Hydrotreating the Low-Boiling-Point Fraction of Biocrude in Hydrogen

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Cite This: Ind. Eng. Chem. Res. 2019, 58, 10210−10223

Hydrotreating the Low-Boiling-Point Fraction of Biocrude in Hydrogen Donor Solvents for Production of Trace-Sulfur Liquid Fuel Zhi-Cong Wang,† Pei-Gao Duan,*,†,‡ Xiao-Jie Liu,† Feng Wang,‡ and Yu-Ping Xu‡ †

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Shaanxi Key Laboratory of Energy Chemical Process Intensification, School of Chemical Engineering and Technology, Xi’an Jiaotong University, No. 28, West Xianning Road, Xi’an, Shaanxi 710049, P.R. China ‡ College of Chemistry and Chemical Engineering, Department of Energy and Chemical Engineering, Henan Polytechnic University, No. 2001, Century Avenue, Jiaozuo, Henan 454003, P.R. China S Supporting Information *

ABSTRACT: The low-boiling-point distillate (LBD) via vacuum distillation of biocrude produced from hydrothermal liquefaction of soybean straw was hydrotreated to produce trace-sulfur liquid fuel. The effects of five hydrogen donor solvents (HDSs), including cyclohexene, cyclohexane, decahydronaphthalene, tetrahydronaphthalene, and Indane, on heteroatom removal efficiency were examined at 350 °C for 2 h with 6 MPa H2 added and 5 wt %(Pt/C)/feed. The LBD to HDS mass ratio was 1:1. HDSs not only could reduce the production of solid and gas and increase the yield of upgraded oil but also could favor denitrogenation, desulfurization, and deoxygenation of the upgraded oil. Among the HDSs examined, decahydronaphthalene showed the best performance for denitrogenation, and tetrahydronaphthalene was the most suitable HDS for deoxygenation and desulfurization. By employing a decahydronaphthalene and tetrahydronaphthalene mixture (w/w, 1:1) as the reaction medium, the effects of temperature (300−450 °C), time (1−6 h), H2 pressure (0.1−10 MPa), and Pt/C loading (0−20 wt %) on the product yields and quality of the upgraded oil produced from hydrotreating the LBD were examined. The upgraded oil was the dominant product fraction under all tested reaction conditions and varied between 76.7 and 87.3 wt %. The HDS mainly acted as a hydrogen transfer agent in the LBD hydrotreatment process, during which the HDS provided the hydrogen for the hydrogenation reaction, and this consumed hydrogen was resaturated by the external hydrogen source. A more positive synergistic effect was observed for the removal of N, O, and S when using the decahydronaphthalene and tetrahydronaphthalene mixture than when using decahydronaphthalene or tetrahydronaphthalene alone. N was the most difficult heteroatom to remove, followed by O and S. Catalyst loading was the most influential factor affecting the N, O, and S removal efficiencies. Under optimal reaction conditions, 93% of N, 95% of O, and 99% of S in the LBD and HDS blend were removed, which corresponded to contents of 0.05 wt %, 0.42 wt %, and 21 ppm in the upgraded oil, respectively. The equilibrium restrictions on denitrogenation, deoxygenation, and desulfurization were essential factors affecting the removal efficiencies of heteroatoms.

1. INTRODUCTION According to statistics, China produces nearly 900 million tons of crop straw every year, and approximately 200 million tons are not properly used.1 There are many ways of crop straw utilization, among which open-air burning is one of the most popular approaches for the straw to return to farmland, commonly adopted by many farmers in the underdeveloped rural areas in China.2,3 Therefore, the Chinese government promulgated the “Atmospheric Pollution Prevention and Control Law of the People’s Republic of China” in 2015,4 and this law prohibited the open-air burning of crop straw. Crop straw is composed of a large amount of organic matter, has small amounts of inorganic matter and water, and is rich in lignocellulose, but crop straw also contains a small amount of crude protein and crude fat. Therefore, if crop straw is collected and applied well, it is also a valuable renewable and © 2019 American Chemical Society

clean energy that brings many benefits to food safety, energy conservation, and environmental hygiene.5 There are many routes of crop straw energy utilization, among which thermochemical conversion (e.g., pyrolysis, hydrothermal liquefaction) is one of the most popular techniques.6 Pyrolysis usually requires a dry feedstock and thus suffers a substantial energy penalty from vaporizing the moisture content of the feedstock. Since many crop straws contain very high moisture when harvested, hydrothermal liquefaction of such biomass feedstocks is attractive from an energy perspective.7 Hydrothermal liquefaction converts crop Received: Revised: Accepted: Published: 10210

March 6, 2019 May 19, 2019 May 28, 2019 May 28, 2019 DOI: 10.1021/acs.iecr.9b01274 Ind. Eng. Chem. Res. 2019, 58, 10210−10223

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Industrial & Engineering Chemistry Research straw into an oily or tarry fluid (usually named biocrude or biooil) via reactions in and with liquid water at elevated temperatures (200−350 °C). This biocrude derived from the hydrothermal liquefaction of crop straw contains twice the energy density of the crop straw themselves.8 However, this kind of biocrude is a complex mixture consisting of hundreds of oxygen-containing organic compounds such as aldehydes, alcohols, carboxylic acids, esters, ethers, furans, ketones, and phenols.9 Therefore, these biocrudes are highly viscous and acidic at ambient temperatures and lack long-term preservation stability. Furthermore, these biocrudes (only when biomass contains N and S) also contain significant quantities of undesirable N and S, which would produce NOx and SOx when burned. Therefore, pretreatment to remove N, O, and S from these biocrudes and to reduce their viscosities are necessary before their use as fuels for transportation without any current engine modifications. Various technologies have been developed for the treatment of biocrudes produced from thermochemical conversion of biomass, 10 among which hydrotreatment is the most commonly studied technique. During the hydrotreatment process, N, O, and S are removed from the biocrude via reactions such as hydrodenitrogenation, hydrodeoxygenation, and hydrodesulfurization. Hydrogenation (saturation) of olefins and aromatics in the biocrude also occurs during this process and increases the saturation of the product oil. To date, hydrotreating entire biocrude derived from different biomass feedstocks is the most commonly examined methods employed by many researchers.6,11−15 To increase the removal efficiency of heteroatoms from the biocrude, catalysts (e.g., Ru/C, Ru/TiO2, Ru/Al2O3, Pt/C, Pd/C, NiMo/Al2O3, and CoMo/Al2O3) are usually introduced in the hydrotreating process. However, many drawbacks, such as high coke deposition, difficult removal of heteroatoms, and fast catalyst deactivation, also accompany the hydrotreating process if the biocrude alone was treated. To overcome these disadvantages, reaction media (e.g., H2O,16 ethanol,17 cyclohexane,18 and tetralin19) are typically employed in the hydrotreating process, and most of these reaction media are hydrogen donor solvents (HDSs). An HDS can reduce mass transfer limitations, effectively restrain the formation of coke precursors, and extract coke precursors from the catalyst pores in situ,14 thus prolonging the lifetime of the catalyst. Moreover, another advantage of using HDSs over hydrogen molecules is that the C−H bond (414 kJ/mol) has lower bond energy than the H− H bond (436 kJ/mol), so the ability of HDS to provide hydrogen ions is higher,20 making the hydrogen ion transfer easier. Moreover, an HDS is also more easily handled than hydrogen molecules. It should be noted that the cost of using HDS might be higher or lower than that using H2 depending on the price of the hydrogen donor itself. Petrochemical plants usually separate crude oil into several simple distillates by atmospheric distillation and then hydrotreat the distillates to usable fuels or chemicals. Since the biocrude derived from crop straw is a very complex mixture consisting of hundreds of compounds with different boilingpoint ranges, preseparation of components by distillation into chemically similar compound classes has also been adopted by many investigators to improve the properties of the biocrude before its subsequent treatment.21−24 This physical separation allows for downstream treatment to occur under conditions appropriate for each fractional cut. However, the high temperatures required for atmospheric distillation produce

additional unwanted products and promote carbonization of the bio-oil.13 In contrast, vacuum distillation requires a lower temperature, which can effectively avoid the thermal degradation of heavy distillates.21,22 In the present study, vacuum distillation (vacuum degree of 0.09 MPa, 25−140 °C, 141−220 °C) was used to separate the biocrude produced from hydrothermal liquefaction of soybean straw into two distillate fractions: low-boiling-point distillate (LBD, 25−140 °C) and high-boiling-point distillate (HBD, 141−220 °C). The LBD was used as the feedstock for the subsequent treatment process. Pt/C (Pt, 5 wt %) was selected as the heterogeneous catalyst because platinum group catalysts performed well in a hydrotreating study of bio-oil.11 Highpressure external H2 was also employed to saturate the dehydrogenation product of the HDSs in situ, ensuring that the whole reaction system is in a hydrogen-rich state. Five HDSs, including cyclohexene, cyclohexane, decahydronaphthalene, tetrahydronaphthalene, and Indane, were screened in terms of N, O, and S efficiencies. Then, using the identified HDSs, the effects of temperature (300−450 °C), time (1−6 h), hydrogen pressure (0.1−10 MPa), and Pt/C loading (0− 20 wt %) on the product distribution and quality of the upgraded oil produced from hydrotreatment of the LBD in the HDS were examined. Finally, the upgraded oil was characterized by gas chromatography−mass spectrometry, thermogravimetric analysis, and elemental analysis.

2. EXPERIMENTAL METHODS Materials. Soybean straw was obtained from Jiashan town, Hunan Province, China, and was dried at 105 °C for 12 h and smashed by a pulverizer to 100 mesh. The Pt/C (Pt, 5 wt %) catalyst was purchased from Zhengzhou Alpha Chemical Co., Ltd. and used as received, and the catalyst has a BET surface area of 419 m2/g and a metal dispersion of 5.4%. Cyclohexene, cyclohexane, decahydronaphthalene, tetrahydronaphthalene, Indane, and dichloromethane with purities of ≥99% were purchased from Aladdin Industrial Corporation (Shanghai, China) and used as received. Deionized water was made in the laboratory. Two stainless steel batch reactors (Zhengxin Instrument Factory, Yancheng, Jiangsu Province, China) with total internal volumes of 1000 and 20 mL were used to perform the hydrothermal liquefaction and hydrotreating reactions, respectively. Figure 1(a) and (b) show a schematic of the reactor for hydrothermal liquefaction of soybeans and a schematic of the apparatus and minibatch stainless steel reactor for hydrotreatment of LBD, respectively. Prior to use, these two reactors were treated with supercritical water at 400 °C for 2 h to remove any internal residues that remained from the manufacture of the reactor parts. The 1000 mL reactor, which was equipped with a mixer and mixer controller, a safety relief valve, and a pressure gauge, was heated by an electric heater with a power of 2.5 kW. A custom-built molten-salt bath consisting of KNO3 and NaNO3 at a mass ratio of 5:4 was used to heat the 20 mL reactor. No mixer was installed for the 20 mL reactor due to limited internal space. The rated working temperature and pressure were 400 °C and 25 MPa for the 1000 mL reactor and 450 °C and 45 MPa for the 20 mL reactor, respectively. Hydrothermal Liquefaction. Here, 150 g of dry soybean straw powder and 400 mL of deionized water were loaded into the 1000 mL reactor and sealed tightly. The residual air in the reactor headspace was replaced with He by flushing the reactor 10211

DOI: 10.1021/acs.iecr.9b01274 Ind. Eng. Chem. Res. 2019, 58, 10210−10223

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Industrial & Engineering Chemistry Research

reactor reached 320 °C. It should be noted that the slow heating rate would result in a biocrude with low quality. So the biocrude used by the present study is not a generic lignocellulosic biocrude. At this point, the reaction time was set to zero. The selection of hydrothermal liquefaction conditions was based on a study by Xu et al.19 The reaction time was 60 min. The pressure inside the reactor was maintained at approximately 12 MPa during the reaction. Once the reaction was over, the reactor was placed into a cold-water bath to quench the reaction. After the reactor was cooled to room temperature, the reactor was removed from the cold-water bath. The pressure inside the reactor was 2 MPa due to the formation of gaseous products during the hydrothermal liquefaction reaction. The gas in the reactor was exhausted prior to opening the reactor. All the materials in the reactor were pureed into a beaker and subjected to filtration by a Buchner funnel. The remaining solid was dissolved in dichloromethane. The reactor was washed with dichloromethane until the extracted solution was light yellow. All the dichloromethane extract was collected together and separated by filtration. The dichloromethane in the extract was vaporized by using a rotary evaporator at 35 °C and a vacuum of 0.09 MPa. The remaining material was the biocrude. Multiple sets of experiments were performed under the same reaction conditions as mentioned above to prepare enough biocrude for subsequent tests. Vacuum Distillation. A vacuum distillation system, which included a 500 mL one-necked flask, a Claisen adapter, an adjustable electrical heater, a thermometer (for monitoring internal reaction temperature), a Graham condenser, a receiver flask, and a vacuum pump, was used to realize the separation of the biocrude and was assembled as shown in Figure 2. The Claisen adapter was fitted with a thermometer. The Graham condenser was attached to the curved arm of the Claisen adapter. The biocrude container was heated by a hair drier to decrease the biocrude viscosity prior to being loaded into the distillation flask. Approximately 100 g of biocrude was fed into the distillation flask by a long-necked funnel. The vacuum pump power was first switched on after each part of the distillation system was tightly sealed. Then, the electrical heater was switched on once the vacuum in the distillation system

Figure 1. (a) Schematic of autoclave reactor for hydrothermal liquefaction of soybean. (b) Schematic of apparatus and minibatch stainless steel reactor for hydrotreatment of LBD.

with He for at least 15 min. The reaction was started by switching on the power. The speed of the mixer was set to 400 rpm. Approximately 80 min later, the temperature inside the

Figure 2. Distillation unit of biocrude. 10212


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Industrial & Engineering Chemistry Research reached 0.09 MPa. The temperature of the distillation flask was adjusted by changing the power of the electrical heater. Tap water was used to condense the effluent material. The reactor was heated from 25 to 140 °C under a vacuum condition of 0.09 MPa and remained at 140 °C until no additional material was eluted. The distillation fraction between 25 and 140 °C was regarded as the LBD. After all the materials eluted at 25− 140 °C were collected, the distillation flask was again heated from 140 °C until all the remaining materials were collected. No more oil fraction was eluted as the temperature increased above 220 °C. The distillation fraction between 141 and 220 °C was defined as the HBD. The solid residue left in the flask after distillation was defined as distillate residue. The yield of each distillation fraction was calculated as its mass divided by the biocrude mass loaded into the distillation flask. Multiple sets of vacuum distillation of the biocrude were performed under the same operating conditions as mentioned above. Hydrotreating the LBD. In a typical run, 3 g of LBD (6 g of LBD was loaded when using the LBD alone), 3 g of HDS, and a certain amount of Pt/C catalyst were loaded into the 20 mL reactor and tightly sealed. This LBD to HDS ratio of 1:1 was selected as the most notable positive synergistic effect between LBD and HDS observed.14 The residual air in the reactor was replaced by H2 and further charged to the desired pressure. The preheated temperature of the molten-salt bath was usually 50 °C higher than the desired reaction temperature, and the temperature was set to the reaction temperature after the loaded reactor was placed into the molten-salt bath. The reaction time was set to zero once the temperature inside the reactor reached the desired temperature. The preheating time was 18−45 min depending on the reaction conditions. A K-type thermocouple was inserted into the thermowell on the top cover of the reactor to monitor the temperature, which was controlled by using a temperature controller. The pressure inside the reactor varied between 9 and 21 MPa depending on the reaction conditions. After the reaction time was exhausted, the reactor was removed from the molten-salt bath and submerged into a cold-water bath. Approximately 30 min later, the reactor was cooled to room temperature, and the water above the reactor was removed by a hair drier. The gas production was estimated by the mass difference before and after venting of the reactor. The reactor was opened after the gas was completely vented. All the materials were carefully transferred into a 50 mL filtering centrifuge tube and submitted to centrifugation at a speed of 10,000 r/min for 10 min. The solid residue stayed in the centrifuge filter, and the liquid at the bottom of the centrifuge tube was the upgraded oil. The solid residue was extracted with dichloromethane to remove the residual upgraded oil to accurately estimate the production of solid residue and dried at 105 °C for 12 h. The dichloromethane extract was vaporized by using a rotary evaporator at 35 °C and a vacuum of 0.09 MPa, and the remaining material was the upgraded oil, which was no more than 10% of the upgraded oil obtained by centrifugation. This upgraded oil was mixed together with the upgraded oil separated by centrifugation. The yield of each product fraction was estimated as its mass divided by the total mass of LBD and HDS loaded into the reactor. Control experiments with the HDS alone were also performed under the same reaction conditions as those with the added LBD and HDS to estimate the recovery yield. Duplicate runs were performed under the same conditions to estimate the uncertainties in the experimental results, and all

the results provided in this paper are the average values of these independent trials. Analysis. The total acid value was tested by a titration method. The titration solution was 0.1 mol/L KOH solution. A mixed solvent consisting of isopropyl alcohol and toluene at a volume ratio of 1:1 was used as the solvent. Phenolphthalein, which had been dissolved in methanol (0.5%, wt./v), was used as the titration terminal indicator. First, 0.2 g of oil and 10 mL of isopropyl alcohol and toluene mixed solvent were added to a 50 mL conical flask and shaken violently to ensure that the oil was completely dissolved in the solvent. Then, a drop of phenolphthalein solution was added to the conical flask. The titrant was loaded into a buret to titrate the oil. The titration end point was reached once the solution became a pink color and did not fade for half a minute. The consumed volume of KOH was recorded and used to quantify the TAN. The moisture content of the oil was detected by a volumetric Karl Fischer titration method. A KF-1B moisture analyzer (Shandong Benchuang Instrument Co., Ltd., Zibo, Shangdong Province, China) was employed to perform the analysis. The one-component titrating reagent (KFR-04) consisted of iodine, and sulfur dioxide dissolved in methanol was used. The Karl Fischer equipment was first calibrated using pure water prior to use. Approximately 0.1 g of oil was dissolved in 40 mL of anhydrous methanol and was titrated by the one-component titrating reagent via a titration buret. The water was quantified based on the consumed volume of the Karl Fischer reagent. The N and S contents in the oil were tested via a TN-3000 chemiluminescence nitrogen analyzer and TS-3000 ultraviolet fluorescence sulfur analyzer (Jiangsu Guochuang Analysis Instrument Co., Ltd., Taizhou, Jiangsu Province, China), respectively. The C, H, and O contents of the oils were determined using a Thermo Fisher Scientific FLASH 2000 organic elemental analyzer (Thermo Fisher Scientific, Waltham, MA, USA). The higher heating values (HHVs) of the oils were calculated by the Dulong formula14 as follows: HHV (MJ/kg) = 0.338 × C + 1.428(H−O/8) + 0.095S

where C, H, O, and S represent their mass percentage contents in the oil. A comprehensive two-dimensional gas chromatographytime-of-flight mass spectrometer (LECO4D Pegasus) (Leco, San Jose, MI, USA) was used to quantify the molecular composition of the oil, and this instrument could be used either for one-dimensional or two-dimensional analysis. Two columns were installed for this analysis system. The first column was a nonpolar Rtx-5SilMS with 30 m × 0.25 mm ID × 0.25 μm film thickness installed in the main oven, and this column separated the compounds based on their boiling points. The second column was a Rxi-17 MS with a 1.10 m × 0.10 mm ID × 0.10 μm film thickness installed in the second oven, and this column separated the compounds based on their polarity. The oil samples were prepared by redissolution in dichloromethane at a concentration of 10 (wt/vol)%. The sample injection volume was 1 μL at a split ratio of 1:5. The inlet temperature was set to 300 °C. The first-dimension column was initially held at 40 °C for 2 min. The first stage of the heating process was ramped to 250 °C at 2 °C/min and held isothermally for 2 min, and then, the temperature was ramped to 300 °C at 2 °C/min and held isothermally for 2 min. The same temperature program was set for the second column. The second-dimension column was initially held at 45 °C for 2 min. The first stage of the heating process was ramped 10213


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Industrial & Engineering Chemistry Research to 255 °C at 2 °C/min and held isothermally for 2 min, and then, the temperature was ramped to 305 °C at 2 °C/min and held isothermally for 2 min. The transfer line temperature was 280 °C. The ion source temperature was 250 °C with a filament bias voltage of 70 eV. The data acquisition rate was 100 spectra/s for the mass range from 35 to 500 amu. The detector voltage was 1500 V. LECO ChromaTOF software recorded and analyzed the data with spectral identification provided by the NIST library. The accuracy and sensitivity of the MS detector were calibrated and tested by using perfluorotributylamine and hexachlorobenzene, respectively. The fuel properties of the upgraded oil were analyzed by a MINISCAN IR VISION mid-infrared fuel analyzer (Grabner Instruments Messtechnik GmbH of Vienna, Austria), which incorporates the advantages of mid-IR and near-IR spectroscopy. The analyzer is supplied fully configured and ready for fuel testing. The spectral scanning range was 7000−450 cm−1, and the wavenumber resolution was less than 0.5 cm−1. Prior to analysis, the sample tank was cleaned with a solvent mixture consisting of acetone, methanol, and toluene at a volume ratio of 1:1:1 and then emptied. In the gasoline mode, a 10 mL oil sample was manually injected into the injection port, and the start button was pressed to start the test. The total analysis time was approximately 5 min.

the biocrude as the high-boiling-point fractions of the biocrude remained in the DR. The LBD and HBD showed some differences in elemental compositions from the original biocrude. The LBD had a higher H content and lower C content than those of the biocrude. Both the LBD and HBD contained less N than the biocrude, as some N remained in the DR. The moisture of both LBD and HBD increased, which resulted in an increase in H and O contents of these two distillates. The LBD had a higher S content (2817 ppm) than the biocrude (2115 ppm), and contrary results were observed for the HBD, suggesting most of the S-containing compounds existed in the LBD. The LBD had similar N and S contents as the LBD produced from the distillation of biocrude produced from hydrothermal liquefaction of peanut straw.11 Both the LBD and HBD had a higher H/C molar ratio and HHV and total acid number (TAN) than those of the biocrude. Different properties of the LBD and HBD from the biocrude would lead to different hydrotreating characteristics of these two different feedstock materials. Control Experiment. Prior to presenting results from hydrotreating the LBD in the HDS, we first report the results from control experiments performed with several HDSs alone at 350 °C for 2 h with 6 MPa H2 added and in the presence or absence of 5 wt %(Pt/C)/LBD. The goal of these experiments was to determine how much HDS could be recovered under the same reaction conditions as those of hydrotreating the LBD and HDS blend as well as the stability of the HDS. It should be noted that the recovery in hydrotreating the HDS blend may be completely different from the hydrotreating the LBD and HDS blend. Nevertheless, these experiments prove at least the reactivity of each HDS and its recovery efficiency after the reaction. Figure 3 shows the results. Clearly, the liquid product was always the predominant fraction for both the catalytic and noncatalytic reactions, and this liquid fraction varied between 70.7 and 86.5 wt % for the catalytic runs and 85.2 and 97.0 wt % for the noncatalytic runs. Higher liquid yield recovery would decrease the cost of this cohydrotreating process. Higher liquid yields were always observed for tetrahydronaphthalene, decahydronaphthalene, and Indane due to their comparatively higher boiling points than those of cyclohexene and cyclohexane. For the same HDS, the liquid product yield produced from the noncatalytic run was 11−15 percentage points larger than that produced with Pt/C due to the high proportion of light fractions in the liquid product produced with Pt/C. These light fractions are easily lost during their handling process. Furthermore, since the liquid products were separated by centrifugation, residual liquid products in the catalyst was inevitable. Mater loss was also inevitable during the recovery process of the liquid products due to the volatility of the HDS. Moreover, some liquid products adhered to the reactor wall, which also decreased the yield of liquid products. No solid was observed for the noncatalytic reactions as all the HDSs were not activated in the absence of Pt/C. For the catalytic reactions, only decahydronaphthalene, tetrahydronaphthalene, and Indane produced 0.35−0.48 wt % solid. All HDSs except cyclohexene produced gaseous products (except the unreacted H2) whether a catalytic reaction or a noncatalytic reaction occurred. A comprehensive two-dimensional gas chromatographytime-of-flight mass spectrometer was used to quantify the molecular compositions of the liquid products produced from hydrotreating the HDS with and without Pt/C. Figures S1−S5 (see S1−S5 in the Supporting Information) compare the total

3. RESULTS AND DISCUSSION Hydrothermal Liquefaction and Vacuum Distillation. Approximately 30 g of biocrude was generated by liquefying 150 g of soybean straw at 320 °C for 60 min, and this amount corresponded to a biocrude yield of 20 wt % which is close to the bio-oil yield obtained from hydrothermal liquefaction of macroalgae and duckweed under similar reaction conditions.25,26 The soybean straw predominantly consisted of cellulose (42.4 wt %), hemicellulose (22.1 wt %), and lignin (18.9 wt %), which were decomposed to form the biocrude intermediates under hydrothermal conditions. At the same time, significant repolymerization of the biocrude intermediates also occurred during the hydrothermal liquefaction process. Therefore, a rather low biocrude yield was observed. The biocrude was a tar-like material that had poor flowability at room temperature because of its high content of highmolecular-weight compounds. This biocrude was also rich in N, O, and S, making it difficult to use in practice. The biocrude was separated into three fractions via vacuum distillation under a vacuum of 0.09 MPa, and these three fractions were named as LBD (25−140 °C), HBD (141−220 °C), and distillate residue (DR) (>220 °C). The LBD, HBD, and DR yields were 21.2 ± 3.0, 34.3 ± 3.0, and 29.4 ± 3.0 wt %, respectively, which are close to the results reported by Capunitan et al.27 Approximately 13 wt % mass of the biocrude was lost during the vacuum distillation process, and most of these lost materials were residual dichloromethane and water. The dichloromethane extract was evaporated under a vacuum condition to remove dichloromethane and residual water. At the end of the vaporization process, most of the dichloromethane and residual water were removed; however, a small amount of dichloromethane and water still existed in the biocrude. These residual dichloromethane and water would be separated during the vacuum distillation process. Therefore, the mass balance of the vacuum distillation process could not reach 100%. Table 1 shows the elemental compositions and other properties of the biocrude, LBD, HBD, and DR. Both the LBD and HBD flowed easily at room temperature relative to 10214


1.01 4.81 1.44 − 2.46

81.32 79.91 82.52 82.91 84.92 84.93 83.87

± ± ± ± ± ± ±

± ± ± ± ±

C

2.44 2.40 2.48 2.49 2.55 2.55 2.52

2.25 2.21 2.28 2.46 2.47

8.01 8.97 9.50 5.68 9.14

± ± ± ± ±

H

10215

Catalyst reuse (400 °C, 4 h, 10 MPa H2, 5 wt %(Pt/C)/feed Once Twice Thrice

0.09 0.08 0.07 0.05 0.05

0.08 0.06 0.05 0.05

0.11 0.08 0.06 0.05

0.19 0.17 0.14 0.09 0.08 0.08 0.08

0.32 0.44 0.35 0.17 0.22

(reuse), LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1) 0.14 ± 0.01 4.8 ± 0.14 86.07 ± 2.58 10.61 ± 0.32 1.51 ± 0.05 0.16 ± 0.01 6.3 ± 0.19 86.28 ± 2.59 10.84 ± 0.33 1.73 ± 0.05 0.14 ± 0.01 8.2 ± 0.25 86.58 ± 2.60 10.76 ± 0.32 1.82 ± 0.05

Catalyst loading/wt % (400 °C, 4 h, 10 MPa H2, LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1) 0 0.18 ± 0.01 9.4 ± 0.28 86.19 ± 2.59 10.40 ± 0.31 5 0.14 ± 0.01 3.8 ± 0.11 87.77 ± 2.63 10.61 ± 0.31 20 0.08 ± 0.01 3.2 ± 0.09 88.36 ± 2.65 9.48 ± 0.28

± ± ± ± ±

± ± ± ±

± ± ± ±

± ± ± ± ± ± ±

± ± ± ± ±

O

1.71 ± 0.05 1.51 ± 0.05 0.74 ± 0.02

2.94 2.73 2.44 1.65 1.51

0.28 0.29 0.30 0.31 0.32

H2 Pressure/MPa (400 °C, 4 h, 5 wt %(Pt/C)/feed, LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1) 0.1 Ar 0.20 ± 0.01 12.4 ± 0.37 86.16 ± 2.58 9.48 0.1 0.39 ± 0.01 9.2 ± 0.28 86.30 ± 2.59 9.61 3 0.25 ± 0.01 7.9 ± 0.24 86.52 ± 2.60 10.07 6 0.19 ± 0.01 4.6 ± 0.14 86.33 ± 2.62 10.32 10 0.14 ± 0.01 3.8 ± 0.11 87.77 ± 2.63 10.61

± ± ± ± ±

2.81 2.14 1.65 1.62

Time/h (400 °C, 6 MPa H2, 5 wt %(Pt/C)/feed, LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1) 1 0.17 ± 0.01 17.2 ± 0.52 85.86 ± 2.58 10.64 ± 0.32 2 0.21 ± 0.01 15.9 ± 0.48 84.51 ± 2.54 10.56 ± 0.32 4 0.19 ± 0.01 4.6 ± 0.14 86.33 ± 2.59 10.32 ± 0.31 6 0.15 ± 0.01 3.8 ± 0.11 85.68 ± 2.57 10.30 ± 0.31

6.34 5.62 4.60 3.01 2.73 2.81 2.64

10.83 14.55 11.52 5.99 7.37

3.80 2.64 2.14 1.62

0.30 0.34 0.33 0.36 0.29 0.28 0.32

0.24 0.27 0.29 0.17 0.27

0.33 0.32 0.32 0.28

± ± ± ±

1.86 0.78 1.10 0.56 0.52 0.35 0.55

75.05 73.58 76.00 82.46 82.23

Temperature/°C (2 h, 6 MPa H2, 5 wt %(Pt/C)/feed, LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1) 300 0.39 ± 0.01 36.3 ± 1.09 82.34 ± 2.47 10.97 350 0.23 ± 0.01 18.4 ± 0.55 83.87 ± 2.52 10.72 400 0.21 ± 0.01 15.9 ± 0.48 84.51 ± 2.54 10.56 450 0.23 ± 0.01 12.4 ± 0.37 87.03 ± 2.61 9.35

± ± ± ± ± ± ±

± 1.35

± 2.85 ± 2.73 ± 2.47

± ± ± ± ± ± ±

61.9 26.1 36.7 18.5 17.3 11.7 18.4

94.9 90.9 82.2 − 45.1

TAN (mgKOH/g)

10.03 11.48 11.14 12.07 9.61 9.42 10.72

0.04 0.01 0.01 0.01 0.01 0.01 0.01

± 0.07

± 0.03 ± 0.14 ± 0.04

Moisture (wt %)

HDS (350 °C, 2 h, 6 MPa H2, 5 wt %(Pt/C)/feed, LBD:HDS=1:1) LBD 1.49 ± LBD+cyclohexane 0.37 ± LBD+cyclohexene 0.35 ± LBD+decahydronaphthalene 0.20 ± LBD+tetrahydronaphthalene 0.23 ± LBD+Indane 0.26 ± LBD+(decahydronaphthalene + tetrahydronaphthalene) 0.23 ±

Biocrude LBD HBD DR LBD+HDS (DHN:THN= 1:1)

± ± ± ± ±

± ± ± ±

± ± ± ±

± ± ± ± ± ± ±

± ± ± ± ±

N

0.02 0.02 0.02 0.02 0.01

0.02 0.02 0.02 0.02

0.03 0.02 0.02 0.02

0.07 0.04 0.04 0.02 0.02 0.03 0.02

0.09 0.08 0.08 0.08 0.04

0.35 ± 0.01 0.57 ± 0.02 0.61 ± 0.02

0.57 ± 0.02 0.35 ± 0.01 0.07 ± 0.01

0.83 0.81 0.63 0.56 0.35

0.74 0.63 0.56 0.53

0.86 0.69 0.63 0.56

2.34 1.22 1.16 0.73 0.77 1.08 0.69

2.84 2.59 2.66 2.77 1.28

± ± ± ± ±

± ± ± ±

± ± ± ±

± ± ± ± ± ± ±

2 2 1 1 1

1 1 1 1

2 1 1 1

4 2 3 2 1 2 1

±7

± 10 ± 14 ±7

44 ± 1 54 ± 1 65 ± 2

83 ± 2 44 ± 1 21 ± 1

165 152 83 60 44

70 68 60 41

329 189 68 42

716 269 368 237 212 243 189

2115 2817 1401 0.07 wt % 1407

S (ppm)

Table 1. Elemental Analysis and Other Properties of Biocrude, LBD, HBD, and Upgraded Oils Produced under Different Reaction Conditions

± ± ± ± ±

± ± ± ±

± ± ± ±

± ± ± ± ± ± ±

0.04 0.04 0.04 0.04 0.04

0.04 0.05 0.04 0.04

0.05 0.05 0.04 0.04

0.04 0.05 0.05 0.05 0.04 0.03 0.04

0.04 0.04 0.04 0.02 0.04

1.48 ± 0.04 1.51 ± 0.05 1.49 ± 0.04

1.45 ± 0.04 1.45 ± 0.04 1.29 ± 0.04

1.32 1.34 1.40 1.42 1.45

1.48 1.49 1.44 1.45

1.59 1.54 1.49 1.30

1.48 1.74 1.61 1.75 1.36 1.33 1.54

± ± ± ± ±

H/C 1.28 1.46 1.50 0.83 1.32

± ± ± ± ±

± ± ± ±

± ± ± ±

± ± ± ± ± ± ±

± ± ± ± ±

0.98 1.02 1.11 1.15 1.21

1.12 1.02 1.10 1.23

1.18 1.25 1.26 1.12

1.21 1.22 1.23 1.31 1.11 1.08 1.21

1.03 1.04 1.11 1.04 1.08

43.97 ± 1.05 44.33 ± 1.11 44.30 ± 1.11

43.67 ± 1.04 44.54 ± 1.12 43.27 ± 1.02

42.13 42.41 43.18 43.62 44.54

43.71 43.26 43.62 43.37

42.81 43.18 43.26 42.47

40.67 42.39 42.97 44.72 41.93 41.65 43.18

34.87 35.08 37.19 34.91 39.53

HHV (MJ/kg)

Industrial & Engineering Chemistry Research Article


Article


Figure 3. Recovery of HDS (cyclohexene (CHE), cyclohexane (CHA), decahydronaphthalene (DHN), tetrahydronaphthalene (THN), and Indane (IDA)) processed at 350 °C for 2 h with added 6 MPa and 5 wt % (Pt/C)/feed or without catalyst.

than that produced by all the LBD and HDS blends due to its high content of high-molecular-weight compounds. The gaseous products mainly consisted of unreacted H2 together with a small amount of CO2 and CH4. Table 1 shows the TAN, moisture, elemental composition, and other properties of the upgraded oils produced from the LBD alone and the LBD +HDS blend. The upgraded oils produced from the LBD and HDS blends had a lower moisture content than those produced from the LBD alone, suggesting the polarity of the upgraded oil produced from the LBD and HDS blend is smaller than that produced from the LBD alone. The upgraded oils produced from the LBD and HDS blends had a TAN of 11.7−36.7, which is almost half that produced from the LBD alone (61.9), suggesting the acidic materials in the LBD were more easily removed in the presence of HDS. Compared with the LBD and HDS blends, the upgraded oil produced either from the LBD alone or the LBD and HDS blends had a higher H content and lower N, O, and S contents, indicating denitrogenation, desulfurization, and deoxygenation occurred during the reaction. Furthermore, the upgraded oils produced from the LBD and HDS blends had lower N, O, and S contents than those produced from the LBD alone, indicating that the HDS favored the removal of N, O, and S, again implying that the integration of HDS and H2 is effective for the upgrading of LBD. Among all the HDSs examined, the upgraded oil produced from the LBD and tetrahydronaphthalene blend had the lowest O and S contents of 2.73 wt % and 212 ppm, respectively, while the upgraded oil produced from the LBD and decahydronaphthalene blend had the lowest N content of 0.73 wt %. Nitrogen is the most difficult removal heteroatom in the bio-oil.16 All the upgraded oils had a higher H/C molar ratio and HHV than the original LBD and HDS blends, implying that the hydrogenation reaction occurred during the reaction. The N and S contents are very important indicators of the oil quality; since the LBD and tetrahydronaphthalene blend produced an upgraded oil with the lowest O and S contents and the LBD and decahydronaphthalene blend produced an upgraded oil with the lowest N content, one would expect the combination of tetrahydronaphthalene and decahydronaphthalene to be promising for the removal of N, O, and S. Therefore, tetrahydronaphthalene and decahydronaphthalene at a mass ratio of 1:1 were selected as the reaction mediums for the hydrotreating of the LBD in the optimization of reaction parameters.

ion chromatograms of the HDSs and their corresponding liquid products processed with and without Pt/C. In the absence of Pt/C, all HDSs except cyclohexane were stable, as no other peaks than the HDS alone were observed. Cyclohexene was unstable at 350 °C and 6 MPa H2 and converted to cyclohexane, benzene, and 3,3-dimethyl-1-butyne. In the presence of Pt/C, cyclohexane and decahydronaphthalene were stable, whereas cyclohexene, tetrahydronaphthalene, and Indane were unstable under the reaction conditions due to their deficiencies in hydrogen. These unsaturated HDSs would convert to their saturated counterparts, which are desirable for the hydrogenation reaction, as the hydrogen in the HDS is more reactive than the external H2. Effect of Parameters on Product Yields and Properties of Upgraded Oils. The LBD and HDS blends were converted into upgraded oil, gas, and solid residue via processing with added Pt/C and H2 at 350 °C for 2 h. We first discuss the influence of different HDSs on the product yields and properties of the upgraded oils. Figure 4(a) shows the effects of different HDSs on the yields of the product fractions. Clearly, the upgraded oil was always the dominant fraction among the three product fractions in the presence or absence of the HDS. Compared with the upgraded oil produced by the LBD alone, the yield of upgraded oil produced from the LBD and HDS blends is always higher, indicating that the presence of HDS favored the production of upgraded oil by reducing the production of solid and gaseous products as indicated in Figure 4(a). A solid yield of 7.5 wt % was observed when processing the LBD alone, and this yield is almost double that of processing all the LBD and HDS blends. This result is desirable, as the HDS suppresses coke formation during this hydrogenation process, and this suppression effect had been demonstrated by many previous studies.28,29 The HDS could provide hydrogen free radicals during the reaction process, which could block the macromolecular polycyclic aromatic hydrocarbons (PAHs) free radicals formed in the thermal cracking process, thus inhibiting coke formation. Furthermore, the HDS could dilute the PAHs free radicals in the reaction system, reduce the chance of collision with each other, and inhibit the coke formation. All the HDSs except cyclohexene resulted in closed upgraded oil yields of ∼87 wt %. A slightly lower upgraded oil yield of 83.1 wt % was observed for cyclohexene, which might be due to the highest solid and gas production. A gas yield of 5.8 wt % was obtained when processing the LBD alone, and this yield was also higher 10216


Article


Figure 4. Products yields produced at different reaction conditions: (a) LBD:HDS=1:1, 5 wt % Pt/C, 350 °C, 6 MPa H2, 2 h. (b) LBD:HDS (DHN:THN=1:1)=1:1, 5 wt % Pt/C, 6 MPa H2, 2 h. (c) LBD:HDS (DHN:THN=1:1)=1:1, 5 wt % Pt/C, 400 °C,6 MPa H2. (d) LBD:HDS (DHN:THN=1:1)=1:1, 5 wt % Pt/C, 400 °C, 4 h. (e) LBD: HDS (DHN:THN=1:1)=1:1, 10 MPa H2, 400 °C, 4 h. (f) LBD:HDS (DHN:THN=1:1)=1:1, 5 wt % Pt/C (reuse), 10 MPa H2, 400 °C, 4 h; cyclohexene, CHE; cyclohexane, CHA; decahydronaphthalene, DHN; tetrahydronaphthalene, THN; and Indane, IDA.

observed at 450 °C due to the accelerated cracking reactions. A similar product yield trend as a function of temperature was also observed when hydrotreating the LBD of biocrude produced from HTL of peanut straw.14 It should be noted that lower upgraded oil yield might be desirable as the upgraded had lower N, O, and S contents. Table 1 provides the TAN, moisture, elemental composition, and other properties of the upgraded oils obtained at different temperatures. The moisture mainly came from the original moisture and hydrodeoxygenation of O-containing compounds in the LBD, which decreased from 0.39 to 0.23 wt % as the temperature increased from 300 to 350 °C, indicating the polarity of the upgraded oil decreased with increasing the temperature. Further increases in temperature had no effect on the moisture content of the upgraded oil. The TAN also decreased with increasing temperature due to the promotion of hydrodeoxygenation of the O-containing acid substance, and

The effect of temperature was first examined at 2 h, 6 MPa H2, 5 wt %(Pt/C)/feed, LBD:HDS (tetrahydronaphthalene and decahydronaphthalene=1:1)=1:1 with changing temperatures from 300 to 450 °C. Figure 4(b) presents the results. The upgraded oil was the dominant product at all temperatures examined and decreased from 87.3 to 76.7 wt % as the temperature increased from 300 to 450 °C. A comparatively lower upgraded oil yield was observed at 450 °C due to the high coke and gas production. Furthermore, the upgraded oil produced at severe temperatures contained a large amount of low-boiling-point fractions which are easily lost during the product recovery process. Denitrogenation, desulfurization, and deoxygenation also contributed to the decrease in upgraded oil yield. The solid yield slightly increased from 3.2 to 4.2 wt % as the temperature increased from 300 to 450 °C because polymerization reactions prevailed at severe temperatures, as did the gas yield. An 11.1 wt % gas yield was 10217


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Figure 5. Composition of biocrude, LBD, treated HDSs, and upgraded oils produced under different reaction conditions (% of Total Peak Area by GC-MS).

this finding was consistent with the variation trend of the O content of the upgraded oil. Increasing the temperature increased the C content and decreased the H content of the upgraded oil, indicating that the aromaticity of the upgraded oil increased, which was evidenced by the increased aromatics content in the upgraded oil as shown in Figure 5. All the N, O, and S contents in the upgraded oil decreased with increasing temperature, which were reduced in the proportions of 34.9%, 57.4%, and 87.2%, respectively, as the temperature increased from 300 to 450 °C, indicating that N was the most difficult atom to remove due to the average C−N bond energy being the largest, followed by O and S. This phenomenon was also observed for the upgrading of another bio-oil.30 At the highest temperature of 450 °C, the upgraded oil had the lowest N, O, and S contents of 1.62 wt %, 9.56 wt %, and 42 ppm, respectively. The increased C content and decreased H content of the upgraded oil resulted in a decreased H/C molar ratio with increasing temperature, again indicating that the content of aromatic compounds in the upgraded oil increased. In addition, better flowability of the upgraded oil was observed at higher temperatures due to the high proportion of low-boilingpoint fractions. Compared with the upgraded oil produced from either the LBD+decahydronaphthalene or LBD+tetrahydronaphthalene blends, the upgraded oil produced from the LBD+(decahydronaphthalene and tetrahydronaphthalene) had lower N, O, and S contents, indicating a positive synergistic effect between decahydronaphthalene and tetrahydronaphthalene during the removal of N, O, and S. Since the lowest N, O, and S contents were observed at 450 °C, the lowest upgraded oil yield was also obtained at that temperature. Overall consideration, in view of the yield and quality of the upgraded oil, 400 °C was selected as a suitable temperature. The effects of reaction time on the product yields and upgraded oil quality were examined at 400 °C, 6 MPa H2, 5 wt %(Pt/C)/feed, and LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1 with changing the reaction time from 1 to 6 h. Figure 4(c) shows the results. The upgraded oil yield decreased from 83.4 to 80.3 wt % as the reaction time increased from 1 to 6 h due to the increased solid and gas yields. At a longer reaction time, the probability of polymerization of PHAs increased, resulting in an increase in the solid yield. The solid yield increased from 3.4 to 6.3 wt % as the

reaction time increased from 1 to 6 h. The same yield trend as a function of reaction time was also observed for gas due to continuous cracking reactions. A slight decrease in viscosity of the upgraded oil was observed at a longer reaction time due to an increase in the proportion of low-boiling-point fractions.31 Table 1 provides the TAN, moisture, elemental composition, and other properties of the upgraded oils obtained at different reaction times. The TAN decreased from 17.2 to 3.8 as the reaction time increased from 1 to 6 h due to the reduced content of acidic compounds. The moisture content of the upgraded oil varied between 0.15 and 0.21 wt % with changing reaction time due to the formation of H2O during the hydrodeoxygenation reaction and original moisture in the LBD. The C and H contents of the upgraded oil remained almost constant at 85 wt % regardless of whether the reaction time was changed or not. A slight decreased in the H content of the upgraded oil was observed with increasing the reaction time. Therefore, a slight decrease in the H/C molar ratio of the upgraded oil was observed at a longer reaction time, indicating that the content of aromatics in the upgraded oil increased. Increasing the reaction time continuously decreased the N, O, and S contents of the upgraded oil, which decreased from 0.74 wt %, 2.81 wt %, and 70 ppm to 0.53 wt %, 1.62 wt %, and 41 ppm, respectively, as the reaction time increased from 1 to 6 h, indicating a longer reaction time was favorable for denitrogenation, desulfurization, and deoxygenation. However, the HHV of the upgraded oil remained at 43 MJ/kg for all examined reaction times. Therefore, in view of the yield and quality of the upgraded oil, 4 h was selected as suitable reaction time. For hydrogenation reactions, the initial hydrogen pressure is a very important factor affecting the product yield and quality of the upgraded oil. Therefore, the effect of hydrogen pressure was examined at 400 °C, 4 h, 5 wt %(Pt/C)/feed, and LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1 with changing hydrogen pressures from 0.1 to 10 MPa. For a comparison purpose, one additional reaction with 0.1 MPa Ar added was also performed. Figure 4(d) shows the results. Compared with the results obtained with 0.1 MPa Ar added, slightly higher upgraded oil and lower solid yields were observed with 0.1 MPa H2 added, suggesting that the presence of hydrogen could suppress coke formation via inhibiting the formation of coke precursor. The gas yield remained at around 10218


Article


catalyst is. The previous study indicated that increasing the catalyst loading promoted the removal of N, O, and S in the forms of NH3, H2O, and H2S, respectively.33 Furthermore, more CO2 and CH4 were formed via water−gas shift and methanation reactions at a higher catalyst loading. Therefore, a lower upgraded oil and higher gas yield was observed at a higher catalyst loading. By increasing the catalyst loading, more coke precursors would stay in the catalyst pores, and this effect would increase the possibility of polymerization reactions. Therefore, a slightly higher solid yield was observed at a higher catalyst loading. The increase in solid and gas yields occurred at the expense of reduced upgraded oil yields. Moreover, adding too much catalyst would bring difficulty in upgraded oil recovery via centrifugation. Table 1 lists the TAN, moisture, elemental composition, and other properties of the upgraded oils obtained at different Pt/C loadings. Increasing the catalyst loading is an effective way to reduce the moisture content of upgraded oil as the moisture content of the upgraded oil was decreased from 0.18 to 0.08 wt % as the catalyst loading increased from 0 to 20 wt %. A significant decrease in TAN from 9.4 to 3.8 was also observed as the catalyst loading increased from 0 to 5 wt %. Further increasing the catalyst loading to 20 wt % only slightly decreased the TAN. Clearly, catalyst loading had a significant effect on the elemental composition of the upgraded oil. Increasing the catalyst loading increased the C content and decreased the H, N, O, and S contents of the upgraded oil. In the absence of Pt/C, the upgraded oil also had lower N, O, and S contents than the initial feedstock, indicating that heat and hydrogen played an important role during the hydrotreating reaction. Increasing the catalyst loading further decreased the N, O, and S contents of the upgraded oil. At the highest catalyst loading of 20 wt %, the upgraded oil had the lowest N, O, and S contents of 0.07 wt %, 0.74 wt %, and 21 ppm, respectively. Possibly, physical adsorption was also responsible for the N, O, and S removal. The H/C molar ratio also decreased with increasing catalyst loading, implying that the content of aromatic compounds in the upgraded oil increased. Catalyst loading had little effect on the HHV of the upgraded oil, which remained at 44 MJ/kg for all catalyst loadings. The catalyst activity maintenance was examined at 400 °C, 4 h, 5 wt %(Pt/C)/feed, 10 MPa H2, and LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1. Since the Pt/C catalyst mixed with coke after the hydrotreating reaction, the Pt/C alone could not be separated from the solid residue. Therefore, activity maintenance was evaluated by employing the overall solid residue (catalyst and coke) recovered from the reactor at the end of a previous run as the catalyst in a subsequent run. We did nothing to the recovered solid residue prior to its subsequent use. Figure 4(f) shows the results. The upgraded oil yield increased from 82.7 to 88.3 wt % as the catalyst was used from the fresh state to its third reuse due to decreased catalyst activity. The solid yield was insensitive to catalyst recycling and remained at approximately 3.7 wt % during these three runs. The gas yield showed a modest decrease after the catalyst was reused for the second and third times. These results are consistent with previous results for crude algal bio-oil hydrotreated with Pt/γ-Al2O3.34 Table 1 lists the TAN, moisture, elemental composition, and other properties of the upgraded oils obtained from these three runs. Catalyst reuse had little effect on the moisture content of the upgraded oil as the moisture content always stayed at around 0.14 wt %. However, the TAN increased from 4.8 to 8.2 from

7.6 wt % whether the reaction was performed with 0.1 MPa Ar or an H2 atmosphere. The upgraded oil yield increased from 78.9 to 82.7 wt %, while the solid and gas yields decreased from 5.4 and 7.6 wt % to 3.7 and 5.6 wt %, respectively, as the hydrogen pressure increased from 0.1 to 10 MPa. In the hydrotreatment process, hydrogen is expected to react with the reactive oil intermediates fragments and stabilize them before they undergo polymerization and condensation reactions, which will lead to coke and char formation.32 Therefore, one would expect a lower solid yield at a higher initial hydrogen pressure. The upgraded oil produced at a higher hydrogen pressure also presented good flowability, which was favorable for the subsequent recovery of the upgraded oil, and thus a higher upgraded oil yield was also observed at a higher hydrogen pressure. Higher pressure is unfavorable for gas formation according to Le Chatelier’s principle, and thus, the lower gas yield was observed at a higher hydrogen pressure.31 Table 1 lists the TAN, moisture, elemental composition, and other properties of the upgraded oils obtained at different hydrogen pressures. The moisture content of the upgraded oil decreased with increasing initial hydrogen pressure due to the increase in the content of nonpolar compounds in the upgraded oil. The TAN also decreased from 9.2 to 3.8 as the hydrogen pressure increased from 0.1 to 10 MPa. A lower TAN was also observed when the reaction was performed at 0.1 MPa H2 relative to that at 0.1 MPa Ar, indicating that external H2 is favorable for deoxygenation. Compared with the upgraded oil produced under a 0.1 MPa Ar atmosphere, the upgraded oil produced under a 0.1 MPa hydrogen atmosphere had higher C and H contents and lower N, O, and S contents, although these differences were small, however. During the treatment process under an Ar atmosphere, the consumption of hydrogen during the hydrogenation reaction was solely contributed from the HDS. However, as the reaction proceeded, the HDS became a hydrogen-deficient substance, which was unfavorable for the hydrogenation reaction. In contrast, under a hydrogen atmosphere, the reacted HDS could be hydrogenated by the external hydrogen and returned to its initial status. Furthermore, the external hydrogen could also contribute hydrogen to the reaction. Therefore, the C and H contents increased, while the N, O, and S contents decreased with increasing the initial hydrogen pressure. The H/C molar ratio of the upgraded oil increased from 1.33 to 1.46 as the hydrogen pressure increased from 0.1 to 10 MPa due to saturation of unsaturated alkanes and aromatics in the oil, and this effect resulted in an increase in the HHV of the upgraded oil. At the highest hydrogen pressure of 10 MPa, the lowest N, O, and S contents of 0.35 wt %, 0.51 wt %, and 44 ppm were achieved, respectively. Therefore, in view of the yield and quality of the upgraded oil, 10 MPa was selected as the hydrotreating pressure. It should be noticed that higher initial pressure would increase the cost of this process. However, if the unreacted hydrogen could be recovered and reused in a new run, it will greatly reduce the cost of this process. The effects of catalyst loading on the product yields and upgraded oil quality were examined at 400 °C, 4 h, 10 MPa H2, and LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1 with changing catalyst loadings from 0 to 20 wt %(Pt/C)/feed. Figure 4(e) shows the results. Increasing the catalyst loading slightly decreased the upgraded oil yield and increased the solid and gas yield. The higher the catalysts loading is, the higher the residual upgraded oil in the spent 10219


Article


Supporting Information compare the total ion chromatograms of the biocrude, LBD, and upgraded oils produced under different reaction conditions. Different oil samples show different intensity scales in the chromatograms, although all the oil samples had the same concentration. However, the relative content of each individual compound is independent of the peak intensity of the oil samples because this content is simply the ratio of each peak area to the total areas of all compounds detected. The differences in the molecular compositions of the biocrude, LBD, and HBD are clearly evident in Figure S6 in the Supporting Information which compares the total ion chromatograms for these three oils. For the biocrude, the peaks are mainly concentrated in three retention time ranges of 4− 40, 74−94, and 109−119 min. The LBD mainly contains compounds with retention times less than 40 min, while the HBD mainly consists of compounds with retention times greater than 40 min. In contrast, the total ion chromatograms of upgraded oils produced with different HDSs or without an HDS showed many differences with each other, as shown in Figure S7 in the Supporting Information. No solvent peaks were observed when using cyclohexene and cyclohexane due to the 4 min solvent delay. Compared with the upgraded oil produced without an HDS, the presence of an HDS affected the component distribution of the upgraded oil. Figures S8− S12 in the Supporting Information show the effects of temperature, time, initial hydrogen pressure, catalyst loading, and catalyst reuse on the component distributions and their relative contents in the upgraded oil. Figure 5 provides a summary of the relative amounts of different compound classes in the biocrude, LBD, and upgraded oils produced under different reaction conditions. The relative contents in Figure 5 are expressed in terms of the percentage of the total peak area in the total ion chromatogram that corresponds to the identified compounds within a given compound class. Ocontaining compounds are the dominant fraction for the biocrude and LBD, and ketones and phenols are the two most prevalent oxygenated compounds. Li et al.37 studied the thermochemical conversion of rice straw in different solvents and showed that using water as the reaction medium the main components of biocrude are ketones and phenols. Biswas et al.38 used corn cobs, rice straw, wheat straw, and rice husks as raw materials to prepare pyrolysis oils and found that the pyrolysis oils had similar compositions and contained many Ocontaining compounds (phenols, ketones, carboxylic acids, etc.). Compared with the LBD, the upgraded oils produced with or without HDS had a higher saturated hydrocarbon content and lower amounts of O-containing, N-containing, and S-containing compounds. Compared with the upgraded oil produced without an HDS, the upgraded oil produced with HDS had higher hydrocarbon content and lower contents of O-containing, N-containing, and S-containing compounds, indicating that the HDS promoted denitrogenation, deoxygenation, and desulfurization. In a temperature range of 300−400 °C, the saturated hydrocarbon and aromatic contents in the upgraded oil increased, while the unsaturated hydrocarbons, O-containing compounds, and N,O-containing compounds decreased with increasing temperature, which is consistent with the elemental analysis as shown in Table 1. A significant decrease in saturated hydrocarbon content and an increase in aromatic content were observed as the temperature increased from 400 to 450 °C. The N-containing compound content always increased with increasing temperature, which is not

the employed fresh catalyst to the third reuse catalyst due to the activity loss of the Pt/C. The N, O, and S contents of the upgraded oil increased as the catalyst was reused for the second and third times. Since the LBD contained a certain amount of S-containing compounds, poisoning of the Pt noble metal by the sulfur in the LBD must have occurred, thus weakening the catalytic activity of the Pt/C.35 The S contents of the upgraded oils produced during these three runs were always lower than that produced without catalyst, indicating that the recycled catalyst still showed some activity for S removal. Of course, this desulfurization might be due to adsorption rather than the catalytic reaction.35 The catalysts recovered from the first run seemed to have no catalytic effect on O and N removal, as the upgraded oil produced from the second reuse of the Pt/C had a higher O content and an equal amount of N relative to those of the upgraded oil produced by the noncatalytic reaction. Elemental analysis of the solid after the third reuse of the catalyst suggested that this spent catalyst contained 1.67 wt % N, 3.56 wt % O, and 0.23 wt % S. We suspect that the N and O in the spent catalyst would enter the upgraded oil again, which resulted in an increase in the N and O contents of the upgraded oil. As mentioned above, the lowest S content of 21 ppm in the upgraded oil was achieved; however, this value is still higher than the required limit for China V diesel.36 The main reason for the difficult elimination of N, O, and S from the upgraded oil produced in a single run of hydrotreating the LBD is that equilibrium restrictions for denitrogenation, deoxygenation, and desulfurization existed due to the formation of NH3, H2O, and H2S in the reaction system. Therefore, the key to realizing deep hydrotreating of the LBD is to break the equilibrium restrictions for denitrogenation, deoxygenation, and desulfurization via in situ removal of NH3, H2O, and H2S in the reaction system. One additional experiment was performed at 400 °C, 1 h, 5 wt %(Pt/C)/feed, 6 MPa H2, and LBD:HDS (decahydronaphthalene:tetrahydronaphthalene=1:1)=1:1. After 1 h of reaction, the reactor was cooled to room temperature. The gas in the reactor was replaced with fresh 6 MPa H2. Then, the reaction was performed at 400 °C for 1 h. This process was repeated three times, and the total reaction time was 4 h. Finally, approximately 78.4 wt % upgraded oil and 5.3 wt % solid were achieved, which were both slightly lower than the upgraded oil and solid produced under the same reaction conditions when the reaction was always sustained for 4 h. The upgraded oil had N, O, and S contents of 0.51 wt %, 1.43 wt %, and 47 ppm, respectively, which were all lower than those of the upgraded oil produced under the same reaction conditions when the reaction was sustained for 4 h. Therefore, H2 replacement removed part of the NH3, H2O, and H2S in the reaction system, avoiding their secondary reactions with the upgraded oil. Molecular Characterization of Upgraded Oils. Many of the molecular components in the upgraded oils produced under different reaction conditions were separated and tentatively identified by comprehensive two-dimensional gas chromatography-time-of-flight mass spectrometer. All the upgraded oils contained a certain amount of high-molecularweight compounds that could not elute from the gas chromatography and be detected since a 4 min solvent delay was set to protect the filament, and this solvent delay resulted in that those compounds with a retention time shorter than 4 min could not be detected. Therefore, the data provided are for only a fraction of these oils. Figures S6−S12 in the 10220



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consistent with the elemental analysis, indicating other Ncontaining compounds were not detected in the upgraded oil. No or only very small amounts of S-containing compounds were observed in the temperature range examined. The variation trend in the contents of all categories as a function of time is similar to that of temperature. The initial hydrogen pressure had a significant effect on the relative content of each compound class. The saturated hydrocarbon content significantly increased, while the unsaturated hydrocarbon and aromatic contents decreased with increasing initial hydrogen pressure due to the promotion of hydrogenation reactions. Higher hydrogen pressure resulted in an upgraded oil with lower contents of N,O- and S-containing compounds. Catalyst loading mainly affected the relative contents of aromatics and heteroatom-containing compounds. Increasing the catalyst loading increased the aromatic content and decreased the N-, N,O-, O-, and S-containing compound contents, and these findings were consistent with the elemental compositions of the upgraded oils. Due to the decrease in the catalytic activity of Pt/C, the contents of saturated hydrocarbon, N-, N,O-, and S-containing compound classes increased while the aromatic content increased with the continuous recycling of the catalyst.

AUTHOR INFORMATION

Corresponding Author

*Tel./Fax: +86-029-82665836/+86-029-3987823. E-mail: [email protected]. ORCID

Pei-Gao Duan: 0000-0002-9461-3566 Notes

The authors declare no competing financial interest. Biographies

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CONCLUSIONS The biocrude produced from hydrothermal liquefaction of soybean straw was separated into LBD and HBD by vacuum distillation. The LBD and HBD showed some differences in properties from the biocrude. The LBD was effectively hydrotreated in the presence of HDS. Using centrifugal separation could reduce the loss of the light-end fraction and avoid contamination of the extraction solvent. The presence of HDS can not only reduce the production of coke and gas but can also facilitate the transfer of hydrogen and improve the hydrogenation efficiency during the hydrotreating process. Better removal efficiencies for N, O, and S when using decahydronaphthalene and tetrahydronaphthalene blends were observed compared with those when using decahydronaphthalene or tetrahydronaphthalene alone. The hydrogen in the hydrotreating mainly came from the contribution of the HDS. The S content of the upgraded oil was significantly affected by all the parameters examined, followed by O and N. Under the optimal reaction conditions of 400 °C, 10 MPa H2, 20 wt % Pt/C, and 4 h, the removal efficiencies of denitrogenation, deoxygenation, and desulfurization reached 93%, 95%, and 99%, respectively, and the S content of the upgraded oil could be reduced to 21 ppm. The upgraded oils mainly consisted of saturated alkanes, unsaturated alkanes, and aromatic compounds, and their relative contents could be adjusted by controlling the amount of HDS added. The catalytic activity of Pt/C was reduced after the catalyst was reused. We view that the integration of HDS and hydrogen is an effective way for upgrading the low-boiling-point fraction of bio-oil.

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Article

Zhi-Cong Wang is a master candidate in the laboratory of Pei-Gao Duan at Xi’an Jiaotong University. She received her B.S. from Zhengzhou University in 2019. She has research experience in biofuels production and their upgrading, specifically in direct liquefaction forestry and agricultural residues and upgrading of crude bio-oils derived from forestry and agricultural residues using the thermochemical method.

Pei-Gao Duan is a professor of chemical engineering at Xi’an Jiaotong University. He received his B.S. from Henan Normal University in 2004 and his Ph.D. in 2009 from the East China Normal University. From 2008 to 2010, he worked as a visiting scholar and postdoctoral researcher in the laboratory of Phillip Savage at the University of Michigan. He had served as an associate professor and professor in the Henan Polytechnic University between 2010 and 2019. Research interests of this group mainly involve (1) production of high valueadded chemicals and liquid fuels from biomass, (2) valorization of solid waste, (3) utilization of CO2 and development of new catalyst, and (4) catalytic conversion of syngas to fuel and high value-added chemicals.

ASSOCIATED CONTENT

S Supporting Information *

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.9b01274.

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Total ion chromatograms of the liquid products obtained from hydrotreating the HDS, LBD, and HDS +LBD blend under different reaction conditions. (PDF)

ACKNOWLEDGMENTS We gratefully acknowledge the financial support of the National Natural Science Foundation of China (21776063), 10221


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the United Fund for NSFC and Henan Province (U1704127), the Scientific and Technological Innovation Team of the University of Henan Province (18IRTSTHN010), the Outstanding Youth Foundation for Scientific and Technological Innovation in Henan Province (184100510013), and the Key Scientific Research Projects in Colleges and Universities of Henan Province (18A580003).

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Hydrotreating the Low-Boiling-Point Fraction of Biocrude in Hydrogen

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