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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'
ChemicalSolvent
Pads
GasPartialPressure
Pdes
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I I
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Re ference: Hochgesand, 1970
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
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0 0 0 0000o o o
q C0 0 0 0
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:HE0 0 0 0s a 72 N 0 0~C \0
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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
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a) c . .
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- -T
IW l T I' * i.1%
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34
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a) c~~- C~ Cd cd -
ON O\ 4-' 4-
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Ld d Cd~ d-P a a) cN N UCd 4S .- *H *H 4-'
w a) 00 ifq (~: >-i >-i 11
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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
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bfl * O rO ;4 rO rO ' C= - *H M q -M -, M 0
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000000 0000V' V) <ztN V l1 n n n V--i I I I
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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
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35
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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-
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¢
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.I i: ! I ;
4
J.
.$ ..t.? Ai'
CD _
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T
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IL
4
4
I1
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0'
i
c
I I I i iIll : I
-
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A , . 'CA
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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.
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, "Phase I: The Pipeline Gas Demonstration Plant,Quarterly Technical Progress Report for Period 1 Oct.,1979to 31 Dec., 1979," FE-2542-26, 1980a.
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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.
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Sobocinski, D. P., F. Kurata, A.I.Ch.E. J., , 545 (1959).
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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
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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.
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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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