Production of Clean Fuels by Catalytic Hydrotreating a Low

Article pubs.acs.org/EF

Cite This: Energy Fuels 2017, 31, 11495-11508

Production of Clean Fuels by Catalytic Hydrotreating a Low Temperature Coal Tar Distillate in a Pilot-Scale Reactor Dong Li,*,† Wengang Cui,† Xiangping Zhang,‡ Qinghua Meng,‡ Qiucheng Zhou,‡ Baoqi Ma,‡ Menglong Niu,† and Wenhong Li† †

School of Chemical Engineering, Northwest University, Xi’an 710069, China Yulin Coal Chemical Industrial Upgrading Technology R&D Center, Yulin 719000, China

‡

S Supporting Information *

ABSTRACT: China is one of the largest coal producers in the world and abundant coal tar is produced from coal gasification and carbonization every year. Thus, catalytic hydrotreating coal tar for the production of clean fuel has received substantial attention. In this work, clean liquid fuel was obtained from the catalytic hydrogenation of a low temperature coal tar (LTCT) distillate in a four-stage fixed bed reactor with various catalyst combinations on the pilot scale. Effects of dominant hydrotreating parameters, reaction temperature (290−390 °C), H2 pressure (8−15 MPa), and liquid hourly space velocity (0.2−0.6 h−1), on hydrotreating activity, the intermediate and final products, and chemical components of the hydrogenated oils were evaluated. Meanwhile, a possible reaction scheme for the conversion of alkyl-naphthalenes (AN) and phenols in feedstock was probed. The results showed that the four-stage reacting system was capable of removing sulfur and nitrogen to less than 10 μg/g. Furthermore, after hydrotreating, AN were transformed into decalins, tetralin, and indenes, and alkyl-cycloalkanes (CA) were the main and final products of phenolic compound hydrodeoxygenation (HDO). In addition, coke formation of the spent catalysts was also studied by thermogravimetric techniques, which suggested that the coke deposits are mainly concentrated in the second and third stages of the reactor. The results of this work show that high-quality clean fuels can be obtained through the multistage hydrotreating process with a catalyst gradation technology, which may bridge the gap between fundamental research and industrial production and offer a route for deep processing of LTCT.

1. INTRODUCTION With the petroleum depletion crisis and limited resources of fossil fuels, increasing efforts have been made to develop alternative resources to ensure energy security, such as shale oil, biofuel, and coal-derived liquid.1 In China, coal remains the main source of energy supply and accounts for about 92.6% of the gross reserves of exploitable fossil energy resource. More than 10 million tons of coal tar are produced every year from the coking industry.2 If those resources can be used to produce transportation fuels via hydroprocessing, a considerable amount of potential energy will be obtained.1,3 Nonetheless, due to its high aromatics and asphaltenes contents, especially, rich in heteroatomic compounds (such as S, N, and O),4−7 which limits its application as an energy fuel for combustion and makes it, compared to conventional crude oil, very difficult to upgrade. Considering this, it is necessary to develop new technologies including catalysts, processes, and reactors for the researchers of the world to upgrade alternative resources to liquid clean fuel. Catalytic hydrotreating may be considered as the most convenient way to upgrade a coal tar distillate to produce clean liquid fuels; through hydrotreating processes, heteroatomic compounds and aromatic content can be considerably reduced simultaneously.8 Researchers have conducted extensive studies concerning the hydrotreating processes during coal tar upgrading.5,9−24 Meanwhile, due to the high viscosity and complex structure, researchers have focused on the study of model compounds such as phenanthrene,25 anthracene,26 naphthalenes,27−29 phenols,30−32 and so on, rather than on a © 2017 American Chemical Society

real fraction. Though abundant research on catalytic hydrotreating coal tar has been carried out, the studies still have some shortcomings: (i) Most of the research has focused on one or two kinds of catalysts, not involving the influence of catalyst grading combination technology.12,14,15 (ii) Studies about real reaction behavior of a low temperature coal tar (LTCT) distillate during industrial operation have not been reported. (iii) Most studies have focused mainly on the laboratory scale; the catalyst loading is not more than 50 mL, which means that the flow rate of the feedstock in these previous research works is usually lower than 10 g/h.5,10,12,14−16,20−23,33,34 Therefore, to help build additional knowledge in this area, the present work is to produce clean liquid fuel via catalytic hydrogenation of a LTCT distillate using a four-stage fixed bed reacting system at pilot-plant scale. The catalysts used in this study were characterized by a variety of techniques, and the effects of process variables on the hydrotreating activity, properties of intermediate products, and chemical composition of products were examined in detail. Characteristics of liquid products including gasoline, diesel, and residue as well as stability of catalyst and mass balance were also investigated. Meanwhile, based on the group distributions both in feedstock and products, we also try to speculate on a reaction scheme for AN and phenols in feedstock. The purpose of this work is attempt to bridge the gap between industry and basic research Received: July 10, 2017 Revised: September 6, 2017 Published: September 11, 2017 11495

DOI: 10.1021/acs.energyfuels.7b01971 Energy Fuels 2017, 31, 11495−11508

Article

Energy & Fuels

water. A 50 g portion of Al2O3-SiO2 support was dried at 120 °C for 4 h in advance to remove the surface water. Then, the Al2O3-SiO2 support was impregnated with the aqueous solution at room temperature for 10 h, accompanied by a 40 kHz ultrasonic vibration. After that, the impregnated catalyst was dried at 120 °C overnight, calcined at 500 °C for 5 h. 2.2.2. Preparation of NiMoW/Al2O3-HY Hydrocracking Catalysts. The support of HC catalyst was prepared in a similar method to the HDS catalyst, but 5% of the HY zeolite (Qingdao Huicheng Environmental Technology Co., Ltd., P.R. China.) was added during the kneading process. That is 50 g P-DF-03-LS pseudoboehmite (CHALCO Shandong Co., Ltd., P.R. China.), 30 mL water solutions of HNO3, 5 wt % sesbania powder, etc., were mixed together in a kneader to prepare paste. Then, the paste was shaped into trilobe extrudates with a cross-section diameter less than 1.5 mm and a length of 0.5−1.0 cm on the F-26 Twin Screw extruder. After, extrudates were dried at 120 °C for 6 h and calcined at 550 °C for 5 h in the air flow. NiMoW/Al2O3-HY catalyst was prepared through step impregnation method. First, 50 g Al2O3-HY support was impregnated in an dissolving solution containing 23.6 g ammonium metatungstate ((NH4)6H2W12O40·nH2O, Sinopharm Chemical Reagent Co., Ltd., P.R. China) and nickel nitrate (Ni(NO3)2·6H2O, J&K Scientific Ltd., Beijing, P.R. China). The techniques of impregnation and drying were present in ref 6. Then, the NiW/Al2O3-HY catalyst was impregnated in a solution containing 14.3 g ammonium molybdate (NH4)6Mo7O24· 4H2O for 24 h. Impregnated samples were dried at 120 °C for 6 h and calcined at 500 °C for 4 h in air. Metal contents of all catalysts were measured by a X-ray fluorescence spectrometer (ZSX Primus II, Rigaku, Japan); N2adsorption measurements were performed on a with a Micromeritics ASAP 2010 M instrument at −196.15 °C to determine the specific surface area (SBET) and pore volumes of the catalysts. The BJH method was used to get information on mesopore size distributions. FT-IR was employed to determine the framework vibration of all catalysts by an infrared spectrometer type JASCO FTIR 4100. The surface appearance of catalysts were examined by a SIGMA 500 scanning electron microscopy (Carl Zeiss, Company). The crystalline structure of catalysts were revealed by X-ray powder diffraction (XRD) analysis, which was carried out in a Rigaku DMAX-RB diffractometer using Cu Kα radiation (λ = 0.15406 nm) under 40 kV, 150 mA; these samples were measured in the scanning angle (2θ) range from 10° to 80° with a scan rate of 2°/min. The amount of coke deposited on the spent catalysts during on-stream reaction was studied by temperatureprogrammed oxidation (TPO) performed on a thermogravimetric TGA/SDTA851 equipment (Mettler Toledo), by raising the sample temperature to a final temperature of 900 °C at a rate of 10 °C min−1 in a 20% O2/N2 gas mixture. 2.3. Reaction System and Catalyst Loading. The pilot-plant test was carried out in a continuous four-stage fixed-beds system. A detailed description of the apparatus was depicted in Figure 1. The entire reaction system consists of the feeding unit, the reaction unit, the hydrogen circulation unit, and the product separation and collection units. The feeding unit included a coal tar supply line and a high-pressure hydrogen supply line. The reaction unit was mainly made up of four reactors: HDM, HDS, HDN, and HC reactors. The gas and liquid phase is controlled and measured by a mass flowmeter, an electronic balance, and a precision metering pump, respectively.

by examining the correlation between the product composition of the pilot plant and the catalytic reaction mechanism, in addition, providing a reliable guide to the industrial scale-up test.

2. EXPERIMENTAL SECTION 2.1. Feedstock. The LTCT used in the this study was obtained from the Taida Carbonnification Co, Ltd., pyrolyzed Fugu low rank coal at 550−600 °C, Shaanxi province, China. The distillate >400 °C of the coal tar was known as coal tar pitch, which can damage hydrogenation equipment and was extremely difficult to hydrogenate. Therefore, a distillate with a boiling point 360 °C) × 100% Moilfeed

(3)

M water × 100% Moilfeed

(4)

yield of water =

yield of light hydrocarbon gases Moilfeed − Mliquidoil − M H2 feed = × 100% Moilfeed

(5)

Mliquidoil = (Mgasoline + Mdiesel + M residua + M water) × 100%

conversions of S, N, and O = Figure 2. Temperature program of catalyst presulfiding. 11497

XF − XP × 100% XF

(6)


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Energy & Fuels Where M represents the mass of all kinds of products and feedstock, XF and XP stand for the contents of S, N, and O in feedstock and product, respectively.

SBET and relatively smaller pore size was graded to make it more suitable for the entire reaction system. Furthermore, HDS and HDN catalysts have a roughly equivalent pore diameter and volume. The morphologies and microstructures of all catalyst were analyzed by SEM (see Figure S2 of the Supporting Information). It is clearly seen from the images that the surface of the HDG and HDM catalysts are relatively coarse and there are few macroporous structures. HDS and HDN catalysts presented a more regular appearance. The morphology of HC catalyst is relatively compact due to its material properties. Meanwhile, as shown in Figure S2, no metal aggregation was found on the surface of all the catalysts, indicating that the surface active metal distributions of all the catalysts are uniform. Figure S3 of the Supporting Information showed the FT-IR spectra for HDS, HDN, and HC catalysts at wavenumbers of 600−4000 cm−1. The peak position and shape of these three catalysts are very similar. A large peak appeared at 3856, 3750, 3610, 2355, 1697, 828, and 637 cm−1. Among them, 3856, 3750, and 3610 cm−1 are attributed to the −OH vibrational peaks of alumina surface, which represent the basic, neutral and acid −OH group of the alumina.36,37 In addition, a larger peak appears at 1059 cm−1 in the HC catalyst spectrum, which can be attributed to the stretching vibration of Si−O−Si bond.36 The bands at 828 and 637 cm−1 are the characteristic peaks of the alumina support and are generally belong to the bending vibrations of the Al−OH and Al−O−Al bonds in the framework,38 respectively. Furthermore, the band appearing near 2355 cm−1 has been attributed to stretching vibrations of Al−OH−Si,39 this hydroxyl group has a strong acidity, similar to the cationic zeolite, and its tetrahedral Si4+ is replaced by pentahedral Al3+. It is obvious that, for HC catalyst, the band at 2355 cm−1 was higher than that of HDN and HDS, which indicated that the HC catalyst had more acidic sites. X-ray diffraction pattern of the five catalysts are shown in Figure S4 of the Supporting Information. All the catalysts exhibited obvious diffraction peaks at 2θ = 37.60°, 45.9°, and 67.0°, which belong to the (311), (400), and (440) planes of characteristic diffraction peak of Al2O3 (PDF 10-0425). In addition, no other characteristic peaks such as NiO, WO3, and NiAlO4 were found in all of the hydrogenation catalysts, indicating that the active metal particles were homogeneously distributed over the support and did not exceed the dispersion threshold. The NH3-TPD curve of HDS, HDN, and HC catalysts were displayed in Figure 4a. It can be seen from Figure 4a that HDS and HDN catalysts have an obvious desorption peak at about 120 °C, while the HC catalyst has two desorption peaks centered at around 180 and 400 °C, respectively. This indicates that, compared with HDS and HDN catalysts, HC catalysts have relatively strong acidity.40 In order to obtain the specific

3. RESULTS AND DISCUSSION 3.1. Catalyst Characterization. N2-adsorption−desorption isotherms and corresponding pore size distribution profile of all catalysts are given in Figure 3. According to IUPAC

Figure 3. N2 adsorption−desorption isotherms (A) and BJH pore size distribution curves (B) of five hydrogenation catalysts.

classification, all isotherms are attributed to IV type with hysteresis loops of type H3.35 The shape of this isotherm is typical for mesoporous catalysts with cylindrical type of pores. As shown in Figure 3, all catalysts have a narrow and similar pore size distribution that centered at 18, 15, 10, 8, and 7 nm, respectively. The respective values of BJH pore size distribution and BET surface area are collected in Table 3. The HDG catalyst has the largest mean pore diameter and pore volume while the HC catalyst has maximal specific surface area, which is mainly due to the fact that the coal tar has been subjected to the upstream hydrogenation reaction, macromolecule compounds have been running out. Thus, HC catalyst with larger

Table 3. Data of Pore Size Distribution Curves and SBET of All Catalysts pore volume distribution (vol %) 2

3

type

SBET, m /g

pore volume, m /g

average pore diameter, nm

2−5 nm

5−10 nm

10−20 nm

>20 nm

HDG HDM HDS HDN HC

92.54 131.12 241.96 253.5 331.6

0.67 0.56 0.47 0.43 0.38

15.50 13.47 7.14 6.45 5.47

2.31 20.14 11.12 14.55 18.95

33.26 40.29 63.40 55.87 50.98

40.21 24.09 15.75 14.99 6.99

13.63 6.21 8.87 5.68 4.26

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(that is, THDS = 290−350 °C, THDN = 310−370 °C, and THC = 330−390 °C) was covered in the experiments. To exemplify this, these temperatures were denoted hereinafter as ref-40, ref20, ref, ref+20, respectively. As presented in Table 4, diesel yield decreases with operating temperature, while gasoline fraction increased. However, diesel Table 4. Effect of the Temperature on the Yield and Properties of Producta items

unit

gases (C1−C4) gasoline diesel residue

wt wt wt wt

density (20 °C) viscosity (80 °C) H/C molar ratio O content N content S content

g/mL mm2/s

% % % %

wt % μg/g μg/g

ref-40 °C Yield 0.29 29.2 66.6 4.2 Properties 0.895 3.42 1.54 1.21 43 216

ref-20 °C

ref °C

ref+20 °C

0.94 30.8 66.7 2.5

1.03 32.9 65.0 2.0

2.42 33.5 64.2 1.3

0.857 3.34 1.62 0.23 23 72

0.853 3.32 1.66 0.12 8 12

0.812 3.22 1.72 0.11 3 6

Other experimental conditions: LHSV = 0.3 h−1, PH2 = 13 MPa, and H2:oil volume ratio = 1600:1. a

is the dominant fraction at all reactor temperatures, even if in the highest operating temperature at ref+20 °C. Specifically, when the temperature reached ref+20 °C, the lowest diesel fraction and the highest gasoline fraction were obtained. Moreover, the gases products also increased with an increasing operating temperature, indicated that higher temperatures not only favor gasoline production but also gas (C1−C4) formation, and it also can be deduced that a small part of gasoline and gas were formed by cracking of the diesel oil. As shown in Table 4, the density of products decreased from 0.895 to 0.812 g/mL and for the viscosity from 3.42 to 3.22 mm2/s with a temperature increase in all cases. Additionally, the temperature increase from ref-40 to ref+20 °C increased the H/C molar ratio from 1.54 to 1.72 and suppressed the yield of residue from 4.2 to 1.3%, indicating a greater saturation level and that a few dealkylation and hydrocracking reactions may have occurred. The denitrogenation and desulfurization results as a function of temperature were also illustrated in Table 4. It is observed that increasing temperature from ref-40 to ref+20 °C, the HDS and HDN activities increased from 89.9 to 99.9% and 81.3 to 96.5%, respectively. At ref-40 °C, the sulfur content was reduced to approximately 43 μg/g and, at ref+20 °C, to 3 μg/g; meanwhile, the nitrogen level can be reduced only to 216 μg/g at ref-40 °C, and at ref+20 °C to 6 μg/g, which suggested that, compared to nitrogen compounds in feedstock, most of the Scontaining species were more reactive which can be much more easily removed from coal tar during the hydrogenation process, which is in accordance to literature.15,19,34 Most of nitrogen was found to occur as aromatic compounds (such as pyridines, quinolines, pyrroles, indoles, and carbazoles, and so on), due to the strong C−N bond in aromatic systems,42−45 which are resistant to HDN. Meanwhile, it is generally accepted that HDN of aromatic N-containing compounds occurs via a twostep process involving complete hydrogenation of the aromatic rings followed by irreversible C−N bond scission (hydrogenolysis).45 In contrast, HDS is a one-step process, which does not require to completely saturate a S-containing aromatic ring

Figure 4. NH3-TPD profiles (A) and acid distribution (B) of HDS, HDN, and HC catalysts.

distribution of acidity on these catalysts, the experimental NH3TPD curves were fitted by Gaussian method to obtain multiple desorption peaks (not shown in the Figure 4), in which three peaks were used to represent the weak (400 °C) acidities, respectively. By integrating each peak, three acid centers and the total acid amount were obtained (Figure 4b). As shown in Figure 4b, the order of weak acidity of catalyst is HDS > HDN > HC, and the order of strong acid is HC > HDN > HDS in all catalysts. 3.2. Effect of Reaction Parameters on Hydrotreating Activity. Process variables, such as pressure, temperature, LHSV, and H2:oil ratio, have a remarkable effect on hydrotreating reactions. Numerous studies5,12,15,23,24,34,41 have shown that properly increase the pressure and lower the LHSV are advantageous to upgrade a coal-derived liquids though hydrogenation. In order to ascertain the optimal operating conditions to produce qualified clean liquid fuel with acceptable properties, the effects of these operating parameters on the comprehensive reaction performance have been investigated detailedly, which were estimated by the products yields, properties, and conversions of S, N, and O in the following serial tests. 3.2.1. Effects of Temperature. Hydroprocessing temperature is a key parameter for catalyst performance, which was normally considered the most economic and easiest ways to control the hydrotreating process. Increasing in temperature normally not only increases the rate of the process but also the conversion. In this section, the effect of the temperature was studied at a constant pressure of 13 MPa, LHSV of 0.30 h−1, H2:oil volume ratio of 1600:1, and the temperature range of 290−390 °C 11499


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Energy & Fuels but mainly involves direct scission of the C−S bonds.45 However, high reaction temperature is easy to promote the occurrence of cracking reactions and coking of reactants,46,47 which easily accelerates the catalyst deactivation. Therefore, THDS = 330 °C, THDN = 350 °C, and THC = 370 °C were selected as optimal temperatures in this study. 3.2.2. Effects of Pressure. High hydrogen pressure not only was conducive to increase impurities removal, but also to slow down the catalyst deactivation. The efforts of reaction pressure on the product properties were investigated at stepwise pressures of 8, 10, 13, and 15 MPa, while LHSV = 0.3 h−1, reaction temperature THDS = 330 °C, THDN = 350 °C, THC = 370 °C, and H2:oil volume ratio of 1600:1 were fixed. The results of yield of products with regard to the pressure are shown in Table 5. As the total hydrogen pressure increased

capital expense at the same time, but if the pressure gets too high, it will also result in some security problems and lead to higher operating costs. Given all these considerations, P = 13 MPa was chosen as the optimal pressure. 3.2.3. Effects of LHSV. LHSV is a method for relating the reactant liquid flow rate to the reactor volume at a standard temperature, which indicates how many reactor volumes of feed can be treated in a unit time. In other words, it reflects the contact time of the reactant oil with the catalysts. For the same catalyst bed, a reduction of LHSV means more time for conversion. However, a too small LHSV will increase the hydrogen consumption of the reaction while also reducing the output in unit time, which is not conducive to industrial production. In the light of the characteristics of LTCT used in this study, four typical LHSV values of 0.2, 0.3, 0.4, and 0.6 h−1 were selected and reaction temperature, pressure, and H2/oil volume ratio were fixed at the reference conditions. Table 6 displays the effect of the LHSV on the yield of products. It can be observed that the yield of both diesel and

Table 5. Effect of the Pressure on the Yield and Properties of Producta items gases (C1−C4) gasoline diesel residue density (20 °C) viscosity (80 °C) H/C molar ratio O content S content N content

unit wt wt wt wt

% % % %

g/mL mm2/s wt % μg/g μg/g

8 MPa

10 MPa

Yield 1.03 1.09 26.4 29.3 69.2 66.9 4.4 3.8 Properties 0.923 0.898 3.24 3.27 1.51 1.58 1.45 1.21 80 32 391 226

13 MPa

15 MPa

1.03 32.9 65.0 2.0

0.13 33.9 63.2 2.1

0.853 3.32 1.66 0.12 8 12

Table 6. Effect of the LHSV on the Yield and Properties of Producta properties

0.831 3.33 1.67 0.21 12 40

Other experimental conditions: LHSV = 0.3 h−1, THDS = 330 °C, THDN = 350 °C, THC = 370 °C, and H2:oil volume ratio = 1600:1

a

from 8 to 15 MPa, significant effects on yield and quality characteristics of the hydrotreatment products were observed, i.e. the yield of diesel decreased from 69.2 to 63.2% and gasoline increased from 26.4 to 33.9%, and density and viscosity decreased from 0.923 to 0.831 g/cm3 and from 3.24 to 3.33 mm2/s, respectively. In addition, H/C molar ratio increased in the hydrotreatment products increased significantly from 1.51 to 1.67, indicating enhanced aromatic saturation under higher pressure. Table 5 further illustrates that the N and S contents stepwise but distinctly reduced from 321 and 80 μg/g to 18 and 52 μg/g, respectively. Moreover, there was a marked decrease of N and S contents with raising the pressure from 8 to 13 MPa, whereas the corresponding values underwent a slight decrease as the pressure increased from 13 to 15 MPa. The reason for this phenomenon is considered to be that the S and N heteroatom could be easily removed via hydrogenation in the early phase, but due to the steric hindrance, hydrogenation to remove of S and N atoms form the unconverted S- and N-containing compounds with complex structures was extremely difficult in the later phase.45 Additionally, there was no sign of an equilibrium limitation, even at 15 MPa, which might be ascribed to the higher pressure making much more hydrogen gas sufficiently dissolved into the feedstock and better contact with the catalysts simultaneously. Meanwhile, there also occurred an increase of HDO in the hydrotreatment products. Elevated pressures can improve the heat- and mass-transfer rates and the gas solubility, to handle large gas volumes at less

unit

gases (C1−C4) gasoline diesel residue

wt wt wt wt

density (20 °C) viscosity (50 °C) H/C molar ratio O content S content N content

g/mL mm2/s

% % % %

wt % μg/g μg/g

0.2 h−1 Yield 1.31 33.2 67.7 1.1 Properties 0.828 3.01 1.63 0.18 20 54

0.3 h−1

0.4 h−1

0.6 h−1

1.03 32.9 65.0 2.0

0.99 29.7 66.2 4.1

0.95 28.9 65.4 5.7

0.853 3.32 1.66 0.12 8 12

0.864 3.49 1.57 0.73 25 155

0.873 3.77 1.55 1.34 50 293

Other experimental conditions: PH2 = 13 MPa, THDS = 330 °C, THDN = 350 °C, THC = 370 °C, and H2:oil volume ratio = 1600:1. a

gasoline fractions slightly decreased from 67.7% and 31.2% to 65.4% and 28.9% with increasing LHSV, respectively. And there was an increase in residue from 1.1 to 5.7% at the same time, suggesting that a higher LHSV can not provide enough time for residue conversion. The viscosity, density, and H/C of hydrotreatment products were also decreased with the increase of LHSV. As shown in Table 6, sulfur and nitrogen compounds were removed to a level of less than 10 μg/g in products at a LHSV 0.2 h−1. In addition, nitrogen removal was actually more influenced than sulfur removal by LHSV, which also indicates that HDN, compared to HDS, requires a longer time to function due to HDN needing two reaction steps.15,45 However, taking into account the industrial production costs, a LHSV = 0.3 h−1 was selected as the optimal value. 3.3. Comparison between Intermediate and Final Products. In order to fully comprehend this hydrotreating processes and certify the activity of five hydrotreating catalysts, the intermediate products from each of the reactors were investigated, including distillation range, H/C, S, N, and metal content, at the reference reaction conditions, respectively. The comparison of each of the products is shown in Table 7. As illustrated in Table 7, all of the catalysts have heteroatom removal activity, but the selectivity is not the same. HDG and HDM catalysts were filled in the first reactor, which mainly 11500


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number of decalins, tetralins, and CA appeared in the products, which accounted for 4.59%, 12.57%, and 24.28%, respectively. In addition, the contents of AB and indenes also increased dramaticlly. An interesting phenomenon is that the content of OC compounds was not reduced but increased, indicating that some phenolic compounds were converted into OC species during the HDO process. As demonstrated in Figures 5 and 6, and Table S1, AN and phenols account for nearly 50% of the feedstock except for the 33.27% of alkanes. Therefore, these two substances were chosen as initial reactants in the network. As shown in Figure 6, the bicyclic compound (AN + tetralins + decalins + indenes = 25.29%) in the hydrogenation product, which was close to the content of (AN + indenes = 23.86%) in the feedstock. And the content of these three kinds of bicyclic compounds in the order of tetralins > indenes > decalins, and decalins + indenes + PO equal to 13.01%, which was very close to the content of tetralins 12.57%, indicating that tetralins, decalins, and indenes may be the main hydrogenation products of AN. It is wellknown that polycyclic aromatic hydrocarbons can not be converted to tricyclic saturated alkanes because the saturation of the last ring is much harder than the saturation of the first two rings, and the ring-opening reaction takes precedence over full saturation.44,45,48,49 Based on the previous studies on the hydrogenation of AN model compounds27−29 and chemical compositions above, we try to speculate the hydrogenation pathway of AN in coal tar (Figure 7): First, AN is converted to alkyl tetralins, and then, this step is followed by three pathways (i) hydrogenation to form decalins; (ii) hydrogenation isomerization to indenes and then transformation into to AB by ring-open reaction; (iii) generation of a small amount of PO by ring-opening reaction. From the distribution of the products shown in Figure 6 and Table S2 of the Supporting Information, path ii was more favored during AN hydrogenation. From the distribution of products shown in Figure 6, the AN hydrogenation process favored path ii, while it was uncertain whether the indenes would continue hydrogenation to form AB and that the AB was not further hydrogenated to CA. In addition, as can be seen from Figure 6 and Table S2 of the Supporting Information, the indenes and the decalins are not continuously hydrogenated or isomerized to form the alkyl perhydroindene. According to the GC-MS analysis of Figures 5 and 6 and Tables S1 and S2 of the Supporting Information, the monocyclic compound (phenols + AB + CA = 30.1%) in the hydrotreatment product is substantially equal to that (phenols + AB = 28.55%) in the feedstock at all treatment ranges, suggesting that the phenols and AB are mainly converted to CA after hydrogenation. Meanwhile, cyclohexanol, an intermediate of the phenol HDO, were identified in hydrotreatment products which were in good agreement with the reactions converting phenols into CA. Moreover, the proportion of AB with C10 and C11, the same carbon number as in PO, was less than 1%. Therefore, the conversion of AB to PO during hydrotreatment was likely insignificant. Therefore, based on the model compound of phenols HDO30−32 and the composition of hydrotreatment products, the reaction pathway of the phenolic compounds in coal tar was simply summarized as follows (Figure 7): a part of phenols were converted to AB by direct deoxygenation, and the remainder were hydrogenated to cyclohexanol followed by the conversion of HDO to CA. In addition, no C6 straight-chain alkanes were found in the product, indicating that CA was the final hydrogenation product of phenolic compounds.

Table 7. Properties of Intermediate Products Derived from Different Reactorsa properties

unit

S content N content metal content H/C

μg/g μg/g mg/L

IBP 10%/30% 50%/70% 90%/FBP

°C °C °C °C

first reactor 876 4132 10.56 1.13

second reactor 136 2647 5.43

1.32 Distillation 159 147 188/221 161/200 274/322 253/299 353/386 344/378

third reactor

fourth reactor

56 488 1.01

23 72 0.45

1.54

1.62

104 141/176 249/274 340/386

85 121/162 229/257 348/388

Reaction conditions: LHSV = 0.3 h−1, P = 13 MPa, THDS = 330 °C, THDN = 350 °C, THC = 370 °C, and H2:oil volume ratio = 1600:1.

a

performs raw material preheating and pretreatment, protecting the highly active downstream catalysts from poisoning by carbon and metal deposition. After HDM reaction, metal content was significantly reduced from 93.98 mg/L in the feedstock to 10.56 mg/L. In addition, it is apparent that HDS mainly exists in the second reactor while HDN is in last two reactors. That is because we have done a lot of catalyst screening tests in the early stage of the experiment. As to HC catalyst, in addition to the high nitrogen removal activity, it also has a certain hydrocracking activity so that more gasoline can be obtained. Furthermore, it is pretty obvious that the distillation data of products showed a decreasing tendency from the first reactor to the fourth reactor, indicating that the liquid products were gradually cleaned and that the catalyst grading combination as well as the multistage process were essential for obtaining high qualified clean fuel. 3.4. Chemical Components of the Hydrogenated Oils. The compositions of the both feedstock and hydrotreated products were identified by GC-MS analysis, and the contribution of each component to the overall composition were calculated by area normalization method. The results are illustrated in Figures 5 and 6 and Tables S1 and S2 of the Supporting Information. As displayed in Tables S1 and S2, the organic compounds detected in feedstock and hydrotreated products can be classified into alkanes, phenols, CA, decalins, tetralin, alkenes, alkylbenzenes (AB), phenyl olefins (PO), AN, indenes, other oxygen containing (OC) compounds such as alcohols, ketones, aldehydes, and “others”, and FPA (fluorenes, phenanthrene, anthracenes). As Figure 6 illustrates, phenols, alkanes, and AN are the dominant group components in feedstock from GC-MS analysis, accounting for 25.8%, 33.27%, and 22.7% of the feedstock, respectively, whereas no decalins and tetralin were detected. Most of phenols detected in the feedstock are polymethylphenols with methyl number from 1 to 4. In addition to phenolic compounds, the total relative content of aromatic in the GC/MS detectable species from feedstock are 31.84%, mainly including AB, AN, FPA, and indenes. As exhibited in Figures 5b and 6 and Table S2, after hydrotreating, the phenolic compounds in the products almost disappeared. AN decreased continuously from 20.7 to 1.01%, and the tricyclic aromatics (FPA) reduced from the original 6.66 to 1.66%. The content of alkanes changed not very noticeably increasing from 33.27 to 35.54%, indicating that the hydrocracking reaction was not particularly significant during the whole reaction process. As shown in Figure 6, a large 11501


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Energy & Fuels

Figure 5. Total ion chromatograms of (a) feedstock and (b) hydrogenated product by GC/MS analysis.

FT-IR was also carried out to study differences in the organic groups before and after hydroprocessing. As presented in Figure 8, the absorption bands between 3100 and 3690 cm−1

are attributed to stretching vibrations of the OH groups in coal tar feedstock.37 The peaks at 1460, 1380, and 2920 cm−1 are the absorption peak of methyl and methylene group in aliphatic 11502


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Figure 6. Distribution of group components in feedstock and hydrogenated product from GC/MS analysis.

Figure 7. Reaction scheme of AN and phenols in LTCT. The legend is as follows: HYD hydrogenation; ISOM isomerism; ROP ring opening; DDO direct deoxygenation.

hydrocarbon.15 The peak centered at 1600 cm−1 is caused by CC stretching vibration.5 Meanwhile, the absorption peaks at 822 and 752 cm−1 belong to the vibration of substituted polycyclic aromatic group.6 The FT-IR spectrum of the product is similar to the feedstock, but there are also some obvious differences, including the following: (1) The absorption band between 3690 and 3100 cm−1 almost disappeared, indicating that the OH group has been removed by HDO reaction. (2) The absorption peak intensity of alkanes at 2920 cm−1 increased and the peak of aromatics at 822 cm−1 became weaker, indicating that the content of alkane increased and the aromatic content decreased in the product. (3) Intensity of the peak at 1600 cm−1 was weakened significantly, suggesting that the CC unsaturated double bond was drastically reduced. These results are consistent with the GC-MS analysis mentioned above. 3.5. Properties of Diesel and Gasoline Products. The ultimate goal of the catalytic hydroprocessing a coal tar distillate is to produce a standards-compliant motor fuels. To bring

meaning to the discussion of this process, the fuel properties of those produced from pilot plant experiments have to be compared with standard fuel specifications. Thus, the liquid products, dependent upon the reference operating conditions, were separated into three fractions through the Engler distillation method: gasoline fraction (360 °C), which can be seen in Figure 9. 3.5.1. Characteristics of the Gasoline Fraction. The quality and properties of gasoline fraction were listed in Table 8. The results showed that the sulfur content, corrosion with Cu, and distillation have reached the national standard 93# standard gasoline, but RON did not meet the specification (82.3 instead of ≥93), which mean that it could not be directly used as a motor gasoline. Table S3 of the Supporting Information summarizes the mainly 26 compounds identified in the feedstock by the GC/MS method, and a gas chromatogram of the organic phase is depicted in the Figure S5. As exhibited in Table S3 and Figure S5, substituted alkyl-cyclohexanes, for 11503


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Energy & Fuels Table 8. Quality and Properties of Gasoline Fuel properties

gasoline

density (20 °C)/(g/mL) viscosity (50 °C)/ (mm2/s) H/C molar ratio S content/(μg/g) N content/(μg/g) RON corrosion with Cu (50 °C, 3 h) existent gum/(mg/100 mL) Reid vapor pressure/(kPa) induction period/(min) 10% 50% 90% FBP

Figure 8. FT-IR spectra of feedstock and product oil.

0.788 3.8 2.1

GB19147-2011 not limited not limited not limited ≤150 not limited ≥93 ≤1

6 12 82.3 1

2.1 ≤5 33.5 ≤72 501 ≥480 distillation/(°C) 80 ≤70 119 ≤120 187 ≤190 203 ≤205

test method GB/T 1884 GB/T 265 element analyzer GB/T 380 GB/T 17674 GB/T 5487 GB/T 5069 GB/T 8019 GB/T 8017 GB/T 8018 GB/T 6536

Table 9. Quality and Properties of Diesel and Residual properties

example methyl-, ethyl-, 1-ethyl-3-methyl-, propyl-cyclohexanes, and so on, are the dominant group components which accounting for more than 75% of the total peak area. Among these materials, ethyl-cyclohexane appeared as the most outstanding one with a content of 16.06% in gasoline. Furthermore, AB (e.g., m-xylene, o-xylene, and 1,2,3-trimethyl-benzene) also prevailed with a content of about 18%. These substituted cyclohexanes and AB may originate from the corresponding phenols with different branched chains. 3.5.2. Characteristics of the Diesel Fraction. As listed in Table 9, compared to gasoline fraction, diesel complys with all of 0# diesel specifications which could be directly used as motor fuels without further upgrading. Furthermore, occurrence of more than 260 compounds constituting over 90% of the total samples were detected in diesel fraction by using a GC-MS/FID method, and the chromatogram of the organic phase is depicted in Figure S5 of the Supporting Information. Saturated ring-containing hydrocarbons (e.g., tetradecahydroanthracene, decalin, alkyl-naphthalenes, etc.) as well as straightchain paraffins (e.g., heptadecane, octadecane, hexacosane, etc.) prevailed in the diesel fraction. Besides, substituted and partially saturated naphthalenes were also identified with a comparatively low content. In other words, this means that the diesel fraction also have a fairly low aromatics content. Reducing heteroatom content (such as S and N) along with aromatic content is generally desirable in regard to the diesel quality, a high aromatic content in diesel fuels not only contributes significantly to the formation of environmentally harmful

diesel

GB19147-2009

residual

test method

density (20 °C)/ (g/mL) viscosity (50 °C)/ (mm2/s) H/C molar ratio

0.848

0.81−0.85

0.915

GB/T 1884

4.4

3.0−8.0

5.8

GB/T 265

1.98

not limited

1.65

S content/(μg/g) N content/(μg/g)

10 22

≤350 not limited

13 24

cetane number flash point/(°C) solidifying point/(°C) corrosion with Cu (50 °C, 3 h) acid value/ (KOH mg/mL) carbon residue/(wt %) ash/(wt %)

57.1 60 −4.2 1

≥45 ≥55 ≤0 ≤1

140 12.6

element analyzer GB/T 380 GB/T 17674 GB/T11139 GB/T 261 GB/T 510 GB/T 5069

5.12

≤7

6.42

GB/T 258

8.31 1.15

GB/T 268 GB/T 508

364 390 449

GB/T 6536

10% 50% 90%

HDS > HDM > HDN, which is mainly due to the different amount of catalyst active metal loading (see Table 2). The third weight loss in the TG profiles corresponding to DTG peak centered at 500 °C is ascribed to the oxidation of coke.50,52 Considering the amount of coke formed during 400 h of on-stream operation, the observed amount of coke follows the trend: HDS (6.2%) > HDN(5.8%) > HC(1.9) > HDM (1.7%) > HDG (0.9%), indicating that the coke is mainly formed with the second and third stages of the reactor. The acidity of the catalyst is directly related to the formation of carbon deposition, and the acid strength directly affects its cracking activity. The NH3-TPD characterization analysis of the catalyst shows that the HC catalyst has a particularly high acidity (see section 3.1), but, as shown in Figure 11, its carbon content is lower than that of HDN and HDS catalysts, indicating that the coal tar has been greatly improved after the hydrogenation of the first three reactors, as evidenced by the conclusions in section 3.3. However, an interesting phenomenon is that HDG and HDM catalysts are filled in the first-stage reactor, but their coke content is lowest, this is mainly due to the fact that the active metal loading of the HDG and HDM catalysts is relatively low and the pore structure is single and mostly large pores (see Table 3), which significantly reduces the chance of condensation reactions of olefins or fused aromatic hydrocarbons at their active sites, resulting in a corresponding reduction in the probability of carbon deposition.

gasoline and diesel has been maintained at around of 32%, 65% respectively. Meanwhile, the values of the S, N content in the liquid products did not show a distinct decrease. The results showed that all of catalysts used in this study have a good stability activity as well as a long life-span. In addition, mass balance was calculated over reactor during the 1 day run at the reference conditions. The results were exhibited in Table 10, it can be observed that several gas species (i.e., CO, CO2, H2S, NH3, C1−C4 hydrocarbons, etc.) occurred and a large proportion of input hydrogen was consumed by reacting with coal tar. Depending on the calculations results, there is an acceptable error of about 1.32% between the total feedstock and the effluents.

4. CONCLUSIONS In this work, a LTCT distillate was hydrotreated in a four-stage fixed bed to produce clean liquid fuels at pilot-plant scale. The impacts of process variables on catalytic activity were investigated via independently varied reaction pressure, temperature, and LHSV. The results have shown that impurities (such as S, N, and other metal impurity) in LTCT distillate could be promisingly removed as the increases in the reaction pressure and temperature as well as decreases in the LHSV. Meanwhile, 30.80% of gasoline and 66.70% of diesel were obtained at optimized operating conditions. Phenols, alkanes, and AN were three most abundant compounds in the feedstock, and after the hydrogenation process, phenols and AN were converted to CA and decalins, tetralin, and indenes,

Table 10. Mass Balance over Reactor during the Run (per Day)a feedstock (g)

a

effluents (g)

coal tar

H2

oil

water

coke

H2

C1−C4

others

2589.57 total in

155.88 2745.45 (g)

2432.4 total out

201.4 2709.13 (g)

4.4 error

58.76 1.32%

9.56

2.61

Reaction conditions: LHSV = 0.3 h−1, PH2 = 13 MPa, THDS = 330 °C, THDN = 350 °C, THC = 370 °C, and H2:oil volume ratio = 1600:1. 11505


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Figure 11. TPO/TG-DTG plots of the spent hydrotreatment catalysts used in pilot-scale study.

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ACKNOWLEDGMENTS The financial support of this work was provided by the National Natural Science Foundation of China (21646009), Shaanxi Province Science and Technology Co-ordination Innovation Project Planned Program (2014KTCL01-09), Shaanxi Province Department of Education Industrialization Training Project (14JF026; 15JF031), and Young Science and Technology Star Project of Shaanxi Province (2016KJXX-32).

respectively. The chemical composition and properties of the gasoline and diesel distillates show that high quality clean fuel oil can be obtained through the four-stage hydrotreatment in a pilot-scale test. All of the catalysts used in this study have good stability activity as well as long life-span during a runtime of 400 h hydrogenation experiment. In addition, coke deposits are mainly concentrated in the second and third stages of the reactor.

■

■

ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.energyfuels.7b01971. Figures S1−S5 and Tables S1−S3 as mentioned in the text (PDF)

■

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]; Telephones: +86-18681859699 (Mobile), +86-029-88373425 (Home). ORCID

Dong Li: 0000-0002-4578-0595 Notes

The authors declare no competing financial interest. 11506

NOMENCLATURE LTCT = low temperature coal tar GC/MS = gas chromatograph/mass spectrometer LHSV = liquid hourly space velocity NIST = Institute of Standards and Technology AB = alkyl-benzenes PO = phenyl olefins AN = alkyl-naphthalenes OC = other oxygen containings HDO = hydrodeoxygenation HDS = hydrodesulfurization Hdn = hydrodenitrogeneration HDM = hydrodemetalation DDO = direct deoxygenation ROP = ring-opening HYD = hydrogenation DOI: 10.1021/acs.energyfuels.7b01971 Energy Fuels 2017, 31, 11495−11508

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ISOM = isomerization HC = hydrocracking HDA = hydrodearomatization RON = research octane number BTX = benzene, toluene, xylene

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