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111. Furthermore, the copper chromites also show limited activity for the decarbonylation step. The results show that of the catalysts tested, catalyst B (Harshaw 0203 copper chromite) combines both highest activity and selectivity. Conclusion Copper metal has been shown to be the active species for dehydrogenation of methanol to methyl formate, and the degree of dispersion of copper is important in the selection of the most active catalyst. Copper chromite catalysts are particularly effective, as a high copper surface area compensates for a negative contribution of the copper chromite support to dehydrogenation activity. A similar support effect is observed in the case of magnesia. This reduction of activity may result from electronic interaction between the catalyst and the support, or adsorption of hydrogen. The major yield-reducing reaction is consecutive decarbonylation of methyl formate which appears to be catalyzed by copper, chromia, and magnesia. However, the decarbonylation reaction is largely suppressed in the copper chromite catalysts. Acknowledgment Funding of this project under the Australian Research Grants Scheme is gratefully acknowledged. Support was provided under the National Energy Research, Develop-
Dev. 1984, 23,388-393
ment and Demonstration Programme administered by the Commonwealth Department of National Development. Registry No. CH30H, 67-56-1; HCOOCH3, 107-31-3; copper, 7440-50-8; copper chromite, 12018-10-9.
Literature Cited Boerma. H. "Scientiflc Bases for the Preparatlon of Heterogeneous Catalysts", Int. Symp. Societe Chimique de Belglque, Brussels, 1975, A71. Bond, G. C. "Catalysls by Metals"; Academlc Press: London 1962. Charles, E.; Robinet, P. U.S. Patent 2504497, 1950. Chono, M.; Yamamoto, T. Shokubal 1981, 23(1), 3. Choudhary, V. R.; Srinivasan, K. R. J . Chem. Tech. Bbtechnol. 1983, 33A, 271. Dowden, D. A. 4th Int. Congr. Catal. Moscow, 1968. Evans, J. W.; Wainwrlght, M. S.; Brldgewater, A. J.; Young, D.J. Appl. Catel. 1983?7, 75. Frankaerts, J.; Froment, G. F. Chem. Eng. Scl. 1964, 79, 807. Hlgdon, B. W.; Hobbs, C. C.; Onore, M. U. U.S. Patent 3812210, 1974. Hirose, A,; Takezawa, N.; Shimokawabe, M.; Takahashi, K.; Kobayashl, H. Appl. Catal. 1982, 4(2), 127. Ivannikov, P. Ye.; Zherko, A. V. J . Appl. Chem. (USSR) 1933, 6 , 1148. Kawamoto, K. Bull. Chem. Soc.Jpn. 1961, 34, 161. Komarewsky, V. I.;Coley, J. R. J . Am. Chem. SOC. 1941, 63, 700. Lawson, A.; Thomson, S. J. J . Chem. Soc.1964, 1861. Marsden, W. L.; Wainwright, M. S.;Frledrich, J. B. I n d . Eng. Chem. Prod. Res. D e v . 1980, 79, 551. Mlyajlma, E.; Yasumori, I. Shokuball966. 8, 10. Rlekert, L. Ber. Bunsenges, Phys. Chem. 1965, 69(6),499. Takahashl, K.; Takezawa, N.; Kobayashl. H. Appl. Catal. 1982, 2 , 383. Walker, J. F. "Formaldehyde", ACS Monograph Series, Relnhoid New York, 1964.
Receiued f o r review November 14,1983 Accepted February 7, 1984
Process Studies with a Promoted Transition Metal-Zeolite Catalyst Henry W. Pennllne,' Robert J. Gormley, and Rlchard R. Schehl Pittsburgh Energy Technology Center, U.S.Department of Energy, Plffsburgh, Pennsylvanls 15236
The conversion of synthesis gas to gasoiine-range hydrocarbons was investigated wlth a cobalt-thoria-zeolite catalyst. The coprecipitated transition metal and promoter were intimately mixed with ZSMd zeolite and then extruded with an alumina binder. Tests were conducted in a gradientless reactor, where initial results of high yields of gasoline-range hydrocarbons with a low ratio synthesis gas led to an extenshre study of the catalyst. The effects of support, temperature (220-320"C),pressure (1.14-2.17MPa), and feed gas composttion (lHp/1CO-3H2/1CO) on catalyst activity, stability, and product selectivity are discussed. Analyses of the deactivated bifunctional catalyst are also reported.
Introduction Since the middle 1970'9, much interest has been generated in producing liquid transportation fuels from coal. More specifically, various new indirect liquefaction catalytic systems have been investigated for the production of liquid fuels from hydrogen and carbon monoxide mixtures. Second generation gasifiers produce low ratio hydrogen to carbon monoxide mixtures (0.6-0.7), and it is a potential economic advantage to indirect liquefaction schemes to use this synthesis gas directly without a costly shift step (Hildebrand and Joseph, 1979). The Mobil M-Gasoline Process converts methanol to gasoline over the shape-selective acidic zeolite ZSM-5 (Chang, 1983). However, synthesis gas with a high hydrogen to carbon monoxide ratio of at least 2 is required to produce the methanol in this process.
Recent catalyst work has combined a Fischer-Tropsch synthesis function with a shape-selective acid function to convert synthesis gas directly to liquid fuels (Chang et al., 1979; Caesar et al., 1979; Rao et al., 1980). Most conventional Fischer-Tropsch catalysts produce a wide spectrum of hydrocarbon products (Henrici-Olive and Olive, 1976). However, if combined with a shape-selective catalyst, a more selective hydrocarbon product can be produced. In the present investigation, a medium pore zeolite, ZSM-5, was combined with a promoted transition metal to convert low ratio hydrogenjcarbon monoxide synthesis gas to gasoline. Catalyst Preparation The catalyst consisted of a cobalt-thoria coprecipitate that was intimately mixed with zeolite, ZSM-5. This was then extruded with alumina to form the final product. The
This article not subject to US. Copyright. Published 1984 by the American Chemical Society
Ind. Eng. Chem. Prod. Res. Dev., Vol. 23, No. 3, 1984 389
ratio of 18 g of thoria/100 g of cobalt was attempted from past work reported in the literature (Storch et al., 1951). Thoria is known to increase the activity of the cobalt and also shift the hydrocarbon product selectivity to a higher molecular weight fraction (Pennline et al., 1983; Rao et al., 1983). In this particular preparation, an aqueous solution of cobalt nitrate and thorium nitrate, which were present in concentrations of 9.1 and 0.9 wt %, respectively, was heated to 70 OC. A 10 wt 90aqueous solution of sodium carbonate, which is a stoichiometric excess for the precipitation, was also heated to 70 "C. The sodium carbonate solution was slowly added while the nitrate solution was stirred and the precipitate formed. This mixture was boiled for a few minutes and then filtered and washed with hot water until the Ybrownring test" indicated that no nitrate was present in the filtrate. The coprecipitate was dried overnight at 110 "C,crushed, and sieved through 200 mesh (74 pm). The ammonium form of ZSM-5 was fabricated by following example 23 in the patent of Argauer and Landolt (1972). After the synthesized zeolite was washed and dried, it was then calcined a t 538 "C in air. The sodium form of the ZSM-5 was ion-exchanged with ammonium chloride, washed, dried, crushed, and sieved through 200 mesh (74 pm). The sodium content in the ZSM-5 was kept below 0.1 w t %, and the final silica/alumina ratio was about 30/1. Scanning electron microscopic results indicated the particle size of the ZSM-5was about 1 pm. The dried NH4-ZSM-5 was mixed with the dried coprecipitate and rolled overnight for intimate mixing. Initially, this mixture was pelleted for testing. However, due to the friability of the pellets, further testing was conducted with catalyst that was extruded with uncalcined Catapal SB alumina. The structural integrity of the extrudates was greatly enhanced by the alumina binder. Microreactor and internal recycle reactor tests of the pelleted vs. the extruded catalyst indicated that the addition of the alumina binder did not significantly affect the catalyst behavior. Extrudates were ' / 8 in. diameter cylinders with random lengths of approximately ' / 8 in. and the final catalyst composition was 12.5 wt % cobalt, 2.0% thoria, 10-15% alumina, and 57.1% ZSM-5.
Experimental Section. The studies were conducted in a stirred-tank reactor system, as described by Berty (1974). The catalyst extrudates, usually in 18-g batches, were loaded onto a stainless steel screen in a 2 in. diameter basket. Synthesis gas enters the reactor system after fmt being dehumidified in a silica gel trap and then purified with activated carbon. The flow is metered and controlled by a mass flowmeter, and hydrogen can be blended with the lHz/lCO synthesis gas via a similar apparatus. Products exit the reactor via a downward sloping heated line (200 " C ) and enter a hot trap (200 "C), where heavy hydrocarbons, if produced, are condensed. Lighter products are condensed in watercooled or air-cooled traps. The product gas is metered by a wet test meter, and volumetric gas samples are injected into a gas chromatograph that can analyze hydrocarbons up to Cq. Initial studies concluded that (1) bulk heat and mass transfer limitations were eliminated at the 1240 rpm speed of the reactor impeller; (2) the reactor did not contribute to the reaction; and (3) reproducibility of tests was attainable throughout the experimental schedule. The catalyst was brought to synthesis conditions in an identical manner for each test. Initially the reactor was pressurized to 2.17 MPa with hydrogen. The activation procedure began by flowing hydrogen over the catalyst at a weight hourly space velocity (WHSV) of 0.12 h-' while
it 2
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rapidly heating to 200 "C. After maintaining this temperature level for 2 h, the catalyst was heated to 350 "C for 21 h under the hydrogen flow. It has been reported by Kibby et al. (1980) that a 350 "C reduction greatly enhances the hydrogen adsorption capacity of a cobalt catalyst. Afterwards the catalyst temperature was reduced to 250 "C, and then the pressure was decreased to 0.79 MPa. At these conditions, the synthesis gas flow rate was incrementally increased over an hour until the design space velocity for the test was reached. Care was taken during this hour period to prevent temperature runaway. After this induction step, the pressure was increased to operating conditions, and afterwards the temperature was increased (10 OC/h) to synthesis conditions. Trap drainings, flow, and gas analyses were done on a 24-h basis for the material balance determination. Unless otherwise stated, tests in this study used a lHz/lCO feed gas. The gaseous and liquid products were characterized by various analytical techniques. Gas exiting the reactor system was analyzed for hydrogen, carbon monoxide, carbon dioxide, and hydrocarbons up to C4 by gas chromatography. The liquid condensate in the trapping system was collected and separated into an aqueous fraction and an oil fraction. The aqueous phase was analyzed by mass spectroscopy to detect oxygenates, and the water content was determined by the Karl Fischer reagent technique. The liquid hydrocarbon samples were characterized by simulated distillation ASTM D-2887 to determine boiling range distributions, by fluorescent indicator adsorption ASTM D-1319 to determine the functionality of the liquid oil, and by bromine number ASTM D-1159 to corroborate the olefin content. Infrared and 'H N M R studies were also performed on the oil fractions. Relative amounts of terminal, trans-internal, and 8-branched-terminalolefins were determined by infrared absorption. The methyl-tomethylene ratio from 'H NMR was used as a qualitative indication of branching. The wax fraction reported was entirely an artifact of the trapping system and ranged in form from a waxy oil to a hard wax depending on catalyst support and process conditions. This wax fraction was analyzed by simulated distillation and bromine number.
Results and Discussion One of the first experiments in the process study was a life test conducted at 280 "C, 2.17 MPa, and 0.89 h-' WHSV of 1H2/1C0synthesis gas. Figure 1and data listed in Table I for test 1-39at different times on stream indicate that catalyst deactivation is significant. The initial high (H,+ CO) conversion of 83.8% decreases after 417 h on stream to 56.6%. The hydrocarbon product distribution also shifts to a lighter fraction with time, as noted by the increase in percent CHI (23.4 to 42.4) and by the decrease in C,+ weight percent (61.0 to 38.5), which corresponds to the decreasing liquid product yield. The functionality of the liquid oil was approximately constant throughout the test, with a high olefin content (76%) and low content
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of paraffins (20%)and aromatics (4%). Aqueous analyses indicated that more than 99% of this fraction was water. It was concluded from this test that deactivation was a key factor, and subsequently all comparisons would be done after the same time on stream using a different catalyst charge for each test. The elucidation of the role of the zeolite function as well as the promoted transition-metal component of the bifunctional catalyst at the temperature of 280 "C was pursued. Three individual experiments-tests 2-19, 1-42, and 1-39-were conducted at 280 OC, 2.17 MPa, and at identical space velocities of 1H2/1CO synthesis gas, based on only the weight of the promoted transition metal. In test 1-42, coprecipitated cobalt and thoria were added, in the same proportion as the ZSM-5 based catalyst in test 1-39, to calcined Catapal SB y-alumina, and this mixture was then extruded with uncalcined alumina. Test 2-19 consisted of 6.06 g of tableted cobalt-thoria coprecipitate without ZSM-5 or calcined alumina and without alumina binder. Results of the testa are shown in Table I. Conversions between the ZSM-Bbased catalyst in test 1-39 and the alumina-based catalyst in test 1-42 are almost identical, indicating that the promoted transition metal is responsible for the synthesis activity. The conversion in test 2-19 was about half the conversion of either of the other two tests and is probably due to the smaller dispersion of the promoted transition metal in this catalyst than in the other two catalysts. The hydrocarbon product distributions of the cobalt-thoria tablets and cobalt-thoria alumina extrudates are of a broader range than the ZSM5-based catalyst. Catalysts from both tests 2-19 and 1-42 exhibit a gamut of products with a high fraction of heavier materials as compared to the ZSM-&based catalyst. Simulated distillation indicates only 58 and 68% of the liquid oil falls in the gasoline range for tests 2-19 and 1-42, respectively, but 89% is in the gasoline range for test 1-39. The higher fraction of light material may be caused by cracking reactions on the acidic zeolite, as reported by Weitkamp and Jacobs (1983). The liquid product oil from the zeolite-based catalyst was highly olefinic (76%)as compared to the oils from the coprecipitate and alumina-based catalysts (33% and 43%). Infrared studies of the olefinic fraction in the oils indicate that the ratio of internal olefins to terminal olefins is greater with the zeolite-based catalyst. As detected by infrared and corroborated by 'H NMR studies, branching is much more prevalent with the ZSM-5-based catalyst. These isomerization reactions appear to be related to the acidic function of the zeolite. Test 2-16 with an alumina-based catalyst and test 1-43 with a zeolite-based catalyst were conducted at identical process conditions; the temperature was 320 "C. Although the testa showed similar conversions, analysis of the liquid oil from the zeolite-based test indicates 95% boiled in the gasoline range as compared to 78% from the aluminabased catalyst. The functionality of the oils was extremely different. Again, mostly a paraffinic fraction (52%) and low aromatic content (3%)were found with the aluminabased catalyst. Quite differently,the ZSM-5-based catalyst exhibited a high concentration of aromatics (46%),which is attributed to the acidic nature of the zeolite. Mass spectroscopic analysis of this aromatic fraction indicates that 4.8 vol % is toluene, 31.2% C8 alkylbenzene, 34.4% Cg alkylbenzene, 17.7% Clo alkylbenzene, 4.1% Cll alkylbenzene, and 7.8% other cyclics. The effect of reaction temperature was investigated with the cobalt-thoria ZSM-5 catalyst in the five separate tests 3-2,2-10, 1-41, 1-44, and 1-43 at 220,260,280,300, and 320
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Figure 3. The effect of reactor hydrogen to carbon monoxide ratio on hydrocarbon selectivity. 40
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Figure 2. The effect of temperature on conversion.
Table 11. Analysis of Fresh and Used Catalysts test 1-41 fresh
"C,respectively. With the exception of the test at 220 "C, the hydrocarbon selectivity to lighter products increased with temperature, as indicated by the increase in the methane fraction and decrease in the C5+ fraction. In Table I from the simulated distillation analysis, the liquid product boiling in the gasoline range increases from a low of 33% at 220 "C to a high of 95% at 320 "C. The functionality of the product oil is paraffinic at the lower temperatures of 220 and 260 "C and is very similar to the results from the coprecipitated and alumina-based cobalt-thoria tests. The test at 300 "Cyielded a high olefin to paraffin ratio not too different functionally from the 280 "C test. However, at 320 "C, the aromaticity of the liquid oil was very high (46%), and this can be attributed to the acid-catalyzed dehydrocyclization reactions within the zeolite framework (Derouane and Vedrine, 1980). Infrared studies of the olefin fraction of the oil indicate an increase in the intemal olefin to terminal olefin ratio with increasing temperature and negligible terminal olefins formed above 300 "C synthesis temperature. The liquid oil consists of nearly linear chains at 220 "C,but the degree of branching increases with temperature, as detected by both infrared and 'H NMR analyses (see Table I). Unfortunately, at the elevated temperatures, the rate of deactivation is quite significant, as seen in Figure 2. As expected, initial conversions are greater at the higher temperatures, but the rate of deactivation is also greater at the higher reaction temperatures. Accelerated carbon or coke formation at the higher temperature could explain the rapid deactivation. The effect of total pressure was investigated in tests 1-41 and 1-45. As expected, lowering the system pressure decreases the partial pressure of the reactants and thus decreases the conversion. The pressure decrease also shifts hydrocarbon selectivity to a lighter product, as evidenced by an increase in the methane fraction, a decrease in C5+ fraction, and an increase in the fraction boiling in the gasoline range (90% vs. 70%). The functionality of the liquid oil remained the same in both cases, with a high olefin-to-paraffin ratio and little aromatic production. The feed ratios of hydrogen to carbon monoxide in tests 1-39,1-47, and 1-46 were 111, 211, and 311, respectively. The product distribution in Table I for the latter two cases is primarily methane with light paraffins. It should be noted that the temperature of the test with 311 synthesis gas was lowered from 280 to 220 "Cand a highly saturated (96%) liquid product was formed, with the major product still being methane. The catalyst in all these tests in the
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1.3 2.1
