 0 0 0E Eq CM0 O O o~

O~~~- N O OCO V3 N 0 0 0 O O O O

~:2~ O O O OO O 0 O

a) c . .

So CQl C\1 Cr

I I I I

N C V O,

I\0 c- c- -\O O' O'- O~

0 0 c'- CO

o \0 0 0 0

O 0 0 0C; C; C~ ;'O O O O

n n N OnO O O OO O O Oo 0 0 0o; 0 0N C n Cn \0o \) 0 c'-

O 0 0 0iN i- NM n

O 0 C\l n

00 0 0V- --I n -w n m m)CO N CO cO

o 0 0 0;__> 4 wr6

o N N N

o 0 0 0C- \o o n, 0 C)- 0I .o o o o

t0 0 0 0e C" " co 0i v- CN T- CMli o o o o* * 0 0 0 0

co \0 C \0o C C'-o o 0 00 0 0 0o 0 0 0O o 0 o oo o o oo o 0 0

o~ -; o;(O c O 00I4 C n -0

0s 0 C Y'ON, ,-4 ON, O._ N N CnIqO O O O0 0 0 0

- -T

IW l T I' * i.1%

_-

.~

34

O ON ON

o7\ C\

* Cd

a) c~~- C~ Cd cd -

ON O\ 4-' 4-

o a)U2 _ -a) a) S

Ld d Cd~ d-P a a) cN N UCd 4S .- *H *H 4-'

w a) 00 ifq (~: >-i >-i 11

n\ \

\_ COON - '\-C'--

* a) .

cd~ ONON a) 3

O ON a) H

_ O CO ') a)

$s cqcd 0 0 *H Hc) Cdo 000-PHM

C- cr\

U) C.'-U-- C--a) OOCd Cd N N N N N c N *H+- O O, r O\ O\ O\ O\ 4-a) . i-I # d i-i i-I ~-I i-I Cd v-f

- ' Cd

-P HH r- HH -H C H ) . ) ) 0 ) ~> Oa) a) a) cd a) a a) U)

bfl * O rO ;4 rO rO ' C= - *H M q -M -, M 0

C.- 0 D O C OCY N ON0?C -.I

0", 0 -.I ~01ON co NC ON

0 0 v-) 0

0000I I I

· .-"t . O.C) 0 o £ Cdo--j co OC\ '--

I I Io d-t \0 cOn C\l C\ O

oN \ O 0 N 0·O-gj -~m~- N\O ONH NO000 - 0I I I I

\D C\? C>1

0 0 Oo000

O 0oC')\0 co0 °c'V-4 V) V'

I

0 C\l

0 0 O o co I

o 0 0 0 d )I I

C- C- V, 0 N H \ ON\ CO

C .- CO \ ' O O \cn C nN co T-0 [,-- CcO O0 ~ 0 0.4 .-. 0

CO 0 \ CO N V)' c O

000000 0000V' V) <ztN V l1 n n n V--i I I I

o~~~~ ~ ~ 0 o o0 0 C0 C

O~ O~ O~ O O O O O C) OOOOOOOOOOOOO0OOOOO00O

C-\ o00 0 0 -. o0 0 0 0 0 0 o 0(O O (O O A ( (H (O (O (O (O (O (: O O (

I

0 N = U N CM C\l Ne\0 cN NO= NO X : : N \ NO N O X: CN CN 0 CZ CN

C, = V Z V 2 C. C. M 0 X V Z C.) V :M: O V t) Z =

:m = t' m: = m0000000 ,0 \0 \0 0 0(n n n n n n n 0 0 0 0 0 0 = ' ,00000000 : : = = : = N N N O 0 0 \ 0 rOV V V V 0 V V V m =:r = = t m V 0 V V 0 )

HH

E

0EIEq D0

O;ZN Ui) -

H tr>E q

I

E O

~HZtq:E

*,rCdP-4

-'

Za)0pS0C-

aNC\

o O U

a) U: CC -)

o ,la ) H

-P N 0 a.· H r oo i o-4 U2>0 o

U~\O cO\0QJG\\Q

'-4

OOH

· \ l Cd-4 a)

*H +'_ a)

E - o

0 *H AE- S )

- C\ 1-IO\ co oD-.I-C\ --O ocl-0 0o0 ' '~~~~~~0 0 m -0C\l

I

o1-40-

C\.l 1Lf\Cno0 - CWl

I I II I I

n V)

00· o0o o

o o0 o o 0 O · o

o a a a 00000\.O0 00000 - 00 0

i cs oI

U) _ N Oj C~l O O O C0~4=:a: 0 0

0 0 C\ 0 C'JeC~io

35

V~ON

o

0

o

0

-iN-P

HH

0

0

0

I-

o

C\Z

a)EH

HHO Eqqr V20 E-1

0Io

Z I'0 Z

PI-I'HU E-

wE

i OV2O RO

EH

Z H,H

'- R

H~

mq

cc0ooc0

0

coCO\ON

00

\0

0\I

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-p

a)

0

00V

36

IV. SIMULATION OF THE RECTISOL PROCESS

Basis for the Process Simulation

The proprietary nature of the Rectisol process restricts

the amount of process data disclosed in the literature. The

information presented in the CONOCO Commercial Design Report

consists solely of the feed and product compositions and a

process flowsheet. While the process flowsheet shows detail

of the sequence of operations and the destination of all in-

termediate streams, no specifics are provided for the process

operating conditions. On account of the lack of actual oper-

ating conditions, it was impossible to develop a simulation

that would match the CONOCO design. Therefore, this work

adopted the approach of developing a viable Rectisol design

based on the CONOCO design flowsheet and feed composition.

The feed gas composition, pressure and temperature displayed

in Table 3 were taken for the process model.

The literature search located typical absorption temp-

erature and pressure ranges only. Detailed simulation of the

other process units using ASPEN established the remaining

operating conditions necessary to achieve the desired separ-

ation.

To minimize the computer time and cost associated with

simulating a complex cyclic process such as Rectisol, a sim-

plified process model was assembled. In this simulation, a

series of simple unit operations such as mixers, splitters,

flashes, heaters and pumps represent the Rectisol process.

Rigorous simulation of the staged separation units is performed

37

Table 3

COMPOSITION OF FEED GAS TO RECTISOL PROCESS

Component lbmol/hr

Hydrogen 1909.7Carbon Monoxide 4052.9Carbon Dioxide 299.1Methane 471 1Ethane 18.9Ethylene 2.2Propane 4.7Propylene 2.2Butane 1.8Butylene 1.6Nitrogen 41.7Hydrogen Sulfide 131.0Carbonyl Sulfide 8.0Carbon Disulfide 0.1Thiophene 0.4Hydrogen Cyanide traceNaphtha 23.1Water 13.8

Total 6982.3

Temperature, OF 95Pressure, psia 427

38

on a stand-alone basis, external to the overall flowsheet.

The results of such rigorous simulations established the splits

to be used in the simplified model, where the units are modelled

by separator blocks. A block flow diagram of the overall

flowsheet simulation is shown in Figure 8, indicating the

correspondence between the model and the actual design flow-

sheet. Appendix A contains the computer input and output for

the simplified Rectisol model.

The development and design philosophy employed in each

of the major process sections is discussed in detail below.

Absorption Section

The hydrogen sulfide absorber consists of two separate

sections; a prewash section, which removes all the water and

naphtha from the synthesis gas, and a main wash section, where

the bulk of the hydrogen sulfide absorption occurs. Each

section was simulated independently using RADFRAC, which is

ASPEN's rigorous model for multistage vapor-liquid fraction-

ation. The results of these simulations were used to set the

degree of absorption in the overall flowsheet simulation in

which two mixers and two separators (PREWASH1, PREWASH2, H2S-

ABS1, H2SABS2) model the absorption section.

In designing the absorbers, the designer must establish

the number of trays, the temperature of the entering gases

and solvent, and the operating pressure. For a given degree

of absorption, these parameters then dictate the required

solvent rate. Hochgesand (1970) indicates that for the Rect-

isol process the methanol circulation rate greatly affects the

process economics, since it influences the size of all process

I

p i 0 239-s

'4

-1 -- - -- . .....

&| s--3--- -- g

CL_ r1 . ' ? w 8 3

. ...... .... ; X/ X7--Lh .0 I WI,!Ge~~~~~~~~~~~~~~~~~~~~~~~~~~~O

- ~~~~~~~IQC4 ~ X ~i

a'

4t1 1

A .1I~~~~~

-4 I

.1 I

C

~1 I'

A.

I

iI

II

p a

O CD

¢- P

O E<E-I D0) Ha5 or : O

I HnV E4OD V

W 2 ! . ! , I =

i I

- . - I

I

- -- --- TI

I'( I

11

-,,1

I

41o I 11

I- t I

F

--- jI

40

equipment, pumping costs, regeneration costs etc.. The design

of the absorption section, therefore, sought to minimize the

methanol solvent rate. Since high pressure favors physical

absorption, the absorbers are designed to operate at the maxi-

mum pressure compatible with the other units. This pressure

of 421 psia is the pressure of the gasifier overhead less an

allowance of 6 psi for the pressure drop through the heat ex-

change train. According to Hochgesand (1970), economics rarely

justify compressing the feed gas further. With the same con-

sideration in mind, the absorbers are designed to operate at

low temperature with the recirculating methanol cooled to -75 °F

and the feed gas cooled to -30° 0F. Sasol report such temper-

ature levels for their Rectisol operation, which appear to be

typical (Hoogendoorn, 1957). Further, although this study

did not examine the detailed heat exchange between the feed

and desulfurized gas, a heat balance shows that the feed may

be cooled by heat exchange to -30° F without supplemental

refrigeration.

Absorption columns may contain packing or trays, although

there is some evidence in the literature to suggest that trayed

columns are more common for Rectisol (Maddox, 1977; Herbert,

1956). In the current design, the absorbers are designed as

trayed columns. The pressure drop through the column is based

on the rule-of-thumb of 0.15 psi pressure drop per stage

(Treybal, 1968).

The design for the prewash absorber features six theoreti-

cal trays and a solvent rate set at 25 percent greater than

the theoretical rate required to absorb all the naphtha. This

follows the recommendation of Kohl and Riesenfeld (1979).

Given the objective of the prewash absorber to recover the

naphtha with as little hydrogen sulfide as possible, the meth-

anol rate is restricted to the minimum that ensures complete

naphtha absorption. In order to improve the performance of

the RADFRAC model for the prewash absorber, the feed is pre-

flashed so that any liquid may bypass the absorber. This

preflash simulates what would actually occur in the bottom

of the absorber.

The gas from the prewash absorber proceeds to the main

wash section where all the hydrogen sulfide and part of the

carbon dioxide are removed. The methanol solvent rate to the

main wash section was determined so that the degree of carbon

dioxide removal agreed with the information presented in the

CONOCO Commercial Design Report, while achieving complete

removal of hydrogen sulfide. However, the basis for the ab-

sorber feed composition did not account for the small amount

of recycled hydrocarbon gas. This means that the predicted

carbon dioxide gas removal in the overall simulation is not

in complete agreement with the CONOCO commercial design.

It is unusual for an absorber to contain less than six

or more than ten theoretical trays (Perry and Chilton, 1973).

A sensitivity study to the number of trays in the main wash

absorber indicated that an increase from six to ten trays re-

sults in a reductions of only 2 percent in the required meth-

anol rate. Figure 9 shows a plot of carbon dioxide in the

outlet gas versus solvent rate and number of trays. The rel-

ative insensitivity of the required amount of solvent to the

42

Figure 9

EFFECT OF SOLVENT RATE AND NUMBEROF TRAYS ON CARBON DIOXIDE ABSORPTION

02 in Feed)er

5200Methanol Solvent Rate, lbmol/hr

260

d~ 220

o0

d 180

q,

; 140C\2

ro0- 0O l

. D-

3000 3400 3800 4200 4600

43

tower size justified designing the main wash absorber with

six trays.

The output from simulating both the prewash and main wash

absorbers rigorously is included in Appendix B. In the final

simulation, the main wash and prewash absorbers operate with

3100 lbmol/hr and 330 lbmol/hr of circulating methanol respec-

tively.

Flash Regeneration Section

In the flash regeneration section, a fraction of the ab-

sorbed gases are recovered from the rich solvent by a reduction

in pressure. The pressure of downstream processes dictate the

final flash pressure. In this simulation, the prewash flash

regenerator flashes the prewash solvent to 30 psia, allowing

sufficient pressure head to feed the water wash column with-

out further compression. Similarly, the main flash regenerator

is designed to flash the main wash solvent to 30 psia in two

stages, with an intermediate pressure of 75 psia. Since the

vapor from the first flash is recycled to the absorption sec-

tion, the intermediate pressure must minimize the amount of

hydrogen sulfide and carbon dioxide desorbed while recovering

as much of the desired hydrocarbon gases as possible. An

intermediate pressure of 75 psia represents an appropriate

compromise. The three flashes are simulated as two-phase

adiabatic flashes (PRWSFLSH, FLSHRG1, FLSHRG2).

Naphtha Extractor

In the naphtha extractor, the methanol-water-naphtha-gas

stream is allowed to separate into two liquid phases to permit

removal of the naphtha-rich phase from the system. The feed

44

consists primarily of the bottoms from the prewash flash re-

generator to which a water-rich stream is added to effect the

liquid-liquid separation. The water source is the bottoms

from the water wash column. In order to ensure that the

naphtha extractor operates without any of the water icing,

the prewash flash bottoms is preheated to 50°F. The design

pressure is 30 psia, 5 psi higher than the water wash column

to which any vapors are directed. At these conditions, the

physical properties model predicts a phase separation into

two liquid phases which are in reasonable agreement with liq-

uid-liquid experimental data for the methanol-benzene-water

ternary. The plot in Figure 10 illustrates the comparison

between the predicted separation at 38°F and experimental

data at 77°F. Unfortunately, published data were not located

at the actual separation temperature.

In the overall model, the naphtha extractor is represented

by a feed mixer (NEXMIX) followed by a three-phase adiabatic

flash at 30 psia (NEXSEP).

Azeotrope Column

The objective of the azeotrope column lies in removing

the residual naphtha and dissolved gases from the methanol-

water mixture. Preliminary simulation of the column using

RADFRAC suggested that a good separation is possible in this

column, yielding an uncontaminated methanol-water stream.

This simulation indicated that all the gas components and

naphtha are distilled in the overhead along with a small

portion of the methanol, while all the water and remaining

methanol are recovered in the bottoms. As the name of the

d)a)

45 N

a)

wIu 1 9 MI4 1-1 a)

I C

46

column suggests, the overhead may be expected to contain ben-

zene and methanol in their azeotropic compositions (61.4 mole

methanol at 128.8°F). Further simulation of the column with

RADFRAC demonstrated that in the overhead methanol and ben-

zene may indeed reach their azeotropic composition on a meth-

anol plus benzene basis.

Development of a feasible design for the azeotrope col-

umn involved setting the distillate rate, the reflux ratio,

the number of theoretical trays, and the feed entry point.

Seeing that the split between the overhead and the bottoms

was well defined in the preliminary study, the distillate

rate could be established purely by material balance consider-

ations. The behaviour of the column proved to be highly sen-

sitive to the reflux ratio. The amount of material in the

rectifying section of the column depends directly on the reflux

ratio. Since such a large percentage of the feed (94 mole

percent) ultimately appears in the bottoms product, a high

reflux ratio is necessary to ensure adequate vapor flow through-

out the column. In the final design, a reflux ratio of 2.0

leads to satisfactory operation. While this ratio appears to

be substantially higher than the minimum reflux ratio, diffi-

culties were encountered in obtaining a converged RADFRAC

solution at lower reflux ratios.

A revealing method of interpreting tray-to-tray output

uses the idealized McCabe-Thiele techniques (Russell, 1981).

This tool provides valuable insights despite the non-ideality

of the particular system. If methanol and benzene are iden-

tified as the two components which require separation, and

47

all the compositions are normalized on a methanol plus ben-

zene basis, a McCabe-Thiele x-y diagram may be prepared. Fig-

ure 11 displays such a diagram for the upper portion of the

azeotrope column, where compositions are plotted on a loga-

rithmic scale for clarity. The equilibrium compositions were

taken directly from the RADFRAC output. This type of plot

identified pinch points, the minimum reflux ratio, and indic-

ated the optimum feed location during the column design. In

the final design, the feed enters at the third stage. Figure

11 suggests that a slight improvement in the column efficiency

may be realized with the feed located at the second stage.

The diagram also shows that the physical properties model pre-

dicts an azeotrope composition close to that reported. Since

the overhead lies at this composition, it can be seen that

increasing the number of trays would not improve the separa-

tion.

The feed quality plays a significant role in the column

design, as well. Given the fixed distillate rate of 45 lbmol/hr

a simplified material balance analysis, assuming constant molar

overflow, predicts the maximum percentage of vapor permitted

in the feed for the column to operate as desired. For in-

stance, with a reflux ratio of 2.0, the feed must be greater

than 82 percent liquid in order to meet the desired distillate

rate. Higher degree of vaporization results in insufficient

vapor in the stripping section. In the actual process flow-

sheet, the azeotrope column feed is preheated by heat exchange

with the waste water from the methanol-water column. The final

temperature of the waste water was taken to be 118°F, as

48

o

I~~~~4z -,e7 _ -I I is

I ; e I I,

m 4ffll

1-H0 H +Hg

N

I; I I i i ' i I I '0Io

4t..

o

0

o0

:E0aio

H¢C)a) Ox4 rG

H-

Eq4

¢

b

.I i: ! I ;

4

J.

.$ ..t.? Ai'

CD _

0

0N:

\0\0XO

, j .

I-_ I

T

0N

IL

4

4

I1

, i i i i.! I I I

, , o InI IvWg

0'

i

c

I I I i iIll : I

-

3a10Y

i i

) q o

i 2 0i I I i10 -g D P 0 1D,

imq N

%t0

(2)i

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w

¢

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. I!---- - I"t~

lp

v

_ _

- I. . I 'li I' 1·

I

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I I I.

I

I I

iI

!15,

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It I tf- . E

. I

-[ - - -

4~D

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,l

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A , . 'CA

I -, . I I

49

indicated on the CONOCO flowsheet, for this simulation. This

heat exchange results in an all liquid azeotrope column feed

at a temperature of 96°F.

The rigorous simulation of the azeotrope column is shown

in Appendix C. This simulation required 40 outside loop iter-

ations to achieve a converged solution. The difficulty in

converging is a consequence of the non-ideality of the system

in which the outside loop variables for RADFRAC, namely K-

values and enthalpy, are highly composition dependent.

Methanol-Water Column

The methanol-water stream from the azeotrope column bot-

toms is separated in the methanol-water column. In this sim-

ulation, the column is designed to yield a 99 percent pure

methanol distillate. A RADFRAC model was used to develop

the column design, and to examine the sensitivity to various

design parameters. In addition, a McCabe-Thiele diagram pro-

vided initial estimates of the number of stages required to

effect the desired separation at a given reflux ratio. This

same diagram showed that the minimum reflux ratio occurs at

a value of approximately 0.9.

Perry (1973) indicates that setting the reflux ratio 30

percent higher than the minimum reflux ratio generally mini-

mizes the combined operating and fixed costs. This general

rule was followed in chosing a reflux ratio of 1.2 for the

column design. At such a reflux ratio, the desired methanol-

water separation requires 25 theoretical trays. The final

design introduces the feed on the nineteenth stage from the

top which is the optimal stage according to the McCabe-Thiele

technique.

50

In the RADFRPAC model of the column, a total condenser

is included on the overhead. In the actual flowsheet, however,

the overhead vapors proceed directly to the hot regeneration

column to act as reflux. Since the methanol overhead and the

hot regeneration column bottoms differ very little in comp-

osition, the hot regeneration column is, in essence, acting

as the total condenser.

Detailed output from the RADFRAC simulation is located

in Appendix D. In the overall flowsheet, the column is

represented by a mixer and separator block (MWMIX, MWSEP)

with the reflux rate calculated by the Fortran paragraph

REFLUX.

Hot Regeneration Column

The purpose of the hot regeneration column is to strip

the main wash solvent completely of the acid gases remaining

after fash desorption. Methanol from the methanol-water col-

umn is also fed to this column. The methanol vapors, combined

with those generated in the reboiler, strip the acid gases

from the solvent.

The column is designed with 10 theoretical stages with

the main wash solvent and methanol distillate feeds located

on the third and sixth stages from the top, respectively.

The column was simulated with RADFRAC at reflux ratios ranging

from 0.5 to 2.0. These simulations indicated that it is poss-

ible to strip essentially all the acid gases from the methanol

at reflux ratios of 1.0 or greater. At the same time, the

overhead vapor contains less than 0.07 mole percent methanol

vapor. Increasing the reflux ratio from 1.0 to 2.0 does not

51

improve the separation significantly. However, at a reflux

ratio of 0.5, the separation is inadequate, resulting in approx-

imately 13 mole percent methanol in the overhead. The overhead

column temperature varies greatly with the reflux ratio. At

ratios exceeding 1.0, the overhead contains practically no

methanol and must be cooled to -40°F or less in order to

condense sufficient material for reflux. The temperature is

low enough so that the acid gases require no further refrig-

eration prior to the water wash column. On the other hand,

at low reflux ratios, the overhead may contain a significant

percentage of methanol vapor and, therefore, the column can

be refluxed by condensing the methanol. For instance, at a

0.5 reflux ratio, methanol accounts for 98 percent of the

condensate which is condensed at 97°F. Under these circum-

stances, a significant amount of the methanol is lost with

the acid gases. Some of this may be subsequently recovered

in the water wash column. Thus, the design of the hot regen-

eration column must consider the trade-off between a high

condenser duty requiring refrigeration on one hand, and in-

creased methanol losses on the other. Such a trade-off lay

beyond the scope of this work. For the purpose of this sim-

ulation, a reflux ratio of 1.0 was selected where the methanol

losses are negligible.

The detailed RADFRAC simulation of this column established

the splits for the separator (HTRGSEP) representing the column

in the overall flowsheet. The RADFRAC model is included in

Appendix E.

52

Water Wash Column

Methanol vapor is recovered from the combined acid gas

streams in the water wash column. The water-rich stream from

this unit is sent to the naphtha extractor to effect the liq-

uid-liquid separation. The rate of water addition was deter-

mined from experimental data so that the subsequent liquid-

liquid separation in the naphtha extractor would yield a meth-

anol-water phase containing less than 1 mole percent benzene.

The results from simulating this column as a four stage ab-

sorber with RADFRAC (see Appendix F) demonstrate that such a

water rate is sufficient to ensure no methanol carryover with

the acid gas. In the simplified flowsheet, the water wash

column is modelled with a separator block (H20WSH) which sep-

arates the feed according to the distribution predicted by

the rigorous model.

Heat Exchangers

The naphtha extractor feed/regenerated methanol, azeotrope

feed/waste water, and hot regenerator feed/regenerated methanol

heat exchangers are each modelled using two heaters. The heat

duty calculated in one heater is passed to the second heater,

which represents the other side of the exchanger, as an input

stream. The split in duty between the two regenerated methanol

coolers is determined so that the naphtha extractor feed (pre-

wash flash bottoms) is preheated to 50°F to ensure operation

of the naphtha extractor without any icing.

The absorber feed gas cooling train is handled by a series

of heaters. The simulation does not model the detailed heat

exchange with the purified gas. An enthalpy balance, however,

53

indicates that the heat removed from the feed gas in cooling

the gas to -30°F is sufficient to heat the purified gas to

67°F. The CONOCO Design Basis shows that the purified gas is

heated to 77°F after heat exchange. Thus, the absorber feed

temperature may be achieved through heat exchange with the

purified gas without any supplemental refrigeration.

Methanol Make-up

A methanol make-up stream contributes to steady operation

of the hydrogen sulfide absorbers by maintaining the methanol

circulation at the design rate. The make-up stream replaces

any methanol lost from the recycle stream. In the simulation,

a Fortran paragraph (MAKEUP) calculates the required make-up

rate.

Flowsheet Convergence

To facilitate the convergence of the overall flowsheet

which contains several recycle loops, the regenerated meth-

anol recycle stream is broken. More specifically, the meth-

anol solvent to the absorption section (RGMETC) is treated

as a feed stream, while the regenerated methanol from the hot

regeneration column (RGMETC2) is considered as a product stream.

This simplification is possible since the stream composition

is well defined and the stream rate is maintained at the de-

sign rate by the methanol make-up stream. Such treatment

permits the flowsheet to be divided into two subsystems which

can be converged separately.

The flowsheet is converged with an ASPEN generated tear

stream set using the Wegstein convergence method. Three

54

material streams and two heat streams must be converged to

converge the entire flowsheet.

55

V. RECTISOL PROCESS SIMULATION RESULTS

The process simulation provides a stream-by-stream heat

and material balance which is shown in the simulation output

in Appendix A.

The predicted composition of the desulfurized gas is com-

pared with the composition from the CONOCO Design Basis in

Table 4. The comparison shows that the ASPEN model predicts

a higher yield of each gas constituent. This discrepancy

stems from the fact that the methanol circulation rate was

established based on the fresh feed gas composition which ne-

glected the recycled hydrocarbon gas, as was mentioned earlier.

Therefore, the degree of absorption predicted by the ASPEN

model is lower than that accomplished by the CONOCO design.

To improve the agreement between the simulation and the design,

the methanol circulation rate in the simulation could be in-

creased.

The acid gas and the desulfurized gas are related by

material balance since virtually all the gases are recovered

in these two streams. Therefore, the lower degree of absorp-

tion shown by the ASPEN model leads to a lower acid gas re-

covery as compared to the CONOCO design. Table 5 presents

the acid gas composition for both the design and the simula-

tion.

The composition of the regenerated methanol calculated

by the rigorous simulation of the hot regeneration column is

shown in Table 6. The model predicts that it is possible to

regenerate the methanol to 99.8 percent purity, contaminated

only by 0.2 percent water and 2 ppm hydrogen sulfide.

56

Table 4

DESULFURIZED GAS COMPOSITION

Component ASPEN Modellbmol/hr

CONOCO Designlbmol/hr

HydrogenCarbon MonoxideCarbon DioxideMethane - ButaneHydrogen SulfideNitrogenWaterNaphthaMethanol

Table 5

ACID GAS COMPOSITION

Component ASPEN Modellbmol/hr

CONOCO Designlbmol/hr

HydrogenCarbon MonoxideCarbon DioxideMethane - ButaneNitrogenHydrogen Sulfide & Carbonyl

SulfideWaterNaphthaMethanol

0o.34.2

125.62.90.0

134098 01.8o0.o

Table 6

REGENERATED METHANOL

Component ASPEN Modellbmol/hr

Hydrogen SulfideWaterMethanol

1909.44048.7173.4499.6

0.041.6o0.0

o0.o0.1

1907.74037.4

159.2485.9

0.041.60.00.00.0

2.015.5

139.916.60.1

138.75.60.00.0

0.138.84

3817.90

57

Table 7 shows the heating and cooling duties calculated

by the process simulation. The heat balance predicts that

the refrigeration duty to cool the recirculating methanol is

10.86 MMBtu/hr. The total predicted cooling duty, including

both refrigeration and cooling water duty, is 16.80 MMBtu/hr.

This is lower than the duty quoted for the CONOCO design.

The deviation may be attributed to the differences in the de-

sign basis for the simulation and the actual CONOCO design.

In order for the simulation to match the degree of carbon di-

oxide absorption in the CONOCO design, a higher methanol circ-

ulation rate or a lower absorption feed gas temperature is

required. Both of these design changes would raise the re-

frigeration duty. As Table 7 demonstrates, the calculated

heating duty is significantly higher than that for the CONOCO

design. The heating duty represents the steam heat required

to reboil the azeotrope column, the methanol-water column and

the hot regeneration column. As such, the heating duty is

highly sensitive to the specific column design.

58

Table 7

COOLING AND HEATING DUTIESFOR THE RECTISOL PROCESS

ASPEN ModelMMBtu/hr

CONOCO DesignMMBtu/hr

Refrigeration andCooling Water Duty

Solvent RefrigerationHot Regeneration Column CondenserAzeotrope Column Condenser

Steam Heating Duty

Azeotrope Column ReboilerMethanol-Water Column ReboilerHot Regeneration Column Reboiler

10o864.311.63

16.80 18.76

4.2210.7510.10

25.07 11.96

59

VI. CONCLUSION

The process model described in this report represents a

viable Rectisol process, based on the CONOCO design. The

model demonstrates that the synthesis gas may be desulfurized

and the methanol solvent may be regenerated with reasonable

operating conditions.

It is difficult to assess the validity of the computer

model in view of the lack of published process data. One

point of comparison lies in the predicted compositions for

the desulfurized gas and acid gas. The current basis for the

methanol circulation rate, which is not consistent with the

CONOCO basis, renders this comparison meaningless. Unfortu-

nately, this project was concluded, due to lack of computer

funds, before a valid comparison could be made at an increased

methanol circulation rate. However, a minor adjustment of

the main wash solvent rate to approximately 3400 lbmol/hr

from the current 3100 lbmol/hr is expected to bring the model

predictions in line with the design basis.

The Rectisol computer model was developed as a flexible

tool which may be applied in future design work or sensitivity

studies. The effect of a design modification to one of the

major process units may be examined using the rigorous stand-

alone model. The results of this simulation may then be trans-

ferred to the overall simulation to assess the impact of the

change on the process as a whole. In addition, the model may

be employed to consider the trade-offs between such variables

as the absorption temperature, the degree of regeneration,

and the methanol circulation rate. An enhancement of the

60

Rectisol model by linking it to the ASPEN Cost Estimation

Subsystem and Economic Evaluation Subsystem would permit these

design changes to be evaluated from an economic standpoint.

LITERATURE CITED

Boston, J. F., Energy Laboratory and Chemical EngineeringDepartment, M.I.T., Cambridge, MA, Personal communication.

Boston, J. F., P. M. Mathias, "Phase Equilibrium in a ThirdGeneration Process Simulator," Presented at Second Inter-national Conference on Phase Equilibria and Fluid Prop-erties in the Chemical Process Industries, West Berlin,March, 1980.

Britt, H. I., R. H. Luecke, Technometrics, 15, 233 (1973).

Clarke, E. C. W., D. N. Glew, Canadian J. Chem., 49, 694 (1971).

CONOCO, Inc., "Phase I: The Pipeline Gas Demonstration Plant,Design and Evaluation of Commercial Plant, Vol. 2,Process Design and Preliminary Project Engineering Design",FE-2542-28, V2, 1980.

, "Phase I: The Pipeline Gas Demonstration Plant,Quarterly Technical Progress Report for Period 1 Oct.,1979to 31 Dec., 1979," FE-2542-26, 1980a.

Cook, M. W., D. N. Hanson. B. J. Alder, J. Chem. Phys., 26 (4),748 (1957).

Dechema Chemistry Data Series, "Vapor-Liquid Equilibrium DataCollection," Published by Dechema, Deutsche Gesellschaftfur Chemisches Apparateweses, Eds: D. Bahrens, R. Eckermann,(1977).

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Evans, F D., R. Battino, J. Chem. Thermo., 3 (6), 753 (1971).

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Hochgesand, G., Ind. Eng. Chem., 62 (7), 37 (1970).

Hoogendoorn, J. C., J M. Salomon, Brit. Chem. Eng., 2, 238 (1957).

Houghton, G., A. M. McLean, P. D. Ritchie, Chem. Eng. Sci.,6, 132 (1957).

"International Critical Tables of Numerical Data," NationalResearch Council, New York: McGraw-Hill, 1933.

Kirk, B. S., W. T. Ziegler, Adv. Cry. Eng., 10, 160 (1965).

Kohl, A. L., F. C. Riesenfeld, "Gas Purification," Houston:Gulf Publishing Co., 1979.

Kretschmer, C. B., J. Nowakowska, R. Wiebe, Ind. Eng. Chem.,3-, 506 (1946).

Maddox, R. N., "Gas and Liquid Sweetening," John M. Campbell,1977.

Mathias, P. M., "A Versatile Phase Equilibrium Equation ofState," submitted for publication, 1981.

Nonhebel, G., "Gas Purification Processes for Air PollutionControl," London: Newnes-Butterworth, 1972.

O'Connell, J. P., "Some Aspects of Henry's Constants and Un-symmetric Convention of Activity Coefficients," PhaseEquilibria and Fluid Properties in the Chemical Industry,ACS Symposium Series, 60, (1977).

Perry, R. H., C. H. Chilton, "Chemical Engineer's Handbook,"Fifth Edition, New York: McGraw-Hill, 1973.

Prausnitz, J. M., "Computer Calculations for MulticomponentVapor-Liquid and Liquid-Liquid Equilibria," EnglewoodCliffs, New Jersey: Prentice-Hall, Inc., 1980.

Rousseau, R. W., J. N. Matange, J. K. Ferrell, A.I.Ch.E. J.,27 (2), 605 (1981).

Russell, W., Badger America, Inc., Cambridge, MA, Course notes,1981.

Sander, W., J. Chem. Soc., 121, 2667 (1922).

Scholz, W. H., Adv. Cry. Eng., 15, 406 (1969).

Siedell, A., "Solubilities of Inorganic and Organic Compounds,"New York: Van Nostrand, 1952.

Sobocinski, D. P., F. Kurata, A.I.Ch.E. J., , 545 (1959).

Spano, J. ., C. K. Heck, P. L. Barrick, J. Chem. Eng. Data,13 (2), 169 (1968).

Stephen, H., T. Stephen, "Solubilities of Inorganic and OrganicCompounds," Oxford: Pergamom Press, 1963.

Toyama, A., P. Chappelear, T. Leland, R. Kobayashi, Adv. Cry.Eng., 7, 125 (1961).

Treybal, R. E., "Mass Transfer Operations," Second Edition,New York: McGraw-Hill, 1968.

Wiebe, R., V. L. Gaddy, Am. Chem. Soc. J., 6, 76 (1934).

Whiting, W. B., J. M Prausnitz, "Equations of State forStrongly Nonideal Fluid Mixtures. Application of theLocal Composition Concept," Presented at the AIChESpring National Meeting, Houston, 1981.

Yorizane, M., S. Sadamoto, H. Masuoka, Y. Eto, Kogyo KagakuZasshi, 72, 2174 (1969).

, S. Yoshimura, H. Masuoka, Kagaku Kogaku, 34, 953(1970).

BIBLIOGRAPHY

Rectisol Process

Herbert, V. W., Erdol und Kohle, , 77 (1956).

Hochgesand, G., "Rectisol and Purisol," Ind. Eng. Chem., 62(7), 37 (1970).

Hoogendoorn, J. C., J. M. Salomon, "Sasol: World's LargestOil-from-Coal Plant," Brit. Chem. Eng., 2, 238 (1957).

Linde, W., G. Ranke, H. Jungfer, "Ammonia," Fertilizer Scienceand Technology Series, New York: Marcel Dekker, vol. 2,233, 1973.

Maddox, R. N.,"Gas and Liquid Sweetening," John M. Campbell,1977.

Partridge, L. J., "Production of Ammonia Synthesis Gas byPurification and Shift Conversion of Gas Produced fromCoal," The Chemical Engineer, 5, 88 (1979).

Scholz, W. H., "Rectisol: A Low-Temperature Scrubbing Processfor Gas Purification," Adv. Cry. Eng., 15, 406 (1969).

Acid Gas Absorption

Kohl, A. L., F. C. Riesenfeld, "Gas Purification," Houston:Gulf Publishing Company, 1979.

Treybal, R. E.,"Mass Transfer Operations," Second Edition,New York: McGraw-Hill, 1968.

Physical Properties Models

Davis, J. J., R. I. Kermode, "Wilson Parameters for the System:H2,N2,CO,C02,CH4,H2 S,CH30H and H20," Fuel ProcessingTechnology, 1, 247 (1977/1978).

Mathias, P. M., "A Versatile Phase Equilibrium Equation of State,"Submitted for publication, 1981.

Prausnitz, J. M., "Computer Calculations for MulticomponentVapor-Liquid and Liquid-Liquid Equilibria," EnglewoodCliffs, New Jersey: Prentice-Hall, Inc., 1980.

Prausnitz, J. M., "Molecular Thermodynamics of Fluid-PhaseEquilibria-" Englewood Cliffs, New Jersey: Prentice-Hall, Inc., 1969.

Reid, FR. C., J. M. Prausnitz, T. K. Sherwood, "The Propertiesof Gases and Liquids," Third Edition, New York: McGraw-Hill, 1977.

Soave, G., "Equilbrium Constants from a Modified Redlich-KwongEquation-of-State," Chem. Eng. Sci., 27, 1197 (1972).

APPENDIX A

Overall Rectisol Model

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