Heavy oil catalytic upgrading under methane environment: A small pilot plant evaluation

Fuel 258 (2019) 116161

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Full Length Article

Heavy oil catalytic upgrading under methane environment: A small pilot plant evaluation

T

Peng Hea, Shijun Menga, Yang Songb, Bei Liuc, Hua Songa,

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a

Department of Chemical and Petroleum Engineering, University of Calgary, 2500 University Dr NW, Calgary, Alberta T2N 1N4, Canada Guangxi Sino-Green Energy and Environmental Technology Ltd., China c PetroChina Guangxi Tiandong Petrochemical Co., Ltd., China b

ARTICLE INFO

ABSTRACT

Keywords: Heavy oil Methane activation Upgrading Catalysis Pilot plant test

The effects of variable factors such as temperature, pressure, weight hourly space velocity (WHSV), catalyst regeneration and reaction time on catalytic upgrading of heavy oil under methane environment over Ag-Ga/ HZSM-5 catalyst are evaluated by small pilot plant tests. The viscosity, TAN and average molecular weight of the heavy oil samples are notably reduced accompanied with a low olefin content after the catalytic upgrading. Methane incorporation to product molecules results in increased liquid product yield. The simulated distillation and compositional analysis results illustrate the conversion of resin and asphaltene components to small molecules. After the upgrading at 430 °C and 5 MPa, the percentage of gasoline content is increased to 17% from 3% in the feedstock, and the percentage of diesel fraction is increased to 27% from 21%. The percentage of light end volatiles is increased from 4.4% to 22.2%. During the catalytic upgrading, a higher temperature and pressure benefits the oil quality improvement and methane participation while maintaining the high liquid yield and low coke generating rate. Outstanding stability of the product oil samples is witnessed by comparing the products obtained at variable reaction time. A higher WHSV results in slightly compromised performance of the catalytic upgrading of heavy oil under methane. A positive effect of catalyst regeneration is witnessed. Stable oil quality from the upgrading process is observed when catalysts have been regenerated twice or more. The knowledge obtained in this study would benefit the potential industrial application of this innovative technology at larger scales and guide the further reaction condition optimization.

1. Introduction Nowadays, as the primary fossil fuel resource, petroleum is considered the most important power source and the resource of hydrocarbon chemicals for industry production and consumption of human society. As the climate change due to CO2 emission from fossil fuel attracts more attention, the production of more valuable chemicals from crude oil becomes a more important approach in the long term. Great efforts have been made to convert fossil fuel into chemical feedstocks like benzene, toluene and xylene. This approach is particularly promising for heavy oil, which contains more aromatic species than conventional crude oil. On the other hand, the large scale of heavy oil potential reserve makes it an alternative source of crude oil to supply hydrocarbon chemicals and power the world. For instance, the remaining proven reserves of heavy oil in Canada is 166.7 billion barrels by the end of 2018 according to the statistics released by Natural Resources Canada, making it the third largest oil reserve in the world,

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after Venezuela and Saudi Arabia. The heavy oil production (mined and in situ) totaled about 2.8 million barrels per day in 2017. However, the value of heavy oil is often lower than that of conventional petroleum deposit due to its high viscosity, high asphaltene content, low hydrogen to carbon ratio, and high impurities content [1]. A typical heavy oil in Canada has an average density of 1.0077 g/cm3 and a dynamic viscosity of 2 × 105–2 × 106 cP at atmospheric conditions [2]. The Western Canadian Select (WCS) price of heavy oil hit 10.3 USD per barrel in November 2018, while the West Texas Intermediate (WTI) price was around 50 USD per barrel in the same period. The large price gap between conventional petroleum and heavy oil has attracted much attention on the utilization of heavy oil. Its conversion to valuable chemicals provides a new method to increase the value of heavy oil, which often contains a large fraction of aromatics and can be used as a better resource to produce chemical feedstocks rather than fuels. In this way, the potential carbon dioxide emission resulting from fuel combustion is also reduced, which is critical as the

Corresponding author. E-mail address: [email protected] (H. Song).

https://doi.org/10.1016/j.fuel.2019.116161 Received 2 August 2019; Received in revised form 26 August 2019; Accepted 4 September 2019 Available online 12 September 2019 0016-2361/ © 2019 Elsevier Ltd. All rights reserved.

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society becomes more worried about the greenhouse effect. A conventional approach to upgrade heavy oil is via thermal cracking [3], which produces smaller molecules along with lots of coke, followed by hydrotreating using hydrogen produced by the steam reforming of natural gas [4]. Another widely used process is adding hydrogen and the catalyst directly into cracking unit to perform catalytic hydrocracking [5]. Some unsaturated olefin molecules from cracking are converted into saturated species by the incorporation of hydrogen molecules. Meanwhile, the sulfur, nitrogen and metal species are also removed from product oil. However, this process must be operated at high hydrogen pressures (14–28 MPa). Besides the requirement of high pressure, this process needs to consume a large quantity of hydrogen, which is not naturally available but produced from methane steam reforming and introduces extra energy consumption over another series of devices. The pressurization of hydrogen also consumes energy and result in additional CO2 emission in the life cycle of this process. These costs will decrease the overall economic and environmental profit of the formed products from heavy oil upgrading. On the other hand, hydrogen is mainly produced from the steam reforming of natural gas. If methane, the principal component of natural gas, is directly used as the hydrogen source for heavy oil catalytic cracking without the steam reforming process, we can avoid the aforementioned costs and environmental issues. Moreover, the operating cost of the upgrading process could be further reduced if a lower reaction pressure (e.g., 5 MPa) can be engaged. Similar to heavy oil, the value of natural gas is also heavily underestimated. The past decade has witnessed the boosted production of natural gas in North America. The production of natural gas in US is increased from 70.6 billion cubic feet per day in June 2008 to 98.4 billion cubic feet per day in June 2018. As a result, the price of natural gas declined dramatically. In June 2018, the price of Henry Hub natural gas was US $2.97 per million British thermal units (MBTU), which is dramatically lower than the price of $12.69 per MBTU in June 2008. Due to the gaseous nature of natural gas, its transportation from the reservoirs to the end users heavily depends on the pipeline construction. The transportation cost impose a significant and even dominant impact on the price of natural gas. In 2018, the average price of natural gas in Canada (Alberta), US (Henry Hub), Germany and Japan are 1.12, 3.13, 6.62 and 10.05 USD, respectively [6]. End users at different locations around the world have to pay dramatically different prices for natural gas, demonstrating that transportation makes up a huge fraction on the cost of natural gas. Therefore, the conversion of methane from gas phase to liquid phase commodities has attracted much attention to boost the profitability of the natural gas industry. The catalytic transformation of methane into more valuable carbon-containing products has provided an effective alternative for the use of natural gas. However, the inert structure of methane impedes its conversion into more value-added chemicals and fuels [7]. In its current utilization scenario, methane must follow multistep conversion strategies via syngas and/or methanol before it can be transformed into higher hydrocarbons, resulting in increased costs and impeding its commercial potential. Oxidative coupling of methane to convert it into more valuable chemicals such as ethylene and methanol has been extensively studied [8–12], but the product yield is limited by the better reactivity of the desired products than methane. Under non-oxidative environment, the activation of methane requires a high temperature, often above 700 and even 1100 °C [13]. Nevertheless, the hydrocarbon moiety produced during heavy oil cracking may assist methane activation at much lower temperatures, i.e., 400–600 °C [14–16]. In recent years, the catalytic heavy oil upgrading has drawn researchers’ attention. Eshraghian et al. [3] has researched on the pyrolysis of heavy oil. They found that one third of gas product is methane. In the presence of catalysts, methane produced in this process will add into the product oil and improve its quality. Ajumobi et al. [17] has studied the catalytic upgrading using catalysts containing Ce, Zr, Fe and Co. They have drawn a conclusion that the oxygen atoms and acid sites

on the surface of catalyst are crucial for heavy oil cracking and upgrading. The upgrading of heavy oil with formic acid in the presence of supercritical water results in lighter liquid product with less coke formation [18]. Formic acid is employed as the hydrogen donor in this process to replace H2 to make the upgrading more economically applicable. The catalytic heavy oil upgrading under methane environment has also been intensively investigated by our research group. It is found that methane will interact with the small molecule intermediates produced from heavy oil cracking over the Ag-Zn/HZSM-5 [19]. This process will decrease the viscosity and TAN of product oil and increase its H/C ratio, which results in a better quality and improved value of the product oil. A series of control experiments using styrene and butylbenzene as model compounds have been conducted to prove methane incorporation and the critical reaction intermediates. Unsaturated species formed upon heavy oil molecule cracking may facilitate methane activation and react with the activated methane species. These positive effects are even more significant when Ag-Ga/ HZSM-5 is employed as the catalyst to upgrade heavy oil under methane environment [20]. Isotopic labelling study has also been employed to track the incorporation of methane into the product molecules over various catalysts such as Mo/HZSM-5 [21] and Ag-Ce-Mo/ HZSM-5 [22]. Hydrogen and carbon atoms are enriched at the phenyl ring and benzylic sites, indicating the participation of methane in the aromatization process. The incorporation of methane enhances the formation of aromatic products, which is critical for the conversion of heavy oil molecules to aromatic products with an increased product yield. The activation and conversion of methane with other feedstocks such as paraffins, olefins and biomass have also been extensively explored in the same research group [23–26], illustrating the methane participation reaction network in the complex reaction system of heavy oil upgrading. The commercial application of heavy oil upgrading under methane environment, however, requires more careful optimization in terms of reaction parameters such as temperature, pressure and catalyst conditions to make the process more profitable. For instance, it is observed that an increased temperature is beneficial for the aromatization of hydrocarbon molecules. When the reaction temperature is 450 °C and the weight hourly space velocity is 3 h−1, the conversion of n-hexane is 77% and the selectivity of benzene could reach to 89% [27]. Ancheyta et al. [28] reported that an increased hydrogen pressure in heavy oil upgrading promoted the improvement of light liquid fractions at the expense of the heavy and residue fractions. On the other hand, executing the reactions using continuous reactors instead of batch reactors makes it more accurate to investigate the stability of the catalysts in a long term [29,30]. The reaction scale may also impose an impact on the catalytic performance. Therefore, even though the catalytic performance study using a pilot plant has to consume more feedstock and energy, this type of testing at larger scales still has strong necessity to give more accurate information for guiding its eventual industrial applications [31,32]. Inspired by the results from the lab scale study, the performance of heavy oil upgrading under methane environment is evaluated using a fixed bed continuous reactor with a maximum processing capacity of 20 L/day in this work. Ag-Ga/HZSM-5 has demonstrated a good catalytic performance in heavy oil upgrading under methane environment [20]. The participation of methane in the reaction improves the product oil quality in terms of reduced viscosity and coke yield along with increased gasoline and diesel content in the product oil. The catalysts modified by 1wt% Ag and 1wt% Ga showed a better performance than other combinations. Therefore, they are employed as the active metal species in this study. The quality of the product oil in terms of viscosity, density, total acid number (TAN), average molecular weight, water content and stability are characterized to obtain a comprehensive understanding of the heavy oil upgrading process. More direct insights regarding the evolution of heavy oil molecules during the catalytic upgrading are achieved by the distillation profile, gasoline and diesel fraction, and SARA (saturate, aromatic, resin and asphaltene) 2

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composition analysis. Liquid product yield and coke generation rate are also monitored. The stability of the catalysts, critical for industrial application, is appraised by performing the reaction for 7–10 continuous days.

major components of the reactor system are shown in Scheme 1 and Fig. S1. 500 mL catalyst is loaded into the constant temperature section of the reactor. To control the position of the catalyst, ceramic balls are loaded above and under the catalyst bed. The crude heavy oil is filled into a raw material tank with heating and temperature control. This tank is heated up to 50 °C to make sure the crude heavy oil can flow into the reactor during the heavy oil injection by pump. The reactor temperature is controlled by electric heating devices. There are six thermocouples inserted to the reactor at various sections aligning well with six heating zones of the surrounding electric heater. The temperature values are sent to the corresponding heating controllers to modify the power to achieve the designated temperature. Six thermocouples are employed to make sure the temperature is uniform throughout the whole reactor. To record the amount of oil intake, an electric scale is put under the raw material tank. The pressure of reactor is maintained at a certain level with a methane flow. The methane flowrate is controlled and measured by a mass flow controller. After liquid-gas separation, the product oil enters the product tank. The gas product is also cooled by circulating water at room temperature to condense more liquid product. From the product tank and product container, the product oil is collected and transferred into a product barrel. The weight change of product barrel after collection is regarded as the amount of oil production. When reaction is done, nitrogen is introduced to replace methane in reactor system. After cooling, the operators open the reactor, take out the catalyst and measure the coke generation rate.

2. Experimental 2.1. Feedstock and chemicals A Merey Crude Oil feedstock was acquired from PetroChina Guangxi Tiandong Petrochemicals and used as received without further treatment. Toluene with a purity of > 99.5 wt% (provided by VWR International) was used to isolate any possible toluene-insoluble solids in bitumen and its upgraded product (i.e. possibly coke and catalyst). Heptane (anhydrous, 99%, Sigma-Aldrich) was exclusively employed to determine asphaltenes content of various oil samples. 2.2. Catalyst synthesis The ammonium ZSM-5 zeolite with SiO2/Al2O3 molar ratio of 23 and specific surface area of 425 m2·g−1 was purchased from Alfa Aesar and calcined at 600 °C for 5 h in air to attain the H-type ZSM-5 for further use. The 1wt% Ag-1wt% Ga/HZSM-5 was prepared by incipient wetness impregnation of HZSM-5 with AgNO3 (99.0%, Sigma Aldrich) and Ga(NO3)2·xH2O (99.9%, Alfa Aesar) aqueous solution, dried in the oven at 92 °C overnight, followed by calcination at 600 °C for 3 h in ambient air. The resultant catalyst was formed to a cylindrical shape with a diameter of 2 mm and length of 1 cm for use through engaging mixture of LUDOX AM-30 colloidal silica, citric acid and methylcellulose as the binder. The fresh catalyst is denoted Catalyst-1. After the reaction, the catalyst was regenerated by washing with toluene and calcination at 600 °C in air for 5 h. The regenerated catalysts are denoted Catalyst-2 and Catalyst-3 when they are used for the second and third time.

2.4. Characterizations The water content of liquid sample produced from each run was determined using Karl Fischer titration (Metrohm 870 Titrino Plus) through averaging the results collected from at least three independent measurements. The total acid number (TAN) of liquid sample produced from each run was measured using a Metrohm 848 Titrino Plus through averaging the results collected from at least three independent measurements. The density values of the oil samples are obtained using an Anton Paar DMA 4500 M density meter at 15.6 °C. The compatibility & stability feature of the collected oil samples was analyzed by engaging a TriboAmix® spot tester designed by TRIBOMAR GmbH and following a testing procedure defined in ASTM D4740. The viscosity of oil sample

2.3. Catalytic performance test The heavy oil upgrading facility is a fixed bed continuous reactor system with maximum heavy oil processing capacity of 20 L/day. The

Scheme 1. Schematic diagram of reactor system. 3

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determined by 1H NMR study of the oil samples dissolved in D8-toluene. The NMR peak area percentage due to olefin contents was calculated. NMR spectra of standard samples prepared by mixing 1-decene with crude oil without olefin content were also collected in a similar manner to calibrate the olefin content values. The Transmission Electron Microscopy (TEM) spectra were acquired on a JEOL2200FS TEM instrument operated at 200 kV. TEM images were acquired in Bright Field. The sample was first dispersed in ethanol and sonicated for about 2 min. A 10 μL droplet was placed on honey carbon grids before the TEM images were recorded. The crystalline phase compositions of prepared catalysts were examined by X-ray diffraction on a Rigaku Multiflex diffractometer with Cu Kα irradiation at a voltage of 20 kV and current of 40 mA in the 2θ of 5–60°. The BET surface areas were measured using ASAP 2420 from Micromeritics. Samples were outgassed under a vacuum at 250 °C for 2 h and then brought to −196 °C via immersion in a liquid nitrogen bath. Total surface area was calculated using multipoint BET analysis. Pore surface area and pore volume were calculated using Barrett–Joyner–Halenda (BJH) analysis. Surface acidity measurements were performed by NH3TPD using ~300 mg samples in a Finesorb-3010 chemisorption analyzer. Prior to measurements, the samples were activated in Ar at 600 °C for 30 min with a ramp rate of 10 °C/min. All samples were then cooled to 100 °C for adsorption of ammonia, performed using a flow of 30 sccm of 5% NH3/Ar for 30 min. After flowing Ar at 30 sccm for 30 min to remove any physically adsorbed NH3, TPD was carried out by ramping to 600 °C at 10 °C/min increments and holding for 30 min. A thermal conductivity detector (TCD) determined the amount of desorbed NH3. The Transmission Electron Microscopy (TEM) spectra were acquired on a JEOL2200FS TEM instrument operated at 200 kV. TEM images were acquired in Bright Field. The sample was first dispersed in ethanol and sonicated for about 2 min. A 10 μL droplet was placed on honey carbon grids before the TEM images were recorded. The crystalline phase compositions of prepared catalysts were examined by X-ray diffraction on a Rigaku Multiflex diffractometer with Cu Kα irradiation at a voltage of 20 kV and current of 40 mA in the 2θ of 5–60°. The BET surface areas were measured using ASAP 2420 from Micromeritics. Samples were outgassed under a vacuum at 250 °C for 2 h and then brought to −196 °C via immersion in a liquid nitrogen bath. Total surface area was calculated using multipoint BET analysis. Pore surface area and pore volume were calculated using Barrett–Joyner–Halenda (BJH) analysis. Surface acidity measurements were performed by NH3-TPD using ~300 mg samples in a Finesorb-3010 chemisorption analyzer. Prior to measurements, the samples were activated in Ar at 600 °C for 30 min with a ramp rate of 10 °C/min. All samples were then cooled to 100 °C for adsorption of ammonia, performed using a flow of 30 sccm of 5% NH3/Ar for 30 min. After flowing Ar at 30 sccm for 30 min to remove any physically adsorbed NH3, TPD was carried out by ramping to 600 °C at 10 °C/min increments and holding for 30 min. A thermal conductivity detector (TCD) determined the amount of desorbed NH3.

was determined by employing a Viscolead One Series L viscometer fabricated by Fungilab at temperature of 60°F (15.6 °C) controlled through an incorporated chiller. Average molecular weight was measured using a K-7000 vapor pressure osmometer manufactured by KANUER. The coke generation rate was obtained by recovering the catalyst after the reaction. The catalyst was first washed by toluene to remove heavy oil residue. The catalyst was then dried at 200 °C to remove toluene. The mass of the dry catalyst was recorded. It was then calcined at 600 °C for 5 h in air to remove coke. The mass of the regenerated catalyst was recorded. The mass change is considered due to coke deposit. The coke generation rate is calculated by the coke deposit mass divided by the mass of the regenerated catalysts and reaction time. The asphaltene contents of various oil samples were determined according to ASTM standard D 6560-00. Briefly, small amount of the sample was mixed with heptane and the mixture was heated under reflux. The resultant precipitated asphaltenes, waxy substances, and inorganic material were collected on a filter paper. The waxy substances were then removed by washing with hot heptane in an extractor. After removal of the waxy substances, the asphaltenes were separated from the inorganic material by dissolution in hot toluene. The obtained extraction solvent underwent evaporation and the received asphaltenes were finally weighed for calculating asphaltenes percentage. SARA analysis was executed by a modified ASTM D2007 method for a smaller sample size to determine the Saturate, Aromatic, Resin and Asphaltene contents in the oil samples. Asphaltene contents are composed of polyaromatic molecules containing oxygen, nitrogen and sulfur atoms, making them polarizable. Asphaltene cannot be dissolved in paraffins but are dissolved in toluene. Besides asphaltene, oil samples also contain maltene contents including saturates, aromatics and resin. Saturates such as paraffins and naphthenes are nonpolar. Aromatics with one or more aromatic rings are slightly polarizable. Resin is polarizable due to the heteroatoms in the molecules. But resin is soluble in paraffins thanks to the smaller phenyl ring units and longer alkyl chains [33]. Asphaltenes was removed from the sample with n-pentane at ratio of 40:1 solvent to sample. The asphaltenes were removed by filtration with an 8 μm filter paper, and the maltenes were collected and dried with a rotovap. Between 1.25 and 1.5 g of maltenes was charged on the top of a two column set containing 40 g silica gel and 10 g Attapulgus clay (lower column) and 20 g Attapulgus Clay (upper column). The silica gel was sized to between 28 and 60 mesh. The column was eluted sequentially with 150 mL of n-pentane, followed by 312 mL of toluene. The clay and silica are separated, the clay is eluted with 100 mL of 60:40 methanol:dichloromethane, followed by 50 mL of methanol. The silica was refluxed with 100 mL toluene for 1 h. The solvents i.e., pentane, toluene and methanol:dichloromethane (DCM) were evaporated in the rotary evaporator at specific temperature and vacuum pressure. Afterward, all three fraction were dried under a gentile nitrogen current to constant mass. The n-pentane fraction gave saturates, the toluene fraction gives aromatics and the methanol:DCM fraction gave polars. Light end loss was calculated form simulated distillation, knowing that material boiling below 200 °C would be lost in the drying step. Other loss was the difference of the sum of all other fractions and loss from 100%, typically materials which did not come off of the column, along with handling losses. Simulated distillation curves were obtained using a GC 8890 manufactured by Agilent equipped with a 30 m HP-5 column and a FID detector. The signal was fitted by a software manufactured by Separations System to obtain the mass fraction of crude oil components at different boiling points. Solid content was determined by removing the contents soluble in toluene. A small amount of the sample was placed on a filter paper, which was washed with hot toluene in an extractor. After that, the filter paper was dried at 120 °C in air to remove toluene. The mass change of the filter paper is calculated and attributed to the solid content in the oil samples. The olefin content was

3. Results and discussion The catalytic upgrading of heavy oil is carried out under a series of various conditions to investigate the effect of several reaction parameters such as temperature, pressure, weight hourly space velocity (WHSV), catalyst regeneration and reaction time. There results would benefit the potential application at larger scales. The reaction conditions in this work are summarized in Table 1 as Entry 1–8. The reaction time is 168 h for all of them. It is noticed that liquid product yields are above 99% in all of them, and even exceed 100% under some conditions, suggesting the incorporation of methane to the liquid product. The coke generation rates are also calculated in these reactions, which are all maintained at a low level, below 0.0015 g/(gcat·h). The primary characterization results in terms of viscosity, density, TAN, olefin content, average molecular weight, asphaltene content, solid content and water content of the heavy oil feedstock and the product oil 4

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samples are shown in Tables 2–4. The SARA and simulated distillation results of selected samples are tabulated in Tables 5 and S1. A detailed discussion on the effects of different factors are given in the following sections.

Table 1 Reaction conditions of the heavy oil upgrading under CH4 environment. Entry

Catalyst

T/°C

P/MPa

Liquid yield/ wt.%

Coke Formation rate/g/ (gcat·h)

1 2 3 4 5 6 7 8*

Catalyst-2 Catalyst-1 Catalyst-2 Catalyst-3 Catalyst-2 Catalyst-3 Catalyst-3 Catalyst-3

400 410 410 410 420 430 430 430

5 5 5 5 5 5 10 5

104.1 101.5 102.3 102.6 99.8 99.5 99.8 99.1

0.00117 0.00132 0.00120 0.00123 0.00131 0.00145 0.00140 0.00142

3.1. Effect of reaction temperature The upgrading of the heavy oil is carried at variable temperatures from 400 to 430 °C to explore the effect of temperature. The viscosity of heavy oil feedstock is 3866 cP. The viscosity of the product oil is significantly reduced after the reaction, indicating the cracking of large molecules. After the reactions at 400 and 410 °C, the viscosity is reduced to 795–1147 cP (Table 3), and further reduced to 128–177 cP at 420–430 °C (Table 4). Besides the viscosity, the average molecular weight is also lowered from 228 g/mol to 210, 195, 192 and 185 g/mol after reactions at 400, 410, 420 and 430 °C. The larger reduction of these parameters at higher temperatures demonstrates a better cracking of heavy oil molecules at higher temperatures as expected. On the other hand, at 400 and 410 °C, the liquid product yield is 101.5–104.1%, which drops to 99.1–99.8% at 420 and 430 °C. The elevated temperature improves the conversion of heavy oil to light contents without sacrificing much of the liquid yield, though more gas products are obtained at elevated temperatures. Under all the reaction conditions, the solid content remains at a low level, indicating a low coke generation rate in these reactions. However, the coke generating rate is increased from 0.00117–0.00132 h−1 at 400 and 410 °C to 0.00131–0.00145 h−1 at 420 and 430 °C (Table 1). Though the coke generation rate is still low, it may impose a negative impact on the upgrading process in the long term. Therefore, the test was not performed at higher temperatures. By comparing the parameters of feedstock and product oil tabulated in Tables 2–4, it is noticed that density is slightly reduced, which is more significant at higher temperatures since the cracking of heavy oil molecules is more significant. However, the density reduction is not dramatic, probably due to the high density of aromatic products than paraffins due to the strong interaction between the phenyl rings. The TAN values, on the other hand, are greatly reduced after the reaction and a greater reduction is witnessed at higher temperatures. After reaction at 400 °C, the TAN value is reduced from 2.28 mg KOH/g to 1.55 mg KOH/g, which is further reduced to 1.21, 1.13 and 0.85 mg KOH/g at 410, 420 and 430 °C. These observations illustrate that the conversion of acid groups in the catalytic reaction under methane environment is promoted by a higher temperature. The reaction between acid groups such as –COOH and methane during the reaction would produce H2O molecules, resulting in an increased water content, as shown by the water content values in Tables 2–4. The water content increment is more dramatic at higher temperatures, in line with the lower TAN values. The effect of reaction temperature is also explored by the compositional analysis. Heavy oil feedstock and the product oil samples are composed of very complicated fractions. To better interpret the compositional change during the reaction, the components are summarized to four parts such as saturate, aromatic, resin and asphaltene (SARA). The SARA analysis results of the feedstock and the product oil samples obtained at 400 (Entry #1) and 430 °C (Entry #6) are listed in Table 5. In the feedstock, the mass percentages of asphaltene content is 10.2%. The asphaltene content in the SARA analysis is determined as insoluble in pentane with diameter > 8.0 μm, which can be dissolved by toluene. After the reaction at 410 °C, the mass percentage of asphaltene is reduced to 6.2%. The mass percentage of resin is reduced from 18.4% to 14.7%. On the other hand, the percentage of light end volatiles is increased from 4.4% to 15.6%. The light components such as benzene, toluene and xylenes (BTX) are included in the light end volatiles based on the simulated distillation results. The notable increment in volatile percentage demonstrates the conversion of heavy oil feedstock to the light hydrocarbons. The production of light hydrocarbons is even more

* WHSV of 1 h−1 in Entry 1–7 and WHSV of 2 h−1 in Entry 8. Table 2 The properties of heavy oil feedstocks. Property

Raw Oil

Viscosity (15.6 °C)/cP Density (15.6 °C)/g/mL TAN/(mg KOH/g) Olefin content/wt% (1-decene) Asphaltene content/wt% Average molecular weight/(g/mol) Water content/% Solid content/%

3866 0.9645 2.28 0 10.4 228 0.33 0.01

Table 3 The properties of product oil obtained at 400 and 410 °C. Reaction Entry

1

2

3

4

Catalyst T/°C Viscosity (15.6 °C)/cP Density (15.6 °C)/g/mL TAN/(mg KOH/g) Olefin content/wt% (1-decene) Asphaltene content/wt% Average molecular weight/(g/ mol) Water content/% Solid content/%

Catalyst-2 400 1147 0.9621 1.55 0.91 5.7 210

Catalyst-1 410 976 0.9602 1.38 0.92 4.8 203

Catalyst-2 410 803 0.9588 1.25 0.95 4.3 197

Catalyst-3 410 795 0.9590 1.21 0.96 4.2 195

0.21 0.03

0.17 0.02

0.16 0.03

0.19 0.03

Table 4 The properties of product oil obtained at 420 and 430 °C. Reaction Entry

5

6

7

8

Catalyst T/°C Viscosity (15.6 °C)/cP Density (15.6 °C)/g/mL TAN/(mg KOH/g) Olefin content/wt% (1-decene) Asphaltene content/wt% Average molecular weight/(g/ mol) Water content/% Solid content/%

Catalyst-2 420 177 0.9563 1.13 1.31 3.4 192

Catalyst-3 430 128 0.9530 0.98 1.68 3.1 185

Catalyst-3 430 142 0.9551 0.85 1.42 3.2 188

Catalyst-3 430 135 0.9558 1.02 1.79 3.4 190

0.22 0.04

0.24 0.03

0.27 0.04

0.26 0.02

Table 5 SARA analysis results of the raw oil and product oil samples. SARA

Raw Oil

1

6

Light End/wt% Saturates/wt% Aromatics/wt% Resins/wt% Asphaltenes/wt%

4.4 26.1 36.6 18.4 10.2

15.6 25.3 37.0 14.7 6.2

22.2 23.5 36.6 13.6 4.2

5

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conclusion is obtained when looking at the simulated distillation results of other samples. Therefore, lots of high value chemicals including benzene, toluene and xylenes are probably in these light end volatiles as reported from SARA analysis. These phenomena show that the formation of valuable chemicals with small molecules and boiling points below 230 °C including benzene, toluene and xylenes significantly counts for the greatly increased fraction of volatiles. In the feedstock, the percentage of gasoline components with a boiling point below 210 °C is about 3%, and the percentage of gasoline plus diesel components with a boiling point below 380 °C is about 24% (Table S1). It is also shown that half of the components have a boiling point above 513.8 °C, while a quarter of them are distillated only at above 640.6 °C. After the reaction at 400 °C (Entry #1), the percentage of gasoline and diesel components is increased to 37%. The boiling points to distillate 50% and 75% of the components are reduced to 459.4 °C and 612.4 °C. These phenomena illustrate the conversion of heavy oil molecules to lighter components during the catalytic upgrading under methane. The reaction at 430 °C (Entry # 6) witnesses more remarkable improvements. The percentage of gasoline and diesel is greatly increased to 44% including 17% gasoline and 27% diesel. The boiling points to distillate 50% and 75% of the components are greatly reduced to 415 °C and 575.6 °C, i.e., a reduction of more than 100 °C. These phenomena reveal that the conversion of heavy oil molecules to lighter components is greatly accelerated by the lifted temperature. The impact of reaction temperature on the stability of the products is also explored, which is examined by spot test (Fig. 2). By comparing the images of the product oil samples after the spot test, it is illustrated that homogeneous spots are obtained from all the reaction conditions. Therefore, we can claim that the product oil samples are stable after the catalytic upgrading under methane environment at temperatures from 400 to 430 °C.

Fig. 1. Simulated distillation results of raw oil and product oil samples.

dramatic at a higher temperature. After the reaction at 430 °C, the mass percentage of asphaltene content is further reduced to 4.2%, while that of resin is reduced to 13.6%. The conversion of asphaltene and resin increased the percentage of volatiles to 22.2%. The amount of volatiles is increased by almost 4 times, demonstrating an even greater commercial application potential of the upgrading at 430 °C. The asphaltene contents of the product samples are determined by following ASTM standard D6560-00 (Tables 3 and 4). A similar trend is observed over the samples obtained under other conditions. A significant reduction of asphaltene content demonstrates the conversion of asphaltene contents during the catalytic upgrading. This change is more dramatic at 420–430 °C than that at 400–410 °C, indicating a high temperature is critical for the conversion of asphaltene component. The simulated distillation results of the feedstock and the product oil samples are tabulated in Table S1, where the temperatures required to distillate different mass fractions of oil components are listed. The corresponding distillation curves are shown in Fig. 1. It is clearly demonstrated that the curve due to the feedstock is located at lower end, which moves up after the reaction at 400 °C. After the reaction at 430 °C, the curves move even higher, indicating that more heavy fractions are converted to light products. The mass fractions of components with variable boiling points are available from Tables 5 and S1. By comparing the fraction of light end volatiles (4.4%) and the simulated distillation results of the feedstock, where the boiling points of the first 4% and 5% component are 228.2 and 244.4 °C, respectively, the boiling point of these light end volatiles should be close to 230 °C. A similar

3.2. Effect of reaction pressure In the heavy oil upgrading reactions, a liquid product yield above 100% is observed, indicating the incorporation of methane molecules into the product molecules, which is critical for the potential application in the industry of this novel upgrading process. Since methane is introduced in the gas phase, its pressure may influence the extent of methane participation. Therefore, the effect of methane pressure is explored in the small pilot plant test. The study is carried out using Catalyst-3 as the catalyst under pressures of 5 and 10 MPa, respectively, displayed as Entry #6 and #7 in Table 1. The positive effect due to higher methane pressure on the liquid product yield is witnessed. As the pressure increases from 5 to 10 MPa, the liquid product yield increases from 99.5% to 99.8%. On the other

Fig. 2. Stability test of product oil samples from Reaction Entry 1–8. 6

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hand, the viscosity of the obtained product oil is increased from 128 to 142 cP, indicating the incorporation of methane to the product oil molecules to have stronger interaction between the product molecules. This might be because the improved aromatization process facilitated by methane incorporation enhanced the interaction between the phenyl rings of the aromatic species. A higher pressure also benefits the inhibition on coke formation. The coke generating rate is reduced from 0.00145 to 0.00140 h−1 when the pressure is escalated from 5 MPa to 10 MPa (Table 1). The olefin content in the product oil is also related to the methane pressure. At 5 MPa, the olefin content in the product oil is 1.68 wt% (equivalent to 1-decene). When the reaction is operated at 10 MPa, the olefin content of the obtained product drops to 1.42 wt% (equivalent to 1-decene). A plausible explanation is that the incorporation of methane to the product oil molecules would saturate the unsaturated bonds. It is also possible that the presence of methane facilitates the conversion of unsaturated intermediates towards aromatics in the upgrading process. Both of the possibilities may contribute to the increased liquid yield and average molecular weight at higher pressures, confirmed by the product characterization results (Table 4). Besides olefin content, the product oil quality is also improved in terms of TAN value. After the reaction at 5 MPa, the TAN of the product oil is 0.98 mg KOH/g. When the reaction is executed at a higher pressure of 10 MPa, the TAN of the obtained product drops to 0.85 mg KOH/g. This phenomenon shows that a higher methane pressure promotes the conversion of acid groups in the heavy oil molecules by reacting with methane. These changes show the participation of methane improves the quality of the product oil, which is further enhanced by a higher methane pressure.

3.4. Effect of catalyst regeneration In the practical application of catalytic heavy oil upgrading under methane environment proposed in this work, catalyst regeneration would be carried out multiple times in the life time of the catalyst to recover its catalytic performance. Therefore, it is critical to have a catalyst that has a stable catalytic capacity upon catalyst regeneration, which is investigated in the present work. To evaluate the catalytic performance of the catalysts upon regeneration, the products from catalyst used for the first, second and third time, i.e., Catalyst-1, Catalyst-2 and Catalyst-3, are compared. The catalysts used for the second and third time were regenerated prior to the reaction. The reactions are conducted at 410 °C and 5 MPa. By comparing the heavy oil upgrading products from these reactions, it is observed that the product from Catalyst-2 and Catalyst-3 demonstrates lower viscosities, average molecular weight values, TAN values and asphaltene contents, while the densities and solid contents are similar (Table 3). The viscosity is reduced from 976 cP after reaction using Catalyst-1 to 803 and 795 cP using Catalyst-2 and Catalyst-3 as the catalyst. A similar trend is observed when comparing the average molecular weight of the product oil samples. When Catalyst-1 is employed as the catalyst, the average molecular weight of the product oil is 203 g/mol. When the regenerated catalysts, Catalyst-2 and Catalyst-3, are charged in the reactor, the average molecular weight values of the product oil samples drop to 197 and 195 g/mol. The TAN value is reduced from 1.38 to 1.25 and 1.21 mg KOH/g, while the asphaltene content is reduced form 4.8% to 4.3% and 4.2%. The density change is from 0.9602 to 0.9588 and 0.9590 g/cm3. These phenomena suggest that the catalytic performance of the catalyst is improved after regeneration for the first time, from Catalyst-1 to Catalyst-2. However, the change is minimal after the second regeneration. The parameters such as viscosity, TAN, olefin content, asphaltene content and average molecular weight are similar in the two product oil samples obtained over Catalyst-2 and Catalyst-3, indicating that the change of the catalyst is not significant between these catalysts. One plausible explanation is that some metal species from heavy oil might deposit on some sites in the catalyst during the reaction. These metal species may act as additional active catalytic sites in the regenerated catalyst, resulting in further improved catalytic performance. By comparing the porous and acidity properties of Catalyst-1 and Catalyst-2 (Table 6), it is noticed that the surface areas are similar after the regeneration, while the acidity is enhanced. The increased number of the acid sites is attributed to the metal species deposited on the catalyst surface during the reaction. This result suggest that some metal sites originating from the heavy crude deposit on the catalyst surface during the reaction, creating more acid sites after the catalyst regeneration. These sites may occupy the inner pores of the zeolite structure, which is evidenced by the increased microporous surface area. This phenomenon is also supported by the suppressed XRD peaks due to (0 1 1), (2 0 0) and (0 2 0) planes (Fig. S2), which are closely related to the inner pores of the zeolite. After the second round of reaction and regeneration, the sites in the catalyst that could be occupied by the metal species are almost taken. As a consequence, no more metal species can deposit on the catalyst to further change the catalytic activity of the catalyst. The catalytic performance of the heavy oil upgrading is thus stable after that. Another process that may induce this

3.3. Effect of WHSV Most of the reactions are carried out at a WHSV of 1 h−1. To explore the effect of WHSV, which may benefit the application of this innovative technology at larger scales, an increased WHSV of 2 h−1 is employed during the upgrading process. The temperature and pressure are maintained at 430 °C and 5 MPa, respectively. First, it is noticed that the liquid product yield from heavy oil upgrading is reduced from 99.5% to 99.1% (Table 1). At a larger WHSV, the mass of liquid feedstock in a unit time is increased. The effect of methane incorporation to the product oil that takes place over the catalytic sites of the charged catalyst on the product yield is not as significant due to the larger denominator and thus the reduced residence time. On the other hand, the cracking of heavy oil to form gas products is less dependent on the catalytic sites. As a consequence, the product yield is lower at a larger WHSV. Similarly, the TAN, olefin and asphaltene contents in the product oil obtained at a larger WHSV are higher (Table 4). TAN is increased from 0.98 to 1.02 mg KOH/g at WHSV = 2 h−1. At the same time, olefin content is increased from 1.68 to 1.79 wt% (equivalent to 1-decene wt %), while asphaltene content is increased from 3.1 wt% to 3.4 wt%. The conversion of acid groups, saturation of olefins and conversion of asphaltene molecules are closely related to the catalytic sites and the participation of methane. More feedstock molecules would compete over the catalytic sites at a larger WHSV. Therefore, the percentages of acidic functional groups, olefin and asphaltene species converted during the upgrading are lower, even though the amount of the converted species may actually be larger. Due to the suppressed conversion of large molecules including asphaltene over the catalyst, the viscosity and the average molecular weight of the obtained product are slightly higher comparing with the product from WHSV of 1 h−1 (Table 3). In summary, a higher WHSV results in slightly compromised performance of the catalytic upgrading of heavy oil under methane. The WHSV in potential industrial application should be further optimized to achieve desired upgrading performance at larger WHSV.

Table 6 The porous and acidity properties of catalysts. Samples

Catalyst-1 Catalyst-2

7

BET surface area (m2/g)

Acidity (μmol NH3/g)

External

Microporous

Total

135 134

208 213

343 347

427 436

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Fig. 5. The TAN of product oil samples from different reaction time.

Fig. 3. The viscosity of product oil samples at 15.6 °C from different reaction time.

after reaction at 410 °C and 5 MPa CH4 (Fig. S3b), it is noticed that the size of nanoparticles is mainly around 3 nm. These nanoparticles remain on the spent catalyst, suggesting a stable structure of the catalyst surface. The change on the morphology due to coking is not significant and the zeolite fringes are visible, which may be due to the low coking formation on the catalyst. The reaction Entry 6 was further continued for another 3 days, to a total reaction time of 10 days, i.e., 240 h. But no significant change in terms of the aforementioned product properties is observed. Based on these observations, it is concluded that the reaction time does not impose a significant impact on the catalytic performance, which confirms the good stability of the catalyst in the reaction. In other words, the product quality is stable at different reaction time, which is critical for the long term reactions in practical applications.

phenomenon is the sulfation of the metal species of the catalyst. It has been noticed that the presence of metal sulfide often facilitates hydrogen incorporation to the product molecules [5]. The sulfur content is about 4 wt% in the feedstock, which may react with the metal species during the heavy oil upgrading to form metal sulfide [34]. As a consequence, the catalytic performance is improved as the reaction proceeds. 3.5. Effect of reaction time To explore the effect of reaction time on the reaction process, the samples from two upgrading reactions using Catalyst-3 as the catalyst at 5 MPa are collected for each 12 h. The reaction temperatures are 400 and 430 °C in these two reactions. The samples are thoroughly analyzed to demonstrate the upgrading performance at different reaction time. The viscosity, density and TAN are shown in Figs. 3–5. It is clearly observed that these parameters are stable despite some fluctuations. Other parameters such as average molecular weight, asphaltene content and olefin content (not shown) also follow the same trend. To analyze the morphology of the surface, we have collected the TEM images of the catalysts. Since the extruded catalyst is too thick for TEM, the catalyst in powder form was analyzed instead. By comparing the metal nanoparticles on the fresh catalyst (Fig. S3a) and those on the spent catalyst

4. Conclusions The catalytic upgrading of heavy oil feedstocks under methane is studied. The effects of various reaction parameters such as temperature, pressure, WHSV, catalyst regeneration and reaction time are evaluated by carrying out upgrading reactions under a series of conditions. The products are thoroughly characterized to obtain the insights of the reaction process and the changes due to different reaction conditions, which would equip us with crucial knowledge for the potential application at larger scales. The viscosity, TAN and average molecular weight of the heavy oil samples are reduced with a low olefin content after the catalytic upgrading. The asphaltene content, SARA and simulated distillations analysis results illustrate that resin and asphaltene components are converted to small molecules such as ones in gasoline fraction, benzene, toluene and xylenes. After the upgrading, the gasoline and diesel fractions are significantly increased as evidenced by the simulated distillation curves of the feedstock and the product oil samples. This phenomenon is also confirmed by the escalated percentage of light end volatiles in the product oil samples. Outstanding stabilities of the product oil samples are witnessed even after reaction at higher temperatures, evidenced by the stable oil properties of the samples obtained at different reaction time. In this study, it is observed that a higher temperature benefits the oil quality improvement while maintaining the high liquid yield and low coke formation rate. A higher pressure facilities the participation of methane in heavy oil upgrading and its incorporation to the product molecules. A higher WHSV results in slightly compromised performance of the catalytic upgrading under methane, indicating that this parameter should be optimized in the future commercial application. The positive effect of catalyst regeneration is still observed between the first and second regeneration of the catalysts, which becomes minimal after regenerating twice,

Fig. 4. The density of product oil samples at 15.6 °C from different reaction time. 8

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resulting in performance changes that are not significant thereafter. These results would provide us with valuable guidance in the potential applications of this novel upgrading technology at larger scales.

[17]

Acknowledgement

[18]

We gratefully acknowledge the financial supports from PetroChina Guangxi Tiandong Petrochemical Co., Ltd. [19]

Appendix A. Supplementary data

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Supplementary data to this article can be found online at https:// doi.org/10.1016/j.fuel.2019.116161.

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