Upgrading Shale Oil Distillation to Clean Fuel by Coupled

Article pubs.acs.org/EF

Upgrading Shale Oil Distillation to Clean Fuel by Coupled Hydrogenation and Ring Opening Reaction of Aromatics on W−Ni/γAl2O3 Catalysts Hongyan Wang,†,‡ Fei Dai,†,‡ Zengxi Li,*,† and Chunshan Li*,‡ †

College of Chemistry and Chemical Engineering, University of Chinese Academy of Sciences, Beijing 100049, People’s Republic of China ‡ Beijing Key Laboratory of Ionic Liquids Clean Process, Institute of Process Engineering, Chinese Academy of Sciences, Beijing 100190, People’s Republic of China ABSTRACT: To obtain clean liquid fuel, a study was conducted on upgrading shale oil by hydroconversion. Various W−Ni catalysts were synthesized and characterized using XRD, BET, TG, H2-TPR, and NH3-TPD methods. The effects of tungsten content and calcination temperature on the physicochemical properties and activity of catalysts were systematically investigated. W−Ni/Al2O3 with 15 wt % W-loading and 550 °C calcination temperature was selected as the optimized catalyst. The product distribution of the primary aromatics in shale oil affected by hydrogenation, ring opening, and cracking reactions was discussed. Finally, key reaction parameters such as pressure, liquid hourly space velocity, and H2/oil volume ratio were optimized. Characteristics of gasoline and diesel fractions were also measured. because of its high content of heteroatoms.30 Thus, using shale oil as a feedstock for clean motor fuels is a difficult research topic, which is significant to improve energy security at same time. Chishti et al.31 concentrated on hydroprocessing of Kimmeridge Clay shale oil. A high hydrogen pressure of 15.0 MPa and temperature of 400 °C were applied. With increasing process duration (8−56 h) the contents of three- and four-ring polyaromatic hydrocarbons decreased, and the contents of single ring and two ring ones increased. Concentration of nitrogen- and sulfur-containing three- and four-ring aromatic hydrocarbons in the oils was reduced. Catalytic hydrotreatment of different distillation fractions of Kukersite oil was studied by Luik et al.32,33 in a laboratory batch autoclave using Co−Mo and Ni catalysts. Fraction boiling at a temperature higher than 320 °C produced 90−91% of refined oil, whereas the heterocompound content decreased from 66.1 to 24.9% along with decreasing. Mathematical models for kinetics of batch wise hydrogenation of shale oil have been calculated by Johannes et al.34 The model that deduces the coefficients and constants can be applied for quantitative evaluation of catalysts and initial oils for hydrogenation. Upgrading of Estonian shale oil rectification residuum fraction has been studied.35 Three types of commercial catalysts that can be used for hydro-purification, hydrocracking, and universal purposes have been used in catalytic hydro-conversion. The highest yield of the fraction boiling below 360 °C was 82.7%. During shale oil conversion to fuel oil, aromatics were typically hydrogenated to form saturated hydrocarbons, and heteroatoms like N, S, and O were removed. But the product distribution of aromatics hydrogenation, which strongly

1. INTRODUCTION Alternative energy sources, such as biofuel, coal-derived liquid fuel, and oil shale, have been extensively explored worldwide because of environmental concerns and fossil fuel depletion.1,2 Global shale oil reserves are estimated to exceed 4.7 billion tons of recoverable oil. Such a large number of potential energy sources have revived interest in shale oil. China has rich oil shale resources, and if those can be used to produce shale oil (a substitute for crude oil), a considerable amount of potential energy amounting to about 436 million tons of shale oil will be obtained.3,4 Hydroconversion of poor quality fuels, such as coal tar and heavy oil, has been investigated extensively. Hydroconversion involved catalysts,5−9 reactors,10−12 and technological conditions13−15 during the hydro-treatment process. The commercial Mo−Co catalyst supported on γ-alumina under required conditions has a long life span and a strong ability to remove high sulfur degrees.16 Tungsten sulfide catalysts performed well during hydrogenation, hydrocracking, hydrogenolysis, and isomerization processes.17−20 Multimetal (W−Mo−Ni) catalysts were used in catalytic hydro-refining of coal tar to liquid fuel with high hydrodesulfurization (HDS), hydrodenitrogeneration (HDN), and hydrodeoxygenation (HDO) levels.5 Due to the high viscosity and complex composition of fuels, rigorous conditions of over 10 MPa and 400 °C are typically used for hydro-treatment. High hydrogen flow rate and low weight hourly space velocity have also been suggested.21−23 Many compounds in oil, including monocyclic aromatics,24,25 polycyclic aromatics,26,27 and heterocyclic compounds,28,29 have also been extensively investigated separately to optimize catalysts. Despite extensive research on coal tar, heavy oil, and model compounds, the transformation process between fractions of shale oil is yet to be completely understood. Shale oil obtained by pyrolysis of oil shale is an unstable, extremely complex mixture of compounds and is a specific one © 2015 American Chemical Society

Received: May 22, 2015 Revised: July 25, 2015 Published: July 27, 2015 4902

DOI: 10.1021/acs.energyfuels.5b01060 Energy Fuels 2015, 29, 4902−4910

Article

Energy & Fuels influenced the yield and quality of gasoline and diesel fractions, was lack of attention. To further understand this process, the present study focused on reactions between products distribution of the primary aromatics in shale oil and activity of catalysts. The conversion of bicyclic aromatics to cycloalkanes and light ring-open products was investigated in detail. Based on the results, the reaction rules of the primary aromatics in shale oil distillate affected by hydrogenation (HYD), isomerization, ring opening (ROP), cracking, and dealkylation reactions was discussed. The effects of pressure, liquid hourly space velocity (LHSV), and hydrogen/oil volume ratio (H2/ oil) on product composition were also studied. The ultimate purpose of the present study is to obtain of gasoline fraction with high octane number and diesel fraction with high cetane number.

compound was the same, the concentrations of compounds in the sample increased linearly manner with peak area percentage. (ii) C, H, O, S, and N elements were analyzed on an Elementar VARIO ELIII (Germany). (iii) Distillation range analysis of feedstock and hydro-treatment products. The distillation range was determined through the Engler distillation method (standard: ASTM D86). (iv) The properties of gasoline and diesel fractions were determined as follows: density of DMA 5000 (Anton Paar, Austria), viscosity of DMA 5000 (Anton Paar, Austria), cetane number, and octane number (SX-100 K, ASTM D613). 2.2.2. Classification of Shale Oil and Hydrogenation Products. Feedstock chemical components and hydrogenation products were identified by GC-MS, and the contribution of each component to the overall composition was calculated using the NIST GC-MS spectral database. To examine the products distribution in the hydroconversion of shale oil, all the compounds identified were divided into eight classifications based on saturation levels as follows: (1) Cycloalkanes (CA): cycloalkanes with C5−C6, alkyl-cycloalkanes with C7−C11; (2) Bicyclic alkanes (BCA): decalins, methyl-decalins, octahydro-1H-indene, methyl-octahydro-1H-indene; (3) Alkyl-benzenes (AB): alkyl-benzenes with C7−C11; (4) Light naphthenes (LN): cycloalkanes (CA) + alkyl-benzenes (AB) with less than 10 carbon atoms; (5) Phenyl cycloalkanes (PCA): alkyl-tetrahydronaphthalenes with C11−C14, alkyl-1H-indanes with C10-C12; (6) Alkyl naphthalenes (AN): naphthalene, alkyl-naphthalenes with C11−C13; (7) Alkanes: alkanes with C4−C27, light paraffins (LP) with less than 6 carbon atoms; (8) O-containings (OC): other containing compounds. In this study, the hydroconversion of aromatics, mainly naphthalene and alkyl naphthalene, played a pivotal role in the creation of reaction products that directly affect the quality of gasoline and diesel fractions. The transformation from polycyclic aromatics to bicyclic aromatics is easy to achieve,36 which is not discussed in the present study. The following classifications of shale oil and hydrogenation products based on hydrogenation reaction are shown in Figure 1.26,37 (1) Hydrogenated products (HYD) are monoand dicycloalkanes, as follows: CA + BCA + PCA. (2) Isomerization products (ISOM): methyl-indane, methyl-perhydroindanes, and dimethyl-pentalanes, with the isomerization reaction followed by ring opening reaction. (3) Ring opening products (ROP) mainly consisted of alkyl-benzenes, alkyl-cyclohexanes, and alkyl-cyclopentanes containing the same carbon atoms as the aromatic reactant (AB + CA, m + n = 4). In addition, side chain carbon atoms from ROP products of alkyl naphthalene were more than 4. (4) LN group comprises alkyl-benzenes and naphthenes with less than 10 carbon atoms (AB + CA, i + j < 4). (5) The LP group comprises light paraffins formed by dealkylation and cracking reactions, with no more than 6 carbon atoms (LP, x ≤ 5). Hydrogenated aromatic products were present in naphthenic ring openings and cracking products. 2.3. Catalyst Preparation and Characterization. The γ-alumina support was obtained from commercial companies (Wish Chemicals Co., Ltd.), and its properties were listed in Table 2. Catalysts were prepared through ultrasonic impregnation method. The supporter was pretreated in a drying oven at 120 °C for 2 h and was impregnated with an aqueous solution of nickel nitrate [Ni(NO3)2·6H2O], ammonium molybdate [(NH4)6Mo7O24·4H2O] and ammonium metatungstate hydrate [(NH4)6W7O24·6H2O] at required Ni/Mo/W molar ratios, and was treated again at 60 °C for 6 h. Ultrasonic vibration with a frequency of 50 kHz was applied during the impregnation process. Then the Ni, Mo, and W precursors were widely dispersed on the pretreated γ-Al2O3 supporter, using the energy from ultrasonic waves. The impregnation precursor was dried overnight at 120 °C. It was then calcined in the one oven with air condition for 1 h at 300 °C (heating rate 5 °C/min). Afterward, it was calcined at desired temperature for another 6 h. The natural coolingoff process was performed, and catalysts with oxidation state were obtained.

2. EXPERIMENTAL SECTION 2.1. Feedstock. The shale oil distillate fraction boiling below 360 °C was used as the feedstock, which was an industrial fraction produced by a fast pyrolysis process. Shale oil distillate with a boiling point above 360 °C is known as heavy residuum bituminous fraction, which is extremely difficult to hydrogenate and has no economic value. The main properties of the feedstock are listed in Table1. Compared

Table 1. Characteristics of Feedstock characteristics

value

density (g/cm3) viscosity (mPa·s) elemental analysis C (wt %) H (wt %) O (wt %) N (wt %) S (wt %) H/C molar ratio distillation range (°C) IBP 10% 50% 90% 95% distribution of components (wt %) alkylbenzenes (AB) phenyl cycloalkanes (PCA) alkylnaphthalenes (AN) alkanes O-containings (OC)

0.8892 16.23 80.39 11.03 5.85 1.98 0.75 1.65 80 179 293 358 360 14.9 5.70 13.6 57.9 7.90

with fossil fuels, one of the outstanding characteristics of shale oil was its O and N contents, accounting for 5.85 and 1.98%, respectively. The molar ratio of H/C was about 1.65, thereby indicating that aromatics were abundant in the shale oil distillate. 2.2. Feedstock and Hydro-Treatment Product Characterization. 2.2.1. Characterization Methods of Shale Oil and Hydrogenation Products. Feedstock and specific hydro-treatment product samples were analyzed through the following standard methods: (i) GC-MS analysis provided a detailed distribution of chemical components. The samples were analyzed through Agilent 6890N/5975B GC/MS equipped with a FID and a HP-5MS capillary column (30 m × 0.25 mm × 0.25 μm). The chromatographic peaks were identified by comparison with the National Institute of Standards and Technology (NIST) library. Assuming that the response factor of each homologous 4903


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Energy & Fuels

Figure 1. Reaction scheme for naphthalene hydro-conversion (compounds in each step are representatives of several isomers with similar structures).

Table 2. Textural and Acid Properties of Catalysts with Different Active Metals active metal (wt %) catalyst

W

Mo

Ni

SBET (m2/g)

pore volume (cm3/g)

mean pore diameter (nm)

total acid amount (μmol/g)

Mo−Ni W−Ni W−Ni−Mo

15 15 15

5

5 5 5

162.3 197.4 190.5

0.59 0.46 0.36

9.4 7.8 8.1

172.5 193.4 294.2

BET surface area and pore size were measured using the nitrogen adsorption isotherm method (Quanta Chrome Instrument NOVA 2000). The samples were degassed at 300 °C for 6 h period to analysis. X-ray diffraction (XRD) was performed using a diffractometer (X′Pert PRO MPD, analytical Co., Ltd.) with Cu Kα radiation and filtered by a graphic monochromator configured at 40 kV and 40 mA. Temperature-programmed desorption of NH3 (NH3-TPD) was measured on an Autochem II 2920 chemisorption analyzer (Micromeritics, U.S.A.). Temperature-programmed reduction of H2 (H2-TPR) experiments were carried out on the same instrument used on NH3-TPD. Sample (50 mg) was heated from room temperature to 900 °C at a rate of 10 °C/min in H2 (10 vol % H2 in Ar). Themogravimetric (TG) analysis was conducted on a Shimadzu DTG-60H analyzer. Samples were heated from room temperature to 800 °C at a rate of 5 °C/min under air in a flow of 30 mL/min. All catalysts were activated via presulfiding by feeding sulfurcontaining aviation kerosene and hydrogen. Catalysts were presulfided in situ using 2.0 wt % dimethyl disulfide dissolved in aviation kerosene and underwent temperature-programmed procedure. The sulfidation of catalysts was performed under 7 MPa (hydrogen partial pressure), 1.0 h−1 (liquid hourly space velocity, LHSV), and 1000v/v (H2/oil volume ratio) conditions. 2.4. Hydro-Processing Procedure. The hydro-processing of the shale oil was conducted in a continuous fixed-bed device as shown in Figure 2. Reactors were made from a stainless steel tube (the length is 110 mm, inside diameter is 15 mm, outer diameter is 26 mm) that was vertically mounted to the heating furnace. The hydrogenation unit consisted of a preheater, mixer, and hydrofining reactors. Catalysts (with a bulk density from 0.95 to 1.05 g/mL) were placed in the middle of the reactor during thermostatic stage. Hydrogen and feedstock mixture was fed from the top of the reactor after preheating and premixing. The products were separated and collected using water cooler, gas−liquid separator, and fluid reservoir. Liquid product sample were collected after the device achieved a stable state.

Figure 2. Experimental device.

3. RESULTS AND DISCUSSION 3.1. Effects of the Active Metal on Catalytic Activity of γ-Al2O3 Supported Catalysts. 3.1.1. Physicochemical Properties of the Catalysts. XRD patterns of the prepared catalysts with different active metals (Mo−Ni, W−Ni, and W−Mo−Ni) are presented in Figure 3. The catalysts were prepared with 5− 15 wt % active metals and calcined at 550 °C for 6 h. The active metal with crystalline was observed in all catalysts, but only few peaks of molybdenum aluminum complex oxide [Al2(MoO4)3] were exhibited on the W−Ni−Mo/γ-Al2O3 catalyst. The BET surface area, pore volumes, and pore sizes of supported catalysts has a different degree compared with γ4904


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Figure 3. XRD patterns of catalysts with different active metals.

Figure 5. H2-TPR profiles of the W−Ni catalysts with different Wloadings.

Table 3. Distribution of Products under γ-Al2O3 Supported Catalystsa

Table 4. Textural and Acid Properties of the W−Ni Catalysts with Different W-Loadings

products (area, %) group distribution

Mo−Ni

W−Ni

W−Ni−Mo

cycloalkanes (CA) alkylbenzenes (AB) bicyclic alkanes (BCA) phenyl cycloalkanes (PCA) alkyl naphthalenes (AN) alkanes O-containings (OC) HYD (CA + BCA + PCA) ROP (CA + AB − LN) light naphthenes (LN) light paraffins (LP)

9.86 10.1 3.57 13.6 2.79 57.8 2.28 27.0 13.6 6.39 0.65

17.7 7.33 2.39 12.9 0.11 57.4 2.17 33.0 14.4 10.6 0.95

10.1 12.6 1.77 15.5 0.03 58.4 1.6 27.4 11.3 11.4 1.82

Wloading (wt %)

SBET (m2/g)

pore volume (cm3/g)



total H2 consumption (μmol/g)

0 5 10 15 20

230.0 228.7 203.2 197.4 152.1

0.60 0.56 0.52 0.46 0.23

9.2 8.3 8.1 7.8 7.4

544.6 158.3 189.5 193.4 281.2

285.0 426.2 518.4 611.3

Reaction condition: T = 360 °C, P = 10 MPa, LHSV = 0.8 h−1, H2/oil = 1000

a

Figure 6. NH3-TPD profiles of the W−Ni catalysts with different Wloadings.

W−Ni−Mo increased obviously more compared with bimetallic catalysts. 3.1.2. Catalytic Activity of γ-Al2O3 Supported Metal Catalysts. The γ-Al2O3 supported metal catalysts were tested at 360 °C. The results are summarized in Table 3. Alkenes disappeared in the products component, which reacted to saturated alkanes through hydrogenation. The conversion of alkyl naphthalenes (AN) increased in the following order under the investigated reaction condition: Mo−Ni < W−Ni < W− Ni−Mo. However, the W−Ni/γ-Al2O3 catalyst exhibited a clearly higher HYD activity (33.0%) and cycloalkanes selectivity (17.7% CA + 2.39% BCA) than that of Mo−Ni and W−Ni− Mo catalysts, which could be attributed to the higher BET surface area and appropriate total acid amount listed in Table 2. W−Ni−Mo catalyst showed higher cracking and dealkylation activity because of the higher contents of LN and LP, which could be due to the high total acid amount.26 Therefore, the

Figure 4. XRD patterns of the W−Ni catalysts with different Wloadings.

Al2O3 support as shown in Table 2. Catalysts have a BET surface area of approximately 160−200 m2/g and a pore volume higher of more than 0.3 cm3/g. The W−Ni catalyst has maximal specific surface area, and the pore volume decreased slightly, which suggests that active metal of W−Ni catalyst has better dispersion in support surface and holes. The Mo−Ni catalyst has a largest mean pore diameter caused by blocking of some micropores. Total acid amount increased in the following order: W−Ni−Mo > W−Ni > Mo−Ni. Total acid amount of 4905


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Energy & Fuels Table 5. Distribution of Products under the W−Ni Catalysts with Different W-Loadingsa products (area, %)

a

group distribution

0 wt %

5 wt %

10 wt %

15 wt %

20 wt %

cycloalkanes (CA) alkylbenzenes (AB) bicyclic alkanes (BCA) phenyl cycloalkanes (PCA) alkylnaphthalenes (AN) alkanes O-containings (OC) HYD (CA + BCA + PCA) ROP (CA + AB − LN) light naphthenes (LN) light paraffins (LP)

0.43 17.4 0.15 2.93 13.1 58.6 7.39 3.51 11.6 6.20 0.11

7.45 11.7 5.30 10.8 3.43 56.0 5.32 23.6 12.3 6.88 0.82

9.54 9.68 3.44 13.1 0.64 57.9 5.70 26.1 12.5 6.74 1.01

17.7 7.33 2.39 12.9 0.11 57.4 2.17 33.0 14.4 10.6 0.95

12.4 15.5 1.40 10.5 0.00 58.2 2.00 24.3 15.4 12.5 1.60


Table 6. Characteristics of Products under the W−Ni Catalysts with Different W-Loadings composition C (wt %) H (wt %) O (wt %) N (wt %) S (wt %) H/C molar ratio distillation range (°C) IBPa 10% 50% 90% 95% a

5 wt %

10 wt %

15 wt %

20 wt %

83.2 13.1 1.02 1.15 1.53 1.89

83.8 13.5 0.91 1.13 0.66 1.93

83.7 14.5 0.82 0.84 0.14 2.08

84.1 13.9 0.78 1.1 0.12 1.98

62.5 145.3 254.2 267.9 307.4

56.2 138.3 255.6 258.4 300.5

50.2 122.5 248.4 250.0 295.1

40.6 118.5 225.6 249.3 296.8

Figure 8. H2-TPR profiles of W−Ni catalysts with different calcination temperatures.

IBP: initial boiling point.

Table 7. Textural and Acid Properties of the 15 wt % W−Ni Catalysts with Different Calcination Temperatures temperature (°C)

SBET (m2/g)

pore volume (cm3/g)



total H2 consumption (μmol/g)

350 450 550 650 750

89.5 117.8 197.4 122.9 74.2

0.13 0.16 0.46 0.18 0.19

6.1 7.2 7.8 8.0 10.0

464.8 374.4 193.4 179.9 167.1

533.9 530.7 518.4 499.2 514.7

catalysts with four different W-loadings, namely, 5, 10, 15, and 20 wt % are shown in Figure 4. During initial loadings, hardly any peak centered on 2θ of 20−30°, and the intensity of WO3 signal increased at higher W loadings. As W-loading on γAl2O3 were increased up to 10 wt %, the evidence for the presence of crystalline WO3 could be seen. The temperature-programmed reduction profiles (H2-TPR) of the catalysts with different W-loadings are presented in Figure 5. The TPR profiles of these samples presented peaks of hydrogen consumption between 400 and 900 °C, and the corresponding hydrogen consumption increased (Table 4). The support almost has no reduction activity. Total H2 consumption increased as W-loadings increased, showing stronger reduction activity. The first sample hydrogen consumption peaks, regarded as NiO reduction, showed decreasing temperatures from 600 to 500 °C. This finding indicated the different interactions of active metal oxide with

Figure 7. XRD patterns of the W−Ni catalysts with different calcination temperatures.

W−Ni/γ-Al2O3 catalyst was chosen as the ideal catalyst, and several important factors that could affect its activity were discussed in the present study. 3.2. Effect of the W-Loading on Catalytic Activity of W−Ni/γ-Al2O3. 3.2.1. Physicochemical Properties of the Catalysts W−Ni/γ-Al2O3. W−Ni/γ-Al2O3 was selected as the best catalyst for hydro-conversion of shale oil. Effect of the tungsten amount supported on W−Ni/γ-Al2O3 was also examined in this study. XRD patterns of W−Ni/γ-Al2O3 4906


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Energy & Fuels Table 8. Products Distribution of W−Ni Catalysts with Different Calcination Temperatures products (area, %) group distribution

350 °C

450 °C

550 °C

650 °C

750 °C

cycloalkanes (CA) alkylbenzenes (AB) bicyclic alkanes (BCA) phenyl cycloalkanes (PCA) alkylnaphthalenes (AN) alkanes O-containings (OC) HYD (CA + BCA + PCA) ROP (CA + AB − LN) light naphthenes (LN) light paraffins (LP)

5.23 19.5 0.56 7.70 3.28 57.9 5.83 13.5 16.8 7.96 0.32

9.40 14.9 4.14 10.5 2.02 56.1 2.94 24.0 16.2 8.11 0.48

17.7 7.33 2.39 12.9 0.11 57.4 2.17 33.0 14.4 10.6 0.95

13.4 10.3 1.56 14.4 0.97 57.7 1.67 29.4 14.6 9.09 0.95

10.5 12.7 5.30 10.6 0.29 56.9 3.71 26.4 14.4 8.84 1.07


Table 9. Characteristics of the Products under 15 wt % W− Ni Catalysts with Different Calcination Temperatures composition C (wt %) H (wt %) O (wt %) N (wt %) S (wt %) H/C molar ratio distillation range (°C) IBP 10% 50% 90% 95%

350 °C

450 °C

550 °C

650 °C

750 °C

83.1 13.2 1.05 1.34 1.31 1.91

83.8 13.7 0.94 0.99 0.57 1.96

83.7 14.5 0.82 0.84 0.14 2.08

83.5 13.7 0.78 1.15 0.87 1.97

83.2 12.5 1.12 2.64 0.54 1.80

68.5 144.1 269.3 340.6 358.4

65.2 127.6 258.2 320.2 333.4

50.2 122.5 248.4 250.0 295.1

54.8 133.4 243.9 256.5 302.5

33.4 117.6 263.8 295.2 314.5

Figure 10. Distribution of products at different LHSV (ROP = CA + AB − LN, reference conditions: T = 360 °C, P = 10 MPa, H2/oil = 1000).

Figure 9. Distribution of products at different pressures (ROP = CA + AB − LN, reference conditions: T = 360 °C, LHSV = 0.8 h−1, H2/oil = 1000).

Figure 11. Distribution of products at different H2/oil ratios (ROP = CA + AB − LN, reference conditions: T = 360 °C, P = 10 MPa, LHSV = 0.8 h−1).

the support.38 Peaks corresponding to WO3 reduction between 650 and 900 °C were observed in all samples and exhibited reduction in steps (WO3 → WO2.9 → WO2), as stated in previous studies.39 W-loading on γ-Al2O3 increased from 5 to 15 wt %, and H2 adsorption capacity of the W−Ni/γ-Al2O3 catalysts increased significantly. Thus, more active sites were formed on the surface of catalysts. Total H2 consumption amount increased slightly above 15 wt % W-loading. This

result, which was verified by XRD, explains the different H2 consumption level assigned to WO3 crystallites. The quantitative analysis of acid sites on the surfaces of samples was determined by stepwise temperature-programmed desorption of NH3 (NH3-TPD). Peaks in each profile corresponded to NH3 desorption related acid sites on the surface of catalysts. The desorption temperature indicated the acid strength of catalysts. The higher temperature of 4907


Article

Energy & Fuels Table 10. Characteristics of Gasoline and Diesel Fractions gasoline fraction yield (%) density (g/cm3) viscosity (mPa·s) S (ppm) N (ppm) distillation range (°C) 10% 50% 90% 95% octane value (RON) cetane number

diesel fraction

experiment

EN-228-2008

experiment

EN-590-2004

20.4 0.776 0.874 37.0 13.1

0.720−0.775

Upgrading Shale Oil Distillation to Clean Fuel by Coupled

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