Hydrocracking of a Plastics Pyrolysis Gas Oil to Naphtha

significant quantities of light naphtha range liquids, there is also a residual ... f naphtha f gas) was applied to simulate the upgrading of the gas ...
31 downloads 0 Views 191KB Size
Download PDF
586

Energy & Fuels 1997, 11, 586-592

Hydrocracking of a Plastics Pyrolysis Gas Oil to Naphtha H. S. Joo and James A. Guin* Auburn University, Chemical Engineering Department, Auburn, Alabama 36849 Received September 10, 1996X

Pyrolysis of waste plastics is one method currently being investigated as an alternative to landfill disposal of this increasingly large waste stream. Although the pyrolysis process produces significant quantities of light naphtha range liquids, there is also a residual fraction that may be considered as a potential feedstock for upgrading. The objective of our research was to investigate the upgrading potential of the gas oil fraction of a typical plastics pyrolysis liquid using catalytic hydrocracking. In this study, the residual fraction of a liquid produced by the pyrolysis of plastics, containing about 70%, +205 °C gas oil, was subjected to batch hydrocracking reactions. The reactions used a commercial NiMo supported on zeolite-alumina catalyst and gave good conversion to naphtha. A simplified three-lump sequential reaction pathway (gas oil f naphtha f gas) was applied to simulate the upgrading of the gas oil fraction to naphtha and gases. The kinetic model considered the variation in H2 partial pressure in the hydrocracking rate expression and thereby was able to account for conversion and yield differences between large and small batch reactors. The model simulated the experimental data well including experimentally observed trends such as the maximum in naphtha yield at lower temperatures.

1. Introduction The amount of waste plastics requiring landfill disposal has been rapidly increasing in recent years. Currently, around 20% of the volume and 8% of the weight of all municipal solid waste in the U.S. is made up of waste plastics.1 Of the approximately 80 billion pounds of plastics currently produced in the United States, most eventually ends up in landfills, with only 2-3% being recycled.2 In contrast to paper and garbage wastes, most plastics are not readily biodegradable and will remain in the landfill for indeterminate periods. The ever increasing costs of landfill disposal coupled with a significant public resistance to the creation of new waste landfills has led to increased efforts toward finding economically feasible and environmentally acceptable means of recycling these materials. Disposal of these waste plastics by direct incineration would lead to increased greenhouse gas emissions, primarily carbon dioxide, and particulate pollutants, while direct mechanical recycling of plastics is limited by the fact that reconstituted objects generally possess a downgraded appearance in comparison with virgin plastics. Moreover, mechanical recycling produces objects that will eventually require discarding because of brittleness caused by oxidation and exposure to UV light. Several alternative approaches requiring more processing than simple mechanical recycling or incineration have been proposed in order to deal more effectively with the increasing volumes of waste plastics. Some processes seek to effect complete depolymerization to the monomer, and there is some success with these processes, especially with polyethylene terephthalate (PET). Most Abstract published in Advance ACS Abstracts, March 15, 1997. (1) Reisch, Marc S. Chem. Eng. News 1995, May 22, 30. (2) Friedman, S. PETC Review; Pittsburgh Energy Technology Center, U. S. Department of Energy, Vol. 12, Winter; PETC: Pittsburgh, PA, 1996; p 14. X

S0887-0624(96)00151-X CCC: $14.00

polymers, however, will give a mixture of products that must undergo extensive separation to recover a monomer stream, and the depolymerization of postconsumer commingled (mixed) waste plastics to pure monomers does not appear feasible. Thus, other processes for dealing with commingled waste plastics have been proposed including pyrolysis,3-8 gasification,9 and catalytic coliquifaction.10 In a pyrolysis type process, such as the Conrad process,5,6 shredded mixed plastics are heated in the absence of oxygen and depolymerized back into liquids (70-80%) and gases (5-10%). The gases are burned in the process to provide heat needed for the depolymerization and pyrolysis. Although the pyrolysis process produces substantial quantities of light naphtha range (-205 °C) liquids, significant quantities of heavier gas oil range liquids are also produced. The heavier gas oil fraction is potentially valuable as upgrading feedstock for naphtha range material with an end use in the transportation fuel pool. The objective of this research was to investigate the upgrading potential of the gas oil fraction of a typical plastics pyrolysis liquid by catalytic hydrocracking. In general, the hydrocracking technology is attractive from an environmental standpoint because the hydrogenation function of the process inherently yields products with lower concentrations of aromatic components and heteroatoms in comparison with catalytic cracking without hydrogen. Since commercially developed in 1927, hy(3) Marsh, J. A.; Cha, C. Y.; Guffey, F. D. Chem. Eng. Commun. 1994, 129, 69. (4) Williams, S. B.; Taylor, D. T. Fuel. 1990, 69 (12), 1474. (5) Meszaros, M. W. Conrad Advanced Recycling Project. Presented at Recycle ’94, Davos, Switzerland, March, 1994. (6) Virgin, B. Seattle Post-Intelligencer 1993, July 1. (7) Kaminsky, W. Adv. Polym. Technol. 1995, 14 (4), 337. (8) Kaminsky, W.; Schlesselmann, B.; Simon, C. J. Anal. Appl. Pyrolysis 1995, 32, 19. (9) Hydrocarbon Process. 1984, 63 (4), 51. (10) Taghiei, M. M.; Feng, Z.; Huggins, F. E.; Huffman, G. P. Energy Fuels 1994, 8 (6), 1228.

© 1997 American Chemical Society

Hydrocracking to Naphtha

drocracking processes have been applied to variety of feedstocks ranging from light naphtha to very heavy hydrocarbon stocks such as tar sands, shale oil, and deasphalted vacuum residue.11,12 The process generally produces liquids with increased overall hydrogen-tocarbon ratio and decreased average relative molecular mass and boiling range. For example, by use of a twostage upgrading process consisting of hydrotreating followed by hydrocracking, coal-derived middle distillates with an end point of about 400 °C have been successfully hydrocracked to produce 80% gasoline fraction13 and jet fuel with attractive properties for special uses.14 On the other hand, there has been little study of the hydrocracking of waste plastic-derived pyrolysis liquids. In one study involving upgrading of waste plastics liquid, Songip et al.15,16 examined the effects of reaction conditions on the activity and selectivity of REY zeolite catalysts in the catalytic cracking of heavy oil from waste plastics. A residual oil distilled from the product of polyethylene pyrolysis and composed mostly of linear paraffins was used as the feed. The gasoline produced in the upgrading was found to be more environmentally favorable (more isoparaffins and fewer aromatics) than a typical commercial gasoline. To our knowledge, our work is the first to apply hydrocracking, as opposed to catalytic cracking, to the upgrading of a plastics pyrolysis liquid. Hydrocracking of the waste plastics-derived liquid is expected to proceed through a network of complex reactions involving a large number of components. Because it is usually not feasible to develop process models that describe the exact kinetics of each of the individual species, a simplification is made whereby the reaction steps and products are lumped into groups as simple pseudocomponents. For example, an overall kinetic model for catalytic cracking of gas oil without hydrogen was developed by using a three-lump (gas oil, gasoline, and other products such as C1-C4, coke) parallel-consecutive kinetic scheme.17 In this study using a simplified lumped model, Volts et al.17 investigated the catalytic cracking of Mid-Continent gas oil and reported that activation energies for gas oil-to-gasoline and gasoline-to-other products including gas cracking were 31.1 and 50.0 kJ/mol, respectively. In the simplified model, the gas oil cracking was taken to be approximately second order and the gasoline cracking reaction to be first order. In a similar manner, Yui and Sanford18 have applied lumped kinetics schemes involving parallel, consecutive, and combined model pathways for the mild hydrocracking of bitumen-derived coker and hydrocracker heavy gas oils. Their kinetic model used modified first-order kinetics, with power law terms for liquid hourly space velocity (LHSV) and hydrogen partial pressure. Songip et al.15 studied kinetics for catalytic cracking of heavy oil prepared by waste plastics (11) Scott, J. W.; Bridge, A. G. Adv. Chem. Ser., 103, 113 1971. (12) Voorhies, A. and Smith, W. M., Adv. Chem. Ser. 1964, 8, 169. (13) Gutberlet, L. C.; Bertolacini, R. J.; Kukes, S. G. Energy Fuels 1994, 8 (1), 227. (14) Yan, T. Y. Ind. Eng. Chem. Res. 1983, 22 (1), 154. (15) Songip, A. R.; Masuda, T.; Kuwahara, H.; Hashimoto, K. Energy Fuels 1994, 8, 131. (16) Songip, A. R.; Masuda, T.; Kuwahara, H.; Hashimoto, K. Energy Fuels 1994, 8, 136. (17) Volts, S. E.; Nace, D. M.; Jacob, S. M.; Weekman, V. W. Ind. Eng. Chem. Process Des. Dev. 1972, 11 (2), 261. (18) Yui , S. M.; Sanford, E. C. Ind. Eng. Chem. Res. 1989, 28, 1278.

Energy & Fuels, Vol. 11, No. 3, 1997 587

Figure 1. GC simulated distillation chromatograms of original, distillate, and residual plastics pyrolysis liquids: (a) benzene; (b) toluene; (c) ethylbenzene; (d) styrene; (e) naphthalene.

using a simplified lumped model, which was found to be in good agreement with the experimental data. In the present work, we have investigated the catalytic hydrocracking of a residual gas oil range fraction from a plastics pyrolysis process using a commercial Ni/ Mo supported on a zeolite-alumina catalyst. Data obtained using batch tubing bomb microreactors (TBMRs) were used to develop a kinetic model for hydrocracking of the gas oil fraction to produce naphtha and gases. The effect of temperature on naphtha selectivity was examined, and the variation in H2 partial pressure in the batch reactor was considered. Because of its importance as a feedstock for petrochemicals and transportation fuel, naphtha yield was emphasized in this work. 2. Experimental Section Materials. The liquid raw material for this study was a plastics-derived liquid from the Conrad pyrolysis process and was obtained through the courtesy of Conrad Industries, Inc.5,6 This liquid was produced by the pyrolysis of a mixed plastics feed consisting of high-density polyethylene, polypropylene, and polystyrene.5 Owing to the fairly high naphtha content (62.3%), the original pyrolysis liquid was separated into two fractions by distillation at 90 °C and 36 mbar. After distillation, approximately 2/3 and 1/3 of the original plastic pyrolysis liquid were recovered as residual liquid and distillate fractions, respectively, disregarding a small loss of noncondensable vapors. Boiling point distributions of the original liquid and the two fractions were examined by GC simulated distillation. The resulting chromatograms for the three liquids (original, distillate, and gas oil) are shown in Figure 1. As shown in Figure 1, several pure aromatic compounds from the plastics

588 Energy & Fuels, Vol. 11, No. 3, 1997

Joo and Guin

Figure 2. Boiling point distributions of plastics pyrolysis liquids and commercial gasoline. Table 1. Properties of Plastics Pyrolysis Liquids wt % boiling point range (°C)

original liquid

distillate fraction

residual fraction

naphtha (205) average boiling point (°C)

62.3 37.7 199.6

98.9 1.1 131.6

28.7 71.3 272.3

wt % major compounds

original liquid

distillate fraction

residual fraction

benzene toluene ethylbenzene styrene naphthalene total density (g/mL) at 25 °C

6.1 14.0 8.7 17.1 2.6 48.5 0.863

4.3 24.0 15.2 25.6 0.3 69.4 0.842

0.2 0.5 2.7 9.4 5.2 18.0 0.889

elemental analysis (wt %)

C H N S O (by difference) H/C atomic ratio

original liquid

distillate fraction

residual fraction

89.17 10.58 0.0243 0.11 0.14 1.42

87.75 9.75 0.0071 0.05 2.45 1.32

88.63 10.84 0.0290 0.14 0.39 1.47

pyrolysis were identified by GC-MS as benzene, toluene, ethylbenzene, styrene, and naphthalene. These aromatic compounds result from the pyrolysis of the polystyrene in the feed plastics. Zmierczak et al.19 also have reported these compounds as typical products in the catalytic liquefaction of polystyrene. The dotted line in Figure 1 at a retention time of 6.6 min is equivalent to a boiling point of 205 °C, taken here as the end point of the naphtha cut. Figure 2 shows the complete boiling point distributions of the original liquid, distillate, and residual fractions, as well as a commercial gasoline for comparison. The distillate fraction is similar to commercial gasoline in boiling point distribution. However, the residual fraction will require upgrading to the naphtha range. This residual fraction is the feedstock for the upgrading study described in the remainder of this study. Some quantitative properties of the original pyrolysis liquid and the distillate and residual fractions are listed in Table 1, which shows that the original pyrolysis liquid was separated (19) Zmierczak, W.; Xiao, X.; Shabtai, J. Fuel Process Technol. 1996, 47, 177.

by distillation into a light fraction containing 98.9% naphtha and a residual fraction containing 71.3% gas oil. The H/C ratio indicates that all three liquids contain substantial amounts of aromatic and/or olefinic components, as indicated by the major compounds identified, all of which are aromatic. The residual fraction from the plastics liquid contains fewer heteroatoms compared to a traditional gas oil, which typically contains around 1-5 wt % S and 500-3000 ppm N.18,20-21 The high oxygen content of the distillate fraction (2.45%, calculated by difference) was consistent with that obtained by direct elemental analysis (2.93%). We speculate that this high oxygen content might be due to oxidation of olefins during or following the distillation process. Catalyst. The catalyst (Akzo KC 2600: 2.6% Ni, 7.0% Mo on zeolite-alumina) was crushed and sieved to -100 mesh before use. Presulfidation was performed at 300 °C for 2 h with dimethyl disulfide in a separate TBMR charged with 69 atm cold hydrogen pressure. The quantity of sulfur added was based upon the catalyst metals content by use of 250% excess of the stoichiometric amount needed for bulk sulfidation to Ni3S2 and MoS2. Reaction Procedures and Product Analysis. All reactions were performed in 20 and 50 mL 316 stainless steel tubing bomb microreactors (TBMRs), which were agitated in a fluidized sand bath described previously.22 Reactions were carried out with a 3 g residual fraction of plastics liquid at varying catalyst loadings (0.1-1.8 g), reaction times (0.5-1.5 h), and temperatures (370, 400, 430 °C) with 69 atm cold hydrogen pressure. After quenching in water, the hydrogen remaining after reaction was determined by collection of the product gas from the TBMR into a gas sampling bag followed by GC analysis with a 30 ft × 1/8 in. stainless steel column packed with 23% SP-1700 on 80/100 Chromosorb P AW (Supelco). This GC column was unable to resolve the hydrogen from the methane peak. However, the detector response for hydrogen was determined to be 26 times that of methane on a molar basis. Thus, neglect of the methane would result in only a 5% error in the amount of hydrogen determined, even for an equimolar mixture. On the basis of these considerations, the methane peak was neglected in the hydrogen analysis. After analysis of residual hydrogen, the amount of gas produced from hydrocracking reactions was calculated by subtraction of the remaining mass of hydrogen left after reaction from the total mass of gases. The liquid products from the reaction were analyzed by GC according to ASTM D2887 wherein the relation of GC retention time to boiling point is calibrated by a standard n-paraffin mixture and the fractions of naphtha (205 °C) were determined. From these data the yields of gas, naphtha, and gas oil conversion were determined. Elemental analysis (Table 1) of liquids for carbon, hydrogen, sulfur, and oxygen was performed by Huffman Labs, Golden, CO. Nitrogen was analyzed using a Dohrmann DN-1000 automated N analyzer. Major compounds in the liquid were identified by GC-MS and standard additions.

3. Results and Discussion The objective of this work was to examine the feasibility and kinetics of upgrading the gas oil fraction of the plastic pyrolysis liquid (Table 1) to the naphtha range in batch hydrocracking reactions. Figure 3 shows the complete boiling range distributions of the liquid products from four reactions at different catalyst loadings, all at 430 °C, 0.5 h reaction time, and 3 g liquid (20) Larocca, M.; Ng, S.; Lasa, H. Ind. Eng. Chem. Res. 1990, 29, 171. (21) Oballa, M. C., Shih, S. S., Ed. Catalytic Hydroprocessing of Petroleum and Distillates; Marcel Dekker: New York, 1991; pp 379, 408. (22) Joo, H. S.; Guin, J. A. Fuel Process Technol. 1996, 49, 137.

Hydrocracking to Naphtha

Energy & Fuels, Vol. 11, No. 3, 1997 589

Figure 3. Boiling point distributions of upgraded plastics pyrolysis liquid with several catalyst loadings at 430 °C and 0.5 h compared with commercial gasoline: (a) no catalyst; (b) 0.1 g; (c) 0.7 g; (d) 1.0 g.

Figure 4. Hydrogen partial pressure as a function of space time θ at three temperatures in the batch hydrocracking reaction: (open symbols) 20 mL TBMR; (solid symbols) 50 mL TBMR.

feed. The original residual fraction (feed) and a commercial gasoline are also shown for comparison. At 430 °C, which is the highest temperature used in the study, there is some hydrocracking due to thermal reactions alone as shown by curve a in Figure 3. The degree of hydrocracking shown by curve b is about the same as the thermal reaction due to the very low catalyst loading. Curves c and d show that increased upgrading results from the higher catalyst loadings. These two curves show products that meet or exceed the boiling point range of the commercial gasoline. Thermal reaction alone produced 2.5% gas at 430 °C and 0.5 h compared to 13.3% gas production with 1 g catalyst under the same conditions. Thus, about 20% of the gas yield with 1 g catalyst is attributed to thermal reaction. The thermal contribution would be less significant at lower temperatures. From Figure 3, it is obvious that the residual fraction of the pyrolysis liquid can be upgraded to the boiling point range of the commercial gasoline, or even lower if desired. Of course, this does not mean that the upgraded product meets all other specifications, e.g., octane rating, etc. for commercial gasoline. Hydrocracking Kinetics. To quantify the upgrading process in more detail as a function of temperature and space time, a kinetics analysis of the data was performed. For the kinetic modeling of the upgrading process, the complete boiling range distribution of reaction products such as shown in Figures 2 and 3 was lumped into three cuts; gas, naphtha (-205 °C), and gas oil (+205 °C). Owing to the complexity of the reactions involved, the following simple sequential lumped reaction pathway was assumed for this kinetic study:

time, respectively. The space time θ is equivalent to the reciprocal of the weight hourly space velocity (WHSV) commonly used in continuous flow reactions.

k1

k2

gas oil + H2 98 naphtha + H2 98 gases On the basis of this sequential reaction pathway, a kinetic model was developed as in eqs 1 and 2 given below where yA, yB, and PH2 stand for the weight fractions of gas oil (A) and naphtha (B) and hydrogen partial pressure in the gas phase, respectively. The space time was defined as θ ) Wcatt/Wliq where Wcat, Wliq, and t refer to the weight of catalyst and liquid feed, and

dyA ) -k1yAmPH2n dθ

(1)

dyB ) k1yAmPH2n - k2yBmPH2n dθ

(2)

For the simple pseudo-first-order case with m ) 1 and n ) 0, eq 1 can be easily integrated to give the wellknown first-order dependence for gas oil conversion. When this simple first-order model was applied to the data, we found that most sets of data for both the 20 and 50 mL TBMR experiments at 370, 400, and 430 °C fit well. However, the rate constants (slopes of the firstorder plots) were different for the 20 and 50 mL TBMRs, with the rate constants for the 50 mL reactor being higher than the 20 mL reactor at all temperatures. These results indicated a dependence of the pseudo-firstorder rate constant on hydrogen pressure, with the large TBMR rate constant exceeding the small TBMR value presumably because of the greater H2 availability in the larger TBMR. Since the TBMR is a batch reactor, hydrogen partial pressure decreases as reaction proceeds as shown in Figure 4 at three different temperatures in the two TBMRs. The hydrogen partial pressure at reaction temperature in Figure 4 was calculated using the ideal gas law, from the residual moles of hydrogen measured following the reaction. Because the liquid charges were the same in the two TBMRs, there is a greater hydrogen availability and correspondingly slower decline in hydrogen partial pressure in the large TBMR. At 370 °C the hydrogen partial pressure levels out at around 20 atm in the 20 mL TBMR and at 55 atm in the 50 mL TBMR. Since the kinetic rate constant should be independent of the volume of the reactor, the kinetic model must account for the fact that the rate of hydrogen pressure decrease in the 50 mL TBMR is slower than in the 20 mL TBMR, as shown in Figure 4. To model this situation, it was necessary to consider the solution of eqs 1 and 2 for various values

590 Energy & Fuels, Vol. 11, No. 3, 1997

Joo and Guin

Figure 5. Plot of hydrogren partial pressure vs gas oil weight fraction according to eq 4 at three temperatures: (open symbols) 20 mL TBMR; (solid symbols) 50 mL TBMR.

Figure 6. Plot based on eq 3 at three temperatures with m ) 0.5 and n ) 1: (open symbols) 20 mL TBMR; (solid symbols) 50 mL TBMR; (×) excluded points.

of m and n. Equation 1 can be formally integrated to give

f(yA) )

∫yy

A0

A

1 dyA ) k1θ PH2n

yA

m

(3)

To complete the integration of eq 3, the hydrogen partial pressure PH2 was obtained as a function of the gas oil weight fraction yA from the experimental data shown in Figure 5 by fitting the empirical equation:

ln PH2 ) a + byA

(4)

where a and b are constants. As shown in Figure 5, the decline in hydrogen partial pressure is satisfactorily fitted by eq 4. By use of eq 4 for PH2 together with numerical evaluation of the integral in eq 3, the rate constant k1 (slope of f(yA) vs θ in eq 3) can be obtained from the corresponding plots for gas oil conversion as shown in Figure 6 for both the 20 and 50 mL TBMR experiments. For linear plots such as in Figure 6, the square of the correlation coefficient R2, which ranges from 0 to 1 (1 ) perfect fit), is a measure of the goodness

Figure 7. Arrhenius plot for hydrocracking rate constants. Table 2. Kinetic Rate Constants and Activation Energies temp (°C)

k1 (atm-1 h-1)

k2 (atm-1 h-1)

370 400 430 Ea (kJ/mol)

0.027 0.053 0.092 76.9

0.0030 0.0070 0.018 112.1

of the correlation. The R2 values for the pseudo-firstorder model with m ) 1 and n ) 0, e.g., 0.62 at 370 °C, 0.37 at 400 °C, and 0.55 at 430 °C, were very low in comparison with the R2 values for optimal m and n values. At all temperatures, k1 is optimized (highest R2 values are 0.95 at 370 °C, 0.96 at 370 °C, and 0.98 at 370 °C) for m ) 0.5 and n ) 1, and the resulting plot is shown in Figure 6. In this figure a few outlying data points as indicated were excluded from the fitting because of significant deviation from the trend of the other points. However, for completeness, these excluded points are still indicated on the plot. An idea of the reproducibility of the data in Figure 6 can be gained from the fact that the average standard deviation of f(yA) is 0.0019 atm-1. As shown in Figure 6, accounting for the variation in H2 partial pressure allows data for both the 20 and 50 mL TBMRs to lie on the same line, thus resulting in the same rate constant k1. After the value of k1 was obtained, the corresponding value of k2 with the same m and n was obtained through trial-and-error solution of the differential equations with the best fit parameters being given in Table 2. After determination of kinetic rate constants at three different temperatures, Arrhenius plots were constructed as depicted in Figure 7, which yielded activation energies for k1 and k2 of 76.9 and 112.1 kJ/mol, respectively, as shown in Table 2. These activation energies compare quite well with literature values.18,23-24 The energy barrier for naphthato-gas cracking is higher than that of gas oil-to-naphtha cracking. Therefore, the selectivity for cracking to gas increases as the temperature increases. Simulated product distributions (curves) using the k1 and k2 values in Table 2 are compared with the experimental data (symbols) in Figures 8 and 9. Figure 8 shows the product distributions from experimental and simulated data in the 20 mL TBMR. The model (23) Campelo, J. M.; Lafont, F.; Marinas, J. M. J. Catal. 1995, 156, 11. (24) Benito, A. M.; Martinez, M. T. Energy Fuels 1996, 10, 1235.

Hydrocracking to Naphtha

Figure 8. Comparison of model (curves) and experimental (symbols) data in 20 mL TBMR.

fits the data points fairly well up to approximately 65% gas oil conversion at 370 °C and up to 85% gas oil conversion at both 400 and 430 °C. Since the gas oil weight fraction levels off above these values, the possibility of catalyst deactivation was considered. This was done by using different amounts of catalyst at the same space time, θ. The gas oil conversion was almost the same in these experiments (within 6%). In another experiment, a reactor was recharged with H2 using the “spent catalyst”, and good activity (additional gas oil conversion) was observed. These experiments suggest little or no catalyst deactivation. The exact reason for the difference between model and experimental data at higher gas oil conversion is not clear, but since the decreasing H2 pressure is accounted for in the model, the departure may suggest a shift in the reaction model from the simple consecutive model that was used in the fitting. A kinetic model with more steps, e.g., a combined series-parallel model having more lumps and more parameters, could be used to fit the data better, although we have used only a simple two-parameter series model for simplicity. Figure 9 compares the experimental product distributions with the model simulation in the 50 mL TBMR. In the 50 mL TBMR, more than 90% conversion of gas oil at all three temperatures was obtained in contrast to the results in the 20 mL TBMR (Figure 8). In this case the experi-

Energy & Fuels, Vol. 11, No. 3, 1997 591

Figure 9. Comparison of model (curves) and experimental (symbols) data in 50 mL TBMR.

Figure 10. Naphtha weight fraction as a function of gas oil conversion in the 50 mL TBMR model: (‚‚‚) 370 °C; (-) 400 °C; (- - -) 430 °C.

mental data are fairly well fitted by the model except for a few points at high gas oil conversion. The effect of temperature on naphtha yield as a function of gas oil conversion for both the model and the experimental data from the 50 mL TBMR is shown in Figure 10. Although the agreement is not exact, a maximum in the naphtha yield is observed in both the

592 Energy & Fuels, Vol. 11, No. 3, 1997

Joo and Guin

Table 3. Comparison of Observed and Model Values for Maximum Naphtha Yield in 50 mL TBMR temp (°C) 370 400 430

θmax (h) observed model 0.626 0.200 0.134

0.701 0.298 0.158

gas oil conversion (%) observed model 97.4 90.3 91.8

98.4 98.1 96.2

yB (wt % naphtha) observed model 83.1 79.9 74.0

85.6 83.2 76.6

model and the experimental data. Both model and experimental data indicate that the selectivity for naphtha is better at the lower temperature of 370 °C because of the lower activation energy for gas oil-tonaphtha cracking compared to that of naphtha-to-gas cracking as discussed before. A numerical comparison of experimental and model values of gas oil conversion, space time θ, and the weight fractions of naphtha at the point of maximum naphtha yield are compared in Table 3. As can be seen, the values are not in exact agreement because of the differences at high gas oil conversion. However, the trend in model values and experimental data are similar in that higher maximum naphtha yield is obtained with decreasing temperature because of the lower gas production rate. There is, however, an economic trade-off to be considered in that decreasing the temperature in order to obtain higher maximum naphtha yield will in turn require a correspondingly higher space time θ, e.g., higher catalyst loading, increasing reaction time, or reduced feed rate. 4. Conclusion A plastics liquid from a pyrolysis process was separated by distillation into distillate and residual fractions,

with the residual fraction containing greater than 70% gas oil. The hydrocracking of this residual fraction was studied in batch reactions where it was successfully upgraded to the boiling point range of commercial gasoline. A simple consecutive lumped model was fitted to the data for the hydrocracking of gas oil to naphtha and gases. The kinetic model accounted for differences due to unequal hydrogen partial pressures in small and large batch reactors. Simulation results showed that the model fit the experimental data fairly well. Both model and experiment showed a maximum in the yield of naphtha as a function of gas oil conversion with the naphtha yield being higher at lower temperatures and longer space times. The trends shown by the model used in this work may be applicable to other plastics pyrolysis liquids, although the numerical values of the rate constants would likely be somewhat different. The study generally indicates a positive potential for successful upgrading of plastics pyrolysis residual liquids to naphtha range materials suitable for inclusion in the transportation fuels pool.

Acknowledgment. This work was supported by the U.S. Department of Energy under Contract No. DEFC22-90PC90029 as part of the research program of the Consortium for Fossil Fuel Liquefaction Science. The authors are grateful to Conrad Industries for supplying the plastics pyrolysis liquid sample used in this study. EF960151G