Energy & Fuels 2006, 20, 1761-1766
1761
Catalytic Hydroprocessing of Coal-Derived Gasification Residues to Fuel Blending Stocks: Effect of Reaction Variables and Catalyst on Hydrodeoxygenation (HDO), Hydrodenitrogenation (HDN), and Hydrodesulfurization (HDS) Dieter Leckel* Fischer-Tropsch Refinery Catalysis, Sasol Technology Research and DeVelopment, P. O. Box 1, Sasolburg 1947, South Africa ReceiVed January 25, 2006. ReVised Manuscript ReceiVed May 31, 2006
Gas liquors, tar oils, and tar products resulting from the coal gasification of a high-temperature FischerTropsch plant can be successfully refined to fuel blending components by the use of severe hydroprocessing conditions. High operating temperatures and pressures combined with low space velocities ensure the deep hydrogenation of refractory oxygen, sulfur, and nitrogen compounds. Hydrodeoxygenation, particularly the removal of phenolic components, hydrodesulfurization, and hydrodenitrogenation were obtained at greater than 99% levels using the NiMo and NiW on γ-Al2O3 catalysts. Maximum deoxygenation activity was achieved using the NiMo/γ-Al2O3 catalyst having a maximum pore size distribution in the range of 110-220 Å. The NiMo/γ-Al2O3 catalyst, which also has a relatively high proportion of smaller pore sizes (35-60 Å), displays lower hydrogenation activity.
Introduction A comparison of the recoverable hydrocarbonaceous sources known to date clearly shows that the coal reserves exceed by far all other carbon sources available.1 Because of this, the coalbased Fischer-Tropsch (FT) synthesis of fuels has been experiencing a revival in recent years, and China’s and other countries’ current interest in FT technology, through which they are able to tap into their vast coal reserves for the production of fuels and chemicals, underlines these developments. Sasol has commercially operated FT plants since 1955. Currently, two types of FT processes are in operation. The lowtemperature FT process is operated by Sasol at Sasolburg, South Africa, producing predominantly high molecular mass linear paraffins and waxes. Whereas, the Sasol Synfuels plants in Secunda are using the high-temperature FT (HTFT) process with iron-based catalysts for the production of fuels and chemicals. Until recently, for all of the Sasol FT plants, the primary source of syngas was the gasification of coal, using dry-ash Lurgi gasifiers. By mid 2004, coal was replaced by natural gas at Sasol 1, which will also be used to supplement the coal utilization at the Sasol Secunda Synfuels plants to allow for capacity expansion. Previous publications1-3 have described the Sasol FT processes in detail, and only a summary is presented here. Coal is first completely gasified with steam and oxygen in the Lurgi dry-ash gasifiers. The resulting gas is cooled, and gas liquor, tar oils, and condensed tar are separated from the synthesis gas. These coal-based FT plants therefore require a tar refinery to deal with the coal pyrolysis products, and the configuration of * Author tel: +27 16 960-3830, fax: +27 11 522-3975, e-mail:
[email protected]. (1) Dry, M. E. Appl. Catal., A: General 1999, 189, 185-190. (2) Dry, M. E. Applied Industrial Catalysis Publisher: Location, 1983; Vol. 2, Chapter 5, pp 167-213. (3) Dry, M. E. Catal. Today 2002, 71, 227-241.
this refinery is dictated by the nature of these products. High levels of aromatics are obtained which contain heteroatoms such as sulfur, nitrogen, and oxygen. Tar acids are also part of these pyrolysis products, which are rich in oxygenates such as phenols, cresols, xylenols, and naphthols. The naphtha and the middle distillate fractions of the tar are hydroprocessed separately to remove the heteroatoms. In the case of the naphtha, the aromaticity of the components should be retained because of its high octane number. Aromatics in the diesel boiling range provide the required density but have low cetane numbers. Severe hydroprocessing is required to remove refractory heteroatoms in this higher boiling fraction, which also results in the loss of aromaticity. However, this benefits the cetane number and therefore the ignition quality of the middle distillate. Saturation of the substituted monoaromatics present in the feed produces cycloalkanes (naphthenes) in the naphtha boiling range, which have lower octane numbers than the aromatics. These compounds are, however, valuable catalytic reformer feed material. The integration, therefore, of a tar refinery into an FT plant creates opportunities for fuel processing and in downstream fuel blending. At Sasol, the residual fractions (gas liquors, tar oils, and condensed tars) from the coal gasification step have been upgraded since the early 1980s. Two coal tar naphtha hydrotreaters and a creosote (tar diesel) hydrogenation unit are in operation. The latter unit is used for the processing of the higher boiling hydrocarbons to mainly diesel and is operated in the temperature range of 280-380 °C and at a pressure of 18.5 MPa. It consists of a fixed-bed reactor train having four reactors in series. The first two reactors are used as pretreating reactors for hydrodemetalation, while the main hydrogenation function takes place in the last two reactors. As feed for the commercial unit, a mixture of medium creosote (MC), heavy creosote (HC), and residue oil (RO) is used, with the balance being coker gas oils or depitched tar
10.1021/ef060034d CCC: $33.50 © 2006 American Chemical Society Published on Web 07/06/2006
1762 Energy & Fuels, Vol. 20, No. 5, 2006
Leckel Table 1. Catalyst Characterization
catalysta NiMo-1 NiMo-2 NiW-1 CoMo a
MoO3 wt %
WO3 wt %
NiO wt %
22.4
3.6 3.0 7.9
18.8 15.4 13.5
CoO wt %
BET area m2/g
pore vol. cm3/g
5.0
158 225 190 200
0.537 0.490 0.454 0.700
35-60 Å
pore vol. distribution cm3/g 60-110 Å 110-220 Å
>220 Å
0.056 0.137 0.076 n.d.b
0.225 0.280 0.280 n.d.b
0.023 0.061 0.026 n.d.b
0.233 0.012 0.072 n.d.b
With γ-alumina as support for all catalysts. b n.d.: not determined.
acids, depending on the refinery operation. The feed has a low hydrogen-to-carbon ratio and high nitrogen, oxygen, and polynuclear aromatic concentrations. Similar compositions are found in liquids derived from direct coal pyrolysis.4 Coal-derived liquids and fractions derived from shale oil have high nitrogen and oxygen levels, which are usually much higher than those found in most petroleum crudes. Nitrogen-to-sulfur ratios of 2:1 and higher are commonly found in the coal gasification residues produced in a typical HTFT refinery. Because of the high levels of nitrogen and polynuclear aromatics, hydrogen consumption in such coal tar naphtha and creosote hydrogenation units is typically high. This is because the preferred reaction pathway for the hydrodenitrogenation (HDN) reaction involves initially the saturation of the aromatic ring carrying the heteroatom (especially for compounds in which nitrogen is part of the aromatic ring) before C-N bond scission takes place.5-7 Hydrodeoxygenation (HDO) of the organo-oxygen compounds proceeds either by direct oxygen removal or through prior hydrogenation of the aromatic ring.8 Apart from the oxygen-containing compounds already indicated above, other oxygenates found in coal-derived liquids are ketones, aryl ethers, and benzofuranes. Similarly, the hydrodesulfurization (HDS) reaction network proceeds via direct hydrogenolysis or through hydrogenation of the aromatic ring prior to sulfur removal. High hydrogen pressures and low space velocities favor the hydroprocessing of coal-derived liquids8-10 and, combined with higher reaction temperatures, kinetically favor heteroatom removal as well.11 A variety of catalyst systems have been tested in the past for the hydroprocessing of material high in nitrogen and oxygen content. Sulfided NiMo supported on γ-alumina has until now been the catalyst of choice in the Sasol operations. The present study describes the pilot-scale hydroprocessing of coal gasification residues typically associated with an HTFT plant using commercially available CoMo, NiW, and NiMo hydroprocessing catalysts supported on γ-alumina. The study tried to establish if coal gasification residues can be successfully integrated into a HTFT refinery to add value to the final petrol and diesel pool. The effects of operating conditions and catalyst type on the HDO, HDN, and HDS reactions and the total liquid product properties, in particular, their fuel properties, were considered. It should also be established whether the choice of NiMo catalysts for the upgrading of coal-derived tars is the correct one. (4) Teo, K. C.; Watkinson, A. P. Fuel 1990, 69 (10), 1211-1218. (5) Girgis, M. J.; Gates, B. C. Ind. Eng. Chem. Res. 1991, 30, 20212058. (6) Delmon, B.; Froment, G. F. Stud. Surf. Sci. Catal. 1987, 34, 39. (7) Cocchetto, T. F.; Satterfield, C. N. Ind. Eng. Chem. Process Des. DeV. 1976, 15, 272. (8) Berg, L.; McCandless, F. P.; Bhatia, H.; Yeh, A. G. Energy Prog. 1982, 2, 236-238. (9) Garg, D.; Tarres, A. R.; Guin, J. A.; Lee, J. M. Prepr. Pap.sAm. Chem. Soc., DiV. Fuel Chem. 1980, 25 (1), 164-175. (10) Satterfield, C. N.; Carter, D. L. Ind. Eng. Chem. Process Des. DeV. 1981, 20, 538-540. (11) S¸ ahin, T.; Berg, L.; McCandless, F. P. Ind. Eng. Chem. Process Des. DeV. 1984, 23, 495-500.
Experimental Section Reactor System. A bench-scale fixed-bed reactor, operating in down-flow mode, was used for the isothermal studies. The reactor consists of a tube of 47.6 mm i.d. and 1.8 m length. The catalyst (260 mL) was loaded in the middle section of the reactor tube, and sand (35-50 mesh) was used to fill the voids between the catalyst particles to avoid channeling. Inert packing (glass balls) above the catalyst bed was used to preheat the feed up to the reaction temperature. To support the catalyst, a fine mesh grid was welded onto the thermocouple well, which runs through the center of the reactor over its entire length. Six thermocouples were placed inside the well at different lengths along the reactor. 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. The liquid product samples were subjected to the following analyses: (i) American Society for Testing and Materials (ASTM) distillation (D 86, D 1160, and D 2887), (ii) Antek CHNS analyses, (iii) capillary column GC-FID, (iv) UV absorption for the detection of aromatics, (v) bromine number, (vi) acid number, (vii) density, (viii) research octane number (RON), and (ix) cetane number (CN) ASTM D 613. The gas (C1-C4) samples were subjected to GCFID analyses. The phenolic components were also determined by GC-FID analyses. Catalysts. The catalysts used in these studies were commercially available cobalt-molybdenum, nickel-molybdenum, and nickeltungsten supported on γ-alumina (see Table 1). All catalysts were available in the form of extrudates of 1.5 mm diameter. Pore volumes were measured with a Micromeritics mercury porosimeter. The physicochemical properties of the catalysts are presented in Table 1. 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 LN/h flow rate), and the reactor was pressurized to 3.0 MPa. Light cycle oil, in earlier studies, and 2 wt % dimethyl disulfide dissolved in a C9-C11 paraffin mixture, in later studies, were used as sulfiding agents. The catalyst was first wetted thoroughly with the sulfiding mixture using the maximum pump speed. At a decreased pump rate, the temperature was ramped hourly from 125 °C in steps of 25 °C to 250 °C, at which temperature it was kept for 4 h. H2S breakthrough was monitored in the off gas by means of Draeger tubes. A 2000 vppm H2S concentration in the off gas was seen as sufficient in order to proceed in increasing the temperature further to 350 °C, at which temperature it was kept for another 8 h. Thereafter, the reactor temperature was decreased to 240 °C. When sulfiding was completed, the reactor pressure was raised and hydrogen was introduced at the desired feed rate. Feedstock. The feedstock was mainly composed of MC, HC, and RO in a volume ratio of 60:40:20 MC/HC/RO. In industrial practice, coker gas oils and depitched tar acids are added in batches to the commercial feed but were not considered for inclusion in our test work.
Hydroprocessing of Coal-DeriVed Residues
Energy & Fuels, Vol. 20, No. 5, 2006 1763
Table 2. Chemical Class Composition of Creosote and Residual Oil class
medium creosotea
heavy creosoteb
residue oilc
acidsd basese N.N.C.f aromatics alkanes other polar materialg asphaltenes residue total
35.7 6.8 1.7 41.6 8.2 3.8 0.2 0.2 98.2
19.3 7.4 7.5 30.6 27.5 2.1 1.3 1.2 96.6
25.5 1.7 3.3 39.8 21.0 4.1 2.5 1.4 99.3
a bp range of 214-276 °C. b bp range of 236-408 °C. c bp range of 244-456 °C. d Phenols, cresols, xylenols, naphthols, etc. e Pyridines, quinolines, indols, etc. f Neutral nitrogen compounds. g Ketones, alcohols, furans, etc.
Table 3. Properties and Composition of Creosote Feeds properties
F1a
F2b
F3c
F4d
density at 20 °C (kg/L) total N (wppm) total S (wppm) total phenolicse (wppm) Br No (g Br/100 g) ASTM D-86 T95, °C
1.021
1.023
1.015
1.010
55 000 3100 46 100
50 000 2800 39 500
15 400 6400 47 000
9000 3000 46 000
71 345
68 383
n.d. 546
81 390
a,b Feeds F1 and F2 are the 205-420 °C fractions distilled from refinery tar and pitch. c Filtered tar fraction >200 °C. d Feed from commercial creosote hydrogenation plant. e Phenolics include phenol, xylenols, cresols, and naphthols.
The feedstocks were analyzed using the same techniques as for the products. Detailed analyses of the feeds are presented in Tables 2 and 3. Hydrotreatment Procedure. Catalyst activity was monitored by drawing product samples from the reactor after steady-state conditions were reached, typically after a period of 24 h. The following 8 h period was used to collect a representative sample for product analysis. Mass balances of 96-104% were achieved with collection of the tail gas, condensed lighter hydrocarbons, and the liquid product from the reactor stream. Condensed light hydrocarbons were kept refrigerated prior to analysis. The sulfided state of the catalyst was monitored by H2S tail-gas analyses. H2S levels of 200 vppm in the tail gas were found sufficient to maintain the catalyst activity. The concentrations of H2S and NH3 derived from the hydrotreating process were not determined, but that of H2O could easily be determined by separation of the liquid aqueous phase from the organic phase.
Results and Discussion A. Hydrotreating Studies: Determination of Operating Conditions. Studies were done to identify the feasibility of producing fuels and to optimize the reaction conditions for producing gasoline and, more specifically, middle distillates from the Lurgi gasifier oils and tars. Main operating parameters such as temperature, pressure, liquid hourly space velocity (LHSV), and hydrogen-to-oil ratios were varied to ascertain the most suitable operating conditions for the production of fuel fractions with acceptable properties. The feeds studied were, initially, the whole filtered tar fraction (200+ °C) and, then, the individual creosote and residue oil fractions. Also, 50:50 v/v % mixtures of medium and heavy creosotes were investigated, whereas in later studies, residue oil was added to achieve a 3:2:1 volume ratio of MC to HC to RO. Typical results obtained using the CoMo/γ-Al2O3 catalyst while varying the reaction parameters are presented in Tables 4-6. The initial results obtained with the CoMo/γ-Al2O3 catalyst were very encouraging, showing that the creosote material can
be converted into naphtha and distillates in a single pass (see Table 4). Higher temperatures not only favor naphtha but also gas (C1-C4) formation. The produced naphtha ( HDN was found for all catalysts tested. Mann et al.24 and Liaw et al.25 studied the catalytic refining of heavy gas oils and catalytic hydrotreatment of coal-derived naphtha, respectively, and found that oxygen and nitrogen removal were faster with the NiW catalyst than with the NiMo catalyst. This is contrary to the result obtained with the NiMo-1 catalyst investigated in this study. The hydrogenation activity of the catalysts, as reflected by the cetane number in Table 8, was found to be the highest for the NiMo-1 catalyst and lowest for the NiMo-2 catalyst. The HDN and more specifically the HDO activity is higher for the NiMo-1 catalyst, although the HDS activity was found to be slightly lower. The hydrogenation activity of the NiW-1 catalyst was lower compared to that of the NiMo-1 catalyst and similar to that of the NiMo-2 catalyst. However, its HDS activity appeared to be the highest for all three catalysts tested. The NiMo catalysts contained approximately similar metal loadings and almost identical Ni/Mo atomic ratios, but the hydrogenation activity of the two catalysts is different, as shown in Table 8. Assuming that the metals are dispersed uniformly, the data suggest that a portion of the active metal in the NiMo-2 catalyst is located in the smaller pores (see Table 1) and is, as a result, not effectively used because of diffusional restrictions. The NiMo-2 catalyst has a higher proportion of pores in the 35-60 Å range. This could explain the lower hydrogenating activity of the NiMo-2 catalyst compared to the NiMo-1 catalyst. Maximum pore volumes for the NiMo-1 catalyst are in the ranges 60-110 and 110-220 Å. For the NiW-1 catalyst, the highest pore volume is found in the 60-110 Å range. Both the NiMo-2 and the NiW-1 catalysts show very low fractions of pore volumes in the 110-220 Å region. It appears therefore that oxygen removal is enhanced by the presence of larger pores. However, this was not the case for nitrogen removal, which could be related to the molecular composition of our feed. Katzer and Sivasubramanian26 have reported that, as the concentration of the active metal is increased, the catalyst activity increases through a maximum and then decreases. With reference to Table 8, the diesel cetane number showed the highest values (i.e., highest aromatics saturation) in the case of the NiMo-1 catalyst. Stiegel et al.,27 who hydrotreated solvent-refined coal in creosote oil with NiMo on γ-alumina catalysts, found a (24) Mann, R. S.; Sambi, I. S.; Khulbe, K. C. Ind. Eng. Chem. Res. 1987, 26, 410-414. (25) Liaw, S.-J.; Keogh, R. A.; Thomas, G. A.; Davis, B. H. Energy Fuels 1994, 8, 581-587. (26) Katzer, J. R.; Sivasubramanian, R. Catal. ReV.sSci. Eng. 1979, 20 (2), 155-208. (27) Stiegel, G. J.; Tischler, R. E.; Polinski, L. M. Ind. Eng. Chem. Prod. Res. DeV. 1983, 22, 411-420.
1766 Energy & Fuels, Vol. 20, No. 5, 2006
Leckel
Table 7. Effect of Temperature on Naphtha and Diesel Product Properties Produced over a NiW/γ-Al2O3 Catalyst at a Pressure of 17.5 MPa, a Liquid Hourly Space Velocity of 0.5 h-1, and an H2-to-Oil Ratio of 1500:1 LN/L Using Feed F3 temp. °C
yields, wt % H2O
gases C1-C4
naphtha C5-200 °C
diesel >200 °C
naphtha RON
naphtha aromatics
diesel cetane number
333 385 408 445
3.9 5.0 3.3 4.2
0.9 1.7 3.5 3.7
21.2 33.1 34.3 39.8
76.0 67.5 62.5 52.3
64 58 57 59
13 4 5 7
34 47 45 45
Table 8. HDO, HDS, and HDN of Creosote Feed F3 over Various Catalysts Using a Temperature of 380 °C, a Pressure of 17.5 MPa, a LHSV of 0.5 h-1, and an H2-to-Oil Ratio of 1500:1 LN/L catalyst type
phenolicsa in TLP wppm
sulfurb in diesel wppm
nitrogenc in diesel wppm
diesel CN
NiMo-1 NiMo-2 NiW-1
18 124 112
27 10 2
64 84 67
48 42 43
a Initial value of phenolics in feed F3: 47 000 ppm. b Initial value of sulfur in feed F3: 6400 ppm. c Initial value of nitrogen in feed F3: 15 400 ppm.
relationship between pore size and HDN as well as HDO activity. High concentrations of the active metal resulted in higher hydrogenation activity, and high metal loadings and large pores enhanced the heteroatom removal activity. Oxygen could be more easily removed from solvent-refined coal than nitrogen. Higher temperatures and lower space velocities favored deoxygenation. Decreasing the pore size, as reported by Katzer and Sivasubramanian,26 increases the pore diffusion resistance and also restricts larger molecules from reaching the active sites of the catalyst. Song et al.28,29 found during the upgrading of coal-derived oil sands, asphaltenes, and heavy coal liquids that maximum HDO was obtained with a catalyst having a median pore diameter (MPD) of 150 Å and that higher HDO activity was observed at MPD values of 150 and 290 Å. The catalyst with the highest HDO activity had a pore size distribution between 100 and 200 Å. This is very much in line with the observations made in this study. Conclusions This study confirmed that pyrolysis products, resulting from the coal gasification of a high-temperature Fischer-Tropsch (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, 619628.
plant, can be successfully refined to fuel blending components by the use of severe hydroprocessing conditions. Temperatures of 350-400 °C and operating pressures of at least 12.5 and at maximum 20 MPa, combined with low space velocities (i.e., 0.25-0.5 h-1 LHSV), ensure the deep hydrogenation of refractory sulfur, oxygen, and nitrogen compounds. The CoMo catalyst produced naphtha fractions with high RON values, while the RON of the naphtha fraction produced with the NiW catalyst was significantly lower. This is the result of the higher hydrogenation activity of the NiW catalyst. The middle distillate cetane number improved at higher pressures because of the saturation of the aromatics. The cetane number increase, when processing the feed between 15 and 20 MPa and a temperature of 430 °C, was found not to be as significant as the increase in cetane number observed when operating between 7.5 and 12.5 MPa at the same temperature. Operation at hydrogen-to-liquid ratios higher than 1500:1 LN/L proved not to be of significant benefit. HDO, particularly the removal of phenolic components, and HDN were achieved at levels greater than 99% using NiMo on γ-Al2O3 catalysts. Maximum HDO activity appeared with a NiMo catalyst having maximum pore sizes in the range of 110220 Å. The reactivity of heteroatom removal follows the order HDO > HDN for the NiMo/γ-Al2O3 catalysts having larger pore sizes, while the reactivity for the NiMo/γ-Al2O3 catalysts also having a relatively high fraction of smaller pores (between 35 and 60 Å) follows the order HDN > HDO at the conditions applied. Acknowledgment. Permission from Sasol Technology Research and Development to publish this work is appreciated, as well as the support from Sasol Infrachem and the Analytical Techniques Department for the supporting analyses of feed and products. The contributions of Dr. L. C. Ferreira, G.G. Swiegers (21/78, 28/79, and 32/80,), K. Kriel, and A. Brodziak (69/85 and 70/85) are also acknowledged. I thank in particular Dr. C. P. Nicolaides for the valuable technical discussions. EF060034D
