PROCESS AND CATALYSTS FOR THE PRODUCTION OF MOTOR FUELS FROM SHALE OILS

Abstract: Shale oils having a high sulfur content are hydrotreated in a two stage process. The shale oil feed is passed along with hydrogen to a reactor (3) where it contacts a catalyst in a first stage (12) to effect desulfurization. The feed is then passed to a second stage (13) where it contacts a second catalyst to effect denitrogenation. The catalyst in the first stage is cobalt and/or nickel and molybdenum supported on a rare earth oxide-modified alumina. The catalyst in the second stage is cobalt and/or nickel, molybdenum, and a promoter on a zeolite and rare earth oxide-modified alumina support.

Description: PROCESS AND CATALYSTS FOR THE PRODUCTION OF MOTOR FUELS FROM SHALE OILS Field of the Invention This invention relates to a process and catalysts for the production of motor fuels from shale oils that are difficult to treat by known processes because of their chemical composition, in particular because of their high sulfur and nitrogen concentration, an example of such shale oils being those extracted from the bituminous marls deposits in the southern part of Israel. Background of the Invention A variety of oil shale materials have been identified as a source for the production of crude oil substitutes. They are regarded as a very large reserve of synthetic oils. Considerable efforts have been made in the art to develop shale oil hydroprocessing technology using conventional hydroprocessing catalysts. Such processes are described, for instance, in R.F. Sullivan et al., DOE Report FE-2315-25, April, 1978 and A.O. Braun et al., Aero Propulsion Laboratory Report, Wright-Patterson Air Force Base, U.S.A., AFWAL-TR-81-2063, May, 1981. A more complex shale oil processing technique is described in P.F. Lovell et al., Hydrocarbon Processing, 60(5), 125 (1981). The key problem in the said hydroprocessing techniques is the hydrodenitrogenation (hereinafter HDN) of shale oils to lower the nitrogen content. HDN is much more difficult than hydrodesulfurization (see T. C. Ho, "Hydrodenitrogenation Catalysis", in Catal. Rev. - Sci. Eng., 30(1), 117, 1988).Therefore a process giving deep HDN, required for motor fuel production (less than 0.02 wt% nitrogen in hydrotreated shale oil,) should fit the requirements for sulfur content in corresponding fuels. Furthermore, according to P.J. Nat, Erdol und Kohle-Erdgas- Petrochemie, 42(11). 447 (1989), an increase in the nitrogen content in feedstock from 0 to 0.2% wt. requires an increase of 100°C to maintain the same hydrocracking conversion with Ni-Mo- or Pd-zeolite catalysts, so that for a high hydrocracking efficiency the nitrogen content of the feedstock should be below 0.02% wt. So the HDN is a problem also in the case that the hydrotreated shale oil is to be used as a feedstock for the hydrocracking process in order to increase the yields of light gasoline and jet fuel. Table I lists the characteristics of shale oils of various origins and it is seen that the Israeli-Negev shale oils have particularly high sulfur, nitrogen and Ramsbotton carbon contents. Table I CHARACTERISTICS OF ISRAELI, AMERICAN, AUSTRALIAN AND BRAZILIAN SHALE OILS Shale Oil Israeli- American- American- Australian- Brazilian- characteristics Negev Occidental Colorado Rundle Ira ti -Gravity, 10.7 23/8 20.6 — 20.6 API 0 -Elemental Composition: Carbon, wt.% 78.93 84.82 83.83 85.40 ~ Hydrogen wt. % 9.27 1 1.83 1 1.72 1 1.50 — Oxygen, wt. % 1.53 1.40 1.31 0.90 — Nitrogen, wt. % 0.94 1.32 2.13 1.10 0.93 Sulfur, wt.% 7.28 0.64 0.75 0.90 0.82 H/C (atomic) 1.41 1.67 1.68 1.61 — Arsenic, ppm 30 11 34 9 — Iron, ppm 35 41 — 42 — Nickel, ppm 40 61 — 0.4 — Ramsbottom carbon, wt.% 5.23 1.22 1.40 — — Distillation, °C IBP 1 16 143 56 100 — 40% 321 343 — — ~ 50% 332 — 425 — 370 FBP 573 — 573 400 485 IBP and FBP are hereinafter Initial Boiling Point and Final Boiling Point , respectively. In Table I the data relating to American Occidental shale oil were taken from European Patent No. 005091 1; those relating to American Colorado shale oil, from P.F. Lovell et al. Hydrocarbon Processing, 60(5), 125, (1981); those relating to Australian Rundle shale oil, from T.G. Harvey et al, Ind. Eng. Chem. Proc. des. Dev., 25, 521, (1986); and those relating to Brasilian Irati shale oil, from G.L.M. Souza et al, Ind. Eng. Chem. Res., 31 , 2127, (1992). Table II shows the results of the hydrotreatment of various shale oils according to prior art techniques, and evidences the very difficult problem involved in achieving deep HDN. With commercial catalysts, the desired final nitrogen level below 0.2% wt. was not achieved even under very severe conditions, such as temperature above 400°C, in fixed bed or fiuidized bed reactors. Only lowering the distillation end point ofthe shale oil to 400°C and operating at 400°C yielded a product containing 0.005% wt. nitrogen, as reported by T.G. Harvey et al., Ind. Eng. Chem. Proc. des. Dev., 25, 521 (1986). EP 005091 1 discloses the conversion of shale oil with experimental CO- Mo-Al-USY catalyst at 415-430°C to <0.001% wt. nitrogen. Table II Hydrotreating of shale oils: comparison of existing results Hydrotreating conditions Shale oil Catalyst LHSV P, psi T C Nitrogen Reference origin h 1 content in hydrotreated shale oil, ppm Commercial: Colorado Ni-Mo-Al 0 6 2200 408 500-1000 [5] Colorado Ni-Mo-Al 0 4 1650 405 500 [7] (Shell-324) Colorado Ni-Mo-Al 1.0 1500 430 9000 [16] (Harshaw 4303E) Colorado Co-Mo-Al 0 5 1000 440 6000 [17] Paraho Ni-Mo-Al 0 9 2078 416 1220 [18] (Am Cyanamide HDS-3A) Occidental 0.9 2078 416 380 m Tosco 0 9 2078 416 3660 Geokinetics 0 9 2078 416 1400 Rundle)* Ni-Mo-Al 1 0 2000 400 50 [9] Rundle)* Ni-W/Al-Si 1 0 2000 400 3000 Irati Ni-Mo- 1 0 1800 400 4092 [19] Al**) Experimental: Colorado Co-Mo-Al- 1 0 1000 440 1995 [17] F Occidental Co-Mo-Al- 0 5 1800 415- <10 [ 10, 1 1] zeolite 430 USY-Cr 2 O 3 LHSV, hereinafter, stands for Liquid Hourly Space Velocity The numbers in the reference column refer to the following publications [5] R F Sullivan, E Strangeland, C E Rudy, D C Green, H A Frumkin, DOE Report FE-2315-25, April 1978, [7] P F Lovell, M G. Fryback, H E Rolf, J P Schwedock, Hydrocarbon Processing, 60(5), 125, (1981), [10] A M Tait, A L Hensley, Prepr ACS, Div Fuel Chem. 27(2), 187, (1982), [ 1 1 ] A M Tait, T D Newitt, A L Hensley, EP N 005091 1 , (1981 ), [16] D P Montgomery, Ind Eng Chem Prod Res & Dev , 7(4), 274 (1968), [17] D B Benson, L Berg, Chem Eng Progr , 62(8), 61, (1966), SUBSTITUTE SHEETT (RULE 26) [18] N.L. Mukherjee, Canad. J. Chem. Eng., 65, 966 (1987); [19] G.L.M. Souza, J.C. Alfonso, M. Schmal, J.N. Cardoso. Ind. Eng. Chem. Res., 31, 2127, (1992). The data available in the prior art dealing with the hydrotreating of shale oils show that deep HDN can be achieved only under highly severe conditions. Continuous commercial operation of the first hydrotreating stage in a fixed bed reactor under such conditions would be difficult because of the rapid deactivation of the catalysts by coke deposits. The aforesaid difficulties are greatly intensified when dealing with shale oils having a high sulfur concentration, and particularly, such a concentration combined with a high Ramsbottom carbon level and high nitrogen content. It is a purpose of this invention to provide a process for the catalytic hydrotreatment of such shale oils for the purpose of obtaining desirable HDN degree, in particular, HDN degree corresponding to a nitrogen content below 0.02% wt. It is another purpose of this invention to provide such a process which involves moderately severe conditions, in particular, a substantially lower temperature than the treatments known in the prior art. It is a further purpose of this invention to provide catalysts for carrying out the aforesaid process. Other purposes and advantages of the invention will appear as the description proceeds. Summarv of the Invention The process according to the invention is characterized in that the starting shale oil feedstock, in particular a feedstock having high sulfur, Ramsbottom carbon and nitrogen content, is treated in two stages, the first stage being a deep desulfurization with wide pore Co/Ni-Mo-RE-Al catalyst and the second stage being a deep hydrodenitrogenation (HDN) with a Co Ni-Mo-zeolite catalyst. Hereinafter, RE stands for Rare Earth elements. It has been found that the first deep desulfurization stage reduces the average molecular weight and therefore the distillation pattern of nitrogen-containing substances, making said substances easily hydrodenitrogenated in the second stage. The process parameters which characterize the two stages are the following. First stage: Temperature from 360 to 400 °C, and preferably from 370 to 380 °C. LHSV not less than 0.5, and preferably from 0.5 to 3.0 h ~ l , and more preferably, from 0.8 to 1.2 h" 1 . Hydrogen pressure from 1000 to 3000, and preferably from 2000 to 2500 psi. V H2 from 300 to 5000 and preferably from 2000 to 3500 Nl/1. Second stage: Temperature from 380 to 400 °C, and preferably from 380 to 390 °C. LHSV not less than 0.5 h " , and preferably from 0.5 to 3.0 h"' , and more preferably, from 0.8 to 1.2 h" L Hydrogen pressure from 1000 to 3000, and preferably from 2000 to 2500 psi. VH 2 from 300 to 5000 and preferably from 2000 to 2500 Nl/1. Preferably, the shale oil is demetallized before submitting it to the aforesaid two treatment stages. In a preferred form of the invention, the two stages are carried out in fixed bed reactors with two different catalysts loaded in one reactor in series or in two reactors in series loaded with different catalysts . The shale oil feedstock, fed to the first stage, may be mixed with an hydrocarbon organic solvent, preferably chosen from among hydrocarbons or mixtures of hydrocarbons boiling out at less than 105°C, from 50 to 105°C, e.g. cyclohexane or light naphtha, in amounts from 30 to 90 vol. %, e.g. shale oil: light solvent = 2:8, by volume, in order to reduce the viscosity of the shale oil and increase its diffusion in zeolite pores of the second stage catalyst. It allows to reduce the nitrogen content in the hydrotreated shale oil to less than 0.0005% wt.. On the other hand, the use of such a solvent invokes the need for solvent recycling, and therefore may not be desirable. The decision whether to use it or not, therefore, will depend on the circumstances of each specific operation. The catalysts used in the first stage comprise: NiO or/and CoO and M0O3, the content of each of the component oxides being: NiO from 0.0 to 3.0 %, CoO from 0.0 to 3.0 %, provided that the sum of NiO and CoO is at least 0.5%, and M0O3 from 8 to 15 %, all percentages being by weight calculated on the whole of the catalyst. The oxides are supported on an AI2O3 substrate modified by RE-oxide, and the ratio of their combined weight to the support weight varies from 0.09 to 0.20, and preferably from 0.15 to 0.17. The packed density of the catalyst g/cc varies from 0.30 to 0.90, and preferably from 0.35 to 0.40. The surface area, in m^/g varies from 150 to 300 and preferably from 200 to 260. The pore volume of the catalyst in cc/g varies from 0.70 to 1.20, and preferably from 0.80 to 1.00. The average pore diameter, in Angstroms, varies from 1 10 to 140, and preferably from 1 15 to 135 The alumina support, containing 1 -3 wt% of RE oxide, must have large pores, e g , an average pore diameter of 140-180 Angstroms, preferably around 160 Angstroms, and can be prepared by known preparation methods, such as that described by R K Oberlander in "Applied Industrial Catalysis", B E Leach Ed Acad Press, v 3, 1984, p 64 It is impregnated, to prepare the catalyst, with nickel, cobalt and molybdenum salts in solution, by impregnating methods that are generally known - see W M eely, P Jerus, E K Dienes and A L Hausberger, "Preparation Techniques for Hydrotreating Catalysts and their Influence on the Location of the Metal Oxides and Performance", Catal Rev -Scie Eng 26 (3 & 4), 485, 1984 - and the impregnated substrate support is then dπed and calcined The catalyst used in the second stage comprises a zeolite stabilized on the same alumina-RE support The zeolite-containing support is impregnated to prepare the catalyst with promoter solution, containing salts of Cr, P, Zr or Ti, and then with Co or/and Ni and Mo salts in solution by impregnation methods that are generally known The zeolites are chosen from among Faujasite type, preferably zeolite Y with SιO 2 /Al 2 O 3 mole ratio from 4 to 6, preferably from 4 5 to 5 5 in hydrogen form with sodium content less than 0 5% wt , preferably less than 0 2 % wt The zeolite content in the catalyst varies from 10 to 40% ww, and preferably from 25 to 35% ww A promoter, e g Cr, P, Zr or Ti oxide can be added to the zeohte catalyst in amount from 2 to 6% ww, and preferably from 4 to 5% ww The catalysts comprise CoO or/and NiO and MoO 3 , the amounts of each of those oxides being CoO - from 0 to 6%, NiO - from 0 to 6%, the combined amount of CoO and NiO being at least 3%, and MoO - from 12 to 18%, all percentages being by weight calculated on the whole of the catalyst The oxides are supported on the zeohte-alumina-RE support, and the ratio of their combined weight to the supports weight vanes from 0 20 to 0 45 and preferably from 0 22 to 0 28 The packed density of the catalyst g/cc varies from 0 40 to 0 90, and preferably from 0 45 to 0 53 The surface area, in m 2 /g vanes from 250 to 400 and preferably from 280 to 320. The pore volume of the catalyst in cc/g varies from 0.60 to 0.75, and preferably from 0.65 to 0.73. All catalysts, for both stages, were sulfided before reaction by known sulfidization methods: see. e.g., McCulloch in "Applied Industrial Catalysis", B.E. Leach Ed. Acad. Press, v. l , 1983, p. 70. The ratio between the volume of the catalysts and the volume of the feed in the two stages, in the case of fixed bed reactor, varies within the following limits: 3 3 First stage: 0.5 to 3.0 and preferably 0.8 to 1.2 m /h per m of catalyst Second stage: 0.5 to 3.0 and preferably 0.8 to 1.2 m 3 /h per m 3 of catalyst. The residence time of the feed in the two stages varies, for the first stage from 0.3 to 2.0. and preferably from 0.85. to 1.25 hours, and for the second stage, from 0.3 to 2.0, and preferably from 0.85 to 1.25 hours, and correspondingly, the LHSV for the first stage is not less than 0.5 h " , and preferably varies from 0.5 to 3.0 h"*, and more preferably, from 0.8 to 1.2 h ~ l , and for the second stage is not less than 0.5 h "1 , and preferably varies from 0.5 to 3.0 h' and more preferably, betwen 0.8 and 1.2 h-i . Brief Description of the Drawings In the drawings: - Fig. 1 is a schematic illustration of a fixed bed reactor pilot plant used in carrying out an embodiment of the invention; Detailed Description of Preferred Embodiments The embodiment hereinafter described refers to a process carried out in a fixed bed reactor pilot plant, such as illustrated in Fig. 1. Fig. 1 is a schematic illustration of a hydrodesulfiirization plant. The hydrotreating of shale oil was studied in the pilot plant apparatus shown in Figure 1. The shale oil (individual or mixed with a light hydrocarbon solvent) to be treated was charged to a reservoir (1). The liquid from reservoir (1) passed by way of metering pump (2) to the mixing zone of the tubular reactor (3) to join a flow of hydrogen from cylinder (4) that passed a mass flow controller (5). After mixing with hydrogen, the shale oil passed the heating zone 1 1 of the reactor (3) and then to two hydrotreating zones 12 (the first-stage zone) and 13 (the second-stage zone) charged with the same volume of different catalysts (see above). Before the first stage hydrotreating zone, appropriate amount of silica particles 20-35 mesh, indicated at 14, were charged to improve the liquid distribution. Reactor (3) consists of a 25 mm internal diameter vertical tube 1.15 meter long with axial thermocouple pocket (not shown). It was heated by two individually and automatically controlled electric heaters (6) and (7), each arranged to heat a hydrotreating zone of the reactor. From the bottom of the reactor (3) there emerged a two-phase fluid which passed high-pressure separator (8) and then low-pressure separator (9), where the liquid products were separated from the gases (H 2 , H 2 S, NH 3 ), that were subsequently passed to a discharge line containing flow measurement equipment (10) and analytical equipment (not shown) and vented to the atmosphere. To prevent blockings of the tube at the reactor outlet with ammonium sulfide crystals, water was pumped immediately to the reactor outlet point (not shown) to dissolve this salt and was collected subsequently in the high-pressure separator (8). The crude shale oil feedstock was the Israeli Negev shale oil, the characteristics of which are listed hereinbefore in Table I. The catalyst used for the first step of the process was a DMN catalyst prepared in the following manner. An alumina support was prepared by the preparation method of Wakabayshi et al., described in "Applied Industrial Catalysis, v.3, 1984, p. 92, which consists of pH oscillating precipitation from an aqueous solution of aluminum nitrate by hydroxide. The cake was impregnated with La(NO 3 ) 3 water solution, washed, extruded into 1.2 mm diameter pellets, dried at 120°C for two hours and calcined at 550°C for five hours. The extruded pellets were impregnated with nickel nitrate, cobalt nitrate and ammonium molibdate in ammonious solutions (15% NH 4 OH, 2.5% NiO or CoO, and 10% MoO 3 ).The impregnated pellets were dried at 120°C for two hours and calcined at 550°C for three hours. Three catalysts - DNM- 1,2,3 - were prepared. Their characteristics are set forth in Table III. Table III Characteristics of the first-stage hydrotreating catalysts Catalyst DMN-1 DMN-2 DMN - 3 Chemical composition, % ww: NiO LS 5.6 — CoO — — 1.8 M0O3 12.0 11.7 12.8 La 2 O 3 1.7 1.6 1.7 Al 2 O 3 balance balance balance Catalysts properties: packed 0.35 0.51 0.39 density, g/cc surface area, 240 220 230 m 2 /g pore volume, 0.88 0.65 0.80 cc/g Average pore 130 110 135 diameter, A For the second step of the process, catalysts designated as HTN were prepared. They contained HY zeolite, LZ-Y62 (manufactured by Linde AG) stabilized in an alumina support. The zeolite was mixed with the alumina cake, prepared by the pH oscillating precipitation method, hereinbefore mentioned, impregnated with La(NO 3 ) 3 water solution, extruded into 1.2 mm diameter pellets, dried at 120°C for two hours and calcined at 550°C for five hours. The extruded pellets were twice impregnated with (NH 4 ) 2 CrO 4 water solution (4.1% Cr 2 O 3 in water) or H 3 PO 4 water solution (3.2% P 2 O 5 in water) and then with cobalt nitrate (or nickel nitrate) and ammonium molibdate in ammonious water solution (15% NH 4 OH, 8.5% CoO (or NiO) and 20% MoO 3 ). After final impregnation, the catalyst was dried at 120°C for two hours and calcined at 550°C for five hours. Five second-stage catalysts HTN- 1,2,3,4,5 were prepared. Their characteristics are set forth in Table IV. Table IV Characteristics of the second-stage hydrotreating catalysts Catalyst HTN-1 HTN-2 HTN-3 HTN-4 HTN-5 Chemical composition, % ww: NiO 4Ji 3.8 4.0 4.1 — CoO — ~ — — 4.5 MoO 3 16.2 16.3 18.6 18.1 18.8 Cr 2 O 3 3.0 — 4.8 4.8 5.5 P 2 O s — 4.5 — — « Zeolite HY 22.6 23.0 29.0 36.5 28.5 A 2 O 3 lance balance balance balance balance Catalysts properties: packed density, 0.46 0.40 0.51 0.45 0.39 g/cc surface area, 280 270 300 340 230 2 /g pore volume, 0.69 0.70 0.63 0.74 0.80 cc/g All catalysts were sulfided before reaction by elemental sulfur in a tubular reactor with hydrogen flow following a standard procedure: see U.S. P. 4,177,136 . Five catalysts were so prepared. The composition of the catalyst was measured by ED AX (Microscope GEM-35, Geol. Co., Link System AN-1000, Si-Li-Detector), the surface area and pore volume were measured by BET (ASTM 3663-84) and by water adsorption at room temperature respectively. The characteristics of the zeolite catalyst are listed in Table IV. The fixed bed reactor plant used in this embodiment of the invention, is the one illustrated in Fig. 1 The process was carried out in the following way. The DMN catalyst was loaded in the upper, first-stage hydrotreating zone of the reactor and the same volume of HDN catalyst was loaded in the lower, second-stage hydrotreating zone of the reactor located immediately after the first-stage zone. The shale oil feedstock was demetallized with commercial Ni-Mo-Al catalyst at 250°C, hydrogen pressure 1500 psi, LHSV = 1 h _1 and V H2 = 500 Nl/1. This treatment reduced the content of Ni, Fe and As with insignificant HDS and HDN (see Table V hereinafter). The runs were carried out using demetallized shale oil - pure or mixed with cyclohexane, as a feedstock. All the runs were carried out with 2200 psi hydrogen pressure varying the temperature in both hydrotreating zones, LHSV and V,, 2 In cases when the shale oil was mixed with cyclohexane, the latter was separated from hydrotreated shale oil by distillation before estimation of the sulfur, nitrogen content and distillation patterns of the products. The results are listed in Table VI hereinafter. When the temperature in the two hydrotreating zones was kept at 380°C and the shale oil was mixed with cyclohexane, varying the composition and physical parameters of DMN and HTN catalysts influenced the HDN and HDS degree of the feedstock. The best results were obtained with catalysts DMN-1 - HTN-5: 70 ppm sulfur and 2.5 ppm nitrogen in the hydrotreated shale oil (Table VII). Using pure shale oil without solvent and the same catalysts combination requires higher temperatures for deep HDN/HDS of the feedstock. The nitrogen content < 0.02% wt. could be obtained at temperatures 380-400°C at the first stage and 390-400°C at the second stage. Table VI Shale Oil Characteristics Original Demetallized Gravity, API 0 10.7 18.2 Density (15°C), g/cc 1.085 0.994 Kinematic viscosity, Cst: at 40°C 17.54 9.73 at 50°C 12.05 7.24 Elemental composition, % wt: 78.93 80.69 C 9.27 1 1.98 H 7.28 7.03 S 0.94 0.90 N 1.53 — O Solids, % wt. 2.05 — Distillation, °C: 1 16 I l l IBP 332 328 50% 573 569 FBP Asphalthenes, % wt.: insoluble in heptane 4.17 — insoluble in toluene 0.09 — Metals content, ppm Ni 40 <5 Fe 35 <5 As 30 <1 Table VII *) Calculated on basis of all the liquid fed to the reactor and volume of the two catalysts loaded in the reactor. Table VII (continued): Hydrotreating of shale oil: test results Catalysts: First Stage DMN-1 DMN-1 DMN-1 DMN-1 DMN-1 Second Stage HTN-3 HTN-4 HTN-5 HTN-5 HTN-5 Test Conditions: Temperature first stage 380 380 370 380 400 second stage 380 380 380 390 400 LHSV, h " '*) 1 1 0.5 0.5 0.5 V H2 , Nl/1 500 500 1500 1500 1500 Amount of cyclohexane solvent in the feedstock, % vol. 80 80 0 0 0 Characteristics of hydrotreated shale oil: Density (15°C), g/cc 0.826 0.824 0.820 0.825 0.830 Sulfur content, ppm 160 120 160 120 100 Nitrogen content, ppm 2.0 6 250 80 30 *) Calculated or l basis of a ill the liquic fed to the reactor and volume o the two catalysts loaded in the reactor. It can be noted that, operating as hereinbefore described, and using the described catalysts, the optimal SOR (Start Of Run) temperature for the hydrotreatment in the first stage is 380°C, and the optimal LHSV for fuels production is 0.5 h'L Hydrogen pressure can be desirably kept at 2200 psi (150 atm) and V = 1500 Nl/1. In this way, one obtains a demetallized, denitrogenized, and desulfurized shale oil - a product containing less than 5 ppm of Ni and Fe, less than 0.02% ww nitrogen and less than 0.03% ww sulfur. This product could be used for production of motor fuels by further distillation: 10-25 volume % gasoline boiling out at <180°C, 20-35% volume jet fuel boiling out in the range 160-270°C or 70-80 volume % diesel fuel boiling out in the range 160-380°C. This product could be also used as a feedstock for hydrocracking process that will increase the yield of light motor fuels - gasoline and jet fuel. The pilot plant was operated for 300 hours with DMN-l-HTN-5 catalysts loaded in series in conditions shown in Table VII - second column from the right, T(DMN-l) = 385°C, T(HTN-5) = 395°C, no solvent addition, LHSV=0.5h-' , V 1I2 =1500 Nl/1. The hydrotreated shale oil was distilled to separate gasoline fraction IBP- 160° jet fuel fraction 160-270°C and diesel fuel fraction 160-380°C. The yields of those three fuels were 12, 26 and 75% vol., conformably, calculated on the basis of hydrotreated shale oil. The quality characteristics of those fuels are compared with the specifications requirements in Tables VIII-X. As could be seen from the characterization results, diesel and jet fuel fit the specifications requirements (diesel fuel fits even the requirement for <0.05 wt.% sulfur that will be introduced in the near future). Gasoline has comparatively high initial and 10% boiling points that does not allow to pass the Reid vapor pressure test, and low octane number. This light fraction is a good component for mixing with petroleum naphtha as a feed for the reforming process that will increase the octane number and add light products to the final gasoline. Table VIII Quality of gasoline distilled off from hydrotreated Israeli shale oil Characteristics Testing Specification Tested method requirements values IS-90 Distillation patterns, °C ASTM D-86 IBP not limited 78 10% <70°C 92 50% <125°C 124 90% <180°C 169 FBP <215°C 195 Reid vapor ASTM D-323 6.5-8.5 5.5 pressure, psi Total sulfur % mas. ASTM D-4294 0.15 0.01 1 Total nitrogen, % mas. ASTM D-4629 not limited 0.005 Corrosion with Cu ASTM D-130 la la Existent gum, ASTM D-381 mg/l 00 ml - before washing not limited 15 - after washing <5 <1 Density (15°C) ASTM D-1298 0.715-0.780 0.7639 RON ASTM D-2699 91-96 72 Table IX Quality of jet fuel distilled off from hydrotreated Israeli shale oil Characteristics Testing Specification Tested values method requirements ASTM D-1655 Distillation patterns, °C ASTM D-86 IBP 10% <204°C 146 20% not limited 169 50% not limited 177 90% not limited 202 FBP not limited 252 Residue, % <300°c 295 Looses, % <1.5 1.3 <1.5 1.2 Density (15°C), g/cm J .ASTM D-1298 0.775-0.840 0.8092 Aromatics content, ASTM D-1319 <25 22 % mas. Smoke point, mm ASTM D-1322 >20 24 Mercaptanes content, IP-342 <0.003 0.0014 % mas. Olefins content, ASTM D-1319 not limited 0.5 % mas. Total Sulfur, % mas. ASTM D-4294 <0.3 0.014 Total nitrogen, % mas. ASTM D-4629 not limited 0.0098 Flash point, °C IP-170 >37.8 42.2 Freezing point, °C ASTM D-2386 <-47 -47 Viscosity at -20°C, cst ASTM D-445 <8 6.2 Corrosion with Cu ASTM D-130 <la 1.9 Corrosion with Ag IP-227 <1 0 Existent gum, mg/l 00 g ASTM D-381 <7 3.7 Water separation ASTM D-1094 <2 (lb) 3(lb) WISM ASTM D-3948 not limited 88 Seibolt color ASTM D-156 not limited 16 Table X Quality of diesel fuel distilled off from hydrotreated Israeli shale oil Characteristics Testing method Specification Tested requirements values IS - 107 Distillation patterns, °C ASTM D-86 IBP 10% not limited 156 20% not limited 193 50% not limited 209 90% not limited 257 FBP <357°C 342 Residue, % not limited 365 Looses, % not limited 1 0 not limited 0 2 Density (15°C), g/cm ' ASTM D-1298 <0 870 0 8361 Color ASTM D-1500 <3 5 2 0 Flash point, °C ASTM D-93 >66 56 Total sulfur, % mas ASTM D-4294 <0 2 0 013 Total nitrogen, % mas ASTM D-4629 not limited 0 020 Viscosity at 40°C, cst ASTM D-445 2 5-6 0 2 52 Pour point, °C ASTM D-97 <5°C summer -3 <-6°C winter Corrosion with Cu ASTM D-130 <1B 1 Ramsbottom carbon, ASTM D-189 <0 2 0 07 % mas Acidity, mg KOH g ASTM D-974 <0 7 0 014 Ash, % mas ASTM D-482 <0 01 0 007 Sediments by extr , ASTM D-473 <0 01 0 001 % mas Accelerated oxidation mg/l 00 ml ASTM D-2274 - precipitate not limited 1 1 - burned not limited 0 46 - total not limited 1 56 CFPP (Cold filter IP-309 <4°C -1 plug point), °C Diesel index ASTM D-976 ≥50 50 While embodiments of the invention have been described by way of illustration, it will be apparent that the invention can be carried out by persons skilled in the art with many modifications, variations and adaptations, without departing from its spirit or exceeding the scope of the claims

Claims: CLAIMS 1. Process for the hydrotreatment of shale oils, particularly shale oils having a high sulfur content, characterized in that the starting shale oil feedstockis treated in two stages, the first stage being a deep desulfurization with wide pore Co/Ni-Mo-RE-Al catalyst and the second stage being a deep hydrodenitrogenation with a Co/Ni-Mo- promoter-zeolite RE-A1 catalyst. 2. Process according to claim 1, wherein the two stages are carried out, after demetallization, under the following conditions: First stage: SOR Temperature from 360 to 400 °C. LHSV not less than 0.5 h" 1 . Hydrogen pressure 1000 to 3000 psi. V H2 from 300 to 5000. Second stage: SOR Temperature from 380 to 400 °C LHSV not less than 0.5 h' 1 Hydrogen pressure from 1000 to 3000 psi. V H2 from 300 to 5000 Nl/1. 3. Process according to claim 2, wherein the two stages are carried out under the following conditions: First stage: SOR Temperature from 370 to 380 °C. LHSV from 0.5 to 3.0 h" ] . Hydrogen pressure from 2000 to 2500 psi. V H2 from 2000 to 3500 Second stage SOR Temperature from 380 to 390 °C LHSV from 0 5 to 3 O h" 1 Hydrogen pressure from 2000 to 2500 psi V H2 from 2000 to 2500 Nl/1. 4 Process according to claim 3, wherein the LHSV is from 0.8 to 1. 2 U ' 5 Process according to claim 1 , wherein the two stages are carried out in fixed bed reactors 6 Process according to claim 5, wherein the two stages are carried out in one bed reactor loaded with two different catalysts is series 7 Process according to claim 5, wherein the two stages are carried out in two reactors in series loaded with different catalysts 8 Process according to claim 1, wherein the shale oil feedstock is mixed with an organic solvent 9 Process according to claim 8, wherein the organic solvent is a hydrocarbon or a mixture of hydrocarbons boiling out at less than 105°C 10 Process according to claim 1 , wherein the wide pore Co/Ni-Mo-RE-Al catalyst, used in the first stage, comprise NiO and/or CoO, and M0O3 1 1 Process according to claim 10, wherein the contents of the component oxides are NiO from 0 0 to 3 0 %, CoO from 0 0 to 3.0 %, provided that the sum of the NiO and CoO is at least 0 5%, and M0O3 from 8 to 1 % ww on the whole of the catalyst 12. Process according to claim 10, wherein the oxides are supported on an Al O3 substrate modified by RE-oxide.. 13. Process according to claim 10, wherein the ratio of the combined weight of the oxides to the support weight is comprised between 0.09 and 0.20 . 14. Process according to claim 1, wherein the Co/Ni-Mo-promoter-zeolite RE-Al catalyst used in the second stage comprises a zeolite stabilized on an alumina support modified by RE-oxide. 15. Process according to claim 14, wherein the zeolites are chosen from among Faujasite type zeolites. 16 - Process according to claim 12, wherein the zeolites are chosen from among zeolites Y with SiO 2 /Al 2 O 3 mole ratio from 4 to 6. 17. Process according to claim 14, wherein the zeolite content of the second stage catalyst varies from 10 to 40% ww. 18. Process according to claim 14, wherein a promoter is added to the second stage catalyst. 19. Process according to claim 14, wherein the promoter is chosen from among Cr, P, Zr or Ti oxides. 20. Process according to claim 1 , wherein the Co(Ni)-Mo-promoter-zeolite-RE-Al catalyst, used in the second stage, comprises CoO and or NiO and MoO 3 . 21 Process according to claim 14, wherein the contents of the metal oxides are CoO from 0 to 6%, NiO from 0 to 6%, provided that the sum of CoO and NiO is at least 3%, and MoO 3 from 12 to 18% ww on the whole of the catalyst 22 Process according to claim 14, wherein the metal oxides are supported on a combined substrate containing zeolite and alumina modified by RE-oxide 23 Process according to claim 1 , wherein all catalysts are sulfided 24 Process according to claim 1, wherein the ratio between the volume of the catalysts and the volume of the shale oil feedstock varies within the following limits First stage 0 5 to 3 0 Second stage 0 5 to 3 0 25 Process according to claim 1, wherein the ratio between the volume of the catalysts and the volume of the shale oil feedstock varies within the following limits First stage 0 8 to 1 2 Second stage 0 8 to 1 2 26 Process according to claim 1 , wherein the residence time of the feed in the two stages varies, for both stages from 0 3 to 2 0 hours 27 Process according to claim 1, wherein the residence time of the feed in the two stages vanes, for both stages from 0 85 to 1 2 hours 28 Catalyst system for the hydrotreatment of shale oils, comprising a deep desulfurization catalyst, which comprises NiO and/or CoO, and M0O3 on a support, and a deep hydrodenitrogenation catalyst comprising CoO or NiO and Mo 2 O 3 , and a zeolite stabilized on a support 29. As a component of a catalyst system for the hydrotreatment of shale oils according to claim 27, a deep desulfurization catalyst , which comprises NiO and/or CoO, and M0O3 on a support. 30. Catalyst according to claim 29, wherein the support is alumina modified by RE- oxide. 31. Catalyst according to claim 29, wherein the content of the component oxides is: NiO from 0.0 to 3.0 %, CoO from 0.0 to 3.0 %, provided that the sum of the NiO and CoO is at least 0.5%, and M0O3 from 8 to 14 % ww on the whole of the catalyst. 32. Catalyst according to claim 29, wherein the ratio of the combined weight of the oxides to the support weight varies from 0.09 to 0.20 . 33. Catalyst according to claim 32, wherein the ratio of the combined weight of the oxides to the support weight varies from 0.15 to 0.17 . 34. Catalyst according to claim 29, having a packed density from 0.30 to 0.90 g/cc, a surface area from 150 to 300 m^/g, a pore volume from 0.70. to 1.20 cc/g, and an average pore diameter from 1 10 to 140 Angstroms. 35. Catalyst according to claim 34, having a packed density from 0.40 to 0.90 g/cc, a surface area from 200 to 260 m^/g, a pore volume from 0.80 to 1.00 cc/g, and an average pore diameter from 1 15 to 135 Angstroms. 36. Catalyst according to claim 30, wherein the alumina support has an average pore diameter of 140-180 Angstroms. 37. As a component of a catalyst system for the hydrotreatment of shale oils according to claim 27, a deep hydrodenitrogenation catalyst comprising CoO or NiO and Mo 2 O 3 , and a zeolite stabilized on a support. 38. Catalyst according to claim 37, wherein the support is an alumina support modified by RE-oxide. 39. Catalyst according to claim 37, wherein the zeolites are chosen from among the zeolites of the Faujasite type 40. Catalyst according to claim 37, wherein the zeolite content of the catalyst varies from 10 to 40% ww. 41. Catalyst according to claim 37, wherein the zeolite content of the catalyst varies from 25 to 35% ww. 42. Catalyst according to claim 37, comprising a promoter. 43. Catalyst according to claim 42, wherein the promoter is chosen from among Cr, P, Zr and Ti oxides. 44. Catalyst according to claim 37, wherein the contents of the metal oxides are: CoO from 0.0 to 6.0%, NiO from 0.0 to 6.0%, provided that the sum of CoO and NiO is at least 3.0%), and MoO 3 from 12 to 18% ww on the whole ofthe catalyst. 45. Catalyst according to claim 37, wherein the ratio of the combined weight of the Ni, Co and Mo oxides varies 0.20 to 0.45. 46. Catalyst according to claim 45, wherein the ratio of the combined weight of the Ni, Co and Mo oxides varies 0.22 to 0.28. 47. Catalyst according to claim 37, having a packed density from 0.40 to 0.55 g/cc, a surface area from 250 to 400 rn^/g, and a pore volume from 0.60 to 0.75 cc/g. 48. Catalyst according to claim 47, having a packed density from 0.45 to 0.53 g/cc, a surface area from 280 to 320 m^/g, and a pore volume from 0.65 to 0.73 cc/g. 49. Catalyst system according to claim 28, the components of which have been subjected to sulfidization. 50. Process for the hydrotreatment of shale oils, substantially as described and exemplified. 51. Catalysts for the hydrotreatment of shale oils, substantially as described and exemplified.

Inventors: GIVONY DANY, HERSKOWITZ MORDECHAY, LAICHTER SARIT, LANDAU MIRON V, GIVONY, DANY, HERSKOWITZ, MORDECHAY, LAICHTER, SARIT, LANDAU, MIRON, V.
Applicants: GIVONY DANY, HERSKOWITZ MORDECHAY, LAICHTER SARIT, LANDAU MIRON V, PAMA ENERGY RESOURCES DEV LTD, UNIV BEN GURION, BEN-GURION UNIVERSITY OF THE NEGEV, GIVONY, DANY, HERSKOWITZ, MORDECHAY, LAICHTER, SARIT, LANDAU, MIRON, V., PAMA (ENERGY RESOURCES DEVELOPMENT) LTD.
Application year: 1996
Main classification: C10G65
IPC codes: B01J23, B01J29, C10G65