I n d . Eng. C h e m . Res. 1987,26, 183-188
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Performance of Dual-Reactor System for Conversion of Syngas to Aromatic-Containing Hydrocarbons Raghunandan L. Varma and Narendra N. Bakhshi* Catalysis & Chemical Reaction Engineering Laboratory, Department of Chemical Engineering, University of Saskatchewan, Saskatoon, Saskatchewan, Canada S 7 N OW0
Joseph IF'. Mathews Department of Chemical Engineering, Monash University, Clayton, Victoria, Australia 3168
Siauw H. Ng Energy Research LaboratorieslCANMET, Department of Energy, Mines and Resources, Ottawa, Ontario, Canada K1A OG1
The product stream from the synthesis gas reaction over cobalt-nickel-zirconia catalyst (FT catalyst) was modified by HZSM-5 zeolite in a dual-microreactor system a t 101.3 kPa and H 2 / C 0 = 1. The effect of operating conditions on steady-state product distribution was investigated by varying the HZSM-5/FT ratio and the HZSM-5 reactor temperature. T h e aromatics selectivity in total hydrocarbons increases significantly either by increasing the temperature of the HZSM-5 reactor or by increasing the HZSM-5/FT ratio from 1 t o 4, when the temperature of the HZSM-5 reactor is maintained below 300 "C. At higher temperatures, however, cracking reactions become prominent, resulting in a decrease in the selectivity of the liquid hydrocarbons and aromatics. In addition, a remarkable increase in ethylene selectivity is observed with a n increase in the HZSM-5 reactor temperature. This is attributed to the favorable shift in reaction kinetics toward the formation of olefins from oxygenates. The Fischer-Tropsch (FT) process for synthesis of hydrocarbons from carbon monoxide and hydrogen has been extensively studied and reviewed (Storch et al., 1951; Vannice, 1976; Dry, 1981; Anderson, 1984). The process yields a broad spectrum of hydrocarbons ranging from methane to waxes. Significant amounts of oxygenates are also produced. The product distribution consists mainly of normal paraffins and linear olefins with little or no aromatics. The F T synthesis is essentially a polymerization process (Anderson, 1984; Biloen and Sachtler, 19811, which results in low selectivity to the desired hydrocarbons. The effluent products require expensive separation schemes and upgrading operation to improve the octane number of liquid products used as gasoline (Dry, 1982). Recently, a new approach has been used to produce hydrocarbons that have improved selectivity and that are rich in aromatics from synthesis gas in one step (Chang et al., 1979; Caesar et al., 1979; Huang and Haag, 1981; Bruce et al., 1984; Shamsi et al., 1984; Varma et al., 1985). In a single-stage process, the shape-selective acid catalyst such as HZSM-5 is combined with a CO reduction catalyst such as FT catalyst or methanol synthesis catalyst (Chang et al., 1979; Caesar et al., 1979). Although the single-stage process is remarkably effective in producing the gasoline-range hydrocarbons containing a substantial percentage of aromatics, there are some disadvantages associated with this single-stage process. The F T catalyst usually deactivates faster than the HZSM-5 catalyst. Therefore, the FT catalyst may require regeneration or replacement while the HZSM-5 is still active. Furthermore, the regeneration procedures for the F T catalyst are different from those for the HZSM-5 catalyst. In a sequentially arranged two-stage process, the product stream obtained from the F T reactor is further upgraded in a second reactor containing zeolite HZSM-5 (Haag and Huang, 1981; Diffenbach et al., 1983). The main advantages of the two-step process are as follows: (i) considerable flexibility with respect to both reaction conditions as well as product distribution; (ii) operation of each component
a t its optimum reaction conditions (the operating temperature of the F T catalyst is usually lower than that of HZSM-5 by about 100-200 "C); and (iii) easier regeneration and replacement of the spent catalyst. Despite these advantages, little information is available in the literature on the performance of a dual-reactor system. Haag and Huang (1981) and Diffenbach et al. (1983) have reported the product distribution obtained in the dual-reactor system using promoted iron catalysts as the F T component. In this work, the synthesis of hydrocarbons in a dual-microreactor system has been investigated using a cobalt-nickel-zirconia catalyst as the FT component. The effect of operating conditions such as the HZSM-5/FT ratio and zeolite reactor temperature on steady-state product distribution has been reported. The FT catalyst was chosen on the basis of earlier studies (Bruce and Mathews, 1982; Bruce et al., 1983) dealing with several cobalt-nickel-zirconia systems of varying compositions, some of which were found to exhibit excellent activity and selectivity for the Fischer-Tropsch reaction. Also, it was found that these catalysts combined with an excess volume of HZSM-5 in a single-stage reactor can give good yields of liquid hydrocarbons with aromatic selectivities approaching 30-35 wt % of the total hydrocarbons under relatively mild conditions (Bruce et al., 1984; Varma et al., 1985).
Experimental Section Preparation and Characterization of Catalysts. The cobalt-nickel-zirconia catalyst (FT catalyst) was prepared by impregnation of 10% Co over a 5% Ni/Zr02 catalyst used as a support. The support, viz., 5% Ni/Zr02, was prepared by coprecipitation of zirconyl nitrate and nickel nitrate solutions in the presence of excess alkali solution under a nitrogen atmosphere. the resulting cake was repeatedly washed with distilled water and filtered under nitrogen atmosphere until the pH of the filtrate was neutral. This was followed by two washings with acetone and drying in a vacuum oven at 60 OC for 24 h. The
0888-5885/87 / 2626-0183$01.50/0 0 1987 American Chemical Society
184 Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987
-25
-78 I'
Figure 1. Schematic diagram of dual-reactor and analytical system for on-line analysis: I, pressure regulator; 2, valve; 3, gas purifier; 4, control valve; 5, three-way valve; 6, mass flow meter; 7, pressure gauge; 8, F T reactor furnace; 9, HZSM-5 reactor furnace; 10, gas sampling valve; 11,sample loop; 12, thermostated chamber; 13, GC oven; 14, ice-cooled condenser; 15, Carbosieve S column; 16, Chromosorb 102 column; 17,1070OV-101 on Chromosorb W HP column; 18, FID; 19, TCD; 20, reporting integrator; 21, flow splitter; 22, multichannel temperature indicator; 23, temperature-controller indicator; 24, bubble flow meter; 25, vent; 26, FT reactor preheater; 27. HZSM-5 reactor; 28, thermocouples.
impregnated FT catalyst was dried in an oven at 110 "C followed by calcination in a muffle furnace at 500 "C for 24 h. The FT catalyst had the composition 10% Co, 4.5% Ni, and 85.5% ZrO,. The BET surface area of the FT catalyst was 75 m2/g. The zeolite ZSM-5 was prepared in the laboratory according to the method described in the patent literature (Plank et al., 1979). The zeolite was washed well with distilled water, calcined a t 500 "C, and treated with 0.5 M hydrochloric acid for 16 h under reflux conditions to convert it to the proton form. The X-ray diffraction pattern of the zeolite exhibited excellent agreement with the X-ray data for ZSM-5 given by Argauer and Landolt (1972). The weight ratio of Si02/A1,03 in the zeolite was 34, and the surface area of the zeolite was 362 m2/g. Apparatus and Procedure. Fixed-bed continuousflow microreactors were used and were made of 4.6-mm-i.d. 304 stainless steel tubing. The schematic diagram of the apparatus is given in Figure 1. The apparatus consisted of two sequentially arranged downflow reactors; the first reactor contained the FT catalyst and the second reactor contained the HZSM-5 catalyst. In the first reactor, the syngas mixture reacted with the FT catalyst, resulting in the formation of a variety of hydrocarbons and oxygenates. The product mixture obtained from this FT reactor was fed to the second reactor containing the HZSM-5 zeolite. The effluent from the HZSM-5 reactor was passed directly through the GC system for product analysis via heated tubes. The reactor and preheater sections of both of the reactors were surrounded by metal block heaters. The temperatures of both of the reactors were controlled independently within i l "C by individual temperature controllers. The feed gas was a premixed mixture (99.970
pure) of CO and H2(ratio 1:l)containing 20% Ar (by volume) which served as an internal standard for product analysis. The operating procedure was as follows: The F T and the HZSM-5 catalysts were loaded in their individual reactors. The amount of FT catalyst used was 0.1 g, while that of HZSM-5 ranged from 0.1 to 0.4 g. In order to reduce the FT catalyst, hydrogen was passed through the reactors at a rate of 15 mL/min, and the temperatures of both of the reactors were maintained at 450 "C for 16 h. Later, the temperatures of both of the reactors were lowered to the reaction temperatures while hydrogen was still flowing. As the temperature stabilized, the flow of hydrogen was replaced by the premixed synthesis gas at a given flow rate. At these fixed flow rate and reactor temperature conditions, about 20-24 h were allowed to obtain the steady-state product distribution. In all the experiments, the weight hourly space velocity, WHSV, was 2.2 h-I (WHSV is defined as the hourly mass flow rate of CO + H2per unit weight of the FT catalyst), and the pressure was 101.3 kPa. The temperature of the F T reactor was fixed at 250 "C, while that of the HZSM-5 reactor was varied in the range of 250-400 "C. Under these operating conditions, the conversion of carbon monoxide was about 1270, and the yield of carbon dioxide (based on moles of CO fed) was about 2%. The reaction products were analyzed by on-line gas chromatography (Varma et al., 1986). The hydrocarbons C,-C,, were analyzed by using 10% OV-101 on a Chromosorb W H P column and flame ionization detector. The CO, CO,, Ar, and C, were separated by a Carbosieve S column and detected by a thermal conductivity detector. The detailed analysis of C1-C4 hydrocarbons, including separation of ethane/ethylene and propane/propylene, was accomplished by a Chromosorb 102 column. The C4 components were well separated by a Carbopack C/O.19% picric acid column. All the peak areas were measured by Hewlett-Packard electronic integrators. The calibration was carried out using various calibration mixtures and pure compounds. In order to ensure the reliability of the data, a few selected runs were repeated several times. This showed that the experimental results were reproducible within f 5 7'0.
Results and Discussion Hydrocarbon Product Distribution. The product distributions obtained using the cobalt-nickel-zirconia catalyst (FT catalyst) alone as well as combined with the HZSM-5 catalyst in a single reactor system have been presented earlier (Varma et al., 1985,1986). The reported alterations in the product spectrum due to the presence of HZSM-5, viz., the sharp reduction in C,,+ products, significant formation of aromatics, and increased isobutane selectivity in the C4 fraction, were also observed in the present investigation by using the dual-reactor system. Figure 2 gives the steady-state product distributions as a function of HZSM-5 reactor temperature in the range 250-400 "C for various HZSM-5/FT ratios. The selectivities to the C,, Cz, and C3 fractions usually increase with an increase in HZSM-5 temperature, except that the C, selectivity shows a slight minimum at HZSM-5/FT ratios of 2 and 4. The C4selectivity varies little with temperature for all three HZSM-5/FT ratios. The selectivity to the Cs+ fraction generally decreases with an increase in temperature, particularly in the temperature range 300-400 "C. However, it slightly increases in the range 250-300 "C for HZSM-5/FT ratios of 2 and 4. Figure 3 shows that the C5+(aliphatics) selectivity decreases with an increase in temperature of the HZSM-5
Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987 185 60 -
I
I
100 r-1100
I
HZSM-5/FT: 1
\
40-
50
50
c
-,
I
I
I
I
I
I
250 300 350 400 TEMP OF HZSM-5 REACTOR, OC
Figure 2. Hydrocarbon product distribution vs. temperature of the HZSM-5 reactor a t various HZSM-5/FT ratios.
a C;
300
52=520 250
350
400 0
TEMP. OF HZSM-5 REACTOR,OC
Figure 4. C22-/total C2 and CS2-/total C3 fractions in the hydrocarbon product vs. temperature of the HZSM-5 reactor a t various HZSM-5/FT ratios.
ALIPHATICS AROMATICS
0 TOTAL
i
70t
2
15-
I -I
2
4
I
45-
/
I
A
0
2
0
a
15-
I
1
I
I
HZSM-S/FT :1 HZSM-5/FT :2 OHZSM-5/FT :4
A
0
250
300
350
400
TEMP. OF HZSM-5 REACTOR,'C
Figure 3. Aromatic and aliphatic hydrocarbons in the hydrocarbon product vs. temperature of the HZSM-5 reactor a t various HZSM5 / F T ratios.
reactor. Derouane et al. (1978) and Chang et al. (1978) also reported that increasing the temperature in the range 300-400 "C leads to a decrease in the C5+(aliphatics) for methanol conversion over HZSM-5. The effect of variation of zeolite reactor temperature on aromatics selectivity is more pronounced a t higher HZSM-6/FT ratios. Nevertheless, for all the HZSM-5/FT ratios, the aromatics selectivity curve exhibits a maximum in the temperature range 275-300 "C. This is due to the increase in zeolite activity for oligomerization and aromatization reactions
0
250
300
350
400
TEMP: OF HZSM-5 REACTOR,OC
Figure 5. i-C,/total C4 fraction in the hydrocarbon product vs. temperature of the HZSM-5 reactor a t various HZSM-5/FT ratios.
with an increase in temperature up to 300 " C . Above this temperature, cracking reactions become prominent. Thus, increasing the temperature in the range 300-400 "C decreases the aromatics selectivity. Figure 4 shows that with an increase in the temperature of the HZSM-5 reactor, the ethylene selectivity in the C2 fraction increases. On the other hand, the propylene selectivity in the C3 fraction exhibits a broad minimum in the temperature range 275-300 O C and then slightly increases in the temperature range 300-400 "C. Figure 5 shows that the isobutane selectivity in the C, fraction also
186 Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987
2 3 HZSM-WFT RATIO
1
4
Figure 6. Variation of total aromatics selectivity with HZSM-S/FT ratio at various HZSM-5 reactor temperatures.
exhibits a maximum with a variation in temperature. An increase in the HZSM-5/FT ratio results in a decrease in olefin selectivity, particularly that of propylene (Figure 4) and an increase in isobutane selectivity (Figure 5). These effects are due to the increased time of contact with the zeolite catalyst and are in accordance with the reported alterations in product distribution induced by HZSM-5 (Varma et al., 1985). It is noted that the temperature range of 275-300 "C, which corresponds to a maximum in aromatics selectivity (Figure 3), also corresponds approximately to a minimum in C3 selectivity (Figure 4) and a maximum in isobutane selectivity (Figure 5). These results are indicative of the high reactivity of propylene (CSz-)toward the formation of aromatics, with simultaneous formation of isobutane. This is consistent with the reported distribution of hydrocarbons for syngas conversions over FT + HZSM-5 systems (Chang et al., 1979; Bruce et al., 1984; Varma et al., 1985). The effect of the HZSM-5/FT ratio on aromatics selectivity is emphasized in Figure 6. It is seen that increasing the HZSM-5/FT ratio causes an increase in the aromatics production when the temperature of the HZSM-5 reactor is maintained within the range 250-300 "C. This increase in aromatics selectivity is due to increased contact time of the reactants with the zeolite catalyst with a simultaneous decrease in selectivity to C5' aliphatics (Figure 3). This is indicative of the reaction sequences lower olefins C6+ (aliphatics) aromatics
-
-
A similar reaction path was suggested for methanol conversion to hydrocarbons (Chang, 1983). At higher reactor temperatures of 350 and 400 "C, the aromatics selectivity increases with an increase in the HZSM-5/FT ratio from 1 to 2 but slightly decreases when the HZSM-5/FT ratio is increased further from 2 to 4. This drop in aromatics selectivity is probably due to cracking reactions taking place under the conditions of high temperatures and long contact times with the zeolite. The effect of an increase in the HZSM-5/FT ratio on aromatics selectivity is most
Table I. Comparison of Dual-Reactor Results with Single-Stage (Follow Bed and Mixed Bed) Operations system single-reactor" mixed reactorb follow bed bed 4 4 4 4 HZSM-5/FT ratio 1.6 1.6 2.2 2.2 WHSV, h-1 250 250 250 250 FT reactor temp, "C 250 300 HZSM-5 reactor temp, "C 13.0 13.0 12.0 12.1 CO conversion, % hydrocarbon analysis, wt % aliphatics 16.3 18.1 16.8 16.0 c1 2.0 2.1 2.0 2.2 c2 9.8 8.0 9.0 10.6 c3 21.3 21.3 19.2 16.0 c4 23.0 25.0 23.3 9.7 CS+ aromatics 0.3 0.4 0.5 1.9 3.6 4.4 12.4 4.3 A, 8.9 10.9 15.1 10.1 A8 8.2 6.6 4.7 7.3 A9 4.4 5.6 7.3 11.4 Am+ 27.6 total aromatics, wt 70 25.5 29.1 45.5 50.6 53.0 55.2 50.5 total C6+ (incl. arom.) 20 45 27 49 C,2'/total C,, 90 5 16 14 23 C:-/total C3, ?& "Varma et al., 1985. *This study.
dramatic at 300 "C.These results suggest that a favorable combination of the two operating variables, viz., the reactor temperature and HZSM-5/FT ratio, is essential in order to achieve high aromatic selectivities. The production of aromatics was found to be maximum at a HZSM-5/FT ratio of 4 and a HZSM-5 reactor temperature of 300 "C. Under these favorable conditions, the aromatics selectivity was found to be as high as -45% (by weight) of the total hydrocarbons. Comparison of Results with Single-Stage Operation. Varma et al. (1985) have reported the performance of the same catalyst components evaluated in a single-stage operation in "follow bed" and "mixed bed" arrangements. In the follow bed arrangement, the synthesis gas mixture first passed through the FT catalyst followed by the bed of HZSM-5. The two catalyst beds were separated by quartz wool. In the mixed bed arrangement, the FT catalyst was physically mixed with HZSM-5, and the mixture was loaded into the reactor. The dual-reactor results of the present investigation are compared with these singlestage operations in Table I. It is seen that at the reactor temperatures of 250 "C, the total aromatics selectivity is nearly the same as in the follow bed, mixed bed, and the dual-reactor pperations. However, in dual-reactor operation, the aromatic selectivity is significantly increased at the HZSM-5 reactor temperature of 300 "C. This suggests that the process selectivity can be improved by proper choice of operating conditions. Thus, the present study of dual-reactor system demonstrates an improvement in gasoline yield and quality over that obtainable in singlestage operation. This has been made possible by operating the two catalyst systems at their individual optimum temperature levels. In single-reactor studies, higher temperatures (>280 "C) could not be used owing to the limitations of the FT catalyst. High temperatures in the FT catalyst caused formation of methane as the dominant product and eventually led to fast deterioration of the catalyst. In general, the C2-C4 olefin selectivity in the product obtained using the FT + HZSM-5 system has been reported to be lower compared to that obtained using the
Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987 187
FT catalyst alone (Varma et al., 1985). The drop in olefin selectivity is essentially due to the conversion of lower olefins to aromatics over HZSM-6. Over the FT catalyst alone at 250 "C (other conditions being same as used in dual-reactor experiments), the C;-/total Cz and C:-/total C3 ratios were found to be about 55% and 90%, respectively. A comparison of these values with those shown in Figure 4 indicates that the decline in olefin selectivity due to the presence of the HZSM-5 component is observed in the dual-reactor system as well, particularly a t lower HZSM-5 reactor temperatures and higher HZSM-5/FT ratios. The exception to this effect is the high ethylene selectivity at high temperatures (see Figure 4). As the temperature of the HZSM-5 reactor increases, the ratio C,Z-/total C2 increases markedly. For example, at a zeolite reactor temperature of 400 "C, and HZSM-5/FT = 1,the ratio C?-/total Cz attains a value of about 84%, which is significantly higher compared to about 55% obtained with FT catalyst alone. It is noted that the increase in C?'/ total Cz ratio is due to increased ethylene selectivity in total hydrocarbons, rather than due to a decrease in ethane selectivity in the hydrocarbon product. The increase in the C:-/total Cz ratio with an increase in temperature of the HZSM-5 component in the dualreactor system (Figure 4) is in contrast to the previously reported decrease in this ratio with an increase in temperature for the FT + HZSM-5 system in a single-stage operation (Varma et al., 1985, Table 8). In the single-stage system, the decrease in olefin selectivity, particularly that of propylene, with an increase in temperature within the range 250-280 "C was attributed to an increase in the activity of the HZSM-5 component, resulting in an increased aromatic selectivity, with a corresponding increase in the i-C,/total C4 ratio. In addition, methane formation increased with an increase in temperature, due to the FT component of the single-stage system. In the present work involving the dual-reactor system, the FT reactor temperature was fixed at 250 "C. Therefore, any slight variation in methane selectivity with a variation in zeolite reactor temperature or HZSM-5/FT ratio is due solely to the HZSM-5 component. An increase in temperature of the HZSM-5 reactor in the range 250-300 "C resulted in similar effects with regard to aromatic selectivity, i-C,/total C4 and C,2-/total C3 ratios as reported in the case of the single-stage operation system. In a single reactor system, the decrease in olefin selectivity, as observed for ethylene and propylene, can be attributed to the reactions hydrogenation olefins
FT catalyst
paraffins
aromatization
HZSM-5
olefins higher olefins, aromatics, paraffins Ethylene is known to be much more reactive for hydrogenation over the metal catalyst compared to propylene (Rautavouma and van der Baan, 1981). Over the HZSM-5 catalyst, however, propylene exhibits much greater reactivity compared to ethylene whose reactivity is very low (Anderson et al., 1980). This suggests that the observed decrease in ethylene selectivity with an increase in temperature in a single-reactor system is mainly due to the increased hydrogenation of ethylene over the FT component. In contrast, the decrease in propylene selectivity is mainly due to increased aromatization over the HZSM-5 component. In a dual-reactor system, the temperature of the FT reactor remains constant, while that of the HZSM-5 reactor is increased. As a result, the hydrogenation of ethylene over the FT component remains unaffected. Also,
unlike propylene selectivity, the ethylene selectivity does not drop significantly under the increased aromatization activity of the zeolite component at higher temperature. This explains the difference in behavior of the Cz2-/total Cz ratio, as observed between the one-stage and two-stage operations. The main reactions taking place over the HZSM-5 catalyst may be summarized as oxygenates heavy hydrocarbons lower olefins
-HZO
lower olefins
cracking
oligomerization
lower hydrocarbons higher olefins
aromatization
higher olefins aromatics + paraffins In the two-stage process, the increase in olefin selectivity, particularly that of ethylene under high operating temperatures can be attributed to (i) a favorable increase in the rate of olefin formation reaction from oxygenates compared to oligomerization and aromatization reactions and/or (ii) an increased rate of cracking reactions. However, the cracking reactions over zeolite catalysts have been found to produce only trace amounts of ethylene (Abbot and Wojciechowski, 1985; Kobolakis and Wojciechowski, 1985). It is concluded, therefore, that ethylene is formed mainly from the oxygenates present in the FT synthesis product. Lower olefins are known to be intermediates in methanol to gasoline reactions (Chang, 1983). The oxygenates are converted to olefins over HZSM-5 according to the reactions 2CH30H + CH3OCH3 + HzO CH3OCH3 + CzH4 + HzO Similarly, CZHbOH + CzH4 + HzO Ethylene has been confirmed as the initial olefin formed from methanol conversion (Chang, 1984; Chang et al., 1984b). A part of ethylene undergoes homologation and oligomerization to be subsequently converted to aromatics. It has also been reported (Chang, 1984) that the selectivity to light olefins from methanol is favored a t high temperatures. This is due to the differences in the activation energies of olefin formation and aromatization reactions. For the simplified reaction scheme
kl
k2
oxygenates olefins aromatics + paraffins the rate constant kl increases with an increase in temperature, while the aromatization reaction constant k, is relatively insensitive to temperature variation in the range 400-500 "C (Chang et al., 1984a). Our results showing an increase in olefin selectivity with a corresponding decrease in selectivity of aromatics at higher temperatures of the HZSM-5 reactor are in accordance with the above reaction scheme for conversion of oxygenates. It appears that high temperatures in our work lead to increased selectivities to both ethylene and propylene. However, the propylene, being much more reactive than ethylene (Anderson et al., 1979; Garwood, 1983), readily takes part in aromatization reaction. This explains the observed large increase in the ethylene selectivity but not the propylene selectivity with an increase in temperature. -+
Summary and Conclusions Hydrogenation of carbon monoxide was investigated using a cobalt-nickel-zirconia catalyst combined with a zeolite HZSM-5 catalyst in a dual-reactor system. Under the steady-state conditions, an increase in the HZSM-5/FT
I n d . Eng. Chem. Res. 1987,26, 188-193
188
ratio from 1 to 4 generally results in an increase in aromatics production, a decrease in olefinicity of the C2 and C3 fractions, and an increase in isobutane content in the C, fraction at HZSM-5 reactor temperatures in the range of 250-300 "C. However, for HZSM-5 reactor temperatures above 300 "C, cracking reactions become prominent, resulting in a decline in selectivities to aromatics and total C5+ and a shift in the product distribution toward lower hydrocarbons. A higher temperature of the zeolite reactor also results in an increase in olefin selectivity, particularly that of ethylene in the C2 fraction, and a decrease in isobutane in the C, fraction. A comparison with earlier studies using single-stage operation suggests that further improvement in C5+ product selectivity is possible in a dual-reactor system by controlling the individual reactor temperatures within narrow limits. Acknowledgment We express our gratitude to Dr. K. Jothimurugesan for careful review of the preliminary manuscript and several valuable suggestions. The expert technical assistance of T. Olauson is gratefully acknowledged. Funding for this project was provided by DSS Contract ISU-82-00308. Registry No. CO, 630-08-0; Ni, 7440-02-0; Co, 7440-48-4.
Literature Cited Abbot, J.; Wojciechowski, B. W. Can. J . Chem. Eng. 1985,63,462. Anderson, R. B. The Fischer-Tropsch Synthesis; Academic: Orlando, FL., 1984. Anderson, J. R.; Foger, K.; Mole, T.; Rajadhyaksha, R. A.; Sanders, J. V. J . Catal. 1979,58,114. Anderson, J. R.; Mole, T.; Christov, V. J . Catal. 1980,61, 477. Argauer, R. J.; Landolt, G. R. US Patent 3 702 886, 1972. Biloen, P.; Sachtler, W. M. H. Adu. Catal. 1981,30,165. Bruce, L. A,; Hope, G. J.; Methews, J. F. Appl. Catal. 1983,8,349.
Bruce, L. A.; Hope. G. J.; Mathews, J. F. Appl. Catal. 1984,9,351. Bruce, L.; Mathews, J. F. Appl. Catal. 1982,4, 353. Caesar, P. D.; Brennan, J. A.; Garwood, W. E.; Ciric, J. J . Catal. 1979,56,274. Chang, C. D. Catal. Reu.-Sci. Eng. 1983,25,1. Chang, C. D. Catal. Reu.-Sci. Eng. 1984,26,323. Chang, C. D.; Chu, C. T.-W.; Socha, R. F. J . Catal. 1984a,86,289. Chang, C. D.; Kuo, J. C. W.; Lang, W. H.; Jacob, S.M.; Wise, J. J., Silvestri, A. J. Ind. Eng. Chem. Process Des. Deu. 1978,17,255. Chang, C. D.; Lang, W. H.; Silvestri, A. J. J . Catal. 1979,56,268. Chang, C. D.; Miale, J. N.; Socha, R. F. J . Catal. 1984b,90,84. Derouane, E. G.; Nagy, J. B.; Dejaifve, P.; van Hooff, J. H. C.; Spekman, B. P.; Vedrine, J. C.; Naccache, C. J . Catal. 1978,53, 40. Diffenbach, R. A.; Schehl, R. R.; Fauth, D. J. Presented at the Proceedings of the International Conference on Coal Science, Center for Conference Management, Pittsburgh, PA, 1983; p 240. Dry, M. E. In Catalysis-Science and Technology; Anderson, J. R., Boudart, M., Eds.; Springer-Verlag: Berlin, 1981; Vol. 1, Chapter IV, p 159. Dry, M. E. Hydrocarbon Process. 1982,61(8), 121. Garwood, W. E. ACS Symp. Ser. 1983,218, 383. Haag, W. 0.; Huang, T. J. US Patent 4 279830, 1981. Huang, T. J.; Haag, W. 0. ACS Symp. Ser. 1981,152,307. Kobolakis, I.; Wojciechowski, B. W. Can. J . Chem. Eng. 1985,63, 269. Plank, C. J.; Rosinski, J.; Rubin, M. K. US Patent 4 175 114, 1979. Rautavuoma, A. 0. I.; van der Baan, H. S. Appl. Catal. 1981,1,247. Shamsi, A.; Rao, V. U. S.; Gormley, R. J.; Obermyer, R. T.; Schehl, R. R.; Stencel, J. M. Ind. Eng. Chem. Prod. Res. Deu. 1984,23, 513. Storch, H.; Golumbic, N.; Anderson, R. B. The Fischer-Tropsch and Related Syntheses; Wiley: New York, 1951. Vannice, M. A. Catal. Rev.-Sci. Eng. 1976,14(2), 153. Varma, R. L.; Bakhshi, N. N.; Mathews, J. F.; Ng, S. H. Can. J . Chem. Eng. 1985,63,612. Varma, R. L.; Jothimurugesan, K.; Bakhshi, N. N.; Mathews, J. F.; Ng. S.H. Can. J . Chem. Eng. 1986,64, 141. Received for review January 24, 1986 Accepted August 7, 1986
Prediction of Cetane Number by Group Additivity and Carbon-13 Nuclear Magnetic Resonance Timothy H. DeFries* and Doren Indritz* Exxon Research and Engineering Company, Products Research Division, Linden, New Jersey 07036
Rodney V. Kastrup Exxon Research a n d Engineering Company, Corporate Analytical Science Laboratory, Annandale, New Jersey 08801
Cetane number is a measure of ignition quality, specifically ignition delay, of diesel fuel. I t is an engine measure of a kinetic phenomena. While it is typically inappropriate to use a thermodynamic measure, such as molecular structure, to predict kinetic behavior, molecular structure does correlate with cetane number. In fact, we use a group additivity approach to dissect structures and predict cetane number. For pure compounds a simple group counting scheme is used to predict the cetane numbers of normal and branched paraffins and singly substituted alkylbenzenes. T o extend the counting scheme to hydrocarbon mixtures, carbon-13 nuclear magnetic resonance (13C NMR) is used. 13C NMR is sensitive to the local environment, up to three to four carbon atoms away, of each carbon atom. Intramolecular reactions that are important for ignition kinetics imply that molecular fragments of three or four carbon atoms must be considered. We show that group concentrations derived from 13C NMR spectra are useful in predicting the cetane number of hydrocarbon mixtures. Cetane number is a specification for diesel fuel (ASTM, 1980). For more details on what cetane number is and why there is current interest in it, see the paper by Indritz (1985). This study of the effect of molecular structure on
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cetane number uses carbon-13 nuclear magnetic resonance (13C NMR) as a measure of the structural groups present in pure compounds and in 93 hydrocarbon mixtures. Myers et al. (1975), Drugarin and Andru (1979), and Ohuchi et al. (1982) have correlated NMR data with octane number and cetane number: we demonstrate that this 62 1987 American Chemical Society
