Energy & Fuels 2008, 22, 231–236
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Hydrodeoxygenation of Heavy Oils Derived From Low-Temperature Coal Gasification over NiW Catalysts—Effect of Pore Structure Dieter Leckel* Fischer-Tropsch Refinery Catalysis, Sasol Technology Research and DeVelopment, Post Office Box 1, Sasolburg 1947, South Africa ReceiVed August 13, 2007. ReVised Manuscript ReceiVed September 28, 2007
The effect of the pore structure on the hydroprocessing of heavy distillate oils derived from low-temperature coal gasification residues was studied using four NiW catalysts with different pore size distributions. The hydroprocessing was conducted at a pressure of 17.5 MPa, a temperature range of 370-410 °C, and a 0.50 h-1 space velocity. The degree of hydrodeoxygenation (HDO) in terms of phenolics removal was influenced by the catalyst pore structure, with the most preferable peak pore diameter for HDO ranging between 6.8 and 16 nm. The catalyst with the highest volume of pores in the 3.5–6 nm range showed the lowest HDO activity. The apparent activation energies for the HDO reaction varied between 59 and 87 kJ/mol, whereby the lowest values are obtained for the catalysts with a peak pore diameter of 11 and 16 nm.
Introduction The refining of distillates derived from heavy petroleum, tar sands, oil shale crudes, or coal to fuels requires catalytic hydroprocessing to remove the sulfur, nitrogen, and oxygen from the heteroatom-containing molecules.1,2 The molecular composition of these materials with regard to the heteroatoms can vary substantially and may markedly influence the hydroprocessing kinetics. In coal, the oxygen content is generally much higher than the nitrogen and sulfur content, as is reflected in Table 1.3–6 Kinetic hydrodeoxygenation (HDO) studies have been conducted in batch and also in trickle bed reactors.7–12 Investigations of the HDO of industrial feeds, such as gas oils derived from thermal hydrocracking of heavy oils, have been carried out by Furimsky13 using a CoMo/γ-Al2O3 catalyst, and various studies of the kinetics of hydroprocessing of fractions of coal-derived liquids (SRC-II) were done by Gates and Petrakis et al.14–17 using standard commercial catalysts. * To whom correspondence should be addressed. Telephone: +27-16960-3830. Fax: +27-11-522-3975. E-mail:
[email protected]. (1) Gray, M. R. Upgrading Petroleum Residues and HeaVy Oils; Marcel Dekker: New York, 1994. (2) Landau, M. V. Catal. Today 1997, 36, 393–429. (3) Furimsky, E. Catal. ReV. Sci. Eng. 1983, 25, 421. (4) Furimsky, E. Appl. Catal., A 2000, 199, 147–190. (5) Burgess, C. E.; Schobert, H. H. Energy Fuels 1998, 12, 1212–1222. (6) Allen, D. T.; Grandy, D. W.; Jeong, K.-M.; Petrakis, L. Ind. Eng. Chem. Proc. Des. DeV. 1985, 24, 737. (7) Grandy, D. V.; Petrakis, L.; Young, D. C.; Gates, B. C. Nature 1984, 308, 175. (8) Li, C.-L.; Xu, Z.; Gates, B. C.; Petrakis, L. Ind. Chem. Eng. Proc. Des. DeV. 1985, 24, 92. (9) Lee, C.-L.; Ollis, D. F. J. Catal. 1984, 87, 325. (10) Satterfield, C. N.; Yang, S. H. J. Catal. 1983, 81, 335. (11) Odebunmi, E. O.; Ollis, D. F. J. Catal. 1983, 80, 56. (12) Artok, L.; Erbatur, O.; Schobert, H. H. Fuel Proc. Technol. 1996, 47, 153. (13) Furimsky, E. Fuel 1978, 57, 494. (14) Petrakis, L.; Ruberto, R. G.; Young, D. C.; Gates, B. C. Ind. Chem. Eng. Proc. Des. DeV. 1983, 22, 292. (15) Petrakis, L.; Young, D. C.; Ruberto, R. G.; Gates, B. C. Ind. Chem. Eng. Proc. Des. DeV. 1983, 22, 298. (16) Katti, S. S.; Westerman, D. B. W.; Gates, B. C.; Youngless, L.; Petrakis, L. Ind. Chem. Eng. Proc. Des. DeV. 1984, 23, 773.
The possible influence of the catalyst pore structure and molecular structure on hydroprocessing activity was first noted by Hoog18 in 1950. Anderson et al.19 developed a method for quantifying the catalyst pore properties necessary for most active desulfurization (HDS) of petroleum feedstocks, which also satisfied the HDS of coal-derived liquids reported by others.20 It was further reported that large pore bimodal catalysts have shown better hydrogenation activity for converting heavy coalderived liquids than small pore unimodal catalysts.21 Literature as such indicates the dependence of the HDO as well as HDS reaction on pore diameter. The feasibility of converting creosote heavy oils derived from low-temperature coal gasification residues into naphtha and diesel fuel blending stocks was reported previously.22 The study also indicated that maximum HDO was achieved with NiMo/ γ-Al2O3 catalysts having a pore size distribution in the range of 11–22 nm. The present paper describes the hydroprocessing of heavy creosote oils using NiW/γ-Al2O3 catalysts of different pore size distributions. The study investigated the relationship of HDO activity (phenolics removal), diesel yield, and the obtained diesel properties to the specific pore size/distribution of the catalysts. Experimental Section Catalysts. The four NiW/γ-Al2O3 catalysts used in this investigation were prepared at Sasol using co-impregnation of aqueous Ni(NO3)2 · 6H2O (Aldrich) and ammonium metatungstate [(NH4)6H2W12O40 · nH2O, Fluka, g85% WO3] solutions sufficient to load the calculated mass percent of oxide. After calcinations at (17) Grandy, D. V.; Petrakis, L.; Li, C.-L.; Gates, B. C. Ind. Chem. Eng. Proc. Des. DeV. 1986, 25, 40. (18) Hoog, H. J. Inst. Petrol. 1950, 36, 738. (19) Anderson, J. A., Jr. et al. U.S. Patent 2,890,162, assigned to Esso Research and Engineering, June 9, 1950. (20) Sooter, M. C.; Crynes, B. L. Ind. Eng. Chem. Prod. Res. DeV. 1975, 14 (3), 199–204. (21) Tischer, R. E.; Narain, N. K.; Stiegel, G. J.; Cillo, D. L. J. Catal. 1985, 95, 406–413. (22) Leckel, D. Energy Fuels 2006, 20 (5), 1761–1766.
10.1021/ef700493b CCC: $40.75 2008 American Chemical Society Published on Web 11/19/2007
232 Energy & Fuels, Vol. 22, No. 1, 2008
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Table 1. Comparison of Elemental Composition (wt %) of Conventional Crude Oil, Oil Shale Crude, a Coal, and a Coal-Derived Liquid
carbon hydrogen sulfur nitrogen oxygena a
conventional crude oil3,4
oil shale crude3,4
Pittsb. No. 8 Penn state coal5
SRC II heavy distillate6
85.2 12.8 1.8 0.1 0.1
85.9 11.0 0.5 1.4 1.2
84.75 5.66 0.83 1.39 7.37
89.5 7.7 0.4 1.1 2.3
Oxygen by difference.
Table 2. Catalyst Properties NiW-1
NiW-2
NiW-3
NiW-4
physical properties 190 154 197 200 BET surface area (m2/g) peak pore diameter (nm) 11 16 5.9 6.8 pore volume (cm3/g) 0.454 0.736 0.457 0.571 pore volume distribution (vol %) 3.5–6 nm 16.7 6.3 89.5 28.2 7–11 nm 61.7 30.0 2.5 28.0 12–22 nm 15.9 44.6 1.3 13.7 23–35 nm 1.1 5.2 0.2 8.6 >35 nm 4.6 14.1 6.5 21.5 chemical properties NiO 7.9 7.0 7.7 7.0 WO3 22.4 22.0 24.6 23.2 SiO2 4.0 W/Ni atom ratio 0.92 1.01 1.03 1.07 0.88 0.55 0.85 0.67 bulk density (g/cm3) metal oxides loaded (g) NiO 20.0 15.2 19.0 13.6 WO3 57.0 47.8 60.6 45.1
Table 3. Hydrocarbon Molecular Composition of a Commercial Heavy Creosote Oil Diesel Hydrotreater Feed compound class
yield (mass %)
sulfur oxygen nitrogen aromatics (no heteroatoms) paraffins olefins unknowns
0.44 26 2.5 34 7.5 0.6 29
500 °C for 4 h, the catalysts (300 mL, 1.5 mm extrudates) were mixed (1:1) with an inert diluent (silicon carbide) before being loaded into the reactor tube. Table 2 show the physical properties of the catalysts used in this study. Catalyst Activation. The catalysts were predried in nitrogen at 125 °C and 1 bar for 8 h prior to sulfidation. Then, nitrogen was replaced by hydrogen (100 L/h flow rate), and the reactor was pressurized to 3.0 MPa. A solution of 2 wt % dimethyldisulfide (DMDS) dissolved in a C9-C11 paraffin was used as the sulfiding agent. The catalyst was first wetted with the sulfiding mixture, whereafter 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, and a 2000 vppm H2S concentration was seen as sufficient to increase the temperature further to 350 °C, where it was kept for another 8 h. Thereafter, the reactor pressure was raised and hydrogen was introduced at the desired feed rate. Feed Material. A typical composition of compound classes present in a heavy creosote oil hydrotreater feed, analyzed at our laboratories and also used in our previous study,22 is shown in Table 3. Aromatic hydrocarbons containing no heteroatoms are the most abundant compound class identified, followed by oxygenates, paraffins, nitrogen components, olefins, and sulfur components. The unknown fraction is relatively large and due to the highly complex nature of the feed composition. Table 4 shows the properties of the feed used in this study.
Table 4. Properties of the Heavy Creosote Oil Feed Used in This Study property density at 20 °C (kg m-3) bromine number (g of Br2/100 g) total N (wppm) total S (wppm) phenolics (wppm) ASTM D-2887 (°C) 10 vol % 30 vol % 50 vol % 70 vol % 90 vol % 95 vol %
1010 81 9000 3000 46 000 222 251 280 318 372 390
The largest heteroatom fraction of the composite feed shown in Table 3 is oxygenates. Phenolics are the main constituents (about 75%) in this fraction, followed by benzofurans (about 12%), naphthalenols (7%), indanols (about 3%), dibenzofurans (2%), and furanols (1%). Mono- and dimethylphenols are the major compounds in the phenolic fraction. Equipment and Procedures. A bench-scale trickle bed reactor was used for the isothermal studies. The reactor consists of a tube of 27.6 mm inner diameter and 1.8 m in length. The catalyst was loaded in the middle section of the reactor tube, and silicon carbide (0.5 mm size) was used to fill the voids between the catalyst particles to avoid channeling. Inert packing (glass beads) above the catalyst bed was used to preheat the feed up to the reaction temperature. Six thermocouples were placed inside the thermocouple well, which runs through the center of the reactor over its entire length. The reactor was heated electrically by three heater elements placed along the reactor tube. Liquid feed and hydrogen entered concurrently at the top of the reactor. The gas and liquid products were separated in the last section of the reactor setup. The liquid product was collected in a catch pot, and the gaseous light hydrocarbons were passed through a cooling coil at 0 °C. This condensed liquid was collected in a second gas liquid separator. A gas sampling point was installed on the low-pressure side of the reactor system. Catalytic activity was monitored by drawing product samples from the reactor effluent after steady-state conditions were reached, typically after a period of 24 h on stream. The following 8 h period was used to collect a representative sample for product analysis. The sulfided state of the catalyst was monitored by H2S tail gas analyses. H2S levels of 200 vppm were observed to be sufficient to keep the catalyst active. The mass balance was checked with a collection of tail gas, liquid product, and condensed lighter hydrocarbons. Condensed light hydrocarbons were kept refrigerated prior to analysis. Mass balances of between 96 and 102% were obtained for all hydrotreating experiments. Analyses. The following analyses were conducted: (i) CHNS analyses using a Leco 600 analyzer, (ii) bromine number American Society for Testing and Materials (ASTM) D-1159, (iii) ASTM D-2887 distillation, (iv) density ASTM D-1298, (v) aniline point as determined by following the procedures of ASTM D-611 and cetane number according to ASTM D-613. The aniline point is an indication of the aromaticity of the sample. The higher the aniline point, the lower the aromatic content of the sample. In addition, the higher the aniline point, the higher the determined cetane number. The determination of the phenolic hydrocarbon content in feed and products used for the kinetics experiments was done via a caustic extraction procedure. The phenolic components were thereby extracted into the caustic layer, then acidified to a pH 3, and thereafter extracted into a diisopropopylether (Dipe) layer, which was then analyzed by gas chromatography mass spectrometry (GC-MS) on a 60 m FFAP capillary column. Quantification was achieved via GC-flame ionization detector (FID) and area normalization internal standards using Dietz
Hydrodeoxygenation of HeaVy Oils
Figure 1. Diesel/naphtha ratios obtained during hydroprocessing of heavy distillate oil using NiW/γ-Al2O3 catalysts at a pressure of 17.5 MPa, a 0.50 h-1 LHSV, and a hydrogen/oil ratio of 1500:1.
response factors.23 The following GC conditions were used: H2 carrier gas, 0.53 mL/min carrier gas flow, 50 m × 0.2 mm × 0.50 µm Pona column; column profile: initial temperature of 35 °C, 2 °C/min heating rate, final temperature of 290 °C, duration for 60 min, FID detector temperature of 300 °C, and 0.50 µL injection volume. The maximum experimental error determined for the product yields was 3.1% (average of 1.3%), while that for the phenolics removal in the liquid product was 5.3% on average.
Results In this paper, we concentrate on the removal of phenolics, which are amongst the most difficult compounds to deoxygenate, especially hindered phenols, such as ortho-substituted alkylphenols.4,24 The term HDO is synonymous with phenolics removal. Effect of the Temperature. The effect of the temperature was studied at a constant pressure of 17.5 MPa, a 0.50 h-1 liquid hourly space velocity (LHSV), and a hydrogen/oil ratio of 1500:1. The temperature range of 360-425 °C was covered in the experiments. The effect on diesel selectivity is presented in Figure 1 and Table 5. The naphtha fraction increased with an increasing operating temperature, while the diesel fraction decreased, which was observed generally for all catalysts tested. The most significant effect of the temperature with regard to product yields was observed for the NiW-3 and NiW-4 catalysts, where the diesel fraction decreased quite significantly at temperatures exceeding 400 °C. In the one case (NiW-3), the diesel fraction decreased from 69 to 60 mass % and, in the other case (NiW-4), from 64 to 58 mass %. Specifically, with the NiW-2 catalyst, the lowest naphtha fraction and the highest diesel fraction were obtained, while the lowest diesel yields were obtained with the NiW-4 catalyst (Figure 1 and Table 5). The naphtha/ diesel ratios produced by the NiW-1, NiW-3, and NiW-4 catalysts are quite similar, with the NiW-2 catalyst retaining a higher ratio at high operating temperatures compared to the other catalysts. It seems that the catalysts with a more random distribution of pores (NiW-1, NiW-2, and NiW-4) show a rather nonlinear relationship between the diesel/naphtha ratio and reaction temperature. For the catalyst NiW-3, consisting of mainly micropores, a rather linear relationship is observed. The selectivity to diesel at a lower reaction temperature is relatively high for the microporous NiW-3 system, but it (23) Dietz, W. A. J. Gas Chromatogr. 1967, February, 68–71. (24) Rollman, L. D. J. Catal. 1977, 46, 243.
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decreases more rapidly (by 37%) from 370 to 410 °C compared to the other catalysts (i.e., by only 23% for the NiW-2 catalyst). The catalyst NiW-2 with the highest pore volume, the lowest surface area, and the highest peak pore diameter showed the highest selectivity to diesel. The cetane number of the diesel fractions generally increased with an increasing hydrotreating temperature. The NiW-2 and NiW-4 catalysts, having a higher percentage of pore volume residing in the larger pore size range, show a better response to temperature. The cetane numbers increased by 22%, reaching at 409 °C values of 50 and 49, respectively (Figure 2). The microporous catalyst (NiW-3) showed again the least response to temperature with regard to cetane improvement. A temperature increase of 40 °C resulted in a mere 12% increase in the cetane number, from 33 to 37 (see Figure 2), indicating a low activity for aromatics hydrogenation. The aniline points (see Table 5) mirrored this performance, where a temperature increase generally resulted in an increase in the aniline point. Higher aniline points mean lower aromatic concentrations, and the lower the aromaticity of a diesel fuel, the higher is its ignition quality, represented by the cetane number. The HDO activity increased with temperature (Table 5). The highest activity for phenolics removal was observed for the NiW-1 catalyst, while the lowest activity was noticed for the microporous NiW-3 catalyst. Comparing the HDO activity of the catalysts at constant reaction conditions results in a HDO activity order of NiW-1 > NiW-2 > NiW-4 > NiW-3. Figure 3 shows the relation between the phenolics removal in the total liquid product and catalyst pore size during hydroprocessing of creosote heavy oil. The average pore size range of 11-16 nm generally corresponds to high HDO activity. NiW catalysts with a peak pore diameter larger than 16 nm were unfortunately not tested. However, we assume that pores up to 22 nm positively influence the HDO activity of the NiW catalysts, because in our previous study,22 NiMo catalysts up to an average pore size of 22 nm showed significant HDO activity. Upgrading of the creosote heavy oil also showed that, as the catalysts pore size increases, the cetane number increases (Figure 4), which is the result of the heavier molecules, in particular, the larger condensed aromatics, being progressively more hydrogenated. The cetane number did not improve further with catalysts having a peak pore diameter larger than 6.8 nm, and a minimum for the cetane number was observed at the peak pore diameter of 11 nm, specifically at higher temperatures. Kinetic Analysis of HDO Data. Because industrial feeds contain numerous components and the transformation of these compounds over the various catalysts yield compounds with new adsorption characteristics, it is clear that the formulation of an exact reaction rate equation is not feasible. The empirical approach, namely, to use the power-law form of the rate eq 1 was found to be most appropriate rA ) -
dCPh n ) kCPh dt
(1)
where k is the reaction rate constant, t is the space time (reciprocal of space velocity), CPh is the concentration of the phenolics in the bulk phase (in ppm), and n is the order of the reaction with respect to the phenolic (oxygenate) species. The reaction was performed with a large hydrogen excess, and a constant hydrogen partial pressure can therefore be assumed.
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Table 5. Hydroprocessing Results Achieved at a Pressure of 17.5 MPa, a 0.50 h-1 LHSV, and a Hydrogen/Oil Ratio of 1500:1 mass % catalyst temperature (°C)
phenolics in liquid product (wppm)
naphtha C5-165 °C
kerosene 165–200 °C
diesel 200–385 °C
aniline point (°C)
145 34 1
17 19 22
9.1 11 12
67 65 63
41 46 55
179 42 1
15 16 18
12 10 14
71 70 65
54 57 65
2262 425 43
17 19 23
8 10 11
69 67 60
32 41 43
395 26 1
18 19 22
12 14 16
64 63 58
50 61 63
NiW-1 370 386 409 NiW-2 370 386 409 NiW-3 370 386 405 NiW-4 370 387 410
For a first-order reaction describing the reactor as a whole and assuming plug flow, the rate equation can be written as follows:
( )
-ln
CA ) kLHSV C0
(2)
The dependence of k on temperature follows the Arrhenius equation (3) k ) AeEa/RT where A is the frequency factor and Ea is the apparent activation energy. Representative Arrhenius plots for the phenolics (oxygenate) removal based on the assumption of a first-order reaction are shown in Figure 5. The intrinsic apparent activation energies found for the phenolics removal in the total liquid product using a constant pressure and space velocity were determined by fitting the experimental data of Table 5 to eq 2. The corresponding activation energies were calculated directly from the slope of regression lines in the Arrhenius plots and are presented in Table 6. The HDO activation energies found for the NiW catalysts ranged between 59 and 87 kJ/mol, with a value of 58.7 kJ/mol found for the NiW-1, 62.2 kJ/mol for the NiW-2, 86.8 kJ/mol for the NiW-3, and 70.7 kJ/mol for the NiW-4 catalysts, respectively. Discussion Phenolic compounds were removed in excess of over 99 mass % from the heavy creosote oil at 17.5 MPa and a temperature
Figure 2. Cetane numbers determined for the diesel fraction obtained during hydroprocessing of heavy distillate oil using NiW/γ-Al2O3 catalysts at a pressure of 17.5 MPa, a 0.50 h-1 LHSV, and a hydrogen/ oil ratio of 1500:1.
of 370 °C using the NiW-1 and NiW-2 catalysts. The NiW-3 catalyst achieved a 95 mass % removal at the same conditions, whereas the NiW-4 catalyst showed the lowest activity (85 mass % phenolics removal). An increasing operating temperature generally decreased the diesel yield (see Table 5). The lowest diesel yield was observed with the NiW-4 catalyst at 410 °C. This can be explained by the fact that the latter catalyst is based on an acidic SiO2-Al2O3 support (see Table 2), which should result in some hydrocracking activity. The extent of hydrocracking that occurred, however, was very limited because no extensive naphtha formation was observed in comparison to the other catalysts. The presence of weak acid sites or the adsorption of basic nitrogen on the acid site, thereby inhibiting the cracking activity of the NiW-4 catalyst, is most likely the reason for the limited bifunctional activity. The hydrogenation activity of the NiW catalysts, reflected by the values for the cetane number (Figures 2 and 4) and aniline point in the diesel fraction (see Table 5), follows the sequence NiW-2 and NiW-4 > NiW-1 > NiW-3. The better hydrogenation activity of the NiW-2 and NiW-4 catalysts can not be related to the metal concentration, because both catalysts have the lowest concentration of metal oxides loaded on a weight basis (see Table 2). According to thermodynamic investigations, high hydrogen partial pressures and low temperatures favor hydrogenation, while low hydrogen partial pressures and high temperatures
Figure 3. Effect of the pore size on the phenolics removal (HDO) during hydroprocessing of creosote heavy oil with the NiW catalysts at 17.5 MPa, 370-409 °C, 0.50 h-1 LHSV, and a H2/oil ratio of 1500:1.
Hydrodeoxygenation of HeaVy Oils
Energy & Fuels, Vol. 22, No. 1, 2008 235
Figure 4. Effect of the pore size on the diesel cetane number during hydroprocessing of creosote heavy oil with the NiW catalysts at 17.5 MPa, 370-409 °C, 0.50 h-1 LHSV, and a H2/oil ratio of 1500:1.
Figure 5. Arrhenius plots for HDO (phenolics removal) of heavy creosote oil calculated from data in Table 6. Conditions applied were a pressure of 17.5 MPa, temperatures of 370–410 °C, a 0.50 h-1 LHSV, and a hydrogen/oil ratio of 1500:1. Table 6. Apparent Activation Energy Obtained for the HDO (Phenolics Removal) in the Total Liquid Product Using a Pressure of 17.5 MPa, Temperatures of 370–409 °C, a 0.50 h-1 LHSV, and a Hydrogen/Oil Ratio of 1500:1
catalyst
peak pore diameter (nm)
apparent activation energy Ea (kJ/mol)
reaction order n
R2
NiW-1 NiW-2 NiW-4 NiW-3
11 16 6.8 5.9
58.7 62.2 70.7 86.8
1.0 1.0 1.0 1.0
0.9947 0.9932 0.9857 0.9959
favor dehydrogenation or the formation of aromatics.25 Exceeding this critical temperature at given reaction conditions leads to aromatization again, which was not yet observed at our reaction conditions, because aniline points and the cetane number did not deteriorate up to the maximum operating temperature applied. The hydrogen pressure of 17.5 MPa was still sufficient up to the temperature of 410 °C to reside the hydrogenation reaction in the regime of kinetic control and below the thermodynamic limitations. Furthermore, the W/Ni atomic ratios determined for all four catalysts are just about the same (see also Table 2). To be noted, however, is that both catalysts, NiW-2 and NiW-4, have the highest pore volume of pores above a size of 35 nm. These catalysts with larger pores seem to promote the conversion and (25) (a) Hengstebeck, G. J. Petroleum Processing; McGraw-Hill: London, U.K., 1959; pp 276–278. (b) Frase, H. F. Handbook of Commercial Catalysts: Heterogeneous Catalysts; CRC Press: Boca Raton, FL, 2000; pp 301–305.
hydrogenation of bulkier condensed aromatics, thereby improving the cetane number of the diesel fraction. This is consistent with the fact that a reaction that is more influenced by diffusion resistance proceeds more easily with large pore catalysts than a reaction in which the contribution of diffusion resistance is small.26 The reverse holds for catalysts with small pores. The NiW-3 catalyst, which has the highest pore volume of small pores (3.5–6 nm), produced the diesel fraction with the lowest cetane numbers, ranging between values of 33 and 37 (see Figure 2). Pores below 6 nm seem therefore not to be beneficial for the hydrogenation of larger aromatic ring systems. This is due to the fact that the accessibility for large condensed aromatics to the active sites residing inside the pores is facilitated with increasing pore size. Larger molecules have a better access to catalytic sites if the catalyst has larger pores.27 The surface area was found to not be an influencing factor in this study regarding the HDO reaction or the aromatic saturation activity (cetane number improvement). The catalyst activity for phenolics removal (see Table 5) followed the sequence NiW-1 and NiW-2 > NiW-4 > NiW-3. This is a different order than noted for the cetane number. If we assume that the pore size distribution influenced the HDO activity, then it can be concluded that the distribution of pores in the range of 6.8-16 nm is the most preferable for HDO of phenolics (see Figure 3). Similar observations were reported previously,28,29 where the pore size range of 10–20 nm appeared to contribute to the maximum HDO activity of heavy coalderived liquids and oil sand-derived asphaltenes, while larger pores (>20 nm) were beneficial to the conversion of the larger molecules into oils. During hydroprocessing of solvent-refined coal with catalysts having different pore size distributions and metal loadings, it was found30 that catalysts with a bimodal pore size distribution having an average micropore diameter of 12 nm and an average macropore diameter of 430–450 nm appeared to be more effective in removing oxygen than catalysts with an average pore diameter below 9 nm. Metal loading did not affect the oxygen removal. These observations are very much in line with our findings for the preferable pore size range for the HDO reaction, which is different from that for the hydrogenation of the larger condensed aromatics. The HDO apparent activation energies found for the NiW catalysts support the concept of a beneficial pore size range for HDO. The NiW-1 and NiW-2 catalysts having a peak pore diameter of 11 and 16 nm showed the lowest apparent activation energy (between 59 and 62 kJ/mol) for the HDO reaction. The NiW-3 and NiW-4 catalysts, having peak pore diameters outside the “optimal” pore size range, showed higher HDO activation energies, between 71 and 87 kJ/mol. Conclusion HDO of heavy oils derived from low-temperature coal gasification strongly depends upon the pore structure of the NiW/γ-Al2O3 catalysts. (26) Kobayashi, S.; Kushiyama, S.; Aizawa, R.; Koinuma, Y.; Inoue, K.; Shimizu, Y.; Egi, K. Ind. Eng. Chem. Res. 1987, 26, 2245–2250. (27) Sullivan, R. F.; Boduszynski, M. M.; Fetzer, J. C. Energy Fuels 1989, 3, 603–612. (28) Song, C.; Nihonmatsu, T.; Nomura, M. Ind. Eng. Chem. Res. 1991, 30, 1726–1734. (29) Song, C.; Hanaoka, K.; Nomura, M. Energy Fuels 1992, 6, 619– 628. (30) Stiegel, G. J.; Tischer, R. E.; Polinski, L. M. Ind. Eng. Chem. Prod. Res. DeV. 1983, 22 (3), 411–420.
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A pore size range of 6.8–16 nm proved to be the preferable one for phenolics removal. Catalysts with pore sizes below 6.8 nm produced a diesel fraction with inferior cetane numbers, indicating that smaller pore sizes are insufficient for the hydrogenation of larger condensed aromatics. The apparent activation energies determined for phenolics removal support the existence of an optimum pore size range for the catalyst. The lowest HDO activation energies were determined for the catalysts having peak pore diameters ranging between 11 and 16 nm. A NiW catalyst with high HDO activity should preferably have a high volume of pores in the range of 6.8–11 nm. Adding
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pores in the region above 11 nm increases the hydrogenation activity with regard to bulkier aromatics. For the NiW catalyst with the highest pore volume in the 3.5-6 nm range, the lowest HDO activity and the lowest hydrogenation activity for large aromatics were obtained. Acknowledgment. The author gratefully acknowledges the technical contributions of A. Brodziak, K. Kriel, and G. G. Swiegers and appreciates the permission of Sasol Technology Research and Development to publish the work. EF700493B
