Hydrocracking of Iron-Catalyzed FischerâTropsch Waxes - Energy

Energy & Fuels 2005, 19, 1795-1803

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Hydrocracking of Iron-Catalyzed Fischer-Tropsch Waxes Dieter Leckel* Fischer-Tropsch Refinery Catalysis, Sasol Technology Research and Development, P.O. Box 1, Sasolburg 1947, South Africa Received March 30, 2005. Revised Manuscript Received July 7, 2005

Sasol iron-catalyzed Fischer-Tropsch (FT) reactor waxes were hydrocracked at 3.5 and 7.0 MPa hydrogen pressure to fuels using a NiMo on SiO2/Al2O3 catalyst. The aim was to predominantly produce diesel. The effects of feed carbon number distribution, reaction variables, and process configurations have been investigated. High hydrogen-to-wax ratios led to improved C23+ conversion and increased diesel yields. Increasing the liquid hourly space velocities decreased C23+ conversion and increased diesel selectivity. Prehydrogenation of FT reactor wax prior to the hydrocracking reaction improved C23+ conversion; however, it negatively influenced diesel selectivity. The apparent n-paraffin reactivity was found to be chain length dependent increasing with pressure.

Introduction Coal conversion was extensively researched between 1915 and 1945 in Germany, resulting in a major technological effort to develop hydrocracking.1 Thereafter, hydrocracking became less important because the availability of crude oil removed the incentive to convert coal to liquid fuel. Rapid growth for hydrocracking was seen in the early 1960s when companies such as Chevron2 and Universal Oil Products3 (UOP) introduced new hydrocracking processes to enable the manufacture of higher octane gasoline. Since the 1970s, the demand for transportation fuels shifted the product distribution to a broader range of products, particularly middle distillates such as jet and diesel fuel. The demand for higher quality products, driven by both performance criteria and environmental constraints continued to increase, and these changes resulted in a continuous development of hydrocracker catalysts and process technologies. Traditional hydrocracker feeds are crude oil derived; however, some alternatives are FischerTropsch waxes. These waxes are essentially free of aromatics, sulfur, and nitrogen but contain oxygenates and olefins. The straight run diesel resulting from this mainly linear paraffinic material has a high cetane number, though the cold flow properties are poor. Hydroprocessing of these highly paraffinic fractions is necessary, and the processing technology applied is hydrocracking. Conventional amorphous bifunctional catalysts, consisting of transition metals dispersed on acidic supports, have been used in the past for hydrocracking.4,5 Specifically metals with high hydrogenation * Tel: +27 16 960-3830. Fax: +27 11 522-3975. E-mail: dieter.leckel@ sasol.com. (1) Scherzer, J.; Gruia, A. J. Hydrocracking Sci. Technol. Dekker: New York, 1996; pp 9-11. (2) Stormont, D. H. Oil Gas J. 1959, 57 (44), 48-49. (3) Sterba, M. J.; Watkins, C. H. Oil Gas J. 1960, 58 (21), 102106.

function, like platinum, provide a favorable balance with the acidity of catalyst supports.6-15 Important for the hydrocracking and hydroisomerization of FT waxes to fuels is a high selectivity for the isomerization of the long chain n-paraffins. UOP completed a study of FT wax hydrocracking in 1988 sponsored by the U.S. Department of Energy (DOE). In this study, Shah et al.16 studied the performance of amorphous silicaalumina supported catalysts and zeolites for the hydrocracking of ARGE (ARGE ) Arbeitsgemeinschaft Ruhrgebiet) tubular fixed-bed iron reactor wax. Middle distillate yields of up to 82 wt % were achieved. McArdle et al.17 investigated the differences of FT ARGE iron reactor wax and a typical crude refinery feed for hydrocracking and found that temperatures required to achieve a specific conversion are as much as 28 °C lower (4) Archibald, R. C.; Greensfelder, B. S.; Holzman, G.; Rowe, D. H. Ind. Eng. Chem. 1960, 52 (9), 745-750. (5) Flinn, R. A.; Larson, O. A.; Beuther, H. Ind. Eng. Chem. 1960, 52, 153-156. (6) Coonradt, H. L.; Garwood, W. E. Ind. Eng. Chem. Process Des. Dev. 1964, 3, 38-45. (7) Steijns, M.; Froment, G.; Jacobs, P.; Uytterhoeven, J.; Weitkamp, J. Ind. Eng. Chem. Prod. Res. Dev. 1972, 11, 46-53. (8) Weitkamp, J. Ind. Eng. Chem. Prod. Res. Dev. 1982, 20, 654660. (9) Weitkamp, J. J. Appl. Catal. 1983, 8, 123-141. (10) Martens, J. A. Jacobs, P. A.; Weitkamp, J. J. Appl. Catal. 1986, 20, 239-281. (11) Martens, J. A. Jacobs, P. A.; Weitkamp, J. J. Appl. Catal. 1986, 20, 283-303. (12) Giannetto, G. E.; Perot, G. R.; Guisnet, M. R. Ind. Eng. Chem. Prod. Res. Dev. 1986, 25, 481-490. (13) Giannetto, G. E.; Alvarez, F. B.; Guisnet, M. R. Ind. Eng. Chem. Res. 1988, 27, 1174-1181. (14) Weitkamp, J.; Ernst, S. Catal. Today 1994, 19, 107. (15) Calemma, V.; Peratello, S.; Pavoni, S.; Clerici, G.; Perego C. Natural Gas Conversion VI. Stud. Surf. Sci. Catal. 2001, 136, 307. (16) Shah, P. P, Sturtevant, G. C.; Gregor, J. H.; Humbach, M. J.; Padtra, F. G.; Steigleder, K. Z. Fischer-Tropsch Wax Characterisation and Upgrading, Final Report. U.S. DOE: Washington, DC, June 1988; Contract DE-AC22-85PC80017. (17) McArdle, J. C.; Humbach, M. J.; Schoonovver, M. W.; Padtra, F. G. F-T Wax Characterisation and Upgrading. DOE review meeting, Pennsylvania, 1986.

10.1021/ef050085v CCC: $30.25 © 2005 American Chemical Society Published on Web 08/16/2005

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when processing FT ARGE reactor wax as compared to vacuum gas oil. Hydrocracking of Fischer-Tropsch waxes and paraffins using bifunctional catalysts18 follow the principles of ideal hydrocracking.19,20 By combining FT synthesis and hydrocracking, diesel selectivities above 80% are achievable.21 It became more or less accepted that high R-value FT synthesis catalysts are needed for the production of transportation fuels22 and to maximize diesel.23 At Sasol wax hydrocracking research dates back to the early 1970s when selective hydrocracking of lowtemperature Fischer-Tropsch (LTFT) wax was investigated to produce transportation fuels.24-26 ARGE tubular fixed-bed iron reactor wax was hydrocracked under mild conditions and recycling the product heavier than diesel to extinction. A final product distribution of about 80% diesel, 15% naphtha, and 5% gas was obtained. Wax hydrocracking was however not implemented at Sasol because petrol production was seen as a more economically viable at that time. In addition, the LTFT route to diesel was not pursued since diesel specifications were not as stringent as they are today. The high-temperature Fischer-Tropsch process (HTFT) was then implemented at the Sasol 2 and 3 plants producing a product spectrum favoring lighter hydrocarbons and olefins.27 The interest has changed, and revised interest worldwide shifted to the route producing high quality diesel. Shell implemented the wax hydrocracking option at their plant in Bintulu, Malaysia, in 1993 using multitubular fixed-bed reactors. The Shell middle distillate process28,29 (SMDS) converts natural gas to transportation fuels. Paraffins and waxes are byproducts. Exxon now offers a process for the hydrocracking/hydroisomerization of FT waxes to liquid oils.30,31 Sasol developed a process based on the cobalt LTFT slurry reactor technology, the Sasol Slurry Phase Distillate (SPD) process.32 Diesel fuel with a cetane number of higher than 70 is produced having significantly lower particulate emissions than a CARB fuel.33 SasolChevron are presently building a cobalt-based slurry FT plant with downstream wax hydrocracking in Qatar scheduled to come on line in 2006. (18) Guisnet, M.; Alvarez, F.; Giannetto, G.; Pe´rot, G. Catal. Today 1987, 1, 415. (19) Pichler, H.; Schulz, H.; Reitmeyer, H. O.; Weitkamp, J. Erdo¨ lKohle-Erdgas-Petrochem. 1972, 25, 494. (20) Weitkamp, J. ACS Symp. Ser. 1975, No. 20, 1. (21) Eisenberg, B, Fiato, R. A.; Mouldin, C. H.; Say, G. R.; Soled, S. L. Natural Gas Conversion V. Stud. Surf. Sci. Catal. 1998, 119, 943. (22) Raje, A.; Inga, J. R.; Davis, B. H. Fuel 1997, 76 (3), 273-280. (23) Gregor, J. H. Catal. Lett. 1990, 7, 317-332. (24) Dry, M. E. Catal. Sci. Technol. 1981, 1, 159. (25) Dry, M. E. Hydrocarb. Process. 1982, 61 (2), 121. (26) Dry, M. E. The Sasol Fischer-Tropsch processes. In Applied Industrial Catalysis, Vol 2; Leach, B. E., Ed.; Academic Press: New York, 1983; p 167. (27) Dry, M. E. In Methane Conversion; Bibby, D. M., Chang, C. D., Howe, R. F., Yurchak, S., Eds.; Elsevier Science: Amsterdam, 1988; pp 447-456. (28) Sie, S. T.; Senden, M. M. G.; Van Wechem, H. M. H. Catal. Today 1991, 8 (3), 371-394. (29) Eilers, J.; Posthuma, S. A.; Sie, S. T. Catal. Lett. 1990, 7, 253. (30) Eisenberg, B.; Fiato, R. A.; Kaufmann, T. G.; Bauman, R. F. Chemtech 1999, 32-37. (31) Everett, B. M.; Eisenberg, B.; Bauman, R. F. Conference on Natural Gas, March 1995, Doha, Qatar. (32) Espinoza, R. L.; Steynberg, A. P.; Jager, B.; Vosloo, A. C. Appl. Catal. A 1999, 186 (1-2), 13-26. (33) Schaberg, P. W.; Myburgh, I. S.; Botha, J. J.; Roets, P. N. J.; Viljoen, C. L.; Dancuart, L. P.; Starr, M. E. SAE Pap. 1997, No. 972898.

Leckel Table 1. Physical Properties of the NiMo/SiO2-Al2O3 Hydrocracking Catalyst Used catalyst properties NiO, wt % MoO3, wt % SiO2/Al2O3 mass ratio BET surface area, m2/g pore volume, mL/g

3.2 14.9 0.5 192 0.622

This paper describes the hydrocracking of Sasol LTFT tubular fixed-bed (ARGE) and slurry iron reactor waxes to fuels, using a NiMo on SiO2/Al2O3 catalyst. Carbon number distribution, reaction variables, and process configurations have been investigated. Experimental Section Reactor System. A bench-scale fixed-bed reactor, operating in concurrent down-flow mode, was used for the hydrocracking studies. The reactor consists of a stainless steel tube (ASTM A 312 GRTP 316) with an internal diameter of 47.6 mm and 1800 mm length, designed for a pressure of 250 bar and a temperature of 454 °C. The reactor is heated electrically by three heater elements placed along the reactor tube. To support the reactor load a fine mesh grid is welded onto a thermocouple well that runs through the center of the reactor over its entire length. Six thermocouples can be individually placed inside the well to measure the temperature inside the reactor along the catalyst bed. Temperature control is automated. Variance of the temperature control is (0.5 °C. The catalyst is loaded in the center of the reactor tube. Sand (35-50 mesh) was used to fill the voids between the catalyst particles to avoid channelling. Inert packing (glass balls) above the catalyst bed was used to preheat the feed up to reaction temperature. Liquid feed and hydrogen enters concurrently at the top of the reactor. Gas and liquid products are separated in the bottom third of the reactor system. The liquid product is collected in a liquid catch pot, where the gas and light hydrocarbons are passed through a cooling coil at 0 °C. The condensed liquid is collected in a second gas liquid separator. A gas sampling point is installed on the low-pressure side of the second separator. Gas flow was measured with a wet gas meter. Catalyst. A commercial NiMo on silica-alumina hydrocracking catalyst was used for the tests. Some characteristic physical properties of the catalyst are listed in Table 1. BET surface area (by nitrogen physisorption) and pore volume were measured using equipment supplied by Micromeritics. Catalyst Activation. The catalyst was predried in nitrogen at 125 °C and 0.1 MPa for 8 h prior to sulfidation. Then nitrogen was replaced by hydrogen (100 L/h flow rate) and the reactor pressurized to 3.0 MPa. Dimethyl disulfide (2 wt %) in a C9-C11 paraffin mixture was used as sulfiding agent. The catalyst was first wetted thoroughly with the sulfiding mixture using maximum pump speed. At decreased pump rate the temperature was ramped hourly from 125 °C in steps of 25 °C to 250 °C, where it was kept for 4 h. H2S breakthrough was monitored in the off-gas by means of Draeger tubes. A 2000 ppm H2S concentration in the off-gas was seen as sufficient to increase the temperature in steps of 25 °C further to 350 °C where it was kept for another 2 h. Thereafter the reactor temperature was decreased to 240 °C. When sulfiding was completed, the reactor pressure is raised, and the desired hydrogen feed rate is established. The feed is introduced at a lower than operating temperature to avoid an overshooting of the operating temperature due to the exotherm of the hydrogenation reaction. To ensure that the wax does not solidify in the lines during the test runs, all lines to and from the reactor were heated to above the melting point of the wax. Procedure. After the reactor reached steady-state conditions, in general after a period of 72 h, catalyst activity was

Hydrocracking of Fischer-Tropsch Waxes monitored by drawing product samples in 8-h intervals. The following 8-h period was used to collect a representative sample for product analysis. Mass balance was closed with collection of tailgas, liquid product, residue, and condensed lighter hydrocarbons from the reactor offgas stream. Condensed light hydrocarbons were kept refrigerated prior to analysis. The sulfided state of the catalyst was monitored by H2S tailgas analyses. H2S levels of 200 vppm in the tailgas were seen as sufficient to keep the catalyst sulfided. Mass balances of between 97 and 102% were seen as appropriate for all hydrocracking experiments. Control samples were subjected to complete analyses. Operating Conditions. Hydrocracking was done with operating temperature varying between 350 and 380 °C. Pressures of 3.5-7.0 MPa were applied. The liquid hourly space velocity (LHSV) was varied between 0.35 and 1 h-1 while the hydrogen-to-feed ratio was investigated between 500 and 1500 Ln/L. Wax conversions levels of between 25% and 98% were covered in the study. Recycle experiments were carried out with an average combined feed ratio (CFR) of 1.56. The CFR is defined as the ratio of the sum of fresh feed and recycle feed divided by the fresh feed. Recycle operation was simulated in such a way that the hydrocracker bottom residue collected from a once through test run was added batchwise to the fresh feed. Feed. The products from the Sasol LTFT reactors at Sasolburg are gas, condensates, and reactor wax. The average yields of the liquid products from an iron slurry reactor are 38 wt % condensates (hot and cold) and 62 wt % reactor wax. Linear paraffins are the major constituents in the products. There are however substantial amounts of olefins (mainly R-olefins) in the streams, with carbon numbers from C5 to C20, and in addition oxygenates are present. Alcohols and carbonyls are the major components in the oxygenate fraction. The product is essentially sulfur- and nitrogen-free. Low concentrations of aromatics (below 1 wt %) are observed. Cold condensates derived from the iron LTFT process have a carbon number distribution of C1-C30, peaking at around C10 and consist mainly of paraffins (55 wt %), R-olefins (24 wt %), internal olefins (6 wt %), alcohols (8 wt %), isoparaffins and other olefins (6 wt %), and ketones (1 wt %). The iron slurry LTFT reactor wax (RW) consists of molecules with a carbon number between C10 to C120 and peaks at a carbon number of around C30. The reactor wax is predominantly paraffinic, more specifically having a n-paraffin content of about 94 wt %. Isoparaffins and olefins as well as a minor amount of oxygenates constitute the balance. ARGE iron tubular fixed-bed reactor wax consists of hydrocarbon molecules between carbon number C10 and C120 peaking at carbon number C28. The carbon number distribution of the iron tubular fixed-bed, iron-based LTFT slurry reactor waxes and iron slurry reactor wax condensates are presented in Figure 1. The difference in carbon number distribution, meaning the ARGE wax being lighter, is not related to the difference in the design of the reactors. It arises from the different modes of product workup. The product derived from the fixed-bed ARGE reactor is externally fractionated while the product of the slurry bed reactor is fractionated online as result of the reactor operation. Iron slurry reactor products are in general much more olefinic than tubular fixed-bed reactor products. They are also more linear and contain somewhat less oxygenates, specifically alcohols and carbonyls.34 The reactor waxes were analyzed for density at 120 °C (ASTM D-4052), viscosity at 120 °C, melting point, acid number (ASTM D-1386/7), bromine number (ASTM D-1159), esters, alcohols, and carbonyls. Sasol in-house methods based on ASTM methods are used in cases where the method had to be modified according to the waxy feed nature. (34) Jager, B.; Kelfkens, R. C.; Steynberg, A. P. In Natural Gas Conversion II; Curry-Hyde, H. E., Howe, R. F., Eds.; Elsevier Science: Amsterdam, 1994; pp 419-425.

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Figure 1. Carbon number distribution (HT-GC) of ironcatalyzed tubular fixed-bed FT reactor wax (ARGE), ironcatalyzed slurry bed reactor wax, and iron-catalyzed slurry bed reactor condensates. This includes ASTM methods D 1386, D 1387, and D 5622 using a Perkin-Elmer EA 1110 CHNSO analyzer. Hightemperature GC analyses were done on a Restek 15 m, 0.288 mm i.d. and 0.15 µm thickness MXT-1 metal capillary column. Product Analyses. The reactor tailgas was collected in a gas sample bomb that was placed into the offgas stream and analyzed by GC. All liquid reactor effluents were collected and fractionated into a naphtha (C5-C9, C9 ) 170 °C) stream, diesel (C10-C22 ) 170-370 °C). and a residue product boiling above 370 °C (equals C23+). Simulated distillation (SIMDIS ASTM D-2887) was applied to yield the boiling point distribution of the liquid products. GC provided the carbon number distribution that was used to calculate conversion, selectivities, and yields. Terminologies. The following terminologies to calculate conversion, selectivities, and yields were used. Conversion is here understood as true conversion, since according to the definitions used naphtha and diesel already present in the feed is corrected for. The hydrocracking conversion (“true” conversion) was defined as the difference between the weight percentage C23+ material in the feed (wt % C23+Feed) and that in the reaction products (wt % C23+Product) divided through the weight percentage C23+ in the feed (wt % C23+Feed) multiplied by 100. The diesel (C10-C22) selectivity was calculated by dividing the difference of the weight percentage diesel in the product (wt % C10-C22 Product) and the weight percentage diesel in the feed (wt % C10-C22 Feed) by the difference in weight percentage C23+ in the feed (wt % C23+Feed) and the weight percentage C23+ in the product (wt % C23+Product) multiplied by 100. The naphtha (C5-C9) and gas (C1-C4) selectivities were calculated analogously. The yields were calculated by multiplying the conversion with the selectivity. Carbon number distributions were taken from GC analyses.

Results and Discussion The effect of different variables on the hydrocracking of LTFT slurry reactor wax (RW) was investigated. These variables included different types of feeds, prehydrogenation of feed, temperature, liquid hourly space velocity (LHSV), hydrogen-to-liquid ratio, and effect of recycle. Effect of Pressure. The apparent reaction order with respect to hydrogen is usually reported in the literature35-38 to be -1 or slightly higher (-0.85).39 This (35) Weitkamp, J.; Jacobs, P. A.; Martens, J. A. Appl. Catal. A 1983, 8, 123. (36) Steijns, M.; Froment, G. F. Ind. Eng. Chem. Prod. Res. Dev. 1981, 20, 660. (37) Ribeiro, F.; Marcilly, C.; Guisnet, M. J. Catal. 1982, 78, 267.

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Figure 2. Effect of pressure on the C23+ conversion during hydrocracking of iron-catalyzed FT slurry bed reactor wax at 0.55 h-1 LHSV and a H2 to wax ratio 1500/1 Ln/L.

implies that higher hydrogen pressures, necessary for the hydrogenation step, will have an inhibiting effect on the isomerization rate and, therefore, the hydrocracking rate. Meaning, the hydrogen pressure should primarily influence the isomerization activity, which in turn influences the cracking activity. According to the bifunctional hydrocracking mechanism,40,41 the first step in hydrocracking is the formation of olefins at the metal site (dehydrogenation). Since the limiting step is the rearrangement of the secondary carbenium ion, an increase in hydrogen pressure will lead to lower steady-state concentration of n-alkenes and secondary carbenium ions. A lower hydrogen pressure should therefore lead to enhanced isomerization followed by secondary cracking. The effect of pressure was studied at 3.5 and 7.0 MPa using a temperature range of 350-370 °C, a LHSV of 0.55 h-1, and a hydrogen-to-wax ratio of 1500/1 Ln/L. It was shown that higher operating temperatures are needed to achieve the same C23+ conversion at the higher pressure (7.0 MPa). For comparison, to achieve a C23+ conversion of 34%, the operating temperature needed at 3.5 MPa was 350 °C, while the temperature needed at 7.0 MPa was 365 °C to achieve a similar conversion. In other words, at a given operating temperature increasing the pressure led to a decrease in C23+ conversion (see Figure 2). It seems that for wax hydrocracking the C23+ conversion is inversely related to the hydrogen pressure. Vapor-liquid equilibrium compositions at a given temperature could be in addition affected by differences in pressures. Low pressures would thereby result in an enrichment of lighter hydrocarbons in the vapor phase. Adsorption of longer chain hydrocarbons on the catalysts active sites is now favored, resulting in higher C23+ conversion. In general product yields will increase (as seen in Table 2). Since more light hydrocarbons are in the gas phase and not adsorbed on the catalyst, the possibility of readsorption of diesel is given. This could result in decreased selectivities to diesel. (38) Baltanas, M. A.; Vasina, H.; Froment, G. F. Ind. Eng. Chem. Prod. Res. Dev. 1983, 22, 531. (39) Beecher, R.; Voorhies, A., Jr. Ind. Eng. Chem. Prod. Res. Dev. 1969, 8 (4), 366. (40) Weitkamp, J. Am. Chem. Soc., Div. Petr. Chem., Prepr. 1975, 20, 489. (41) Langlois, G. E.; Sullivan R. F. Am. Chem. Soc., Div. Petrol. Chem., Prepr. 1969, 14 (4), 8.

Leckel

Figure 3. Effect of pressure on the carbon number distribution during hydrocracking of iron-catalyzed FT slurry bed reactor wax at 0.55 h-1 LHSV and a 1500/1 Ln/L H2 to wax ratio. Table 2. Hydrocracking of Iron-Catalyzed FT Slurry Bed Reactor Wax at 3.5 and 7.0 MPa Using a 0.55 h-1 LHSV and a H2/Wax Ratio of 1500/1 Ln/L run no. 2/1

2/2

2/3

2/4

2/5

2/6

pressure, MPa 7.0 7.0 7.0 3.5 3.5 3.5 T, °C 360 365 370 350 355 360 C23+ conversion, wt % 26 37 56 34 50 84 yields, wt % C1-C4 1.6 2.2 3.6 2.3 3.8 11 C5-C9 2.1 3.5 8.4 3.4 6.7 13 C10-C22 22 31 44 27 40 63 C23+ 74 63 44 66 49 13 selectivity, wt % C1-C4 6.1 6.2 6.3 6.9 7.5 10 C5-C9 8.1 9.4 15 10 12 15 C10-C22 85 84 79 82 80 75

The latter effect is observed in Table 2 at both pressures however being more pronounced at the low pressure. Diesel selectivities decreased, while gas and naphtha selectivities increased with increasing C23+ conversion. Comparing selectivities at similar C23+ conversion, increasing the pressure did not significantly affect product selectivities (see run nos. 2/2 and 2/4 in Table 2). Figure 3 shows the carbon number distribution of iron slurry reactor wax and the product carbon number distribution of iron slurry reactor wax hydrocracked at 7.0 and 3.5 MPa compared at similar C23+ conversion, 37% and 34%, respectively. The differences in product yields are not very pronounced, varying only a few percentage points. A 15 °C higher operating temperature is needed to achieve the same C23+ conversion at the higher pressure. The decrease in hydrogen pressure results ultimately in a shift to lighter products with increasing naphtha yields at the expense of the diesel yields. This is more apparent at high C23+ conversion. Figure 4 shows the effect of pressure on the diesel-tonaphtha ratio at different C23+ conversion. The dieselto-naphtha ratio at low pressure (3.5 MPa) decreases from a ratio of 7.9 at low C23+ conversion (34%) to a ratio of 4.8 at high conversion (84%). At the higher pressure of 7.0 MPa, the diesel-to-naphtha ratio decreases from 10.4 at 26% C23+ conversion to 6.8 at 66% conversion. The higher pressure resulted in a 20% higher diesel-to-naphtha ratio. In general, higher operating temperatures seemed to lead to more secondary cracking, thus increasing the naphtha yields at the expense of diesel yields.

Hydrocracking of Fischer-Tropsch Waxes

Energy & Fuels, Vol. 19, No. 5, 2005 1799 Table 3. Hydrocracking of Iron-Catalyzed FT ARGE Fixed-Bed Reactor Wax Using Different Liquid Hourly Space Velocities (LHSV) at 3.5 MPa, 370 °C, and a H2 to Wax Ratio of 1500/1 Ln/L C23+ selectivity, wt % yield, wt % LHSV, conv., wt % C1-C4 C5-C9 C10-C22 C1-C4 C5-C9 C10-C22 h-1 0.5 0.60 1.0

57 41 30

7.1 8.5 8.7

21 15 12

72 78 80

4.1 3.5 2.6

12 6.2 3.5

41 32 24

Table 4. Hydrocracking of Iron-Catalyzed FT Slurry Bed Reactor Wax Using Different Hydrogen to Wax Ratios at 7.0 MPa, 370 °C, and 0.5 h-1 LHSV

Figure 4. Effect of pressure on the diesel-to-naphtha ratio during hydrocracking of iron-catalyzed FT slurry bed reactor wax at 0.55 h-1 LHSV and a 1500/1 Ln/L H2 to wax ratio.

Other authors report opposite trends with regards to pressure. Keogh et al.,42 who studied the hydroisomerization and hydrocracking of n-hexadecane over a platinum-promoted sulfated zirconia catalyst, found at constant temperature an increase in conversion with increasing reactor pressure. Cracking yields increased, and the majority of the increase was accounted for as naphtha, the C5-C9 fraction. The distribution of the cracked products in their studies showed an unusual asymmetric distribution. The hydrocracking seemed not to follow the conventional bifunctional hydrocracking mechanism. Venkatesh et al.43 reported the same and concluded that platinum loaded, sulfated zirconium oxide catalysts seem not to act as a conventional bifunctional catalyst for higher (C7+) alkanes. Addition of olefins inhibited hydrocracking of higher alkanes over anion-modified zirconium oxides (AZOs), indicating that the reaction mechanism may not involve a metalcatalyzed dehydrogenation to an olefin. A unimolecular mechanism was proposed, and it was suggested that olefins are unlikely precursors to carbocations. Liu et al.44 identified combined metal-acid sites on small Pt particles anchored to the support by a proton on platinum promoted AZOs. Hydrogen dissociates on the Pt and serves as source of hydride ions which are transferred to adjacent acid sites reacting with surface carbenium ions to form isoalkanes.45 The formation of an unsaturated intermediate is not necessary in this pathway on combined metal-acid ensembles. In addition an increased surface turnover of acidic isomerization is observed. Buchholz et al.46 defined these ensembles as “compressed bifunctional sites”. Effect of Liquid Hourly Space Velocity. Iron fixed-bed ARGE reactor wax was hydrocracked using the NiMo/SiO2-Al2O3 catalyst at a constant pressure of 7.0 MPa, a temperature of 370 °C, and a 1500/1 Ln/L hydrogen-to-wax ratio. The liquid hourly space velocity was varied from 0.5 to 1.0 h-1. Table 3 shows the variation in C23+ conversion, product selectivities, and (42) Keogh, R. A.; Sparks, D.; Hu, J.; Wender, I.; Tierney, J. W.; Wang, W.; Davis, B. H. Energy Fuels 1994, 8, 755-762. (43) Venkatesch, K. R.; Hu, J.; Wang, W.; Holder, G. D.; Tierney, J. W.; Wender, I. Energy Fuels 1996, 10, 1163-1170. (44) Liu, H.; Lei, H.; Sachtler, W. M. H. Appl. Catal. A 1996, 137, 167. (45) Iglesia, E.; Soled, S. L.; Kramer, G. M. J. Catal. 1993, 144, 238. (46) Buchholz, T.; Wild, U.; Muhler, M.; Resofszki, G.; Paa´l, Z. Appl. Catal. A 1999, 189, 225.

C23+ selectivity, wt % yield, wt % H2/wax, conv., Ln/L wt % C1-C4 C5-C9 C10-C22 C1-C4 C5-C9 C10-C22 500 1000 1500

27 35 67

1.9 2.4 3.4

17 19 20

80 79 77

0.5 0.8 2.2

4.8 6.5 13

22 28 52

yields. The major effect of increasing space velocity is a decrease in C23+ conversion. Increasing the liquid hourly space velocity from 0.5 to 1.0 h-1 decreased the C23+ conversion significantly from 57% to 30%. The selectivities to gas (C1-C4) changed marginally. Diesel selectivities increased while naphtha selectivities decreased. Increasing the space velocity from 0.5 to 1 h-1 resulted in an almost 50% lower C23+ conversion. A 40% lower diesel, a 70% lower naphtha, and a 40% lower gas yield was observed. To maintain a constant C23+ conversion when increasing the liquid hourly space velocity, an increase in operating temperatures would be needed. In a typical crude oil refinery, hydrocracking space velocity changes generally do not affect the distillate yields, as long as the reactor temperature is adjusted to keep the conversion constant. Effect of Hydrogen-to-Wax Ratio. The effect of the hydrogen-to-wax ratio was studied at the same operating conditions as used in the previous investigation. In this study the liquid hourly space velocity was kept constant. The ratio of hydrogen to wax varied between 500 and 1500/1 Ln/L (see Table 4). The most obvious effect of increasing the hydrogen-to-wax ratio was the major increase in C23+ cracking conversion. At the ratio of 500/1 Ln/L the cracking conversion was low (27%) and only slightly increased to 35% at a 1000/1 Ln/L ratio. As the ratio was further increased to 1500/1 Ln/L, the cracking conversion almost doubled (67%). High hydrogen flows seem to be beneficial for the hydrocracking of reactor waxes. One could argue that hydrogen uptake by the catalyst and/or the hydrogenation step is mass transfer limited and higher hydrogen flow rates overcome these limitations. The other reason could be the that higher reactivity of the longer chain paraffins and partial evaporation of the lighter part of the feed and product molecules leads to the situation that the liquid phase is enriched in heavier hydrocarbons, as was confirmed by Calemma et al.,15 who hydrocracked longer chain paraffinic material at 5.0 MPa and 340 °C, using a noble metal mesoporous hydroconversion catalyst. They found that a decrease in the hydrogen-to-wax ratio led to a decrease in the conversion rate of the C23+ fraction at given reaction conditions. Higher values of the H2/wax ratio led to a higher percentage of feed being in the vapor phase and such higher concentration of C23+ material

1800


in the liquid phase. The possible influence of partial vaporization of the wax feed on the difference in reactivity is demonstrated in Table 9 where the tubular fixed-bed (ARGE) and slurry bed wax hydrocracking is compared at identical operating temperatures. An approximately 10% higher C23+ conversions at a given temperature is observed when the lighter ARGE wax is used as feed while selectivities to diesel are about 5-10% higher. Hydrocracking the heavier slurry bed wax at identical temperatures resulted in higher gas and naphtha selectivity. Higher operating pressures will lead to more product hydrocarbons being in the liquid phase and as such more hydrocarbons boiling in the naphtha and diesel range might be converted at given reaction conditions, as might be the case in our study with regards to the more severe reaction conditions applied. Although the hydrogen-to-liquid ratio has a big influence on the C23+ conversion, the effect on selectivities to gas, naphtha, and diesel was negligible at the conditions applied. The conversion level is still in a regime where secondary cracking is not favored. Effect of Wax Prehydrogenation. The LTFT iron slurry reactor wax was subjected to mild hydrogenation at temperatures sufficiently low to avoid any isomerization and/or hydrocracking. Some catalyst deactivation observed during the hydrocracking studies with untreated feed was suspected to have been caused by the oxygenates and olefins present in the wax feed. Shell has patented a process47 whereby the reactor wax is hydrogenated before hydrocracking in order to saturate the olefins and oxygenates present. This is to protect the noble metal hydrocracking catalyst used in the second hydroconversion step of the process. Water produced during the hydrocracking reaction was found to reduce catalyst performance. It was further observed that selectivity to hydrocarbon fuels, specifically diesel, improved. Previous studies by Venkatesch et al.,43 using metal-promoted anion-modified zirconium oxides, showed that short chain olefins inhibited the hydrocracking activity for alkanes having the same carbon number. The inhibition was found to be less severe for longer chain hydrocarbons such as n-hexadecane since the difference in cracking activity of an alkane and its corresponding olefin decreased with increasing chain length. This might be catalyst specific since this behavior deviates from the normal bifunctional hydrocracking mechanism. The results of hydrocracking prehydrogenated iron slurry reactor wax feed with the NiMo/SiO2Al2O3 catalyst are presented in Table 5. Prehydrogenation was done with a state of the art hydrotreating catalyst. Table 5 shows that compared to the runs with unhydrotreated wax, prehydrogenation of the wax feed improves the C23+ conversion significantly. At 360 °C operating temperature, a 70% C23+ conversion is achieved using the hydrogenated wax feed while a 15 °C higher operating temperature is needed to achieve the same conversion using unhydrogenated wax. Hydrogenated iron slurry reactor wax requires lower temperatures to obtain the same conversion as compared to unhydrogenated wax. At identical operating temperatures, selectivities to naphtha are generally (47) Shell. EP0583836 A1.

Leckel Table 5. Hydrocracking of Hydrogenated (H) and Unhydrogenated Iron-Catalyzed FT Slurry Bed Reactor Wax at 7.0 MPa, 0.55h-1 LHSV, and a H2/Wax Ratio of 1500/1 Ln/L

T, °C 360 (H) 365 (H) 370 (H) 375 (H) 360 375 380

C23+ selectivity, wt % yield, wt % conv., wt % C1-C4 C5-C9 C10-C22 C1-C4 C5-C9 C10-C22 70 83 99 100 25 69 97

2.7 3.4 3.9 4.8 3.7 4.1 7.6

30 30 38 41 19 20 34

68 66 59 54 76 74 58

1.9 2.8 3.9 4.8 0.9 2.8 7.4

21 25 38 41 4.7 14 33

48 55 58 54 19 51 56

higher for pretreated feed while the selectivity to diesel is higher in the case of unhydrogenated feed. The lower C23+ conversion in the case of unhydrogenated feed is most probably caused by the oxygenates adsorbing strongly on the acid sites thereby suppressing the cracking activity of the acid sites. Selectivities to diesel are clearly lower in the case of the hydrogenated feed at the gain of naphtha selectivity. At the same C23+ cracking conversion higher gas selectivities are observed for the unhydrogenated wax feed. This is the result of higher temperatures applied to achieve the same conversion. Higher temperatures generally increase cracking. Diesel selectivities are higher and naphtha selectivities are lower using unhydrogenated wax. At the same C23+ conversion the diesel-to-naphtha ratio is significantly higher for unhydrogenated wax. Higher operating temperatures to achieve identical conversion could be a disadvantage due to faster catalyst deactivation rates. This behavior might be further explained on the basis of the classical bifunctional alkane hydroisomerization and hydrocracking mechanism. Oxygenates seemed to modify the catalyst properties by inhibiting the catalyst’s conversion and cracking activity, however influencing positively the product selectivities to diesel. Oxygenates can strongly adsorb on acid sites, especially the strong acid sites. This reduces the available number of acid sites (NA) and thereby influences the balance between acid and hydrogenation/dehydrogenation function. The number of metal sites (NMe) stayed constant, and the net effect is an increase in the NMe/NA ratio. The likelyhood is higher for the intermediate alkene to encounter a metal site and to be hydrogenated before successive isomerization to multibranched isomers or secondary cracking occurs. Transformation to multibranched isomers is inhibited, and most likely only one transformation of the primarily formed alkene to a monobranched intermediate is possible before hydrogenation on a metallic site occurs. Thus, the higher diesel formation in favor of naphtha formation and the lower C23+ conversion experienced during hydrocracking of unhydrogenated wax can be explained. Similar observations were reported by Alvarez et al.,48 who hydroisomerized and hydrocracked n-decane over PtHY catalyst with different metal to acid site ratios. The balance between acid and hydrogenation functions influenced the transformation of the alkane. Low cracking conversion was attributed to high ratios of metal site number to acid site number. Light products formation was favored by low metal site number to acid site number ratios. (48) Alvarez, F.; Ribeiro, F. R.; Perot, G.; Thomazeau, C.; Guisnet, M. J. Catal. 1996, 162, 179-189.



Table 6. Hydrocracking Iron-Catalyzed Tubular Fixed-Bed (ARGE) FT Reactor Wax in Recycle Mode Using a Pressure of 3.5 MPa, a 0.5 h-1 LHSV, 360 °C, and a H2/Wax Ratio of 1500/1 Ln/L ratio ARGE wax to C23+

C23+ conv., wt %

C1-C4

ARGE wax 1: 0.92 1: 1.22

42 44 54

1.7 2.7 3.1

yield, wt % C5-C9 C10-C22 5.8 6.3 9.1

34 35 42

diesel to naphtha ratio 5.9 5.6 4.7

Table 7. Hydrocracking Iron-Catalyzed Tubular Fixed-Bed (ARGE) FT Reactor Wax in Recycle Mode at Different Space Velocity Using a Pressure of 3.5 MPa, 370 °C, and a H2/Wax Ratio of 1500/1 Ln/L ratio C23+ diesel to yield, wt % naphtha LHSV, ARGE wax conv., to C23+ ratio wt % C1-C4 C5-C9 C10-C22 h-1 0.5 1.0

ARGE wax 1: 0.78 ARGE wax 1: 1.38

57 59 30 37

4.1 5 2.6 4

12 14 3.5 9.3

41 38 24 35

Figure 5. Conversion of iron-catalyzed FT slurry bed reactor wax hydrocracked at 7.0 MPa, 0.55 h-1 LHSV, and a H2 to wax feed ratio of 1500/1 Ln/L using different temperatures.

3.4 2.7 6.9 3.8

Effect of Recycle Operation. Hydrocracking studies were conducted with tubular fixed-bed iron (ARGE) and slurry reactor wax to investigate the effect of recycling the C23+ residue to the hydrocracking reactor. Recycling was modeled in the way that the C23+ material resulting from once through operation, produced at the same operating condition as applied in the recycle mode of operation, was added in a certain ratio to the fresh wax feed. The major effect observed was increases in C23+ conversion in recycle mode and a reduction in diesel-to-naphtha ratio at a given temperature. ARGE iron reactor wax was hydrocracked using a NiMo/SiO2-Al2O3 catalyst. Table 6 shows the results for an operating temperature of 360 °C and 3.5 MPa hydrogen pressure increasing the recycle ratio. Table 6 shows that recycling the C23+ material increased the overall conversion. The reason for this behavior is that the recycled material is already isomerized, and it is easier to hydrocrack than the fresh wax feed, which has a higher linearity. The diesel formed tends to crack further to naphtha in recycle mode. Gas yields and naphtha yields increased with recycle while the diesel yield decreased. A drop in the diesel-to-naphtha ratio is observed. A further increase in recycle to fresh feed ratios increased the C23+ conversion; however, a more pronounced reduction in the diesel-to-naphtha ratio is the result. Higher temperatures (370 °C, see Table 7) led to an increase in conversion; however, the dieselto-naphtha ratios decreased by almost 50%. Increasing the space velocity from 0.5 to 1.0 h-1 resulted in a significant decrease in C23+ conversion. The combination of high operating temperatures and high recycle ratio decreased significantly the diesel-to-naphtha ratio. When iron slurry reactor wax was hydrocracked at 7.0 MPa hydrogen pressure using a fresh feed to recycle ratio of 1:0.65, a 5 °C lower operating temperature could be applied to achieve the same conversion as in once through mode (see Table 8). However, diesel selectivities decreased at the expense of naphtha selectivities while gas selectivities seemed to be less affected. Effect of Hydrocarbon Chain Length. A typical C23+ conversion for iron LTFT slurry reactor wax hydrocracked with a NiMo/SiO2-Al2O3 catalyst at 7.0 MPa, a LHSV of 0.55 h-1, and a 1500/1 Ln/L hydrogen

Figure 6. Diesel yield achieved at different C23+ conversion during hydrocracking of iron-catalyzed FT slurry bed reactor wax and condensates at 7.0 MPa, 0.55 h-1 LHSV, and a H2 to wax feed ratio of 1500/1 Ln/L. Table 8. Hydrocracking Iron-Catalyzed FT Slurry Bed Reactor Wax in Recycle Mode (Fresh Wax to Recycle 1:0.65) Using a Pressure of 7.0 MPa, 0.55 h-1 LHSV, and a H2/Wax Ratio of 1500/1 Ln/L mode of operation

T, °C

diesel to C23+ selectivity, wt % naphtha conv., ratio wt % C1-C4 C5-C9 C10-C22

once through 367

51

4.1

20

73

3.7

recycle

52

4.3

25

69

2.8

362

to wax feed ratio is shown in Figure 5. In the temperature range of 360-380 °C the C23+ conversion increased from 25% at 360 °C to almost complete conversion (97%) at 380 °C. Figure 6 shows the effect of carbon number distribution on the diesel yield at different C23+ conversion. Higher diesel selectivity is evident for the feed including the lighter material (condensates), which may be due to vaporization. This also implies that the condensates are not overcracked. Yields depend also on conversion and Figure 6 shows that heavier feed material may require more than one cracking event to reach the C10-C22 diesel range. This implies that the probability of naphtha forming is in relation to cracking frequency. Table 9 compares the hydrocracking of ARGE tubular fixed-bed iron reactor wax and slurry bed iron reactor wax at temperatures of 355 and 361 °C. ARGE reactor wax is lighter in nature, having a higher amount of lighter hydrocarbons (C13-C25) relative to slurry bed reactor wax, while the difference in concentration of hydrocarbons in the tail end (C50-C100) is not as

1802


Leckel

Table 9. Comparison of Tubular Fixed-Bed and Slurry Bed Iron-Catalyzed FT Reactor Wax Hydrocracked at 3.5 MPa, 0.55 h-1 LHSV, and a H2/Wax Ratio of 1500/1 Ln/L

wax

C23+ selectivity, wt % yield, wt % T, conv., °C wt % C1-C4 C5-C9 C10-C22 C1-C4 C5-C9 C10-C22

ARGE 355 361 slurry 355 361

47 84 41 77

7.7 7.9 9.2 11

12 17 13 20

82 77 78 70

6.3 6.6 3.7 8.4

5.7 14 6.2 15

38 64 32 54

significant. At the same operating temperature, higher C23+ conversions are achieved with ARGE reactor wax as compared to the slurry reactor wax. Generally higher C23+ conversion is achieved at the higher temperature. Significant is that a 6 °C increase in temperature had a major influence on the conversion showing that accurate temperature control is an important factor in wax hydrocracking. An increase in 6 °C operating temperature (355-361 °C) translated to a 78% increase in C23+ conversion for the ARGE iron reactor wax. A 53% increase in C23+ conversion was observed for the iron slurry reactor wax. The selectivities to gas (C1C4) and naphtha (C5-C9) increased for both wax feeds while the selectivities to diesel (C10-C22) decreased with increasing temperature. Sie49 reported that the reactivity of n-paraffins is expected to be proportional to the number of secondary carbon atoms per molecule, being Cn-6 for hydrocracking and Cn-4 for hydroisomerization. The heavier the paraffin, the more methylene groups are available in the longer chain paraffins to form carbenium ions at the acid sites, the higher the hydrocracking or hydroisomerization activity. Figures 7 and 8 display the apparent conversion for the C40-C80 paraffin molecules, resulting from hydrocracking of an iron LTFT slurry reactor wax with a NiMo/SiO2-Al2O3 catalyst at 0.57 h-1 LHSV and a 1500/1 Ln/L H2/wax ratio, calculated from high temperature GC analyses. It must be stressed that the observation and results are derived from a FT wax, a mixture of paraffins, meaning that individual paraffins were not studied. Since lighter paraffins can be formed from heavier ones, competitive adsorption of products derived from the hydrocracking or hydroisomerization of long chain reactants and uncracked reactants are taking place in parallel and are embedded in our results on the chain length effects. First-order reactions were estimated for the hydroconversion of the iron FT wax paraffins. Calemma et al.,50 who investigated the hydroisomerization and hydrocracking of long chain nalkanes (C16, C28, and C36 paraffins) on Pt/amorphous silica-alumina catalysts, determined first-order kinetics with regards to long chain hydrocarbons. Girgis and Tsao51 used first-order kinetic models for the hydroisomerization and hydrocracking of n-hexadecane. Figure 7 (hydrocracking done at 7.0 MPa) shows that for a C40 paraffin molecule the conversion increased from 40% at 360 °C to 83% at 370 °C. The conversion for a C80 paraffin increased from 87% at 360 °C to 99 at 370 °C. Figure 8 (hydrocracking done at 3.5 MPa) shows similar trends. Dauns and Weitkamp52 found in earlier studies, using a mixture of n-decane and n-dodecane, that a competi(49) Sie, S. T. Ind. Eng. Chem. Res. 1993, 32, 403. (50) Calemma, V.; Peratello, S.; Perego, C. Appl. Catal. A 2000, 190, 207-218. (51) Girgis, M. J.; Tsao, Y. P. Ind. Eng. Chem. Res. 1996, 35, 386.

Figure 7. Effect of chain length on the reactivity of paraffins during hydrocracking of iron-catalyzed FT slurry bed reactor wax at 7.0 MPa, 0.55 h-1 LHSV, and a H2 to wax ratio of 1500/1 Ln/L.

Figure 8. Effect of chain length on the reactivity of paraffins during hydrocracking of FT iron-catalyzed slurry bed reactor wax at 3.5 MPa, 0.55 h-1 LHSV, and a H2 to wax ratio of 1500/1 Ln/L.

tive adsorption takes place, with the longer chain hydrocarbon being the preferred component adsorbed. Compared to the pure alkanes, the conversion of ndecane in the mixture with n-dodecane is more inhibited than the conversion of the longer chain alkane. Longer chain alkanes are converted faster than the shorter chain component. Similar findings of the effect of chain length on the conversion were reported by Zhang et al.,53 who investigated longer chain normal paraffins in the C23-C26 carbon number range. An increase in conversion from 75.9 to 81.8 wt % was calculated for the C23 and C26 paraffins, respectively. Sie et al.,54 who studied the combination of chain length independent FischerTropsch synthesis with chain length dependent cracking, reported the reactivities for hydrocracking C10-C17 FT paraffins to increase sharply with increasing carbon number. The heat of adsorption for n-paraffins is known from literature to increase about linearily with increasing carbon number. As the adsorption coefficient is an exponential function of the heat of adsorption, one expects an approximately exponential increase with increase in carbon number for their respective relative reactivity. Above a certain chain length the effect could level off since the degree of paraffin adsorption is (52) Dauns, H.; Weitkamp, J. Chem.-Ing.-Tech. 1986, 11, 900-903. (53) Zhang, S.; Zhang, Y.; Tierney, J. W.; Wender, I. Fuel Process. Technol. 2001, 69, 59-71. (54) Sie, S. T.; Eilers, J.; Minderhout, J. K. Proc. 9th Int. Congr. Catal. 1988, 743-750.


expected to reach 100%. The result would be a chain length dependent reactivity. Denayer et al.55 concluded that, for hydroisomerization of n-alkane mixtures using Pt/USY zeolite catalysts, observed reaction rates in hydrocracking are strongly dependent on adsorption and that competition and chain length effects can be to a large extent ascribed to physisorption phenomena. It seemed that the paraffin reactivity after C50 leveled off and increased again after C70. Calemma et al.50 reported a similar observation, however, for shorter chain n-paraffins (C16, C28, and C36). In their study an increase in paraffin reactivity from the C16 paraffin to the C28 paraffin was noticed; however, they found a slight decrease for the C36 paraffin with respect to the C28 paraffin. The presence of diffusion phenomena was assumed, though the values of the activation energies of the rate constants did not confirm this observation. It is speculated presently how large paraffinic macromolecules adsorb onto the catalyst surface. Chain folding could occur on the catalyst surface such as described by Lee and Wegner56 for the crystallization behavior of linear and cyclic alkanes with carbon numbers of higher than 100. Since the catalyst support surface is of relatively polar nature, the fairly apolar alkane macromolecules could try to establish a more hydrophobic environment by chain folding or micellar-type of aggregation. The other possibility could be a kind of “tailtoward-the-liquid phase” of arrangement, whereby a part of the alkane molecule is adsorbed at the catalyst support surface, while the rest of the molecule is directed toward the hydrophobic liquid phase. The long alkane molecule could such be only partly in contact with the catalyst active sites in a pore structure, leading ultimately to a different cracking pattern as compared to shorter chain alkanes. Conclusions Fischer-Tropsch feed is different to crude oil because oxygenates adsorb strongly, olefins tend to crack faster on the catalyst, and the high degree of paraffin linearity requires hydroisomerization before cracking. It was consequently observed that recycling C23+ material led to a decrease in diesel-to-naphtha ratio and an increase in C23+ conversion than when operated at the same temperature in once-through mode. The hydrocracked material recycled with the fresh feed is more easily converted since it is already isomerized. If recyle operation is envisaged, the recycle material entry point should be considered further down the hydrocracking reactor. (55) Denayer, J. F.; Baron, G. V.; Jacobs P. A.; Martens. J. A. Phys. Chem. Chem. Phys. 2002, 2, 1007-1014. (56) Lee, K. S.; Wegner, G. Makromol. Chem., Rapid Commun. 1985, 6, 203-208.


Prehydrogenation of the wax feed prior to the hydrocracking reaction resulted in a significant increase in C23+ conversion, accompanied by an increase in naphtha selectivity and a decrease in diesel selectivity. Unhydrogenated iron reactor wax requires higher operating temperatures to achieve the same C23+ conversion compared to the hydrogenated feed. The reason for this is most probably the oxygenates present in the untreated wax, which seem to adsorb strongly on the acid sites. Pressure had an inversely proportional effect on the C23+ conversion. Higher pressure resulted in lower C23+ conversions and a reduction in naphtha selectivity while diesel selectivity increased. The most dominant effect of increased liquid hourly space velocity was an increase in diesel selectivity. Lower space velocities resulted in higher residence times of the feed material over the catalyst. The consequences were higher gas and naphtha yields an indication of an increase in isomerization and secondary cracking. High hydrogen-to-wax ratios led to improved C23+ conversion and a significant increase in diesel yields. Lighter feed and product material, probably due to a higher vaporization at given conditions, is flushed out of the reactor faster resulting in an enrichment of heavier hydrocarbons in the liquid phase. This shows that the vapor-liquid equilibrium plays an important role with regards to conversion, yields and diesel-tonaphtha ratios. Finally, a chain length dependent apparent n-paraffin reactivity was observed. Longer chain paraffinic hydrocarbons showed a higher conversion activity during the hydrocracking of iron FT wax. The conversion reactivity for a given n-paraffin increased with temperature or C23+ conversion and was found to be higher at the lower operating pressure. The reactivity, moving from carbon number 40 to carbon number 80, increased faster at higher pressure for a given temperature or C23+ conversion. At high C23+ conversion and at low operating pressure, the reactivity leveled off. It seems reasonable that lower activation energies are required for longer carbon chain length or in other words that an inverse function of activation energy could exist with respect to carbon chain length. Acknowledgment. Permission from Sasol Technology Research and Development to publish this work is appreciated, as well as the contributions of L. C. Ferreira (23/78), G. G. Swiegers (59/83, 267/96), and B. L. Mothebe (302/97). EF050085V

Hydrocracking of Iron-Catalyzed FischerâTropsch Waxes - Energy

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