340
Ind. Eng. Chern. Process Des. Dev., Vol. 17, No. 3, 1978
Process Variable Effects in the Conversion of Methanol to Gasoline in a Fluid Bed Reactor David Llederman, Solomon M. Jacob, Sterling E. Voltz,' and John J. Wise Mobil Research and Development Corporation, Paulsboro, New Jersey 08066
The conversion of methanol to high octane gasoline was studied in a fluid bed reactor. Some effects of temperature, pressure, and space velocity on methanol conversion and product yields were determined. Catalyst deactivation due to coke formation and steaming were investigated during a kinetic aging test. The activity loss from coking can be regained by oxidative regeneration. Coupling the methanol conversion process with the commercially proven technology for the production of methanol from coal provides an alternate route for the conversion of coal to high octane gasoline.
Introduction Coal and other nonpetroleum materials are expected to become important sources of fuels during the next several decades. In particular, considerable effort is being expended on the development of improved processes for the conversion of coal to gasoline. The Bergius and Fischer-Tropsch processes have been used commercially to produce gasoline from coal. However, no Bergius plants exist today, and the Fischer-Tropsch process is currently being used only in South Africa. The hydrocarbon products obtained from these processes range from methane to high molecular weight compounds. In addition, large amounts of oxygenates are produced. Consequently, the quality of the gasoline product is poor, and extensive processing is required to produce a stable, high octane motor fuel. Coal liquefaction processes such as H-Coal, Synthoil, and Solvent Refined Coal are being developed. Large pilot plants have been in operation for years, and demonstration plants are under construction. Most of the products from these processes are quite deficient in hydrogen, and considerable upgrading will be required to produce a quality gasoline. Studies have shown that methanol can be used directly as a fuel in automotive engines or can be blended with gasoline (Reed and Lerner, 1973; Mills and Harney, 1974; Wigg, 1974; Burke, 1975; Koehl, 1975). Significant problems related to performance and reliability would have to be solved for methanol-gasoline blends to be used in existing vehicles. Modified engines need to be developed for the efficient use of methanol alone. A novel process for the conversion of methanol to gasoline over a zeolite catalyst has been developed (Meisel et al., 1976, 1977; Wise and Silvestri, 1976; Daviduk et al., 1976; Chang et al., 1978).Since the conversion of coal t o methanol is well established commercial technology, coupling of the methanolto-gasoline process with this technology provides a new route for the conversion of coal to gasoline. The gasoline from this process has exceptional product qualities (i.e., high octane). Gasoline stability (i.e., potential gum formation) is acceptable a t reasonable additive levels. Durene concentrations are higher than in petroleum-derived fuels, and because of its high melting point (174 O F ) , durene deposits in automotive fuel systems could be a problem. However, car tests with gasolines containing added durene have shown that concentrations as high as 4 wt % durene have no significant effects on vehicle performance. Methanol-derived gasoline will have lower durene concentrations. Both fixed and fluid bed process concepts have been investigated (Wise and Silvestri, 1976; Chang et al., 1978). Two reactors are used in the fixed bed process. Methanol is par0019-7882/78/1117-0340$01.00/0
tially dehydrated to an equilibrium mixture of methanol, dimethyl ether, and water over a dehydration catalyst in the first reactor. Both methanol and dimethyl ether are converted to high octane gasoline by a zeolite conversion catalyst in the second reactor. The use of two reactors, together with recycle of light gases to the conversion reactor, minimizes the heat removal problem associated with this highly exothermic process. In contrast, a single reactor is used in the fluid bed process. This process is highly selective, and can be represented by the following overall reaction (Chang et al., 1978) nCH30H
-
hydrocarbons
+ nHzO
The hydrocarbons consist primarily of high octane Lasoline (greater than 75 wt %) with smaller amounts of T2G (C, and C4) and fuel gas (C, and Cz). The ultimate gasoline yield can be increased significantly by alkylating C3 and C4 olefins with the isobutane produced in the process. The gasoline is composed predominantly of methyl-substituted aromatics and isoparaffins, and it has an unleaded research octane of over 90. Mechanistic details of the conversion of methanol and other oxygenated compounds to high octane gasoline were recently presented by Chang and Silvestri (1977). The formation of hydrocarbons from methanol was reported by Mattox in 1962; light olefins were obtained in the dehydration of methanol with NaX zeolite. Other workers (Heiba and Landis, 1964;Venuto and Landis, 1968;Topchieva et al,, 1972; Swabb and Gates, 1972) have reported similar results. In 1974, Pearson obtained aromatic hydrocarbons from the dehydration of methanol over P205 a t high temperature. His yield of hydrocarbons was higher than those of previous workers. This paper describes the results of a process variable study of the methanol-to-gasoline conversion in a fluid bed reactor. Some effects of pressure, temperature, and space velocity on product yields were determined, and results from a kinetic aging test in the fluid bed pilot plant are briefly described.
Experimental Section A schematic diagram of the fluid bed unit is shown in Figure 1.The reactor was fabricated from 321 stainless steel (15/8 in. i.d.) and was heated by a five-zone electrical resistance furnace. During operation, charge stock and nitrogen carrier gas were pumped through a preheater coil where the charge was vaporized. The vapor then passed through a doughnut-shaped distributor, with eight holes (0.04 in. diameter) positioned around it, into the bottom catalyst zone, Conversion occurred in the dense fluid bed, which contained four umbrella-shaped
0 1978 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978 341 CATALYST
METHANOL
1W
0
2W
Methanol Processed itb MeOH Lb Cat I
Figure 3. C4+ and Cs+ hydrocarbons (700 O F , 0 psig, 1 WHSV).
t LIQUID PRODUCT
E
l
DISTRIBUTOR’
0
CATALYST DUMP
1w
2w
Methanol Processed ILb MeOHltb Cat. I
A l l Dimensions a r e Inches Figure 1. Schematic of fluidized dense bed pilot plant.
Figure 4. Light paraffins (700 O F , 0 psig, 1WHSV).
c
+
Methanol Processed itb MeOHltb Cat. I
Figure 5. Light olefins and durene (700 O F , 0 psig, 1WHSV). 0
IO0
ZW
Methanol Processed itb hlerlH Lb Cat I
Figure 2. Hydrocarbons and aromatics (700 OF, 0 psig, 1WHSV).
baffles to reduce bypassing. The nitrogen diluent was used to assure that the bed was continuously fluidized; otherwise the distributor might become plugged. Gaseous products were carried out of the reactor through a porous metal filter to a back-pressure regulator (where the pressure was dropped to atmospheric) and then through a condenser (at about 60 OF). The mixed aqueous-hydrocarbon product was collected in a receiver; the gases were analyzed by gas chromatography. After the liquid product was separated and weighed, portions of the hydrocarbon and aqueous phases were analyzed by gas chromatography. Samples of the hydrocarbon layer from selected runs were submitted for octane number, conjugated diolefins, bromine number, specific gravity, and simulated distillation.
All catalysts were regenerated in situ. After an initial nitrogen purge at 650 to 700 O F , the coke was burned off in a gas consisting of 0.5 vol % oxygen in nitrogen. Temperature was gradually raised to 900 or 1000 O F with constant monitoring of temperature and effluent gas composition. The oxygen content was gradually increased until it reached 20 vol %. When the coke was burned off, the catalyst was calcined a t 1100 O F for 4 h. Gas chromatography was used to analyze charge and product streams. The chromatographs were highly automated, and the data were fed directly to a computer system for reduction and storage. The analytical data and material balances were stored in an APL computer system and were readily accessible through time-sharing terminals. Details of the analytical equipment and procedures have been described elsewhere (Stockinger, 1977). Pure methanol was purchased from several commercial sources. In addition, samples of crude methanol were obtained from two commercial methanol plants. Pure methanol was used as the charge stock for the process variable study, and
342
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
Table I. Typical Material Balances in Fluid Bed Pilot Planta
Table 11. Composition of Hydrocarbon Products from Fluid Bed Pilot Planta
Material balance no. 26-01 25-04 25-07 MeOH charged (lb of MeOH/lb 6 71 135 of cat.) Product (wt % of charge) 0.32 0.84 1.13 Methanol 0.02 0.07 0.15 Dimethyl ether 42.40 42.20 42.80 Hydrocarbons Water 56.77 55.81 55.54 0.15 0.28 0.21 99.7 99.2 99.8 Hydrocarbons (wt % of hydrocarbons) Methane 0.93 0.90 0.81 Ethane 0.25 0.17 0.17 Ethylene 4.84 5.64 6.26 Propane 3.65 3.08 2.83 Propene 5.41 5.98 7.21 Isobutane 15.71 13.14 12.30 n-Butane 1.48 1.25 1.47 Butenes 4.64 5.22 5.64 c5+ 63.11 64.62 63.30 a 700 O F , 0 psig, 1 WHSV.
Material balance no. 26-01 25-04 25-07 25-10 MeOH charged (lb of MeOH/lb 6 71 135 203 of cat.) Isopentane 12.36 11.45 11.06 10.07 n-Pentane 0.49 0.43 0.42 0.33 Pentenes 2.19 2.60 3.04 3.12 Cyclopentane 0.44 0.05 0.10 0.07 0.71 0.67 0.63 0.58 Methylcyclopentane n-Hexane 4.06 5.02 5.14 5.42 5.46 6.04 6.03 5.40 Methylpentanes 0.99 1.17 1.18 1.00 Dimethyl butanes 0.26 0.39 0.43 0.47 Hexenes 0.03 0.23 0.09 0.25 Cyclohexane CT-PON 3.31 3.87 4.00 4.29 2.89 3.62 3.46 3.91 Cs-PON 1.64 2.43 2.15 2.20 CS-PON 0.11 0.22 0.43 0.27 Clo-PON 0.00 0.00 0.00 0.00 Benzene 1.38 2.12 2.38 4.03 Toluene 0.16 0.18 0.17 0.17 Ethylbenzene Xylenes 6.05 6.30 5.80 5.50 8.39 8.30 7.30 7.20 Trimethylbenzenes 1.34 1.44 1.30 1.31 Methylethylbenzenes 0.02 0.03 0.02 0.03 Propylbenzenes 1.94 3.38 2.87 3.93 1,2,4,5-Tetramethylbenzene 1,2,3,5,-Tetramethylbenzene 1.52 1.80 1.74 1.02 1,2,3,4,-Tetramethylbenzene 0.56 0.50 0.41 0.45 1.36 1.23 0.97 1.07 Other Clo-benzenes 1.72 0.48 0.43 0.33 C11-Benzenes 1.14 0.02 0.00 0.01 Naphthalenes 2.58 0.67 1.71 1.86 Other aromatics a 700 O F , 0 psig, 1 WHSV; concentrations are wt % of total hydrocarbons.
25-10 203 1.82 0.43 43.20 55.08 0.23 100.8 0.77 0.14 6.54 2.37 7.97 10.78 0.92 6.24 64.27
a sample of commerical crude methanol was used in the kinetic aging test. Developmental zeolite catalysts PC 291-6-3and PC 331-6-1 were used in the process variable study and the kinetic aging test, respectively.
Process Variable Study A study was conducted to determine the effects of key process variables on methanol conversion and product selectivities. The ranges of process variables are as follows: Temperature, 600-900 O F ; pressure, 0-50 psig; WHSV, 0.50-3.5 Ib of MeOH/lb of cat.-h. A wide range of process variables was investigated. In fact, 25 different combinations of temperature, pressure, and space velocity were used to provide sufficient information for potential operating conditions to be defined. Complete material balances and detailed product compositions were determined for each combination of process conditions. Some typical data obtained a t 700 OF, 0 psig (1 atm), and 1 WHSV are shown in Figures 2-5. The yields of hydrocarbons and aromatics were about 42 and 11 wt % of charge, respectively. The concentrations of unreacted methanol and dimethyl ether in the product increased with increasing amount of methanol processed. The yields of C4+ and C,+ hydrocarbons are shown in Figure 3. The C4+ hydrocarbons decreased from 85 t o 80 wt % of hydrocarbons, whereas the C,+ hydrocarbons were about 62 wt % of hydrocarbons. As shown in Figure 4, isobutane was the major light paraffin product; its yield decreased from 16 to 10 wt % of hydrocarbons. Propane was initially about 4 wt % of hydrocarbons and decreased slightly (to 3 wt %). Small quantities of methane, ethane, and n-butane were formed. The yields of light olefins and durene are plotted in Figure 5. Ethene, propene, and butenes increased significantly with increasing methanol processed. Also, durene increased from about 2 to 3 wt % of hydrocarbons. Typical material balances are listed in Table I for these process conditions (700 O F , 0 psig, 1 WHSV). Some of these data are plotted in Figures 2-5. However, it is interesting to observe the total spectrum of products. Both unreacted methanol and dimethyl ether increased with increasing methanol processed. Of course, the primary products were hydrocarbons and water. Only trace quantities of carbon monoxide, carbon dioxide, and hydrogen were formed. The composition of the hydrocarbon products is also given in Table
I; these results were discussed in conjunction with Figures 2-5. The corresponding compositions of the Cs+ hydrocarbon products are given in Table 11. The primary paraffins were isopentane, n-hexane, and methylpentanes, but small amounts of pentenes and hexenes (2-3 wt % of total hydrocarbons) were present. Some C7-Cg PON (paraffins, olefins, and naphthenes) were formed, but no significant quantities of higher PON (Clo+) were found. Methylbenzenes constituted the major aromatic compounds. In most instances, not even a trace of benzene was detected. Under these process conditions, alkylation of benzene would have been expected to occur quite readily. Only small quantities of ethyl- and propylbenzenes were present. The higher alkylbenzenes (Le., propyl and butyl) would be expected to crack catalytically a t 700 OF. Some product yields from another set of conditions (700 OF, 25 psig, 2 WHSV) are plotted in Figures 6-9.The yields of hydrocarbons and aromatics in Figure 6 are similar to those in Figure 2.Also, the C4+ and Cg+ hydrocarbons in Figure 7 are about the same as the yields in Figure 3. Thus, increasing the pressure from 0 to 25 psig and doubling the space velocity had no significant effect on these products. The initial yield of isobutane (Figure 8) was lower under these new conditions, but it declined only slightly during the run. The yields of light olefins were significantly different (compare Figures 5 and 9); more ethene was produced a t the higher pressure and space velocity. The higher pressure is also probably responsible for the greater durene (about 5 wt %). The data shown in Figures 10-13 were obtained a t 750 O F , 25 psig, and 2 WHSV. A comparison of these data with Figures 6-9 provides a direct measure of the effect of increasing temperature on product selectivities. The yields of hydrocarbons and aromatics were practically unchanged (Figures 6 and 10).More unreacted methanol and dimethyl ether was
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
13 tiydrocarbors V '
L
30
r--7l
I
!l re
343
20
5
'.'ethanol Processed I i b ?leOY Lb Cat
,
0
1W
200
Figure 9. Light olefins and durene (700 OF, 25 psig, 2 WSHV).
,Methanol P r a e s i e d I t b MeOHl Lb Cat 1
Figure 6. Hydrocarbons and aromatics (700 OF, 25 psig, 2 WHSV).
- v
i u d roc a r bo n i
q 5
70
0 60 100
Figure 7. C*+ and C5+ hydrocarbons (700 O F , 25 psig, 2 WHSV).
S
t
a
A-A
1W
200
\'lethano1 PrOcessed 113 ueOk Lb Cat
:lethano1 Processed I L b WeOH Lb Cat 1
12
l 0
200
, Figure 10. Hydrocarbons and aromatics (750
OF,
25 psig, 2
WHSV).
I
I
I
Prooane t
n-Butane
>
0
100
x
x
200
Methanol Processed I L 1 W I H Lb Cat 1
Figure E. Light paraffins ( 7 0 0 O F , 25 psig, 2 WHSV).
observed in the product a t the lower temperature. At 700 "F the Cq+ hydrocarbons decreased slightly (83 to 77 wt %) during the run, whereas they remained a t 84 wt % a t 750 O F ; the Cg+ hydrocarbons decreased (67 to 64 w t %) a t 700 O F and increased slightly a t 750 O F (64 to 66 wt %). As would be expected, more isobutane was formed a t the higher temperature, but the yields of the light paraffins were similar (Figures 8 and 12). As shown in Figures 9 and 13, propene was unaffected by temperature, whereas ethene decreased with increased temperature. Durene was lower a t the higher temperature. The effect of temperature on methanol conversion is shown in Figure 14. Methanol concentration (in reactor effluent) is plotted as a function of the reciprocal of corrected space velocity (l/Sc) for temperatures of 650-800 "F. The conversion of methanol increases with both increasing contact time (reciprocal of space velocity) and increasing temperature.
Comparison of the data a t different temperatures indicates that the reaction rate increases approximately by a factor of 1.5 for a 50 "F increase in temperature. This corresponds to an activation energy of about 12 kcal/mol. The space velocity (as shown in Figure 14) has been corrected to account for both catalyst deactivation and inert gas dilution by the following equation.
rS
sc = B
where S C = corrected space velocity, S = experimental space velocity, r = dilution factor, which is (mol of reactant mol of diluent)/mol of reactant, and 7 = relative catalyst activity expressed as a function of initial activity, 70. An exponential deactivation function was used t o relate 7 and 70.
+
344
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
5
f
i
1W
10
I
-
5 -
Methane + Ethane
20
25 prig
200
MWhanol Precessed ILb MeOH'Lb Cat I
0-
I
I
A
I
Figure 12. Light paraffins (750 O F , 25 psig, 2 WHSV).
1!Sc
Figure 15. Effect of pressure on methanol conversion at 700 O F . Propylene
1 75
o Butenes
e Observed
50 Ethylene
-
-
Predicted From Equilibrium
1.25
f
la,
20
075
200
Methanol Prrrersed ILb MeOHlLb Cat. 1
0 50
Figure 13. Light olefins and durene (750 O F , 25 psig, 2 WHSV). 0.25
Merhanol I W t % I
'\ 650'F
0 0
01
02
0 1
04
Figure 16. Comparison of calculated equilibrium and experimental concentrations of methanol and dimethyl ether.
05
0 6
07
0 8
sc
Figure 14. Effect of temperature on methanol conversion at 25
psig.
In the methanol-to-gasoline process, coke formation causes some catalyst deactivation. This deactivation is reversible, and most of the activity can be restored by regeneration. In addition to the loss in catalyst activity due to coke formation, there is some irreversible deactivation related to steaming. In the conversion of methanol to gasoline, the catalyst is constantly exposed to water vapor which causes catalyst deactivation. A number of different runs with fresh (or regenerated) batches of catalyst PC 291-6-3 were made in the process variable study. Various sets of experimental conditions were investigated in each run. Periodically, the catalyst activity was determined a t a particular set of conditions (700 O F , 0 psig, 1 WHSV) to provide a measure of catalyst deactivation.
Figure 15 shows the effect of pressure on methanol conversion at 700 "F. Methanol concentration is plotted as a function of the reciprocal of corrected space velocity a t pressures of 0-50 psig. Again, the conversion increases with increasing pressure. At the same space velocity, the vapor residence time at 50 psig is about 4.4 times as great as that at 0 psig. With first-order kinetics, the reaction rate would be expected to increase by a factor of 4.4in going from 0 to 50 psig. However, the observed increase is only about 1.4, which suggests inhibition of catalytic activity by adsorption of reactants and/or products on the catalyst surface a t higher pressures. It was observed in the experimental study that the concentrations of methanol, dimethyl ether, and water at the reactor exit corresponded to thermodynamic equilibrium values. This phenomenon is illustrated in Figure 16, which is a plot of methanol and dimethyl ether concentrations. The points are the experimental data and the line represents the calculated equilibrium concentrations. The agreement between the experimental and calculated results is very good. As shown in Figure 14 and as described in the preceding discussion, the increase of methanol conversion with increasing pressure is considerably less than would be expected by first-order kinetics. A simple expression of the type (111 + KAPA) has been used to account for inhibition effects caused by the adsorption of reactants and/or products. K A is an adsorption constant and P Ais the sum of the partial pres-
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
345
Table 111. Simulated Distillations Temperature, O F : Pressure, psig: WHSV, on catalyst: IBP
2
5% 10%
81
2 81
133 139 162 196 249 290 333 378 387 490
109 140 150 213 281 287 330 338 386 424
20% 30% 40% 50% 60% 70% 80% 90% 95%
$10,
Figure 17. Plot of methanol concentration versus q5/q50.
700 750 800 800 25 25 25 25 2.1 2.1 2.1 3.2 Boiling range, O F 10
1
100 138 197 277 286 295 334 340 372 396 445
79 94 143 202 281 286 319 335 340 387 437
800 50 3.1 7 78 83 139 194 247 283 293 333 338 386 396
25
sures of methanol, dimethyl ether, and water. Another term, 4/40, has been chosen to account for both inhibition and the change in gas residence time (due to different pressures). I t is defined by the equation
20
E
15
5-
PA^ is the sum of the partial pressures of methanol, dimethyl ether, and water a t a set of arbitrary reference conditions (1.0 atm, 750 O F ) . po is the gas density at the reference conditions. P A and p are, respectively, the sum of partial pressures and gas density at another set of pressure and temperature conditions. Using the ideal gas law, p = P(MW)/RTR,and the reference conditions (1 atm, 750 O F ) , eq 2 becomes (1
+ K A ) P 750 + 460
""= (1+ KAPA) Tf + 460
1
(3)
Tf is the temperature in OF, and S c is the corrected space velocity, which was defined previously. When P = 1.0 atm and Tf = 750 O F (reference conditions), 4/40 = 1/Sc. The numerical value of K A was calculated to be 1.5. The utility of eq 3 is demonstrated in Figure 17 which is a plot of methanol vs. 4/40. The data a t each temperature include several different pressures. As mentioned previously, a number of different runs with fresh (or regenerated) batches of catalyst PC 291-6-3 were made in the process variable study. Approximately 200 lb of methanol/lb of catalyst was processed in each run. The average amount of coke on the catalysts at the end of the runs were 24 wt %. The surface areas of the catalysts decreased about 50% primarily due to coke formation, since most of the surface area was recovered after regeneration. A number of samples of raw gasoline product from the fluid bed pilot plant were evaluated for octane numbers, bromine number, conjugated diolefins, and simulated distillation (gas chromatography). The samples were obtained from a wide variety of process operating conditions. The research octane (R 0) ranged from 92 to 99 and the leaded octane (R 3) was about 99-102. Based on a few samples, the unleaded and leaded motor octanes (M 0 and M+ 3) were about 85 and 90, respectively. The bromine numbers of the raw gasoline samples ranged from 11 to 91, although most of them were between 25 and 40. The concentrations of conjugated diolefins were low (approximately 0.03 to 0.1 mmol of conjugated dienes/g of sample). Some simulated distillation data (by gas chromatograph) products from a variety of process operating conditions are given in Table 111. These boiling ranges are reasonable for gasoline product (Le., based on 90% points).
+
+
+
E
10
I
5
0 Methanal Prccerred (Lb of MeOHitb a1 Cat. I
Figure 18. Methanol conversion during kinetic aging test. Kinetic Aging Test An aging test, consisting of seven complete cycles (54 days on stream), was made in the fluid bed pilot plant under conditions which gave incomplete methanol conversion. The purpose was to quantify catalyst deactivation during individual cycles and from cycle to cycle. The methanol in the reactor effluent provided a convenient measure of changes in catalyst activity. A sample of crude methanol (83 wt % methanol, 17 wt % water) from a commercial methanol plant was used as the charge stock. The operating conditions were approximately 750 O F , 25 psig, and 2 WHSV (lb of methanol/lb of catalyst-h). Methanol conversion was very high during the initial portion of each cycle, but decreased during the cycle. The data from the seven cycles are shown in Figure 18.During the first cycle, the methanol in the effluent increased from about 1to 18 wt %; methanol conversion decreased from 99 to 82 wt %. The other cycles are similar, except that the sixth one was continued for a longer time. The activity loss in a given cycle is primarily related to the formation of coke. Most of the activity is restored by oxidative regeneration. However, some activity is lost irreversibly from cycle to cycle due to steaming and other factors. The methanol conversion a t the start of each cycle is a measure of this change in activity. Figure 19 is a plot of initial cycle activity vs. the amount of methanol processed. The activity of the second cycle was higher than that of the first cycle. This behavior is similar to the longer cycle lengths of second cycles in other aging tests (with no methanol breakthrough). The catalyst activity decreased in subsequent cycles. The activity a t the seventh cycle of the aging test was about 50% of the starting value. The activity loss during the sixth cycle is also shown in Figure 18.
346
Ind. Eng. Chem. Process Des. Dev., Vol. 17, No. 3, 1978
Certain data described in this paper were used to establish operating conditions for process aging tests (i.e., 2 months duration) with complete methanol conversion in the fluid bed unit. The fluid bed technology for the conversion of methanol to high octane gasoline is being advanced in another study which involves the design, construction, and operation of a 4 BPD pilot plant.
1. 0
0 8
,r
2
0.6
2 +
-
z
Literature Cited
04
0.2
n 0
1000
2m
Methanol Praerred itb of MeOHiLb Or Cat. I
Figure 19. Initial catalyst activity in each cycle of kinetic aging
test.
Concluding Remarks The results of this process variable study have demonstrated the feasibility of converting methanol to high octane gasoline in a fluid bed reactor. In addition, some effects of temperature, pressure, and space velocity on methanol conversion and product selectivities were defined. The concentrations of methanol, dimethyl ether, and water in the reactor effluent were in thermodynamic equilibrium. A kinetic aging test with a total of 54 days on stream was made. The unreacted methanol in the reactor effluent provided a means of measuring catalyst deactivation in each cycle and from cycle to cycle. Most of the activity loss related to coke formation was restored by regeneration. However, steaming of the catalyst caused some irreversible deactivation.
Burke, D. P.. Chem. Week, 33 (Sept 24, 1975). Chang, C. D., KUO,J. C. W., Lang, W. H.. Jacob, S. M., Wise, J. J., Siivestri, A. J., lnd. Eng. Chem. Process Des. Dev., 17, 255 (1978). Chang. C. D., Silvestri, A. J., J. Catal., 47, 249 (1977). Daviduk, N., Masiuk, J., Wise, J. J., paper presented at Eleventh intersociety Energy Conversion Engineering Conference, State Line, Nev., Sept 1976. Heiba, E. I., Landis, P. S., J. Catal., 3, 471 (1964). Koehl. W. J., unpublished literature survey, 1975. Mattox, W. J., US. Patent 3 036 134 (1962). Meisel, S. L., McCullough, J. P., Lechthaler, C. H., Weisz, P. B., CHEMTECH, 6, 86 (1976). Meisel, S. L., McCullough, J. P., Lechthaler, C. H., Weisz, P. B., paper presented at the 174th National Meeting of American Chemical Society, Chicago, Ill., Aug 1977. Mills, G. A., Harney, B. A., CHEMTECH, 4, 26 (1974). Pearson, D. E.,J. Chem. Soc., Chem. Commun., 397 (1974). Reed, T. B., Lerner. R. M., Science, 182, 1299 (1973). Stockinger, J. H., J. Chromatogr. Sci., 15, 198 (1977). Swabb, E. A., Gates, B. C., lnd. Eng. Chem. Fundam., 11, 540 (1972). Topchieva, D. V., Kubasov, A. A,, Dao, T. V.. Khimiya, 27, 628 (1972). Venuto, P. B., Landis, P. S., Adv. Catal., 18, 303 (1968). Wigg, E. E., Science, 186, 785 (1974). Wise, J. J.. Silvestri. A. J., paper presented at Third Annual international Conference on Coal Gasification and Liquefaction, Pittsburgh, Pa., Aug 1976; Oil Gas J., 141 (Nov 22, 1976).
Receiued for reuieu, August 12, 1977 Accepted March 20,1978
This work was conducted under the Department of Energy (DOE) Contract No. E(49-18)-1773,which was jointly funded by DOE and Mobil Research and Development Corporation.
Direct Solution of the Isothermal Gibbs-Duhem Equation for Multicomponent Systems Jos6 A. Martinez-Ortir and David B. Manley" Department of Chemical Engineering, University of Missouri-Rolla,
Rolla, Missouri 6540 1
The numerical solution of the relative volatility form of the isothermal Gibbs-Duhem equation for multicomponent systems is carried out by a relaxation technique for two ternary systems at five different temperatures. This direct method of solution eliminates calculating and differentiating the excess Gibbs free energy as required by other indirect methods.
Scope The viability of theoretically determining vapor compositions from experimental vapor pressures over liquid mixtures under isothermal conditions has been well established by numerous investigators, and the various techniques for accomplishing the calculations are given by Prausnitz et al. (1967), Van Ness (1970),and Mixon et al. (1965). The incentive for using this method is great since the cost associated with making vapor composition measurements is eliminated, and as suggested by Manley (1973) and Van Ness et al. (1973), the results may be more accurate. This paper presents an alternative technique to those already established and provides
a set of differential equations linking the experimental P-x data and the relative volatilities. These equations are shown to be soluble by an iterative numerical algorithm similar to that used by Mixon et al. (1965). The resulting technique is similar to those already established and has some of the advantages and disadvantages of each.
Introduction The classical technique for solution of the isothermal Gibbs-Duhem equation as demonstrated by Prausnitz et al. (1967) yields correlation constants for an analytical equation expressing the excess Gibbs free energy as a function of liquid
0019-7882/78/1117-0346$01.00/0 0 1978 American Chemical Society
