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Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 4, 1979
Evaluation of Fluidized-Bed Methanation Catalysts and Reactor Modeling James T. Cobb, Jr. Department of Chemical & Petroleum Engineering, University of Pittsburgh, Pittsburgh, Pennsylvania 1526 1
Robert C. Streeter Bituminous Coal Research, Inc., Monroeville, Pennsylvania 15 146
Performance data have been obtained for several methanation catalysts in both a bench-scale fixed-bed reactor and in a larger-scale fluidized-bed reactor. Three catalysts containing molybdenum were found to give selectivities (water-gas shift reaction vs. hydrocarbon formation)approaching 50 % at 80 % CO conversion. Three catalysts containing cobalt were found to give less than 20% selectivity, making them less useful in reactors where both reactions are desired. The performance of two catalysts containing molybdenum and two containing cobalt in the bench-scale unit were fitted to a simple first-order rate equation. This kinetic model was then used in a simple fluidized-bed reactor model to predict the performance of the larger-scale methanator. A reasonable match between predicted and experimental results was obtained.
Introduction Synthesis gas is the major primary product of all oxygen-blown coal gasifiers. Used as it is produced, it could conceivably be substituted as a fuel gas in many industrial processes. However, synthesis gas cannot practically be mingled with natural gas in pipelines for distribution to the homes of the general public. Therefore a number of processes are being developed for the production of substitute natural gas (SNG) for pipeline use. Most of these processes make use of supported metal catalysts (Mills and Steffgen, 19731, the principal ones including nickel as the major metallic component. The reactor systems, which are being developed, include (Lom and Williams, 1976; Pennline et al., 1977; Seglin et al., 1975) fixed-bed reactors, fluidized-bed reactors, slurry reactors, and tube-wall reactors. This paper describes the operation of and data obtained from a fluidized-bed reactor system using promoted-nickel catalysts. This vapor-phase fluidized-bed methanation process is being developed by Bituminous Coal Research, Inc. (BCR), in conjunction with the BI-GAS coal gasification process. However, the fluidized-bed methanation concept could be extended to other coal gasification processes currently under development. Test Equipment and Procedures Two systems are used to test methanation catalysts at BCR. One system is a small fixed-bed reactor for initial catalyst screening. The other system is a process and equipment development unit (PEDU) for obtaining data that will be used in planning pilot plant tests. The catalyst screening (life test) system is a small fixed-bed reactor consisting of four copper-clad stainless steel tubes, each about 'I8-in. i.d., encased in a heated aluminum block. Typically, catalyst charges of 1 cm3 are reduced under hydrogen and then exposed to a synthesis gas containing 20% CO, 60% H2, and 20% CH, at space velocities in the range of 1500 to 3000 h-' (STP conditions) and a pressure of 1000 psig. For the first part of the test, the temperature is cycled several times between 800 and 1100 O F to determine whether any thermal deactivation occurs; then the temperature is held constant at 800 O F 0019-7882/79/1118-0672$01.00/0
for the remainder of the test to observe any change in activity with time. The total duration of this life test is normally 800 h. A special test was performed to obtain the kinetic data reported later. In this test the temperature was held at 800 OF for about 7 days. It was then dropped to 650 O F , where data were obtained at space velocities of 1500 and 3000 h-l. The temperature was then increased by 25 O F increments through the range of 650 to 800 O F and by 50 O F increments through the range of 800 to 1000 OF, with data being taken at both flow rates at each temperature level. Finally, after exposure of the catalyst to 1000 O F , the temperature was decreased to 800 O F and the range of 800 to 650 O F was again covered in 50 O F decrements. The overall duration of the test was 1570 h. The PEDU reactor, shown in Figure 1,is a 6-in. i.d., 15-ft long vessel with a reaction zone of approximately 8 ft. The vessel contains approximately 33 ft2 of internal heattransfer surface in the form of two vertical finned tube bundles as well as a cooling jacket around the gas inlet zone. Dowtherm G is employed as the heat transfer fluid. An auxiliary feed gas nozzle located near the middle of the catalyst bed provides the capability of splitting the feed to two different zones of the reactor. Although the system was designed to permit recycling a portion of the product gas to the reactor for temperature control, this has been found to be unnecessary in any of the tests conducted to date. In fact, very few problems have been found in controlling catalyst bed temperatures using only one-third of the available finned-tube heat transfer surface, even with the most active catalysts tested. Such problems, when they arose, involved control of only the distributor zone temperatures near the feed gas inlet. These problems were alleviated by the addition of an independently controlled Dowtherm cooling system for the distributor zone of the reactor. With the system as originally designed, all of the Dowtherm for both the distributor jacket and the upper finned tubes passed through a common temperaturecontrol system. As more highly active catalysts were received for testing, it became progressively more difficult to control the distributor zone temperatures without overcooling the upper portion of the catalyst bed. Con0 1979 American Chemical Society
Ind. Eng. Chem. Process Product Gas-
C a t a l y s t Charping
f Di I n g og'in g Z o n e
1
5 Fo.1
1
Roaction
Zone
lntermodiatc
' F e e d Gas
4
Zone
2 Foot
Figure 1. Fluidized-bed methanator.
sequently, an external cooling coil was added to the upper part of the distributor and an additional pump, heater, and heat exchanger were installed as a separate circuit for the heat transfer fluid at, the reactor inlet. This change resulted in much more nearly isothermal operation of the reactor, and in the more recent PEDU tests it has not been uncommon to maintain the distributor Dowtherm temperature from 200 to 250 O F cooler than that of the Dowtherm circulated through the finned tubes. Catalyst samples which show promise during the life test are normally selected for further testing to determine their attrition resistance, bulk density, and fluidization behavior. For the latter purpose, a full-scale 6-in. i.d. Plexiglas model of the PEDU methanator is available, as well as smaller apparatus for determining minimum fluidizing velocities. Nearly all of the catalysts tested have exhibited minimum fluidizing velocities (U,) on the order of 0.01 ft/s. Normal PEDU operation involves superficial velocities at the inlet in the range of 0.08 to 0.18 ft/s, or 8 to 18 times U,. The usual catalyst charge to the PEDU is approximately 1 ft3 (50 to 60 lb) based on the bulk density at minimum fluidizing velocity. As in the bench-scale tests, normal operating conditions are temperatures of 800 to lo00 O F , 1000 psig pressure, and a feed gas composition of 20% CO, 60% Ha, and 20% CH,. A typical PEDU test lasts between 5 and 10 days and includes between 10 and 25 data periods. For the tests reported in this study, conversion of CO was always greater than 5070,usually greater than 70%, and frequently greater than 85%. Variables examined were temperature, pressure, space velocity and feed composition. The H2 to CO ratio in the feed varied between 2 and 3.
1979
673
A t the present time, a computer data acquisition and processing system contains 33 monitor commands and 35 utility programs which make it possible to ascertain the status of some 60 process variables at any time during operation of the system. In addition, these routines can generate interim mass and heat balances based on selected time intervals, and a t the conclusion of each data period a separate report generation package allows the computer operator to generate summary data reports, edit any erroneous data, copy data files, and produce complete mass and heat balances for that period. Details concerning the PEDU test program at BCR have been published in the Proceedings of the Fifth, Sixth, Eighth and Ninth Synthetic Pipeline Gas Symposia (Graboski and Diehl, 1973; Diehl et al., 1974; Streeter et al., 1976; Streeter, 1977). The remainder of this paper provides an analysis of the selectivity and kinetics of the catalysts used in the fluidized-bed methanator and an analysis of that methanator's performance. Selectivity The solids- and tar-free synthesis gas emerging from the gasifier of a typical coal gasification process is characterized by a H 2 / C 0 ratio between 0.8 and 2.0 (Institute of Gas Technology, 1976). Before the raw synthesis gas is fed to the methanation process, steam is added, the water-gas shift reaction is carried out to increase the H 2 / C 0 ratio, and the resulting C 0 2 is removed (Eisenlohr et al., 1975). The usual objective is to obtain a H2/C0 ratio of 3.0, which is the stoichiometric ratio for the CO methanation reaction CO + 3H2 CH4 + HzO (1)
x,.!cp~ G a s Distribution
Des. Dev., Vol. 18, No. 4,
-
Unfortunately, the methanation reaction is not the only overall reaction occurring in the fluid-bed methanator. Two others are observed. One of these is the water-gas shift reaction CO + H20 -.+ COZ + Hz (2) in which the steam produced by reaction 1 competes with the methanation reaction for CO. This competition reduces the yield of CHI and dilutes the final SNG with C02, thus lowering its heating value (Gruber, 1975). The second competing reaction is the ethanation reaction 2CO + 5H2 CzHs + 2H20 (3) +
Selectivity, S , is defined as 100N, S= N , + 2N,
(4)
where N,, N,, and Ne are, respectively, the net moles of C02, CHI, and C2H6appearing in the product. Conversion, X,is defined as 100Np (5) Nf where Nf represents the moles of CO fed to the methanator and Np is the sum of 2Ne, N,, and N,. (Definitions of selectivity and conversion follow those used by Carberry, 1977, p 24.) Figure 2 shows selectivity as a function of conversion for seven catalysts used in 12 PEDU tests. The data points represent periods with a wide range of temperatures, pressures, space velocities, and feed configurations (split feed and single feed operation). Thus, the correspondence indicated is quite remarkable. The difference in selectivity between the two groups of catalysts is striking. Group 1, consisting of catalysts no. 3165, 3234, 3547, and 3615, is characterized by high selectivity at lower conversions.
x=-
674
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 4, 1979 SELECTIVITY, PERCENT
U
60-
N o . 3165
50-
40-
30-
PEDU l e s t No.
. o
*
20-
4
h Catalyst No.
8
3234 3256 3472 3547 3582
11-16 18 19 20
N o . 3615
\ ’
10-
No. 3582 Equilibrium 4’
04 60
r
70
I
80
90
Point
100
CONVERSION, PERCENT
Figure 2. Water-gas shift selectivity of PEDU test methanation catalysts at H2/C0 = 3:l.
These four catalysts all contain molybdenum. Catalyst no. 3165 is molybdenum on alumina, while the other three are molybdenum-promoted nickel catalysts. The catalysts of group 2, consisting of catalysts no 3256, 3472, and 3582, are all cobalt-promoted nickel catalysts. This group is characterized by low selectivity at lower conversions. These results suggest that molybdenum catalyzes the water-gas shift reaction to a greater extent than cobalt and nickel. Interestingly, most of the lines in Figure 2 tend to converge at about 20% selectivity at 100% conversion. This value is appreciably higher than the predicted equilibrium selectivity of 4 % a t 875 O F and 68 atm pressure, for a feed composition of 60% H2,20% CO, and 20% CH4 (based upon values given by Mills and Steffgen, 1973), suggesting that the shift reaction is much closer to equilibrium than the methanation reaction at high conversions. Thus, the experimental data indicate that for a 3:l H 2 / C 0 feed gas containing no C02, no more than about 5 out of every 6 moles (83%) of CO fed to the methanator can be converted to hydrocarbons (predominantly CH4) under conditions approaching 100% utilization of CO, regardless of which catalyst is employed. This limitation is believed to have resulted from the particular catalysts and reaction conditions employed in this study, rather than the characteristics of the fluidized-bed methanator, since more recent tests with a more active catalyst have demonstrated that near-equilibrium conversions are possible. Kinetics The complete rate equation for the conversion of CO in a methanator is quite complex (Seglin et al., 1975). However, those working in the area of reactor development have found it convenient, as well as satisfactory, to use a simple first-order rate equation (Blum et al., 1975; Bridger and Woodward, 1975; Hausberger et al., 1975) -ra = k,C, (6) where -r, is the rate of disappearance of CO, k, is the rate constant, and C, is the concentration of CO. The integrated rate expression for a fixed-bed reactor is given by Levenspiel (1972) as kt7 = -(1 + e,) In (1 - X,)- caX, (7)
where T is the space time, X, is the fractional CO conversion, and e, is the fractional change in volume between no conversion and complete conversion of CO. The data obtained from the special bench-scale test mentioned earlier were analyzed using eq 7 . The results are shown in Figures 3 through 6 as Arrhenius plots of In k, vs. 1/T. When these data were plotted, it became evident that: (1)in some cases the rate constant appeared to be a function of flow rate (i.e., space time), and (2) data taken after the catalyst had been exposed to the highest temperature (1000 O F ) showed a marked departure from those taken earlier in the life test. Consequently, on each graph there are separate lines showing k , vs. 1/T at both the high and low feed gas flow rates as well as for the data after exposure of the catalysts to 1000 O F . The heavy dashed lines on Figures 3 through 6 represent the lines of best fit through all the data points (i.e., at both high and low flows) during the first temperature cycle up to 1000 O F . A similar line could be drawn through all of the data below 800 O F taken after exposure to 1000 OF. Before exposure to 1000 OF, the two cobalt-promoted nickel catalysts (no. 3256 and 3472) were the most active catalysts and the supported molybdenum catalyst (no. 3165) was the least active. The difference between the most active catalyst (no. 3256) and the least active one (no. 3165), however, is relatively small-essentially a difference of a factor of 2.3 at all temperatures. Exposure to lo00 OF led to considerable losses in activity below 800 O F for the supported nickel catalysts, especially no. 3256 which decreased in activity by a factor of about 7 at 650 O F . There is a good possibility that this loss was caused by the agglomeration of the nickel metal crystallites above 800 OF, which is well-known phenomenon. Only the supported molybdenum catalyst, no. 3165, showed little activity loss upon exposure to 1000 O F , indicating a lesser tendency on the part of molybdenum towards reduced dispersion under these conditions. It should be pointed out, however, that no catalyst characterization studies were undertaken to establish the cause of deactivation. In some cases, the kinetic behavior of the catalysts appeared to vary with the flow rate. This was particularly evident with the less active catalysts, no. 3234 and 3165, where higher values of k, were almost consistently obtained at the higher flow rates (Figures 5 and 6). The apparent
Ind. Eng. Chern. Process Des. Dev., Vol. 18, No. 4, 1979 675 TEMPERATURE. F 1000
950
900
050
000
750
700
650 I
c
Iw
-
f
-3 m
',"\
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7.6
R E C I P R O C A L TEMPERATURE, R-1
1
0.6
1
I
0.0
9.0
J
x 104
Figure 3. Arrhenius plot from life test data for catalyst BCR lot no. 3256. TEMPERATURE, F 1000
-3
950
050
900
000
750
700
4 F l o w w 2 5 cc/min m
Flow-40
e
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\
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R E C I P R O C A L TEMPERATURE, R-1
Figure 4. Arrhenius
650
8.6
0.0
9.0
x 104
from life test data for catalyst BCR lot no. 3472.
variation in the rate constant with T is caused by either an incorrect kinetic model, a nonuniform temperature distribution (hot spot) in the bed at low flow rates, mass transfer limitations at low flow rates, or a combination of these causes. Because the effect is relatively small, no attempt has been made to determine its cause. Composite lines plotted in (or suggested by) Figures 3 through 6 can be represented by the logarithmic form of the Arrhenius equation In k = --(Ea/R)(1/T)+ In A (8) Values for the activation energy (E,) and the frequency factor ( A ) are given in Tables I and I1 for the four catalysts before and after the 1000 O F treatment, respectively, as well as correlation coefficients which indicate the degree to which the data fit eq 8. As shown in Table I, initial activation energies on the order of 5 to 9 kcal/mol were obtained. After the samples were heated to 1000 O F , activation energies were larger, in the range of 11 to 15 kcal/mol, as shown in Table 11. In either case, these values may be on the low side; Mills
Table I. Kinetic Data for the Initial Activity of Four Fluid-Bed Methanation Catalysts Calculated from Life Test Data cat. no.
Ea, kcal/mol
3256 3234 347 2 3165
4.63 7.32 8.72 5.65
A , s"
correl. coeff.
2.45 11.43 28.18 1.94
-0.715 -0.909 -0.917 -0.817
Table 11. Kinetic Data for the Steady-State Activity of Four Fluid-Bed Methanation Catalysts Calculated from Life Test Data Ea, correl. A, s-' coeff. cat. no. kcal/mol 3256 3234 3472 3165
14.84 15.02 14.49 11.09
21.7 x l o 2 29.1 X 10' 9.9 x 102 1.0 x 102
-0.990 -0.957 -0.977 -0.944
and Steffgen (1973) cite typical values of 18 to 28 kcal/mol for the methanation reaction with various promoted nickel
676
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 4, 1979 TEMPERATURE, F 1000
950
aoo I
a50
900
I
I
750
700
I
650 I
I
e
0
9
A
-4
-5
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8.8
9.0
R E C I P R O C A L T E M P E R A T U R E , ~ - 1x 1 0 4
Figure 5. Arrhenius plot from life test data for catalyst BCR lot no. 3234. TEMPERATURE, F 9:o
,IO,OO
-5
1
9 ~ o
I
a:o
I
catalysts. On the other hand, workers at the Institute of Gas Technology (1975) derived values of E, = 6.9 kcal/mol, from a rate expression which was and A = 1.06 X reportedly applicable for all pressures and within the temperature range of 525 to 900 OF. Similarly, workers a t Brigham Young University (Bartholomew, 1975) have reported activation energies for promoted nickel catalysts of 12 to 18 kcal/mol for both carbon monoxide conversion and methane formation, although they cautioned that these lower activation energies may have resulted, in part, from mass transfer (diffusional) limitations. Thus, the values of activation energy reported in this paper are in reasonably good agreement with others that have appeared in the literature, although it is generally recognized that pore diffusion and/or film mass transfer effects are probably present at these higher temperatures. Fluidized-Bed Reactor Model The kinetic model, defined by eq 6 and 8 and the composite parameters shown in Table 11, has been used in developing a model for the fluidized-bed reactor. The basic form of the reactor model was determined from an
8po
I
770
7;o
1
1
I
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examination of the physical behavior of the bed of the PEDU methanator. This reactor has an equivalent diameter of 5l/* in. (equivalent to a vessel free of internal heat exchangers and having the same average volume per unit length for catalyst as the actual reactor). The difference between the superficial velocity of the feed and its minimum fluidization velocity is typically between 0.08 and 0.18 ft/s. According to Figure 5.2 of Hovmand and Davidson (1971), this reactor operates well within the bubbling flow regime. Therefore, a bubbling-bed type of reactor model was chosen. The particular model selected was a very simple one-the Orcutt (Orcutt et al., 1962) or Davidson-Harrison (Davidson and Harrison, 1963) Model, which has been shown (Chavarie and Grace, 1975) to provide satisfactory overall performance predictions. This model assumes that the catalyst bed is composed of two phases: a particulate phase having the porosity of incipient fluidization, and a bubble phase containing no solid particles but carrying all the gas other than the minimum required to fluidize the bed. The mathematical treatment involves a large number of variables which define the fluid-mechanical properties
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 4, 1979 ? I E O I C l E D CONVERSION
U
677
PREDlClED CONVERSION
100-
./ 7: -
/
.#
90PEDU l . $ I
No
6
*
Calaly%1N s 3165 IO-
so
/*
so
60
70
no
1
PO
100
OBSERVED CONVERSION
oo
5050
OBSERVID CONVERSION
Figure 7. Predicted vs. observed conversions for PEDU test no. 6.
Figure 9. Predicted vs. observed conversions for PEDU test no. 11.
PREDICTED CONVERSION
U
?REDIClED CONVERSION
too,
--
~
50
60
70
110
90
100
OBSERVED CONVERSION
Figure 8. Predicted vs. observed conversions for PEDU test no. 8.
of the catalyst bed. Carberry has provided a clear description of all the assumptions upon which this model is based (Carberry, 1977, p 560). A preliminary calculation using data from Period 3 of PEDU Test No. 11 was performed to determine the magnitude of the exchange coefficient between the particulate and bubble phases. The value of the rate constant, kt,was approximated from the data in Figure 3, using the results for the catalyst after exposure to 1000 O F . This preliminary calculation showed that the exchange coefficient was very large. Therefore, the gas throughout both the particulate and bubble phases was assumed to be totally backmixed. This essentially reduces the model to that for a well mixed reactor. Because of possible variations in catalyst bed temperatures, and to accommodate the fact that part of the feed can be introduced near the center of the bed, the model was refined to consider the reaction zone as a series of three separate fluid beds. The fraction of carbon monoxide remaining unreacted in each zone could then be calculated from a knowledge of the bed temperatures, superficial feed velocities, and the reaction rate constant; the rate constant was estimated for the designated temperatures by the Arrhenius equation, using the values shown in Table I1 for the activation energy (E,) and frequency factor (A)for each catalyst. Finally, the results for each of the three zones were combined to obtain an overall or “predicted” conversion, which was then compared with the conversion actually observed based on the experimental data. A short computer program was written to effect the three-stage Orcutt model calculations. The results of these calculations are summarized in Figures 7 through 10 for PEDU tests no. 6, 8, 11, and 18; the kinetic values employed in these calculations are listed in Table 111. These kinetic values were adjusted empirically from those shown in Table I1 to obtain a better fit for the data represented
O V E R V E D CONVERSION
Figure 10. Predicted vs. observed conversions for PEDU test no. 18. Table I11 PEDU test no.
cat. no.
Ea kcal/mol
6 8 11 18
3165 3234 3256 347 2
11.5 15.8 13.8 13.7
9
A , s-’ 100 2910 2170 1090
in Figures 7 through 10. However, it should be noted that the values of E, and A employed are within 10% of the experimentally determined values listed in Table I1 as composite values after exposure of the catalysts to 1000 O F . It was felt that these values would best represent the steady-state activities of these four catalysts. Two other catalysts are of particular interest to the overall development project being conducted a t the BIGAS pilot plant site in Homer City, Pa., at which fluidized-bed methanation will be tested on a larger scale. These two catalysts are catalysts no. 3547 and 3631. Both of these are molybdenum-promoted, supported-nickel catalysts. Catalyst no. 3631 is the catalyst which has been charged to the pilot plant methanator and catalyst no. 3547 was the prototype to catalyst no. 3631. Catalyst no. 3547 was tested in PEDU test no. 19 and catalyst no. 3631 was tested in PEDU test no. 22. Both catalysts were expected to have values of kinetic parameters similar to that for catalyst no. 3234 ( E , = 15.8 kcal/mol, A = 2910 s-l). However, the values which gave the best fit of the parity line by the reactor model (Figures 11 and 12) are shown in Table IV, which shows that catalysts no. 3547 and 3631 are essentially identical in their performance in the PEDU. However, their performance is somewhat different from that of catalyst no. 3234. This
Ind. Eng. Chem. Process Des. Dev., Vol. 18, No. 4, 1979
678
V R l D l C l I D CONVERSION
U
OBSERVED CONVERSION
Figure 11. Predicted vs. observed conversions for PEDU test no. 19. VREDICTED CONVERSION
?EDU l e t 1 No. 2 2 Calal~~ No t
3631
60
50
puterized simulation of the BI-GAS process for process development, design or optimization purposes. Some refinements could be made to the model to accommodate changes in mode of operation and additional knowledge gained concerning the hydrodynamic properties of the catalyst bed. The kinetic model, based upon a first-order rate equation to describe the catalyst behavior in the PEDU, also appears to be satisfactory as to its utilization in the reactor model. Further refinements could be made to yield a more mechanistically accurate expression, which would include the effects of non-first-order kinetics and deactivation on CO conversion and of selectivity on product distribution. The selectivity of molybdenum-containing catalysts for the water-gas shift reaction has been shown to be significantly higher than that of cobalt-containing catalysts. This suggests that catalysts containing molybdenum might be more suitable for processes in which the water-gas shift reaction must be combined with methanation. However, this increase in selectivity appears to be coupled with a reduction in overall activity. Further investigation is needed to determine whether this reduction is real or merely the result of an imbalance between particle size and pore structure of the molybdenum-containing catalyst, which has resulted in mass-transfer limitations in those catalysts tested so far.
Acknowledgment OBSERVED CONVERSIOH
Figure 12. Predicted vs. observed conversions for PEDU test no. 22. Table IV
PEDU
E,
3
t e s t no.
c a t . no.
kcal/mol
19 22
3547 3631
13.8 13.9
A , s-'
2000 2000
is probably due to differences in metal loadings, which were substantially higher for catalysts no. 3547 and 3631 than for catalyst no. 3234. The diagonal parity lines in Figures 7 through 12 indicate the reliability of the model; for a perfect correlation, all of the data points would fall on this line. Although there is admittedly some scattering of the points about the parity line for most of the PEDU tests, it is felt that the correlation is sufficiently good to support the validity of the model. It should be recalled that these data were taken over a wide range of operating conditions. For example, in several cases the principal outlying points have been found to be from split-feed periods, suggesting that the model could be improved for data obtained under these circumstances. In other cases, outlying points represent data periods either very early in the PEDU test, presumably before the catalyst had reached its steady-state activity, or near the end of the test after the catalyst had suffered some deactivation. In either of these instances, the kinetic behavior of the catalyst would have been different from that assumed in the modeling calculations.
Conclusions The simple reactor model developed for the fluidized-bed methanator PEDU gives reasonably good parity plots for six PEDU tests using six different catalysts. This model can be used to predict the performance of the pilot plant methanator and as a building block in any com-
The coauthor from the University of Pittsburgh thanks Bituminous Coal Research, Inc., for the opportunity to participate in this development work as a consultant during the summer of 1975.
Literature Cited Bartholomew, C. H., "Alloy Catalysts with Monolith Supports for Methanation of Coal-Derived Gases", Quarterly Technical Progress Report for Period July 23-0ct 22, 1975; prepared for ERDA, Nov 6, 1975. Bium, D. B., Sherwin, M. B., Frank, M. E., Adv. Chem. Ser., No. 148, 149-159 (1975). Bridger, G. W., Woodward, C., Adv. Chem. Ser., No. 148, 71-86 (1975). Carberry, J. J., "Chemical and Catalytic Reaction Engineering", McGraw-Hill, New York, N.Y., 1977. Chavarie, C.. Grace, J. R., Ind:, Eng. Chem. Fundam., 14, 75, 79, 86 (1975). Davidson, J. F., Harrison, D., Fluidized Particles", pp 100-103, Cambridge University Press, Cambridge, England, 1963. M I , E. K., Stewart, D. L.. Streeter, R. C., "Progessin Fluidied-BedMethanatbn", Sixth Synthetic Pipeline Gas Symposium, Chicago, IL, 1974. Eisenlohr, K.-H., Moeller, F. W.. Dry, M., Adv. Chem. Ser., No. 146, 113-122 (1975). Graboski, M. S.,Diehl, E. K., "Design and Operation of the BCR Fluidized Bed Methanation PEDU", Fifth Synthetic Pipeline Gas Symposium, Chicago, IL, 1973. Gruber, G., Adv. Chem. Ser., No. 148, 31-46 (1975). Hausberger, A. L.. Knight, C. B., Atwood, K., Adv. Chem. Ser., No. 148, 47-70 (1975). Hovmard, S.,Davidson, J. F., "Skrg Flow Reactors", h "FlUzatbn", J. F. Davidson and D. Harrison, Ed., p 195, Academic Press, New York, N.Y., 1971. Institute of Gas Technology, "Preparation of a Coai Conversion Systems Technical Data Book. Final Report, October 31, 1974 - April 30, 1976", pp 269-352, ERDA Report No. FE-1730-21 (UC-90). Contract No. E(49-18)-1730, 1976. Institute of Gas Technology, "HYGAS: 1964 to 1972. Volume 3. Part V: Methanation", p 66, R&D Report No. 22, prepared for ERDA, 1975. Kunii, D., Levenspiel, O., Fluidization Engineering", p 112, Wlley, New York, N.Y.. 1969. Levenspiel, O., "Chemical Reaction Engineering", 2nd ed, p 110, Why, New York, N.Y., 1972. Lom. W. L.. Williams.. A. F... "Substitute Natural Gas", .DD . 184-167. Wllev. New York, N:Y., 1976. Mills, G. A,, Steffgen. F. W., Catal. Rev., 8, 159-210 (1973). Orcutt, J. C., Davidson, J. F., Pigford, R. L., Chem. Eng. Progr. Symp. Ser., 58, 1 (1962). Penniine, H. W., Schehl, R. R., Haynes, W. P., "Operation of a Tube Wall Methanation Reactor", 2nd Joint Conference CIC/ACS, Montreal, Canada, May-June 1977. Seglin. L.,Geosits, R.. Franko, 6. R., Gruber, G., Adv. Chem. Ser., No. 146, ~~
1-30 (1975). - -, \
Street&, R. C., "Recent Developments in FluidizecCBed Methanation Research", Ninth Synthetic Pipeline Gas Symposium, Chicago, IL, 1977.
Ind. Eng. Streeter, R. C., Anderson, D. A., Cobb, J. T., "Status of the 61-GAS Program. Part 11. Evaluation of Fluidized-Bed Methanation Catalysts", Eighth Synthetic Pipeline Gas Symposium, Chicago, IL, 1976.
Received f o r review September 25, 1978 Accepted April 20, 1979
Chem. Process Des. Dev., Vol. 18, No. 4, 1979 679
The work described in this paper was conducted under the sponsors~ipof the u.s, D~~~~~~~~~of E~~~~~and the G~~ Research Institute under DOE Contract No. EF-77-C-01-1207. The paper is based in part on presentations given a t the Eighth Synthetic Pipeline Gas Symposium (Streeter et al., 1976) and a t the Second Joint CIC/ACS Conference, Montreal, Canada, 1977.
Residual Enthalpies of High-Boiling Hydrocarbons Gunther Muller and John M. Prausnitz' Chemical Engineering Department, University of California, Berkeley, California 94720
To facilitate design calculations for processes using heavy hydrocarbons, reduced charts are given for estimating enthalpies of high-molecular-weight fluids relative to their ideal-gas enthalpies at the same temperature. The reduced charts follow from an equation of state based on Prigogine's theory of fluids containing large, polysegmented molecules. Instead of critical data, the charts use three welldefined molecular parameters: hard-core volume, potential energy, and number of external degrees of freedom. Estimates for these parameters are provided for heavy alkanes, simple aromatics, and polynuclear aromatics. Significant comparison with experiment is unfortunately not possible because reliable experimental enthalpy data for heavy hydrocarbons are not available.
Recent developments in energy technology have directed attention to heavy hydrocarbons. Prior to the energy crisis of 1973, heavy hydrocarbons ("bottom of the barrel") were used primarily for road construction or other low-grade applications. As a result of changing economic conditions, new processes are under development for utilizing heavy hydrocarbons as raw materials for higher-grade applications, including lubricants, fuels, and feed stocks for manufacture of chemicals. For quantitative engineering design of such processes, it is often necessary to know the physical properties of the heavy hydrocarbons. Since experimental data for these materials are extremely scarce, it is necessary to establish estimation techniques. This work reports a procedure for estimating enthalpies of heavy hydrocarbons [in particular the effect of pressure on enthalpies] including the enthalpy of vaporization. Basic Thermodynamic Relations The enthalpy of a fluid in the ideal-gas state at temperature T i s directly related to the ideal-gas heat capacity cpo according to the well-known relation
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