Ind. Eng. Chem. Res. 1993,32, 2602-2607
2602
Low H2/CO Ratio Synthesis Gas Conversion to Methanol in a Trickle Bed Reactor Sinoto Tjandra, Rayford G. Anthony, and Aydin Akgerman' Chemical Engineering Department, Texas A&M University, College Station, Texas 77843-3122
The performance of a trickle bed reactor for methanol synthesis over a Cu/ZnO/A1203 catalyst was investigated. The pressure was kept constant at 5.2 MPa, and the CO/CO2 ratio was maintained at 9.0 throughout the investigation. The effects of time on stream, space velocity, Hz/(CO CO2) ratio, and liquid-phase flow rate on methanol productivity were studied a t 508,523, and 533 K. The H2/(CO CO2) ratios were set at 0.5,1.0, and 2. The space velocity was in the range 3000-20000 L/(h kg of catalyst) at 523 K and 3000-8000 L/(h kg of catalyst) at 508 and 533 K. The feed gas along with an inert oil was passed over a fixed bed of catalyst particles of size 500-600 pm. The product was mainly methanol with small amounts of water. The catalyst deactivated during the first 100 h on stream and reached a steady-state activity which was maintained for over 1000 h on stream. Methanol and water productivities increased with space velocity and with HZ/(CO + C02) ratio, as well as with temperature. Productivities as high as 39.5 mol of methanol/(h kg of catalyst) were achieved. The results are compared to slurry reactor data a t the same and similar conditions.
+
+
Introduction Methanol is one of the major basic raw materials in the chemical industry. Methanol is the only oxygenate that can be produced directly from synthesis gas with high selectivity. It also provides a single carbon building block for making higher oxygenated chemicals, which cannot be synthesized directly and in high selectivity from synthesis gas (King and Grate, 1985). At present, most of the methanol produced is from synthesis gas derived from natural gas in gas-phase fixed bed reactors that have elaborate design schemes for removal of reaction heat. Due to the highly exothermic nature of the reaction, heat dissipation becomes a severe problem in gas-phase fixed bed reactors, and the maximum productivity achieved cannot exceed 1kg/(h kg of catalyst) (Wentworth and Stiles, 1980). Thus, elaborate schemes such as multibed reactors with cold shot feeds in between the beds are used. Furthermore, the reactors have to be operated at high Hz/(CO + CO2) ratios with CO concentrations not to exceed 12%. If the CO concentration exceeds 12%, runaway reactions will likely occur (Stiles et al., 1991). In addition, the unreacted hydrogen also acts as a heat sink and prevents the runaway reactions. Recently, ChemSystem, Inc. and Air Products, Inc. developeda liquid-entrained catalytic reador (Frank, 1980; Frank and Mednick, 1982; Weimer et al., 1987). The catalyst is slurried in a mineral oil which has high heat capacity. Therefore, the heat of reaction is absorbed by the liquid, which makes the operation nearly isothermal. This type of reactor can operate at a wide range of Hz/(CO + COZ)ratios and without catalyst deactivation at low H d ( C 0 + C02) ratio. This makes it suitable for using synthesis gas derived from coal as feed gas which has low H d ( C 0 + C02) ratio. However, the slurry reactor has an upper limit for catalyst loading and hence operates at low space velocities. I t also has low conversion per pass due to the high extent of backmixing, and difficulties are encountered in separating the catalyst from the slurry because of catalyst attrition and agglomeration. Trickle bed reactors, which have concurrent gas/liquid flow over a fixed bed of catalyst, offer some advantages for methanol synthesis. Trickle bed reactors combine the
* Author to whom correspondence should be addressed.
advantages of both the slurry reactors and gas-phase fixed bed reactors. Similar to gas-phase fixed bed reactors, trickle bed reactors can have high catalyst loading and hence can operate at high space velocities. They also operate at nearly plug flow conditions, leading to higher conversions per pass. Since the bed is fixed, there is no catalyst attrition which permits the use of costly catalysts. Moreover, the trickle bed reactors, like the slurry reactors, employ a suitable liquid phase to absorb heat generated by the reactions that makes the operation nearly isothermal. However, pressure drop in trickle bed reactors may be significant if small catalyst particles are used, and large catalyst particles (in excess of 3 mm) may introduce limitations due to intraparticle diffusion effects. The purposes of this study were to investigate methanol synthesis in a trickle bed reactor and to study the effects of H&CO + COZ)ratio, space velocity, temperature, and oil flow rate on the methanol production rate, conversion of the reactants, and composition of the product. The effect of time on stream on catalyst activity was studied as well. The synthesis was carried out over a Cu/ZnO/ A1203 catalyst in the temperature range 498-523 K and at a pressure of 5.2 MPa. The Hz/(CO + C02) ratio was varied at 0.5, 1.0, and 2.0, and the CO/COz ratio was maintained at 9.0. The results obtained from this study were compared to those obtained in slurry reactors operating at the same conditions. There are not many studies on methanol synthesis using a trickle bed reactor. Pass et al. (1990) carried out methanol synthesis in a trickle bed reactor and observed a higher methanol productivity compared to slurry reactors (Frank and Mednick, 1982;Weimer et al., 1987;Krishnan et al., 1989). It was also shown that trickle bed reactors have better transport properties than slurry reactors. However, the methanol production rate is a function of many variables, namely, H2/(CO + C02) ratio, temperature, type and composition of the catalyst, and even the reduction procedure of the catalyst. Since Passet al. (1990) conducted the study under different conditions from those used in slurry reactor studies, it is not meaningful to compare the two reaction rates. This necessitates the need to perform a study in a trickle bed reactor, using the same catalyst and under the same conditions which were used in a slurry reactor.
0888-5885/93/2632-2602~04.00l0 0 1993 American Chemical Society
Ind. Eng. Chem. Res., Vol. 32, No. 11, 1993 2603 "2 YENT
Back Pressure Regulator C Carbon Oxlder Mixture CD Condenser G Glass Bead Bed GB GuardBcd H Hydrogen HB Herring Block M Mass Flow Controller P Pump PI Pressure [ndicatot R Rencter B
R V Relieve Valve
s
SB
Separator Soap Bubble Meter
VI
Three Way Valve
Figure 1. Trickle bed reactor assembly.
Experimental Apparatus and Procedure Figure 1 depicts the overall process for methanol synthesis in a trickle bed reactor. Feed gas to the reactor consists of pure HZand premixed CO/COz mixtures with a CO/COz ratio of 9.0. This ratio was kept constant throughout the study. All the feed gases are purified by flow through a guard bed which consists of activated carbon and molecular sieve with a particle size of 0.16 cm in order to remove any carbonyls that would poison the catalyst. HZand CO/COz gas flow rates are controlled with Brooks Model 58503 mass flow meters. The feed gas streams are combined at a predetermined H2/(CO + COZ)ratio. To enhance mixing, the gases are passed through a bed of glass beads prior to the reactor. The mineral oil is introduced with a Milton Roy Model DE-1-6OP pump at the reactor inlet, along with the gas mixture. The oil flow rate is maintained at 3 kg/(m2s), which is sufficiently high to ensure complete wetting (Colombo et al., 1976; Sicardi et al., 1980; Dudukovic and Mills, 1986). A relief value, which is set to 10.6 MPa, is placed before the reactor to prevent an uncontrolled pressure rise in the system. The reactor is a 316 stainless steel tube 25 cm in length with 0.96-cm i.d. and 1.20-cm 0.d. It is mounted vertically in a bed of aluminum pellets. The reactor bed is divided into three sections. Prior to and after the 5.6cm catalyst bed are 6.0- and 13.4-cm supporting sections. These are filled with 0.2-cm-diameter glass beads. The reactor is heated by the heating block, and the temperature is controlled by an Omega Model 6100 temperature controller. A thermocouple inserted in the catalyst bed is usedto measure the temperature of the bed. The reactor pressure is maintained at 5.2 MPa with a Grove Model 91W back pressure regulator. Reactor effluent passes through the back pressure regulator where the pressure is reduced to atmospheric pressure prior to the gas-oil separator. The separator is heated to 373-383 K with a heating tape, and the temperature is controlled with an Omega Model 6100 temperature controller. The oil collected a t the bottom of the separator is recycled back to the reactor. After the gas-oil separator, a Gow Mac 550 gas chromatograph equipped with a HP 3969 integrator is installed to check the steady state of the reaction. Methanol, water, and any other products are separated from the unreacted gas after the gas chromatograph by using a series of condensers immersed in a dry ice/acetone bath. The condensate is analyzed off-line on the Gow Mac 550 gas chromatograph. The tail gas is passed through a soap bubble meter to measure the volumetric flow rate before it is vented to the hood. A sample of the tail gas is taken at the sample port after the condensers and
injected into a Carle gas chromatograph, equipped with a Varian 4290 integrator, to analyze the composition of the tail gas. The oil used in this study was supplied by Witco Co. (Witco Freezene 100) and consists of saturated aliphatic and naphthenic hydrocarbons. The average molecular weight of the oil is 349 g/mol, and the specific gravity is 0.71 at 523 K. The catalyst used in this study is a Cu/ZnO/A1203water gas shift reaction catalyst (United Catalyst L-951). Seven grams of the catalyst is crushed to 500-600-pm particle size and put into the reactor. Since the catalyst is partially oxidized, it is necessary to reduce the catalyst prior to methanol synthesis. In situ reduction is carried out accordingto the procedure described by Sawant et al. (1987). The catalyst is reduced by passing 5 % hydrogen in nitrogen with a space velocity of 5000 L/(h kg of catalyst). The oil flow rate is maintained at 3 kg/(m2s), and the pressure is held a t 1.7 MPa. The reactor is heated from room temperature to 393 K and is held there for 1h. The reactor is then heated again to 448 K and held there for 1h. Then the temperature is brought to 478 K, and synthesis gas is introduced into the reactor for 6 h. It should be noted that this reduction procedure may be catalyst specific and may depend on the catalyst loading in g of catalyst/g of oil. Sawant et al. (1987) recommended this procedure for UC EPJ-19 catalyst, and it is at best questionable whether the conditions specified above are the optimum for the catalyst used in this study, UC L-951. However, optimizing the reduction process would result in better activity for the catalyst used in this study. The reactor is heated to 523 K and the pressure is adjusted to 5.2 MPa prior to running the reaction by flowing Nz a t 1 mL/min. Once the reactor temperature reaches constant temperature, the nitrogen flow is stopped and the synthesis gas with a predetermined space velocity and H d ( C 0 + COz) ratio is introduced into the reactor, by adjusting the setting of the mass flow controllers, to start the reaction. The steady state of the reaction a t a particular set of conditions is checked by using the on-line Gow Mac 550 gas chromatograph. If the steady state is reached, the three-way valve, V1, is turned to the position where the effluent gas can flow to the condensers instead of bypassing them. The condensers are dried and weighed prior to a run. The products are then collected in the condensers for a certain period of time. During the collection,a sample of the tail gas is taken and analyzed using the Carle gas chromatograph, and its flow rate is measured. After the run, the three-way valve is then turned to the bypass position. The condensers are weighed to determined the accumulation of the reaction products. Then the reaction can be either terminated or continued at another set of conditions. For all the experiments performed in this study, the atomic material balance closures, defined as outletdinlet, were 0.99 f 0.06 for C, 0.98 f 0.05 for 0, and 0.95 f 0.08 for H, with the exception of experiments a t 523 K and a H d ( C 0 + COZ) ratio of 2, where the hydrogen atom material balance closure was 1.05 f 0.05. The high error may be due to low conversions of hydrogen (15-28% 1.
Results and Discussion The studies were performed at three temperatures, 508, 523, and 533 K and a space velocity range of 3000-2oooO L/(h kg of catalyst) at 523 K and 3000-8000 L/(h kg of catalyst) at 508 and 533 K. The space velocity is defined
2604 Ind. Eng, Chem. Res., Vol. 32, No. 11, 1993
12
~~
0
200
600
400
800
loo0
".W
0
200
Time on Stream (h)
Figure 2. Variation of methanol productivity with time on stream at 523 K, 5.2 MPa, H2/(CO + COz) ratio of 1, and space velocity of 5000 L/(hkg of catalyst).
at a temperature of 298 K and a pressure of 0.1 MPa. As mentioned above, the H d ( C 0 COZ)ratios used were 0.5, 1, and 2. Methanol was the main product, and a small amount of water was formed as a side product. No other product was detected during the investigation, but a trace amount of methane and a trace amount of unidentified product (observed as a peak in Carle GC backflush, suspected to be dimethyl ether) were observed at a Hz/ (CO C02) ratio of 2 and at high space velocities. The observations by Pass (1990) and Al-Adwani (1992) indicated that the catalyst deactivated significantly at the beginning of the synthesis. Therefore, it was necessary to run a base line experiment a t specified conditions after each experiment at different conditions in order to verify that the catalyst maintains its steady-state activity. The reaction was started at the base line conditions which were specified as Hz/(CO COZ)ratio of 1.0, temperature of 523 K, and a space velocity of 5000 L/(h kg of catalyst). These base line conditions were run prior to and after each experimental series. Pass et al. (1990)evaluated the mass-transfer parameters for the reactor used in this study, and it was shown that interparticle and intraparticle mass-transfer resistances were negligible. Further simulation studies verified this assumption (Akgerman and Anthony, 1993). Therefore, in this study mass-transfer limitations are neglected. The catalyst was found to deactivate during the first 100 h and approach a constant activity afterward. No catalyst deactivation was observed even after approximately 1000hof time on stream (Figure 2). Since hydrogen is the limiting reactant at a Hz/(CO + COZ)ratio of 1.0, its conversion behaved like the methanol productivity, which decreased with time on stream and stayed constant after 100 h. At this ratio, the time on stream does not have any significant effect on carbon monoxide conversion, which remained constant at about 17% ,and the methanol productivity remained constant at about 16 mol/(h kg of catalyst). An important observation is shown in Figures 3 and 4;the amount of water in the effluent, although small, increases with time on stream, and correspondingly COz conversion increases with time on stream as well, although the methanol production rate remained constant. Scatter in the COZconversion data is due to measuring the difference of two small quantities; the C02 inlet concentration was 5%. However, since the amount of water in the condensate isvery small, ita effect on methanol productivity is within experimental error. This observation suggests that the COz hydrogenation (either by water gas shift or methanol synthesis reactions) rate increases
+
600
800
loo0
Time on Stream (h)
Figure 3. Increase in water production with time on stream at 523 K,5.2MPa,H2/(CO+ C02)ratioofl,andapacevelocityof5000L/(h kg of catalyst). 0'4
0.3
+
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200
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Figure 4. Increase in C02 hydrogenation rate with time on stream at 523 K, 5.2 MPa, HJ(C0 + Con) ratio of 1, and space velocity of 5000 L/(hkg of catalyst).
with increasing time on stream. This observation is consistant with the findings in a slurry reactor, where an increase in the water content with time on stream was also observed (Al-Adwani, 1992). The first experimental series was run at a time on stream of 231h where the catalyst had achieved aconstant activity. The temperature remained at 523 K. The space velocity was set to 3000, 5000, and 7500 L/(h of catalyst) at the H d ( C 0 + C02) ratio of 1.0 and was set to 3000,5000, and 8OOO L/(h kg of catalyst) a t the other Hd(C0 + COZ)ratios. An experiment at base line conditions was run to check the catalyst activity after each experiment. The temperature was then reduced to 508 K where the space velocity was set to 3000,5000, and 8000 L/(h kg of catalyst) for the Hz/(CO + C02) ratios of 0.5,1.0, and 2.0. A run was then carried out at the base line conditions to determine the steady state of the catalyst activity. The reaction temperature was increased to 533 K, and the space velocity was adjusted to 3000,5000, and 8OOO L/(h kg of catalyst) for the three Hd(C0 + COP)ratios. Again a check for the catalyst activity was made prior to termination of the reaction. These studies took over 1000 h on stream, and experiments a t base line conditions verified that the activity remained constant throughout this period. The used catalyst was unloaded after over 1000 h on stream (due to reactor runaway and catalyst sintering caused by a burned relay in temperature controller), and 7 g of fresh catalyst (batch 2) with a particle size of 500600 pm was put into the reactor. After reduction using the same procedure as in the previous experiment, the
Ind. Eng. Chem. Res., Vol. 32, No. 11, 1993 2605 _"
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Figure 7. Methanol productivity at 533 K and 5.2 m a .
Figure 5. Methanol productivity at 508 K and 5.2 MPa.
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Figure 6. Methanol productivity at 523 K and 5.2 MPa.
reaction was started at the base line conditions. At the base line condition, it was observed that the new batch of catalyst had methanol productivity of 14.9 mol/(h kg of catalyst), which was 6% less than that of the first batch of catalyst. This was probably due to the different heating rate during the catalyst reduction procedure caused by a different wiring scheme for the reactor heater. In order to combine the two experimental series, the catalyst activity must be the same at the base line conditions. Therefore, it was necessary to raise the temperature during the investigation with batch 2 catalyst to 530 K to match the catalyst activity of batch 1 catalyst. A set of experiments was conducted at 530 K and Hz/(CO + COz) at 1 at various space velocities to ensure that the activity of batch 2 catalyst is the same as that of batch 1 catalyst. Experiments were then conducted at high space velocities, 10000-20000 L/(h kg of catalyst) and at various H2/(CO + COz) ratios. Figures 5-7 summarize the methanol productivity at each experimental condition. CO conversion and Hz conversion at 523 K are given in Figures 8 and 9. However, it should be realized that since methanol synthesis takes place through three reactions, two of which are independent, conversions do not have much significance, and productivity is the important parameter. It should be noted that at low Hz/(CO + COd ratios hydrogen is the limiting reactant, and ita conversionis higher whereas at the stoichiometric ratio CO conversion is higher. Nevertheless, the water gas shift reaction compensates for the reactant deficiencies. Thus, conversions do not relate to the conversion of a reactant to methanol. Water content in the product behaves similar to methanol formation rate. More water was produced at higher Hz/(CO+ C02) ratio and higher temperature. Water
,
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Space Velocity, I/(h kg cat)
Figure 8. Effect of the liquid-phase flow rate on methanol productivity. H2/(CO+C02)-0.5
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Space V e l o c i t y , I/(h kg cat) Figure 9. Carbon monoxide conversion at 523 K and 5.2 MPa.
mole fraction in the product increased with increasing space velocity a t each temperature. This implies that the carbon dioxide hydrogenation rate increases at higher space velocities. In order to maintain complete wetting of the catalyst in a trickle bed reactor, the liquid mass flux should be maintained above 2.0 kg/(m2 s) (Colombo et al., 1976). Another batch of fresh catalyst (batch 3) of the same amount and particle size was loaded into the reactor to study the effecta of oil flow rate. Again, the same reduction procedure was followed, and the reaction was started at the base line conditions. Methanol productivity with time on stream at base line conditions for this run was the same as that for batch 2 catalyst. The study was conducted at
2606 Ind. Eng. Chem. Res., Vol. 32, No. 11, 1993
0 This Study (1.0) This Study (0.5) A Al-Adwani, 1992 (2.0) ALAdwani. 1992 (1.0)
14
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I 2
3
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Oil flow rate (kg/m2 s)
Figure 10. Hydrogen conversion at 523 K and 5.2 MPa.
a Hz/(CO + C02)ratio of 1.0 and a temperature of 523 K. The initial oil flow rate was 3.0 kg/(m2 s) and then was adjusted to 2.5, 2.0, and 1.5 kg/(m2 s). A slight increase of methanol production rate was found with decreasing oil flow rate. However, it increased abruptly at the oil flow rate of 1.5 kg/(mZ s) (Figure 10). This is expected since the reactants are in the gas phase. Lowering the oil flow rate to a certain degree will lead to lower liquid holdup and possibly to partial wetting of the catalyst or reduced heat-transfer capacity for the liquid phase. Moreover, there is a possibility of development of dynamic dry spots; i.e., parts of the catalyst which are not wetted both externally and internally form and disappear, which would promote the reaction rate significantly. This was confirmed by the occurrence of runaway reactions when an attempt was made to run the experiment with an oil flow rate of 1 kg/(m2 8 ) ; it was not possible to reduce the flow rate to less than 1.5 kg/(mZs)without runaway. The water productivity increased with decreasing oil flow rate as well (Tjandra, 1992). Since the maximum methanol production rate of gasphase methanol synthesis reaction in fixed bed reactors is about 30 mol/(h g of catalyst), trickle bed reactors, at least in laboratory-scale reactors, outperform gas-phase fixed bed reactors by achieving a methanol productivity as high as 39.5 mol/(h kg of catalyst). In addition, the gas-phase reactors necessitate elaborate design for heattransfer effectiveness, but the trickle bed reactor achieves this productivity without any heat-transfer limitation. A further advantage is operation at low Hd(C0 + COZ)ratios, which is also a limiting factor for gas-phase fixed bed reactors, without any catalyst deactivation. Al-Adwani (1992) carried out methanol synthesis in a slurry reactor over the same catalyst used in this study. Al-Adwaninot only conducted the experiment at the same conditions but also followed the same catalyst reduction procedure. Figure 11 shows that methanol production rate in trickle bed reactors is significantly higher than that in slurry reactors. Some of this difference is probably due to better gas/liquid/solid contact in trickle beds combined with hydrodynamic effects; i.e., trickle beds operate close to plug flow conditions while there is a high extent of backmixing in the slurry reactors. However,this comparison should be taken with a certain degree of skepticism, because the activation procedure used in the slurry and the trickle bed reactors were not exactly the same, and there was a possibility for gas bypass in the slurry reactor system we employed. In addition, a set of experiments we performed in a 300-cm3 slurry reactor (Pass, 1990) also exhibited similar behavior to data of AlAdwani (19921, which were taken in a 100-cm3 slurry
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Space Velocity (I/h kg cat.)
Figure 11. Comparison of methanol productivities obtained in a trickle bed reactor and a slurry reactor operating with the same catalyst at the same conditions.
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Figure 12. Comparison of methanol productivity data obtained in this study to literature data.
reactor. The initial deactivation data as well as the steadystate productivity data taken in 100-and 300-cm3reactors were similar. In the trickle bed reactor, the catalyst activity stabilized in 100 h compared to 170h in the slurry reactor. We believe that the difference in activities presented in Figure 11 is due to the effect of mixing in the catalyst reduction process. Since all the trends in the slurry and the trickle bed reactor systems are the same, a catalyst reduction procedure due to the hydrodynamic effects probably contributed to the differences in activities (Akgerman and Anthony, 1993). Figure 12 compares methanol productivity data a t 523 K obtained in this study with data obtained from other trickle bed reactors (Passet al., 1990) and slurry reactors (Frank and Mednick, 1982; Weimer et al., 1987). The numbers in parentheses in the legend of Figure 12indicate the Hz/(CO + C02) ratio used in each study. Although using the same catalyst, Pass et al. (1990) reported methanol productivity data higher than those in this study. This may be due to different reduction conditions used by Pass et al. and to performing experiments with a short time on stream, probably before the catalyst achieved its steady-state activity. Moreover, the methanol productivity data from this study are about the same as those from Frank and Mednick and from Weimer et al. It should be noted that the catalyst used by them was a catalyst optimized for methanol synthesis, whereas the catalyst we used was a water gas shift catalyst not optimized for methanol. In addition, it was not reported whether their
Ind. Eng. Chem. Res., Vol. 32, No. 11, 1993 2607 data were taken before or after the catalyst reached its steady-state activity. Furthermore, in a trickle bed reactor the catalyst volume in effect is the reactor volume. Hence, space velocity in terms of the catalyst volume or the reactor volume is the same. In a bubble column or a slurry reactor, the maximum catalyst loading of the liquid phase is about 45 % by weight based on efficient catalyst dispersion. Even if the volume due to gas holdup is neglected, catalyst volume in the reactor will be 20% ; with gas holdup it is more like 10%. Thus, if the space velocity is expressed in terms of the reactor volume, that is if GHSV is used instead of WHSV, a reactor 9 times as large would be needed for the bubble column operation if there are no mass-transfer limitations in the trickle bed reactor. In a commercial reactor, larger catalyst particles will be required to minimize the pressure drop in the reactor. For example, in hydrotreatingreactors typically 1/16-in. extrudes are used. Due to diffusion limitations, larger particles will result in productivity reduction. However, when 2.5-mm (1/16-in. extrudes of 4-5-mm length) catalyst particles are used in the trickle bed reactor, the effectiveness factor will be about 0.25 due to methanol diffusion limitations (Akgermanand Anthony, 1993). This implies that in terms of GHSV the productivity in a trickle bed reactor will still be about double that in a slurry reactor of the same volume.
Acknowledgment This study was conducted as part of the project funded by the United States Department of Energy under Contract DE-FG22-89PC89787. The authors greatly appreciate the financial support from DOE.
Literature Cited Akgerman, A,; Anthony, R. G. Novel Reactor Configuration for Synthesis Gas Conversion to Alcohols; Final Report, DE-FG2289PC89787, Pittsburgh Energy Technology Center, 1993.
Al-Adwani,H. A. A Kinetic Study of Methanol Synthesis in a Slurry Reactor Using a CuO/ZnO/Al20a Catalyst. Master of Science Thesis, Texas A&M University, 1992. Colombo, A. J.; Baldi, G.; Sicardi, S. Solid-Liquid Contacting Effectivenese in Trickle Bed Reactors. Chem. Eng. Sci. 1976,31, 1101-1108. Dudukovic, M. P.; Mills, P. L. Contacting and Hydrodynamics in Trickle Bed Reactors. In Encylopedia of Fluid Mechanics; Cheremisinoff, N. P., Ed.; Gulf Pub.: Houston, 1986; p 969. Frank, M. E. ChemSystems Liquid-Phase Methanol Process. Presented at 15th Intersoc. Energy Conversion Eng. Conf., Seattle, 1980. Frank, M. E.; Mednick, R.L. Liquid-Phase Methanol Project Status and Laboratory Results. Preaented at 7th Ann. EPRI Contractor’s Conf. on Coal Liquefaction, Palo Alto, 1982. King, D. L.; Grate, J. H. Look What You Can Make From Methanol. CHEMTECH 1985, April, 244-251. Krishnan, C.; Elliott, J. R., Jr.; Berty, J. M. Continuous Operation of the Berty Reactor for the Solvent Methanol Process. Paper 133e, AIChE Annual Meeting, San Francisco, 1989. Pass, G. G.A Kinetic Model for Methanol Synthesis on a Cu/Zn/ CrzOs Catalyst. Master of ScienceThesis, Texas A&M University, 1990. Pass, G.G.; Holzhauser, C.; Akgerman, A.; Anthony, R.G. Methanol Synthesis in a Trickle Bed Reactor. AZChE J. 1990,36, 10541060.
Sawant,A.;Ko,M.K.;Parameswaran,V.;Lee,S.;Kulik,C.J.InSitu Reduction of a Methanol Synthesis Catalyst in a Three Phase Slurry Reactor. Fuel Science and Technology Znt’l. 1987,47788. Sicardi, S.; Baldi, G.;Specchia, V. Hydrodynamic Modela for the Interpretation of the Liquid Flow in Trickle-Bed Reactors. Chem. Eng. Sci. 1980,35,1775-1782. Stiles, A. B.; Chen, F.; Harrison, J. B.; Hu, X.;Storm,D. A.; Yang, H. X. Catalytic Conversion of Synthesis Gas to Methanol and Other Oxygenated Products. Znd. Eng. Chem.Res. 1991,30,811821. Tjandra, S. Methanol Synthesis in a Trickle Bed Reactor. Master of Science Thesis, Texas A&M University, 1992. Weimer, R. F.; Terry, D. M.; Stefanoff, P. Laboratory Kinetics and Mass Transfer in the Liquid-Phase Methanol Procese. Paper 25d, AIChE Ann. Meet., New York, 1987. Wentworth, T. 0.; Stiles, A. B. U. S. Patent 4,235,799,1980.
Received for review June 15, 1992 Revised manuscript received March 8, 1993 Accepted June 15,1993
