oEPA
United States
Environmental Protection
Agency
Industrial Environmental Research
Laboratory
Research Triangle Park NC 27711
EPA-6OO/7-79-097
April 1979
The Solubility of Acid
Gases in Methanol
Interagency
Energy/Environment
R&D Program Report
-------
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EPA-600/7-79-097
April 1979
The Solubility of Acid
Gases in Methanol
by
J.K. Ferrell, R.W. Rousseau, and D.G. Bass
North Carolina State University
Department of Chemical Engineering
Raleigh, North Carolina 27650
Grant No. R804811
Program Element No. EHE623A
EPA Project Officer: N. Dean Smith
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
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ABSTRACT
A coal gasification/gas cleaning facility has been constructed
at North Carolina State University as part of a study funded by the
Environmental Protection Agency. A major part of this facility in-
volves absorption and stripping of acid gas constituents, and the
description of operations to carry out these processes requires in-
formation on the equilibrium behavior of the constituents with the
system's solvent. Two approaches may be used to obtain this inform-
ation: measuring extensive equilibrium data or developing a thermo-
dynamic model that uses limited data to predict equilibrium behavior.
The latter approach is suggested. The system chosen for study re-
flects current opinion concerning the most important candidate for
industrial acid gas removal systems.
A thermodynamic model was developed to predict phase equilibrium
in the methanol-carbon dioxide-nitrogen-hydrogen sulfide system based
on parameters determined from binary vapor-liquid equilibrium data
available in the literature. An experimental apparatus for obtaining
high pressure vapor-liquid equilibrium data was used to obtain data
for the methanol-carbon dioxide system at -15.0°C and at pressures
froin_5 to 16 atmospheres. Results showed excellent agreement with
published data. Equilibrium data were obtained for the methanol-car-
bon dioxide-nitrogen-hydrogen sulfide system at -15.0°C and pressures
ranging from 8 to 35 atmospheres. Experimental liquid phase composi-
tions and temperatures were used to predict vapor compositions and
totel pressures. Predicted values showed an average deviation from
expfrimental data of 21 percent for vapor compositions and 10 percent
for total pressures indicating the possibility of ternary effects not
acccjnted for by the model.
-------
TABLE OF CONTENTS
Page
ABSTRACT ii
FIGURES iv
TABLES v
1. INTRODUCTION "1
2, FUNDAMENTAL EQUATIONS 5
3. MODEL DEVELOPMENT 10
4. EXPERIMENTAL EQUIPMENT AND PROCEDURE 34
5. EXPERIMENTAL RESULTS 38
6. SUMMARY OF RESULTS 44
LITERATURE CITED 45
NOMENCLATURE 48
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FIGURES
Number Page
1. Solubility of gases in methanol 3
2. P-x Diagram for CH3OH(1 )-C02(2) 20
3. P-x Diagram CH3OH(1)-H2S(2) 22
4. P-x Diagram for CH-OHO )-N9(2) 25
5. P-x Diagram for C02(l)-H2$(2) 27
6. P-x Diagram for C02(1)-N2(2) 30
7. P-x Diagram for H2S(1)-N2(2) 32
8. Experimental Equipment 35
9. Calibration apparatus 37
10. P-x Diagram Experimental CH3OH(1)-C02(2) Data 39
-------
TABLES
Number Page
1. Composition of raw gas from coal gasification 6
2. Binary equilibrium data from literature 7
3. Relationships of binary Margules parameters to multicomponent
parameters 11
4. Binary interaction constants for Redlich-Kwong equation 13
5. Pure component parameters 13
6. Constants for Equation (12d) 16
7. Margules parameters 17
8. Calculated and equilibrium data for methanol-carbon
dioxide at 298°K 18
9. Calculated and equilibrium data for methanol-carbon
dioxide at 258°K 19
10. Calculated and equilibrium data for methanol-carbon
dioxide at 243°K 19
11. Calculated and experimental equilibrium data for
methanol-hydrogen sulfide at 258°K 21
12. Calculated and experimental equilibrium data for
methanol-hydrogen sulfide at 248°K 21
13. Calculated and experimental equilibrium data for
methanol-hydrogen sulfide at 273°K 23
14. Calculated and experimental equilibrium data for
methanol-nitrogen at 298°K 24
15. Calculated and experimental equilibrium data for
methanol-nitrogen at 310°K 24
16. Calculated and experimental equilibrium data for
carbon dioxide-hydrogen sulfide at 293°K 26
17. Calculated and experimental equilibrium data for
carbon dioxide-hydrogen sulfide at 277.6°K 26
18. Calculated and experimental equilibrium data for
carbon dioxide-nitrogen at 273°K 29
-------
TABLES (Contd)
Number
Page
19. Calculated and experimental equilibrium data for
carbon dioxide-nitrogen at 273°K 29
20. Calculated and experimental equilibrium data for
hydrogen sulfide-nitrogen at 300.1°K 31
21. Calculated and experimental equilibrium data for
hydrogen sulfide-nitrogen at 277.7°K 31
22. Comparison of model and experimental equilibrium data
for carbon dioxide-methanol at 258°K -- 40
23. Experimental vapor-liquid equilibrium data for mettianol-carboi
dioxide-nitrogen-hydrogen sulfide at -15.0°C 41
24. Comparison of calculated and experimental pressures and gas
compositions for CCL-H?S-N?-methanol -- 42
VI
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SECTION 1
INTRODUCTION
North Carolina State University, under the sponsorship of the
Environmental Protection Agency, Industrial Environmental Research
Laboratory at Research Triangle Park, is engaged in a research pro-
ject to study the environmental effects of coal gasification. So
that the results of this research will represent what might be ex-
pected from full-scale gasification operations, a highly flexible
and complete coal gasification and gas cleaning facility has been
constructed for the project. The overall objective of the project
is to characterize completely the gaseous and condensed phase emis-
sions from typical coal gasification-gas cleaning processes and to
determine how emissions depend upon various process parameters.
The North Carolina State University facility consists of a con-
tinuous fluidized bed gasifier, devices for removing participates,
condensables, and soluble matter from the raw synthesis gas, and an
acic gas removal system (AGRS). The gasifier operates at pressures
up to 100 psig, has a capacity of 50 Ib coal/hr, and can run with
eitfer steam-02 or steam-air feed mixtures to produce roughly 25 SCFM
of crude synthesis gas. The AGRS is modular in design so that alter-
native absorption processes may be evaluated with a minimal amount of
system modification.
The primary purpose of an AGRS is to remove most of the sulfur
containing compounds and, for some applications, the carbon dioxide.
The processes available to accomplish this employ either chemical or
physical absorption. In chemical absorption, acid gas species react
with alkaline compounds or solvents; the solvent is regenerated by
heating. One of the problems with processes employing chemical ab-
sorption for gas cleanup is that some of the crude gas components such
as carbonyl sulfide often react with the solvents to form decomposi-
tion products which prevent complete regeneration of the solvent. Pro-
cesses employing physical absorption capitalize on the differing solu-
bilities of various gases in liquids. Polar liquids are best suited
for these processes because they have a high degree of selectivity be-
tween acid gases and the valuable gas components. Methanol serves as
the solvent in Lurgi's proprietary Rectisol process. It is a good sol-
vent for hydrogen sulfide, carbonyl sulfide, carbon disulfide and car-
bon dioxide. As shown by plots of equilibrium constants in Figure 1,
methanol will absorb hydrogen sulfide preferentially to carbon dioxide.
The economics and desirability of these two types of gas cleanup
processes depend heavily on their particular application. However, the
use of methanol is particularly attractive and economical for cleanup
systems used in conjunction with coal gasification facilities for some
applications (1). This has been demonstrated by Lurgi in their design
and successful operation of the Rectisol process.
Because of its immediate importance, methanol is the first solvent
that will be studied in the NCSU gas cleaning test facility and solubility
data are needed to evaluate the AGRS performance.
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Figure 1. Solubility of gases in methane 1.
x"103
o
ii
O,
io2
hydrogen
nitrogen
carbon monoxide
temperature *F
-ISO -1OO -5O O 5O 1OO ISO
-------
In the facility the crude gas consists mainly of nitrogen, car-
bon monoxide, carbon dioxide, hydrogen, ethane, water, methane, hydro-
gen sulfide, other sulfur gases, and small amounts of higher hydrocar-
bons. The gas is dehydrated, compressed to 500 psi, cooled, and fed
to a packed absorber column, where it is contacted with cold methanol.
Carbon dioxide, hydrogen sulfide and other sulfur gases are absorbed
from the crude gas which is then discharged to a flare. Methanol en-
ters the absorber at about -30°F. The circulating solvent flows from
the bottom of the absorber through a pressure reduction valve to a
flash tank at 15-25 psig. Liquid solvent leaving the flash tank is re-
generated by stripping carbon dioxide and hydrogen sulfide with nitrogen.
A careful search of the literature determined that the data and
correlations required for an accurate analysis of the methanol-based
process are not available. Binary vapor-liquid equilibrium data for
some of the system components were located, but there appears to be
very little multicomponent data. It is clear that experience and pro-
prietary data or rough estimates based on ideal solution behavior have
been used in the design of existing facilities. It is not feasible to
meet the requirements of the current study using either of the above ap-
proaches; without accurate vapor-liquid equilibrium data it will be al-
most impossible to analyze the operation of the AGRS.
It is not practical to develop an exhaustive data base for multi-
component vapor-liquid equilibrium behavior. An approach that has been
used satisfactorily to describe systems similar to those encountered
here is to develop a thermodynamic model that uses binary equilibrium
data to predict multicomponent behavior.
The_purpose of this study was to begin developing a useful model
for multicomponent equilibrium behavior and to obtain binary vapor-liquid
equilibrium data for systems consisting of carbon dioxide, nitrogen, hy-
drogen sulfide and methanol. The range of temperatures, pressures and
compositions that have been considered in these studies include those
found in typical acid gas removal systems. Both theoretical and experi-
mental parts of this effort have been developed to a point that it seems
appropriate to describe the progress and point out directions for future
efforts. This report does that in its following sections.
SECTION 2
FUNDAMENTAL EQUATIONS
A typical composition of gases to be treated in acid gas removal
systems is shown in Table 1. The nature of these components indicates
that both gas and liquid solutions containing these materials will not
behave ideally at the system conditions; simple thermodynamic expressions
can not be used to describe their equilibrium behavior with methanol. It
is necessary, therefore, either to take extensive multicomponent equili-
brium data or to develop thermodynamic models using binary and perhaps
-------
Table 1. Composition of Raw Gas from Coal Gasification (3).
Component With 02 (Vol. %) With Air (Vol. %)
carbon dioxide, CCu
carbon monoxide, CO
hydrogen, H?
methane, CH.
nitrogen, N~
argon, A
hydrogen sulfide, HLS
carbonyl sulfide, COS
31.
15.
45.
8.
0.2
0.2
6800 ppmv
400 ppmv
10.
19.
22.
0 5
48.
. . .
0.5
150 ppmv
-------
limited multi component equilibrium data. The latter approach was used
in this research.
Systems consisting of C02, H2S, N;> and methanol were chosen for
study. An extensive search of the literature for equilibrium pro-
perties of binary mixtures formed by methanol and these three major
acid gas components is summarized in Table 2. While some miilticompon-
ent data are available in the literature, it often is not in the range
of conditions encountered in the AGRS processes or does not consider
one or more of the significant components of the AGRS mixtures.
Models and Correlations
The objective of thermodynamic models and correlations is to deve-
lop expressions relating temperature, pressure and compositions in co-
existing phases at equilibrium. For example, given a system pressure
and a liquid phase composition, a thermodynamic model could be used to
determine the system temperature and composition of a gas phase in equi-
librium with the liquid. To obtain useful expressions relating these
quantities it is necessary to begin with the fundamental thermodynamic
relationship which states that the fugacities of a component in coexist-
ing phase are equal at equilibrium; for vapor-liquid equilibrium
Two methods can be used to express these fugacities in terms of
systam temperature, pressure and composition of each phase. The first
follows from an exact thermodynamic treatment of the system to give
either of the following equations:
VT.V.n. ,- " T-
i J/i
' \ RT.
r ] (3)
To use Equation 2 a pressure explicit equation of state must be avail-
able and it must describe the system over the limits of the integration.
Equation 3 requires a volume explicit equation of state. This technique
can !)e and is used to estimate gas phase fugacities but, since no single
equa ;ion of state can be used to describe a mixture from ideal gas condi-
tions to liquid conditions at elevated pressures, a second approach is
used to estimate liquid fugacities.
Liquid fugacities may be defined in terms of their deviation from
ideal solution behavior, as given by Raoult's or Henry's laws. The work-
ing equation is
-------
Table 2. Binary equilibrium data from literature
Pressure
Binary Range, Atm
co2
co2
co2
co2
co2
co2
N2-
H2S
H2S
co2
co2
N2-
N2-
N2-
N2-
- CH3OH
- CH3OH
- CH3OH
- CH3OH
- CH3OH
- CH3OH
CH3OH
- CH2OH
- CH3OH
- H2S
- H2S
H2S
H2S
co2
co2
2,33
1 ,30
16,43
2,60
7,69
5,80
35,63
2,10
1
10,89
7,82
3,136
11,204
24,143
37,125
Temperature
Range, °C Source
0,-SO
185, -50
60,15
25
0,75
40,25
60,15
0.-25
30, -30
89,0
91, -48
-45, -73
71, -17
15, -20
25, -40
Yorizane et al (4)
Bezdel and Teodorovi:h (5)
Hemmaplardh and King (6)
Katayama et al (7)
Krichevakii and Lebedova (8)
Ohgaki and Katayama (9)
Hemmaplardh and Kin
-------
where YI defines the deviation of the solution from ideal behavior
and f." 1S the fugacity of the reference state at the system con-
ditions. If the ideal solution is defined in terms of Raoult's law
then
When this definition of the reference state is used
Y.J ->• 1.0 as x. •+ 1.0.
A major difficulty associated with using the Raoult's law definition
of the ideal state is evaluating p* for components having a critical
temperature below the system temperature. In this case a hypotheti-
cal value of the vapor pressure can be estimated by extrapolation of
P* vs. T expressions.
When the ideal solution is defined in terms of Henry's law
fi = ^fi ^inf = Hi (HenrY's constant) (6)
Using this definition of the reference state T. -> 1.0 as x + 0
0 Connell (16) has noted severe difficulties ir\ using Henry's law to
define the reference state for multicomponent systems. Therefore,
Kaoult s law has been used in this research.
Although significant progress has been made in developing tech-
niques to predict activity coefficients from solution theory, it is
almost always necessary to rely on experimental data for the develop-
ment of suitable correlative and predictive routines. The van Laar,
Margules, Wilson (17) and UNIQUAC (18) equations have been used with
varying degrees of success to correlate activity coefficients. In most
cases fairly accurate predictions of multicomponent system behavior can
be obtained by using parameters evaluated by fitting these equations to
data for each binary mixture making up the system.
-------
SECTION 3
MODEL DEVELOPMENT
The primary constituents of methanol -based acid gas removal sys-
tems are COz, ^S, N£ and methanol. A description of the phase equi-
librium behavior of quaternary mixtures of these components is neces-
sary for a quantitative analysis of absorber and stripper operations
in these processes.
The initial approach in this investigation was patterned after
that reportedly used industrially for process development and design
studies with high pressure nonideal systems (19). A modified Redlich-
Kwong equation was used to calculate fugacities of the qaseous compon-
ents, the four suffix Margules form of the Wohl equation was used to
correlate activity coefficients, and the Chao-Seader equatior. (20) was
used to estimate the fugacity of the reference state (pure liquid at
the system conditions).
The Wohl equation gives the activity coefficient of a component
in a mul ticomponent solution as
, = 4 I I Zx.x.x Piik - I I I I x.x x x, ;•:
1 j = l k=l £-1 J k l 1Jk£ i = l j = l k=l f.= l n J k '' ljk'
(7)
where the values of 3--kf are related to binary Margules constants and
a ternary constant; J " the Margules constants are defined for a
binary mixture by the expression
In Y- = x.2 [A.. + 2(A.. - A.. - D..) x. + 3D., x.2] (8)
Yi J iJ Ji ij iJ i ij i l '
Table 3 gives the definition of 8,-jk(, in terms of the binary Margules
constants. Following the observation of Adler, Friend and Pigford (21),
the ternary constants C*-^ were set equal to zero. Nonlinear least
squares techniques can be used to estimate the Margules constants in
Equation 8. Adler, Ozkardesh and Schreiner (22) showed that the tem-
perature dependence of these parameters follow either a directly pro-
portional or inversely proportional relationship with temperature.
The Redl ich-Kwong equation of state as modified by Prausnitz and
Chueh (23) was used with Equation 2 to estimate vapor phase fugacity
coefficients. This equation of state is described in the following re-
lationships.
-------
Table 3. Relationships of binary Margules parameters to multicompon-
ent parameters.
Q
Combination ijkfc
i = j = k = s, 0
i = j = if k A^/4
i = * t J = k (A.. + A... - D...)/6
J = *.* M i C(Akj , AIJ + A.k1) -
where
= Aik
C*., = ternary constant
i JK
, _
"
R T
lL Ob)
a- ? y.y.a.- (9c)
« /. L J iJ -\ t i
? 2 5
R T
(9d)
-------
TT
? 9 R
T
U (9e)
4T
(9g)
(v
All of the mixing rules were taken from Prausnitz (24) except (9g)
which came from Tarakad and Darmer (25). The binary constant, k-j-,
is independent of composition, temperature and density. Table 4
lists the values k-jj used in this study and their sources. The dimen-
sionless pure component constants are characteristic of a particular
substance and are calculated from volumetric data for the pure satur-
ated vapor. All pure component parameters used are listed in Table 5.
An expression for the vapor phase fugacity coefficient of a com-
ponent in a mixture of n components is obtained by combining Equations
2 and 9.
Table 4. Binary interaction constants for Redlich-Kwong equation
k
System ij Source
CH3OH - C02 0.063 (25)
CH3OH - H2S 0.14 (25)
CH3OH - N2 0.14 (25)
C02 - H2S 0.08 (24)
C02 - N2 0.048 (25)
HS - N 0.105 (12)
10
-------
Table 5.
Pure component parameters
Component a ^b
CH3OH
co2
H2S
N?
0.8451
0.4470
0.4340
0.4290
0.2075
0.0911
0.0882
0.0870
*
Tc
(°K)
512.6
304.2
373.2
126.2
*
Pc
(atm)
79.9
72.8
88.2
33.5
*
Vc
(cc/mole)
118.0
94.0
98.5
87.5
*
0)
0.559
0.225
0.100
0.040
t
From Prausnitz (24) except values for methanol which were calculated
from volumetric data of Eubank (27).
From Reid, et al. (26).
n
2 ^ y a
- m
, .
K. V-D V-D RT h
DT1.5.2 T" " V?b" " RT
RT b
The molar volume of the mixture, V, is found by solving Equation 9 for
the largest real root.
Various equations were used to calculate the fugacities of the
reference state, chosen to be the pure liquid. The exact thermodynamic
relationship given in Equation 5 was used for methanol. The vapor pres-
sure for methanol was calculated from the Antoine equation.
log [P*(mm Hg)] = 7.87862 - 1473.11/[T(°K) - 43.247] (11)
The Antoine constants used were determined from experimental vapor pres-
sure data from Eubank (27) for temperatures from -40°C to +30°C. The
molar volume of methanol was calculated from a corresponding states cor-
relation of Chueh and Prausnitz (Reid, Prausnitz and Sherwood (26)).
11
-------
9Z N(P-P*) ,/q
- P = P[l +-±- - 1 ' 02
s
N = (1.0 - 0.89u)[exp(6.9547 - 76.2853 Tr + 191.3060 Tp
- 203.5472 T^3 + 82.8731 Tr4)] (12a
2b)
pc = l/Vc 02c)
Vr(j) = a.. + b. Tr + c. T2 + d.. T^ + 6j/Tr + f . ln(l-Tr) (12d)
The constants for Equation (12d) are given in Table 6. Using the
above correlation for liquid molar volume and evaluating the integral
in Equation 5, the following expression is obtained for the pure li-
quid reference fugacity.
P 9Z N(P-P.*)
V • P,* V «P (gRT^V «' - -^ -- >8/9 - 1]) ("I
si ci ci
For carbon dioxide and hydrogen sulfide, a three parameter reduc-
ed states correlation was used (28)
f°
log _i.= i0g v(°) + a, log v(1) (14)
log v(0) = Bo + B1 Pr + B2 Pr2 - log Pp (14a)
B = -20.651608 + 84.517272 Tr - 15.376424 1^ + 152.65216 Tf3
- 84.899391 T 4 + 24.84688 T^ - 2.9786581 T^6 (14b)
12
-------
If0.8>Tr>0.3,
,0.286-
(He)
If 0.9 > Tr >_ 0.8, B1 = 0.321895 Tr - 0.184316
If 1.8 >.Tr > 0.9, B1 = 58.16962 - 326. 54444 T
+ 775.11716 Tr - 1006.8122 Tr
+• 87.677429 T 6 - 9.2517986 T 7
773.32667 T 4 - 351.56938 T 5
Table 6. Constants for equation (12d)
a . h . r . d . e .
1
0
1
2
JL _i
0.11917 0.009513
0.98465 -1.60378
-0.55314 -0.15793
_1
0.21091
1.82484
-1.01601
JL _1
-0.06922 0.07480
-0.61432 -0.34546
0.34095 0.46795
f .
JL
-0.084476
0.087037
-0,239938
If 0.8 > Tr >_ 0.3 , B2 = 0
If 0.9 > T > 0.8 , B0 = 0.0549369 (0.8 - Tj
r — c '
If 1.0 > T ^0.9 , B2 = 0.673344 x 10"3 - 0.685226 x 10
-2
If T >_ 1.0 , B2 = 0.72203901 - 2.7182597 Tr + 3.984423 T^
- 2.8712448 T 3 + 1.0202739 T 4 - 0.14314712
log
-- log ^
(1
log y
(1)
3P
(14d)
(He)
13
-------
"log vn c = - 660.08698 + 7766.7774 T - 40007.379 T '
U. b r r
+ 116582.6 Tr3 - 209756.24 Tr4
+ 238673.14 T 5 - 167856.45 T 6
r r
+ 66762.602 Tr7 - 11504.984 Tr8 (14f)
rp~ - = - 0.28997623 + 0.96418856 Tr - 1.3344703 Tp
+ 0.82575807 T 3 - 0.18939410 T 4 (14n)
r r v J/
For nitrogen, the Chao-Seader equation with adjusted parameters
was used (29).
f°
log -p1 = 2.7365534 - 1. 981831 0/T - 0.51487289 Tr
+ 0.042470988 T 2 - 0.002814385 Tr3 + (-0.029474696
+ 0.021495843 T) P - log P + <*.
To use the techniques described, activity coefficients were evalu-
ated from literature x-y-P-T data on the six possible binary mixtures
of the four component system: methanol-C02, methanol-^S, methanol-N2,
C02- H2S, C02-N2, and H2S-N2- Using the references listed in Table 3,
Margules parameters and their temperature dependence were determined us-
ing a non-linear least squares procedure as described by Bass (30). The
results of these evaluations are summarized in Table 7 and discussed in
the following paragraphs. Details of computer programs are given by
Bass (30).
14
-------
Table 7. Margules parameters
Binary
CH3OH - C02
CH3OH - H2S
CO, - H0S
l C-
C02 - N2
H2S - N2
(°K)
298.2
258.2
293.2
253.3
300.1
1.1672
1.0694
0.6607
1.2832
0.6165
A-i
0.594C
0.6608
0.5141
0.4359
0.7885
CH3OH -
n
i
0.5056
1.0068
0.5991
2.5135
-0.1280
298.2 0.6317 0.7683 1.6583
Temperature
dependence
A" = A1 T]/T2
A" = A1 T,/T2
A" = A1 T2/T]
A" = A1 T]/T2
A" = A1 (T-,
D" = D1 (Tg/^
A" = A1 T,/T2
D" = D' T2/T1
Model parameters for the methanol-carbon dioxide system were de-
termined from the 298°K data of Katayama et al. (7). The model fit is
illustrated in Table 8. Predictions are compared with the 258°K and
243°K_data of Yorizane et al. (4) in Tables 9 and 10. Vapor phase com-
positions at the latter two temperatures were not available for com-
parison. Average deviations of the predicted equilibrium pressure from
measured values increase with further extrapolations from 298°K but
plots of the model (Figure 2) appear to be asymptotic to the experimen-
tal values at higher concentrations.
Model parameters for the methanol-hydrogen sulfide system were de-
termined from the 258°K data of Yorizane et al . (4); the fit is shown
in Table 11. Vapor compositions were not given and the concentration
of methanol was reported to be too low to be measured. For the pur-
poses of determining activity coefficients, the vapor composition of
methanol was assumed to be 0.0001. Extrapolations to 273°K and 248°K
show an excellent fit to the experimental data except for high hydrogen
sulfide concentrations at 273°K where the assumed vapor composition is
expected to be most in error. Figure 3 graphically compares the model
to Yorizane's data.
15
-------
Figure 2. P-x Diagram for CH3OH(1)-C02(2!
60
50
40
30
2O
10
DATA OF
YORIZANE
A 258'K
DATA OF
KATAYAMA
• 298'K
0.2
O.4 O.6 O.8 l.O
16
-------
Table 11. Calculated and experimental equilibrium data for metha-
nol-hydrogen sulfide at 258°K.
0.165 2 0
0.231 3.0
0.298 3 4
0.367 4.2
0.403 4.4
0.490 5.0
0.585 5.4
0.662 5.8
Average percent deviation in P = 3.7.
Table 12.
0.203
0.290
0.327
0.465
0.582
0.733
Calculated and experimental equilibrium data for metha-
nol-hydrogen sulfide at 248°K.
P
exp
(atm)
2.0
2.5
3.0
3.4
4.0
4.3
Average percent deviation in P = 31
17
-------
Figure 3. P-x Diagram CH,OH(•)-H?S(2).
\J (—
10 r
8
0.2
DATA OF
YORIZANE
• 273-K
A 258 -K
• 248 K
- MODEL
__l 1 1 >
0.4 0.6 0.8 l.O
-------
Table 13. Calculated and experimental equilibrium data for metha-
nol-hydrogen sulfide at 273°K.
0.092 2.0
0.199 4.0
0.329 6.0
0.453 7.5
0.484 8.0
0.608 9.1
0.743 9.8
0.840 10.0
Average percent deviation in P = 7.8.
High pressure vapor-liquid equilibrium data for the methanol-nitro-
gen system are scarce. Model parameters for this system were determined
from the 298°K data of Hemmaplardh and King (6) with the fit illustrated
in Table 14. As shown in Table 15 extrapolations to 310°K produced good
agreement between predicted and experimental vapor phase compositions but
poor agreement between predicted and experimental pressures. Figure 4
shows how the model fits both sets of data. More data are needed before
an improved fit can be obtained. However, the region of primary inter-
est is for pressures at or below 35 atmospheres where the model predic-
tions are expected to be more accurate.
Model parameters for the carbon dioxide-hydrogen sulfide system were
determined from the 293°K data of Bierlein and Kay (10); the resulting
fit is illustrated in Table 16. Extrapolations were made to 277.6°K for
comparison with data from Sobocinski and Kurata (11) with the resulting
fit shown in Table 17 and Figure 5.
19
-------
Table 14. Calculated and experimental equil
ibrium data
for metha-
nol-nitrogen at 298°K.
X2
0.0241
0.0295
0.0483
0.0575
Average
P
exp
(atm)
35.8
35.1
49.8
62.8
Pcalc
(atm)
32.3
37.7
52.8
58.4
percent deviation
DP
(atm)
-3.5
2.6
3.0
-4.4
: in P,
y2
exp
0.9939
0.9938
0.9952
0.9958
7.6; in y,
y2
calc
0.9941 0
0.9934 -0
0.9954 0
0.9957 -0
0.02.
Dy
.0002
.0004
.0002
.0001
Table 15
X2
0.0355
0.0539
0.0559
0.0611
Average
. Calcul
nol-ni
exp
(atm)
35.0
43.7
47.7
48.6
ated and
experimental equil
ibrium data
for metha-
trogen at 310°K.
Pcalc
(atm)
40.0
52.1
53.2
55.8
percent deviation
DP
(atm)
5.0
8.4
5.5
7.2
: in P,
y?
L.
0.9891
0.9909
0.9914
0.9915
15; in y,
y?
^calc
0.9902 0.
0.9919 0.
0.9921 0.
0.9923 0.
0.09.
Dy_
0011
0010
0007
0003
Table 16
. Calcul
ated and
xide-hydrogen
X2
0.980
0.910
0.821
0.715
0.580
0.437
0.261
0.180
exp
(atm)
20.0
25.0
30.0
35.0
40.0
45.0
50.0
52.0
calc
(atm)
19.4
26.4
31.1
34.7
38.9
43.4
48.2
49.7
experimental equil
sulfide
DP
(atm)
-0.6
1.4
1.1
-0.3
-1.1
-1.6
-1.8
-2.3
at 293°K.
y2
exp
0.869
0.685
0.550
0.457
0.370
0.290
0.189
0.130
ibrium data
y?
^calc
0.851
0.633
0.533
0.462
0.381
0.293
0.192
0.147
for carbon dio-
Dy_
-0.018
-0.052
-0.017
0.005
0.011
0.003
0.003
0.017
Average percent deviation: in P, 3.4; in y, 4.1.
20
-------
Figure 4. P-x Diagram for CH-OH(1)-N9(2).
60
50
40
< 30
Q.
20
10
298
31O -
DATA OF
HEMMAPLARDH
• 298 'K
• 31O 'K
- MODEL
' f
0.02 O.04 0.06 O.O8
21
-------
Table 17. Calculated and experimental equilibrium data for cartx n
dioxide-hydrogen sulfide at 277.6°K.
Pexp calc DP y~ y~
^2_ (atm) (atm) (atm) exp __caj£ Dy_
0.967 13.61 14.60 0.99 0.855 0.774 -0.081
0.915 17.0 18.14 1.14 0.660 0.621 -0.039
0.850 20.41 20.88 0.47 0.540 0.532 -0.008
0.757 23.0 23.50 0.50 0.470 0.458 -0.012
0.605 27.22 27.07 -0.15 0.355 0.363 0.008
0.485 30.0 29.89 -0.11 0.214 0.29? 0.078
0.340 33.0 33.15 0.15 0.210 0.210 0
0.290 34.02 34.15 0.13 0.185 0.184 -0.001
Average percent deviation: in P, 2.5; in y, 7.3.
Model parameters for the carbon dioxide-nitrogen system were de-
termined from the 253°K data of Kaminishi and Toriuini (15). Extra-
polations to 273°K were compared to data from the same source. These
results are shown in Tables 18 and 19 and in Figure 6. The best fit
was obtained at pressures below 80 atmospheres.
Model parameters for the hydrogen sulfide-nitrogen system we e
determined from the 300.1°K data of Besserer and Robinson (13). , x-
trapolation to 277.7°K was fit best by having the Margules constants
proportional to temperature to the 1.7 power. The resultinc fits are
shown in Tables 20 and 21 and Figure 7.
Poor fits were observed for both the methanol-nitrogen and car-
bon dioxide-nitrogen systems. Data were not available over a wide con-
centration range for either system. For instance, in the carbon diox-
ide-nitrogen system, data at low nitrogen concentrations allowed an ac-
curate determination of fy] but the lack of data at low carbon dioxide
concentrations prevented an accurate determination of An 2- Thus, the
model fits the experimental data more closely at low nitrogen concen-
trations anc1 consequently at low pressures. This problem was even more
pronounced in the methanol-nitrogen system where the data were even
more limited and had significant scatter.
Good fits were obtained for the methanol-carbon dioxide, metha-
nol-hydrogen sulfide and carbon dioxide-hydrogen sulfide systems.
Among these three systems, the average percent deviation in pressure
was 6.6 and in vapor phase composition was 2.3. In the absence of tern-
ary effects, it is reasonable to assume that the model will provide a
good estimate of total pressure and vapor phase composition for a mix-
ture consisting of these three components. Predictions for the four
component system containing nitrogen is expected to be less reliable.
-------
Figure 5. P-x Diagram for CCUl )-H9S(2).
60
50
40
30
<
tL
20
10 -
DATA OF
SOBOCINISKI, BIERLEIN
• 293 'K
• 277.6 -K
- MODEL
0-2 0.4 0.6 0.8
1.0
23
-------
Table 18. Calculated and experimental equil
bon-dioxide-nitrogen at 253°K.
p P
X2
0.023
0.059
0.095
0.170
0.210
0.246
Average
exp
( a trn )
36.5
52.0
67.5
100.5
115.8
125.7
percent
calc
(atm)
33.7
51.1
63.5
86.5
101.2
117.2
deviation:
DP
(atm)
-2.8
-0.9
-4.0
-14.0
-14.6
-8.5
in P,
y
2exp
0.382
0.509
0.568
0.596
0.582
0.557
8.1 ; in y,
ibrium data for car-
V^
calc
0.411
0.555
0.609
0.668
0.692
0.713
14.
Dy
-0 029
0.046
0 041
0.072
0 1 10
0.156
Table 19. Calculated and experimental equilibrium data
bon-dioxide-nitrogen at 273°K.
ii
0.076
0.133
0.188
0.215
exp
(atm)
73.2
95.0
108.7
112.3
Pcalc
(atm)
77.9
100.3
122.8
135.5
OP
(atm)
4.7
5.3
14.1
23.2
y
exp
0.353
0.395
0.377
0.366
calc
0.400
0.444
0.471
0.487
for car-
D^
0.047
0.049
0.094
0.121
Average percent deviation: in P, n; in y, 21
24
-------
Figure 6. P-x Diagram for C02(1)-N2(2).
140
120
1OO
80
60
4O
20
253 K
DATA OF
KAMINISHt
• 253-K
• 273-K
- MODEL
—I 1 1 1 I
0.2 0.4 O.6 0.8 1.0
25
-------
Table 20. Calculated and experimental
sulf ide-nitrogen at 300. 1°K
X2
0.0071
0.0153
0.0239
0.0317
0.0401
0.0495
0.0570
0.0663
0.0762
0.0844
0.0957
Average
exp
(atm)
34.7
50.7
69.1
86.1
101.9
119.0
135.7
152.9
170.7
187.2
204.1
calc
(atm)
33.3
49.6
66.7
82.1
98.8
117.3
131.9
149.9
168.7
183.9
204,4
percent deviation:
DP
(atm)
-1.4
-1.1
-2.4
-4.0
-3.1
-1.7
-3.8
-3.0
-2.0
-3.3
0.3
in P, 2.4;
equilibrium data for hydrogen
y2
e*£
0.332
0.488
0.576
0.628
0.653
0.674
0.684
0.691
0.695
0.694
0.693
in y, 3.6.
calc
0.352
0.514
0.598
0.642
0.672
0.692
0.703
0.711
0.716
0.719
0.720
Dy_
0.020
0.026
0.022
0.014
0.019
0.018
0.019
0.020
0.025
0.025
0.027
Table 21.
X2
0.0035
0.0102
0.0186
0.0241
0.0314
0.0377
Calculated and experimental
sulf ide-nitrogen at 277.7.
P P
exo calc DP
(atm) (atm) (atm)
17.1 19.7 2.6
33.7 35.7 2.0
52.1 55.9 3.8
67.4 69.0 1.6
85.1 86.2 1.1
102.1 101.0 -1.1
equi 1 ibrium
y y
exp
0.271
0.582
0.698
0.743
0.769
0.791
data
9
calc
0.383
0.622
0.726
0.760
0.788
0.804
for hydrogen
Dy
0.112
0.040
0.028
0.017
0.019
0.013
Average percent deviation: in P, 5.5; in y, 9.8.
26
-------
Figure 7. P-x Diagram for H9S(1)-M0(2).
TOO
277.7
80
60
0.
40
20
DATA OF
BESSERER
• 3O0.1 K
• 277.7 'K
- MODEL
-I 1 1 I (
0.01 0.02 0.03 0.04 0.05
27
-------
It is difficult to estimate the error involved as it was not possible
to test the binary predictions for all of the nitrogen containing sys-
tems in the concentration and pressure range of interest.
To test predictions of the equilibrium behavior of C02-HpS-N2- me-
thanol mixtures it was necessary to obtain x-y-P-T data experimentally.
The apparatus and procedures used to do this are described in the fol-
lowing section. Results of the experimental study are discussed in
Section 5.
SECTION 4
EXPERIMENTAL EQUIPMENT AND PROCEDURE
The experimental apparatus used in obtaining high pressure vapor
liquid equilibrium data is shown in Figure 8. The stainless steel
565-ml equilibrium cell was equipped with baffles and agitated with a
magnetic stirrer. Pressures were measured using a 16 inch Heise gauge
graduated in 0.5 psi increments up to 85 atm. The gauge had a guaran-
teed accuracy of 0.1 percent of full scale. Its calibration was check-
ed using an Ashcroft type 1327 portable deadweight tester. Tempera-
tures were measured using a copper-constantan thermocouple and a digi-
tal thermocouple meter calibrated against known temperatures. Refrig-
eration was provided with a Harris industrial freezer. Control of the
bomb temperature to within 0.1°C was accomplished using a fan, heater,
and Thermistemp temperature controller. All valves were teflon packed
and rated for high pressure use.
Methanol used was Fisher Spectranalyzed Qy with a stated purity of
99.95 percent. It was checked for water content using a gas chromato-
graph and dried when necessary with 3A molecular sieves to less than
0.05 percent water. Coleman grade carbon dioxide with a stated purity
of 99.99 percent and ultra pure carrier grade nitrogen with a stated
purity of 99.999 percent supplied by Air Products and Chemicals were
used. A mixture of 1.5 percent hydrogen sulfide in nitrogen supplied
by Linde Division of the Union Carbide Corporation was used.
Analysis of the samples was done on a Tracor Model 550 gas chro-
matograph equipped with a thermal conductivity cell, temperature pro-
grammer and heated gas sampling valve. A 3 meter long by 0.0032 meter
diameter column packed with Porapak QS was used to achieve the separa-
tions. Compositions were determined using Southern Analytical"s Supe-
grator 3 digital integrator and a Leeds and Northrop strip chart re-
corder. For the methanol-carbon dioxide system, the chromatograph was
calibrated by mixing the components at below atmospheric pressures in
the 10-5, vessel shown in Figure 9. Concentrations were determined from
pressure measurements. For the multicomponent system, calibrations were
done by injection of known amounts of each component.
28
-------
Figure 8. Experimental Equipment.
r\3
4-way valve
temperature
controller
digital
thermocouple
pressure gauge
CO,
N
Mix
heater
capillary
tubing
heated
sample
containers
freezer
equilibrium cell
magnetic stirrer
-------
Figure 9. Calibration apparatus.
TO G.C.
SEPTUM
FLANGE
9.61 VESSEL
MANOMETER
GAS SOURCE
VACUUM PUMP
MAGNETIC
STIRRER
30
-------
Before each series of runs, the equilibrium cell was filled with
approximately 250-ml of methanol and then purged until no air or water
could be detected leaving the cell. The cell was then pressurized
and brought to the desired temperature. The gases of interest were
added to achieve the desired pressure. The contents were agitated for
at least six hours and then allowed to sit unagitated for at least
twelve hours prior to sampling. Samples were allowed to expand through
the capillary tubing into the evacuated sample containers. Sampling
was done quickly and the cell pressure was seldom disturbed by more
than 0.14 atm. The vapor sample container was pressurized to approxi-
mately 1.4 atm with helium and then both containers were heated to 140°C
for five hours. The pressures in the containers were monitored to in-
sure that they did not approach the vapor pressure of methanol. The
contents of each container were analyzed a minimum of three times using
the gas chromatograph.
SECTION 5
EXPERIMENTAL RESULTS
Methanol-Carbon Dioxide
For the purpose of verifying the proposed experimental procedure,
vapor liquid equilibrium data were obtained for the carbon, dioxide-metha-
nol system. The data are compared with those of Yorizane et al. (4) in
Figure 10 and to the model predictions in Table 22. From Figure 10, it
is evident that the liquid phase concentration and pressure measurements
show good agreement with those of Yorizane et al. (4). Table 22 shows
the average percent deviation between calculated and experimental pres-
sures to be 8.8, which is identical to that calculated from Yorizane1s
data. The vapor phase concentration of methanol is higher than expected
and shows significant scatter. It is believed that liquid methanol be-
came entrained with the vapor leaving the equilibrium cell during sampl-
ing. In these experiments pressures were measured to within 0.3 atm.,
temperatures within 0.1°C and compositions within 2 percent. These bi-
nary carbon dioxide-methanol vapor-liquid equilibrium data demonstrate
the utility of the equipment and experimental procedure in obtaining and
analyzing liquid samples. Using the present vapor sampling procedure
unusually high and unpredictable concentrations of methanol can be ex-
pected in the vapor phase. Further experiments should be conducted in
an apparatus modified to improve vapor sampling. At temperatures less
than or equal to -15°C and at pressures greater than approximately 10
atm., the assumption that the vapor is methanol free introduces an er-
ror of less than 1 percent. If this sampling technique does not appre-
ciably affect equilibrium, the equipment can be used to obtain data for
a multicomponent system to serve as a check of the thermodynamic model.
31
-------
Figure 10. P-x Diagram Experimental CH_OH(1)-C02(2) Data.
20
16
5
8
T = 258-K
• DATA OF YORIZANE
• DATA-THIS INVESTIGATION
- MODEL
O.2 O.4 O.6 O.8
l.O
32
-------
Table 22. Comparison of model and experimental equilibrium data for
carbon dioxide-methanol at 258°K.
p p
exp calc
X2 (atm) (atm)
0.085 5.03 5.76 0.73 0.936 0.997 0.061
0.149 8.45 9.24 0.79 0.995 0.998 0.003
0.227 12.13 12.84 0.71 0.980 0.999 0.019
0.349 16.30 17.21 0.91 0.912 0.999 0.087
Average percent deviation: in P, 8.8; in y, 4.6.
Multicomponent Equilibrium Data
Experimental equilibrium data were obtained for C02~HoS-N2
nol mixtures at -15°C. These data are presented in Table 23. ^Preci-
sion of the analytical equipment varied among the system components:
mole fractions of the most concentrated component could be determined
to 1.5»; N2 and CC^ in the liquid could be determined to 2 and 81-, re-
spectively; methanol and C02 in the vapor to within Zc: and 1«, respec-
tively. Confidence in hydrogen sulfide compositions was good at high
pressures, but at low pressures a decrease in sensitivity of the ther-
mal conductivity detector caused ^S determinations to became less pre-
cise.
Equilibrium temperature and liquid composition were used with the
algorithm BULP (described by Bass (30) to calculate system pressure
and gas composition at equilibrium. Comparisons between calculated and
experimental results are given in Table 24. The average deviation be-
tween calculated and experimentally measured pressures is 10.81". Con-
siderable differences between experimental and calculated gas composi-
tions may also be noted. In general, pressure predictions are more ac-
curate at low pressures while vapor composition predictions are more
accurate at high pressures. The average percent deviation in pressure
is only slightly worse than the 6.6 parcent value calculated for all
binaries used in determining and checking the model parameters. How-
ever, the percent deviation in vapor compositions is much worse. While
not conclusive, it is believed that tnis unusually large discrepancy is
due to ternary effects among the components carbon dioxide, hydrogen
sulfide and methanol. Calculations made ignoring ternary effects show
vapor compositions for carbon dioxide and hydrogen sulfide to be larger
than the experimental values. Thus, gas cleanup systems will likely be
overdesigned if calculations are based on vapor-liquid equilibrium pre-
dictions made by considering only binary interactions.
33
-------
Table 23. Experimental vapor-liquid equilibrium data for ir,ethano:-carbon dioxide-nitrogen-
hydrogen sulfidc at -15.0°C.
P(atm)
8.23
9.37
9.75
19.9
27.9
34.6
*CH3OH
0.945
0.927
0.929
0.909
o.i:>-
0.920
Xco2
0.052
U . w u y
0.057
0.033
0.0/0
0.055
X
0.002
r\ A A t)
\J • W «^
0.003
0.007
0.010
0.013
XH2S
0.001
0.001
0.001
0.001
0.001
0.002
ycH3c:;
0.003
0.002
0.002
0.002
0
0
yco2
0.319
0.276
0.270
0.134
0.101
0.035
\
O.C77
0. 722
0.726
0.863
0.0 j'd
0.913
!!2S
0.001
-
O.UD1
0.001
0.001
0.002
-------
Table 24. Comparison of calculated and experimental pressures and gas compositions for C00 - H0S - N0 -
methanol *- <• i
P(atm)
yCH3OH
'CO,
exp
calc
DP
exp calc Dy
exp calc
exp calc Dy
exp calc Dy
8.23 7.77 0.46
9.37 10.77 -1.40
9.75 10.66 -0.91
19.9 19.1 0.8
27.9 23.8 4.1
34.6 28.8 5.8
0.003 0.002 0.001
0.002 0.002 0
0.002 0.002 0
0.002 0.001 0.001
0 0.001 -0.001
0 0.001 -0.001
0.319 0.493 -0.174
0.276 0.466 -0.190
0.270 0.459 -0.189
0.134 0.327 -0.193
0.101 0.234 -0.133
0.085 0.188 -0.103
0.677 0.502 0.174
0.722 0.053 0.192
0.726 0.537 0.189
0.863 0.863 0.192
0.898 0.898 0.134
0.913 0.913 0.103
0.001 0.004 -0.003
0.003 -
0.001 0.003 -0.002
0.001 0.002 -0.001
0.001 0.001 0
0.002 0.002 0
-------
SECTION 6
SUMMARY OF RESULTS
1. A thermodynamic model was developed which successfully corre-
lated binary vapor-liquid equilibrium data for the methanol-hy-
drogen sulfide, methanol-carbon dioxide, carbon dioxide-hydrogen
sulfide and nitrogen hydrogen sulfide systems. The lack of ade-
quate data prevented an adequate correlation for the methanol-ni-
trogen and carbon dioxide-nitrogen systems.
2. Experimental vapor-liquid equilibrium data were obtained for the
methanol-carbon dioxide system at -15.0°C which showed excellent
agreement in pressure and liquid phase composition measurements
to literature data. Vapor compositions were unusually high in
methanol.
3. Vapor phase composition and total pressure predictions made using
the thermodynamic model showed poor agreement with experimental
results. Predicted vapor compositions for carbon dioxide and hy-
drogen sulfide were larger than the experimental values. It is
postulated that this is the result of ternary effects among this
system components.
4. Models that do not require experimental data on ternary effects
should be tested.
5. Additional modifications to the experimental equilibrium cell are
needed to improve the vapor sampling technique.
36
-------
LITERATURE CITED
1. Ranke, Gerhard, "The Rectisol Process for the Selective Removal
of C02 and Sulfur Compounds from Industrial Gases," Chemical
Economy and Engineering Review, 4, 25 (1972).
2. Morrison, James A., "Technical Data Manual for EPA Gas Cleaning
Facility Located at North Carolina State University," EPA Con-
tract Mo. 68-02-2601, 1977.
3. Fisch, E. J. and J. A. Sykes, "Synthetic Fuel Gas Purification
Using Shell Treating Processes," ACS Meeting, Dallas, Texas,
1973.
4. Yorizane, Masahiro, S. Sadamoto, H. Masuoka, and Y. Eto, "Gas
Solubilities in Methanol at High Pressure," Kogyo Kagaku Zashi.
72, 2174-2177 (1969).
5. Bezdel, L. S. and V. P. Teodorovich, "The Solubilities of Carbon
Dioxide Hydrogen Sulfide, Methane, and Ethylene in Methanol at
Low Temperatures," Gazovaia Promsnlenmost, (Moscow), 8, 38-43
(1958).
6. Hemmaplardh, B. and A. D. King, Or., "Solubility of Methanol in
Compressed Nitrogen, Argon Methane, Ethylene, Ethane, Carbon
Dioxide and Nitrous Oxide. Evidence for Association of Carbon
Dioxide with Methanol in the Gas Phase," The J. of Phys. Chem.,
76, 2170-2175 (1972).
7, Katayama, Takashi, 0. Kazunari, G. Maekawa, M. Goto and T. Nagano,
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37
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39
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NOMENCLATURE
DP - Pcalc ' Pexp
DV = ycalc " yexp
f. = fugacity of i
f o = fugacity of i in reference state
H = Henry's constant for i
i
k = binary constant, Equation 9f
ij
n = moles of i
i
p = pressure
p = reduced pressure
r
p = critical pressure of i
ci
p * = vapor pressure of i
R = gas constant
7 = temperature
T = reduced temperature
r
v = molar volume
V = critical volume
V L = liquid molar volume of i
i
v = mole fraction i in gas phase
y\
Greek Symbols
activity coefficient of i
fugacity coefficient of i
fugacity coefficient of saturated i
Pitzer accentric factor
dimensionless pure component constants, Table 6
40
Yi
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TECHNICAL REPORT DATA
(Please read Inuructions on the reverse before completing}
1 REPORT NO.
EPA-600/7-79-097
3. RECIPIENT'S ACCESSION NO.
4. TITLE ANDSUBTITLE
The Solubility of Acid Gases in Methanol
5. REPORT DATE
April 1979
6. PERFORMING ORGANIZATION CODE
7~~AUTHOR(S)
j.K.Ferrell, R.W.Rousseau, and D.G.Bass
8. PERFORMING ORGANIZATION REPORT NO.
9. PERFORMING ORGANIZATION NAME AND ADDRESS
North Carolina State University
Department of Chemical Engineering
Raleigh, North Carolina 27650
10. PROGRAM ELEMENT NO.
E HE 62 3 A
11. CONTRACT/GRANT NO.
Grant No. R804811
•12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF RfcEQRT AND PEJ
Final: 10/76 - 9/78
RIOD COVERED
14. SPONSORING AGENCY CODE
EPA/600/13
15. SUPPLEMENTARY NOTES jERL_RTppro.ectofficer
2708.
Smith , MD- 61 , 919/541-
16 ASSTRACT The report describes a thermodynamic model developed to predict phase-
equilibrium behavior in a methanol/carbon-dioxide/nitrogen/hydrogen-sulfide sys-
tem based on parameters determined from binary vapor/liquid equilibrium data
available in the literature. Model predictions are compared with actual experimen-
tal data. Predicted values show an average deviation from experimental data of 21%
for vapor compositions and 10% for total pressures, indicating the possibility of
ternary effects not accounted for by the model. (The model is to be used in conjunc-
tion with a coal-gasification/gas-cleaning facility constructed at North Carolina
State University as part of a study funded by the EPA. The overall objective of the
project is to characterize completely the gaseous and condensed phase emissions
from typical coal-gasification/gas-cleaning processes and to determine how emis-
sions depend on process parameters. To describe and evaluate operations involved
in the removal of acid gas constituents from the crude synthesis gas, information is
needed concerning the equilibrium behavior of these constituents with the particular
solvent used in the removal unit; hence, the model.)
7.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b,IDENTIFIERS/OPEN ENDED TERMS
COSATi Field/Gioup
Pollution
Carbinols
iSour Gas
Solubility
1 Mathematical Models
Thermodynamics
Coal Gasification
Gas Purification
Pollution Control
Stationary Sources
Methanol
Acid Gases
13B
07C
2 ID
07D
12A
20M
07A,13H
"3 DlSTFt UUTION STA TEMENT
Unlimited
19. SECURITY CLASS (This Report/
Unclassified
21. NO. OF PAGES
46
20. SECURITY CLASS (This page/
Unclassified
22. PRICE
EPA Form 2220-1 (9-73)
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