Optimization of the CAPRI Process for Heavy Oil Upgrading: Effect of

Article pubs.acs.org/IECR

Optimization of the CAPRI Process for Heavy Oil Upgrading: Effect of Hydrogen and Guard Bed Abarasi Hart,† Amjad Shah,† Gary Leeke,† Malcolm Greaves,‡ and Joseph Wood*,† †

School of Chemical Engineering, University of Birmingham, Edgbaston, Birmingham, B15 2TT, U.K. IOR Research Group, Department of Chemical Engineering, University of Bath, BA2 7AY, U.K.

‡

S Supporting Information *

ABSTRACT: Toe-to-heel air injection (THAI) and its catalytic version CAPRI are relatively new technologies for the recovery and upgrade of heavy oil and bitumen. The technologies combine horizontal production well, in situ combustion, and catalytic cracking to convert heavy feedstock into light oil down-hole. The deposition of asphaltenes, coke, and metals can drastically deactivate the catalyst in the CAPRI process. A fixed bed microreactor was used to experimentally simulate the conditions in the catalyst zone of the oil well of CAPRI. In this study, oil upgrading and catalyst deactivation in the CAPRI process were investigated in the temperature range of 350−425 °C, pressure of 20 barg and residence time of 9.2 min. Additionally, a guard bed consisting of activated carbon particles prior to the active catalyst in a microreactor and/or the addition of hydrogen to the gas feed were used to minimize coke formation and catalyst deactivation through respectively removing and hydrocracking the coke precursors. It was found that depending on the upgrading temperature, the viscosity of the produced oil reduced significantly by 42−82% and (American Petroleum Institute) API gravity increased by ∼2 to 7 °API relative to the feedstock of 0.49 Pa·s and 13 °API, respectively. Conversely, the use of hydrogen further increased the API gravity by 2 °API and the viscosity by 5.3%. Notably, the coke content of the catalyst reduced from 57.3 wt % in nitrogen to 34.8 wt % in hydrogen atmosphere. The use of a guard bed increased the API gravity of the produced oil by a further 2° and reduced the viscosity by an average of 8.5% further than achieved with the active HDS catalyst CoMo/alumina.

1. INTRODUCTION Crude oil is currently one of the primary sources of energy globally. As conventional light crude oil production has arguably reached its peak and its production begins to decline, attention has shifted to vast deposits of unconventional oil resources, that is, heavy oil, bitumen, and shale oil, etc. to meet the ever rising energy demand.1 Heavy oil and bitumen are dense, viscous, and difficult and costly to extract, produce, and refine. Conversely, heavy oil and bitumen account for about 70% of the world’s total 9−13 trillion barrel oil resource.2,3 Therefore, the viability of these resources is dependent on recovery and upgrading technology that will convert them to light oil in an economical and environmentally friendly manner. Since heavy oils are often produced in remote areas of the world, transporting the produced oil is challenging, costly, and energy demanding, usually being accomplished by pipeline heating and solvent dilution. However, heavy oil and bitumen cannot be refined by present refineries without upgrading processes to convert them to synthetic light crude oil first so as to meet refinery feedstock specification.4 In the view of this, the major cost associated with heavy oil and bitumen exploitation is the additional cost incurred for the upgrading facility. This is one reason for their lesser exploitation in the past. Since 2005, rising oil prices have made the recovery and upgrading of heavy oils more economic compared to conventional fuels. This has greatly increased investment in the production of extra heavy oil and natural bitumen to supplement conventional oil supplies, raising the production levels to more than 1.6 mb·d−1 or just under 2% of world crude oil production.5 In 2009 this increased to 2.3 mb·d−1 or less than 3% of the world © 2013 American Chemical Society

demand and is projected to meet about 10% of the world crude oil demand in 2035.6 To achieve these projections innovation is needed in technology to increase recovery levels and increase quality of the recovered oil by upgrading. Steam based technologies such as cyclic steam stimulation (CSS) and steam assisted gravity drainage (SAGD) have been the most successful and commercialized techniques for heavy oil and bitumen recovery and upgrading. These processes rely on reducing the heavy crude oil and bitumen viscosity by heating the oil to improve oil flow from the reservoir to the production well.7,8 However, limitations imposed due to geology of the oil reservoirs1 and heavy environmental footprint, where 2−10 barrel of water has to be injected as steam for every barrel of oil produced (SAGD),9 are the potential disadvantages. Moreover, the produced oil from the aforementioned technologies needs further upgrading and expensive diluent in large quantities for transportation to refineries. To reduce costs and environmental footprint, a high percentage recovery of oil in place and upgrading at the well head or even in situ within the well are desirable. This need is further augmented by the fact that surface upgrading processes are expensive to set up and energy intensive to run with extensive emission and negative environmental impact.10 Special Issue: NASCRE 3 Received: Revised: Accepted: Published: 15394

March 1, 2013 April 24, 2013 April 24, 2013 April 24, 2013 dx.doi.org/10.1021/ie400661x | Ind. Eng. Chem. Res. 2013, 52, 15394−15406

Industrial & Engineering Chemistry Research

Article

THAI-CAPRI is relatively a new technology simultaneously incorporating down-hole in situ catalytic upgrading with thermally enhanced oil recovery. THAI integrates in situ combustion with advanced horizontal well concepts, whereby a small fraction of the reservoir oil is burnt to mobilize the heavy oil.11 Thermal cracked heavy oil produced during the process not only aids oil recovery but also has the added benefit of upgrading the oil.10 CAPRI is the catalytic extension to THAI developed in collaboration with the Petroleum Recovery Institute (PRI), where the objective is to achieve further heavy oil upgrading in situ by placing an annular layer of catalyst around the perforated horizontal producer well to create a downhole catalytic reactor.12,13 The thermal-cracking reactions of THAI, taking place ahead of the combustion zone, that is, the coke zone, creates the precursor for CAPRI. Oil upgrading is thought to occur by a combination of carbon-rejection (thermal cracking) and hydrogen addition reactions at the surface of a hydroconversion (HCT) or hydrotreating (HDT) catalyst.12 From laboratory scale experiments of the THAICAPRI process Xia et al.10 reported recovery levels of 79% of the original oil in place (OOIP), API upgrading of 23 degrees and viscosity reduction of as low as 20−30 mPa.s using Lloydminster heavy crude oil as the feed and CoMo/alumina as the catalyst. Detailed studies of the technology have been reported elsewhere.12,14,15 Subsequently the CAPRI process was further investigated using a set of microreactors to replicate underground upgrading conditions and to optimize catalyst type, oil and gas flow rate, temperature, and pressure.14 Shah et al.16 pointed out a number of shortcomings of the technology, such as asphaltenes, coke, and metal deposition drastically deactivates typical refinery HDS and HDM catalysts when extra heavy oil/bitumen from Athabasca is used. In the current study, as an investigation to overcome these limitations of the THAI-CAPRI technique, an HDS CoMo/alumina was used along with a number of modifications to the catalytic bed and reaction media, such as the use of a guard bed and hydrogen atmosphere. A guard bed involves packing inert porous particles above the catalyst in order to adsorb or filter coke precursors from the feed before they reach the catalyst. The addition of hydrogen to the feed was intended to augment catalytic hydroconversion and hydrocracking reactions in order to achieve a higher level of upgrading than would be achieved with inert gasesalbeit at low partial pressure. The source of hydrogen is expected to be gasification and/or water gas shift reactions.17 A number of temperature regions have been reported in a THAI-CAPRI reservoir18 and to investigate whether catalytic activity can be maintained at lower temperatures compared to the previously optimized temperature of 425 °C,14 experiments were conducted at lower temperatures in the current study.

Table 1. Properties of the THAI Feed Oil property

value

density at 20 °C (g·cm−3) viscosity at 20 °C (mPa·s) API gravity (deg)

0.9801 ± 0.002 490 ∼13

2.2. Experimental Apparatus. In the THAI process air (and initially steam for raising temperature) is injected and once the combustion is started the oil flows downward in the mobile oil zone (MOZ) into the perforated horizontal production well (Figure 1). This is the area where most of

Figure 1. Schematic of the THAI-CAPRI process11.

the thermal cracking is believed to be happening. Hence the process moves from the “Toe” position to the “Heel”. The oil in the MOZ remains hot upon reaching the horizontal production well due to heat gained from the combustion zone. Moreover the horizontal production well is always kept hot to keep the oil flowing. In the CAPRI process, the production well is packed with an annular catalyst layer. As the hot oil passes through the annular catalytic layer, catalytic upgrading is believed to happen. From Figure 1 a down-flow microreactor was used in the laboratory representing a cylindrical core of 1 cm taken in a radial direction through the annular layer of the catalyst surrounding the producer well as represented in Figure 2.

2. EXPERIMENTAL DETAILS 2.1. Feedstock and Catalysts. The heavy crude oil was supplied by Petrobank Energy and Resources Ltd. from its Whitesands THAI pilot trial at Christina Lake, Alberta, Canada. The properties of the feedstock are presented in Table 1. A hydrotreating catalyst CoMo/alumina of quadra-lobed shaped extrudate (Akzo) was used in this study. The composition and microstructural properties of the used catalysts such as specific surface area, pore volume, and pore diameter were determined by Brunauer−Emmett−Teller (BET) technique and are presented in the Supporting Information, Table 1S.

Figure 2. Scaling the field CAPRI section to laboratory model representation with catalyst bed volume of 12.86 cm3.

A flow diagram of the laboratory scale experimental rig used in this study is shown in Figure 3. The experimental setup was built and commissioned at the School of Chemical Engineering, University of Birmingham, UK. The dimensions of the reactor are an inner diameter of 1 cm and a length of 41 cm. A downflow microreactor was used to ensure the complete wetting of 15395

dx.doi.org/10.1021/ie400661x | Ind. Eng. Chem. Res. 2013, 52, 15394−15406


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Figure 3. Schematic diagram of the CAPRI experimental setup.16

the catalytic bed with the help of gravity. The experiment was initiated by turning on the furnace and the trace heating and setting the controls to the desired temperatures and pressures. Once the operating conditions of temperatures and pressures were achieved in the reactor, the oil flow metering valve was opened manually to initiate the oil flow from a feed tank pressurized by nitrogen gas. The flow metering valve was adjusted until the desired flow rate of 1 mL·min−1 was achieved. The pressure in the feed tank was kept 5−10 barg higher than the packed microreactor inside the furnace to ensure oil flow through the reactor. The THAI feed oil (i.e., heavy crude oil) was delivered to the packed catalyst bed in down-flow mode to ensure flow by gravity and complete wetting of the catalytic bed. The THAI feed oil was fed into traced heated lines having a set temperature of 280 °C and into the furnace after achieving the desired experimental temperature. The furnace provides isothermal conditions along the active section of microreactor. Two gases, N2 or H2 (THAI gas: 80% N2, 13−14% CO2, 3% CO, 4% CH4 was used by Shah et al.16 and presented in Figure 3 for reference purposes only), were used as the reaction media to simulate the combustion gases expected in a real THAICAPRI reservoir. These gases were mixed with the THAI feed oil in the mixing chamber and the gas−oil mixture passed through the reactor in concurrent flow. The gas−oil mixture of the partially vaporized THAI feed oil flowed downward through the voids of the packed catalyst bed where it underwent cracking reactions aided by heat consumption. The product stream coming out of the reactor passed through a back pressure regulator (Swagelok Co. UK), which regulated and maintained a constant pressure of 20 barg in the reactor. The upgraded products which include light oil and gases were passed to the gas−liquid separator, where the light oil was collected, while the gaseous products were flashed off, and either vented or sent to a refinery gas analyzer (RGA) for concentration and compositional analysis by gas chromatography. The analyses were performed periodically during each experimental run. The light oil sample was drained from the gas−oil separator initially every 20 min for the first hour, every 30 min for 4 h, and thereafter every 40 min, and the collected oil was analyzed using techniques listed in section 2.3. The

experimental conditions used in this study are listed in Table 2. A hydrotreating catalyst CoMo/alumina of quadra-lobed Table 2. Operating Conditions in the Experiments feed flow rate (mL·min−1) catalyst inventory (g) pressure (barg) reaction temperature (°C) WHSV (h−1) gas-to-oil ratio (mL·mL −1)

1 6 20 350−425 9 200−500

shaped extrudate (Akzo) was used in this study. The composition and microstructural properties of the used catalysts such as specific surface area, pore volume, and pore diameter were determined by Brunauer−Emmett−Teller (BET) technique and are presented in Supporting Information, Table 1S. 2.3. Product Analysis. Density of the feed and produced oils were measured using an Anton Parr DMA 35 portable density meter (Anton Paar GmbH, Austria) at 15 °C and reported in kg·m−3. API gravity was calculated using eq 1, 141.5 API gravity = − 131.5 (1) SG Where SG represents specific gravity The Bohlin CVO 50 NF rheometer (Malvern Instruments Ltd., United Kingdom) was used to measure the viscosity of the THAI feed oil and produced oils. All viscosity measurements were performed at 20 ± 0.1 °C. Aluminum parallel plate geometry was used. The diameter of the used plate is 40 mm and it was made of aluminum material with a polished surface. The parallel plate gap size was set at 150 μm and a shear rate of 100 s−1 was used. ASTM-D2887 provides a comprehensive boiling range distribution of carbon numbers of petroleum and its distillates. For this reason simulated distillation (SIMDIS) based on an Agilent 6850N GC and calibrated in accordance with the ASTM-D2887 was used to characterize the feed and produced oils. Agilent 6850N Network gas chromatograph system J&W 125-10 is fitted with a DB-1 10 m length, 530 μm ID, and 2.65 15396



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μm film thickness capillary column. Prior to the injection of the sample, the feed and produced oil samples were diluted with carbon disulfide (CS2) in a ratio of 1 to 10. The gas products were analyzed using an Agilent 7890A Refinery Gas Analyzer (RGA) to determine the volume percentage of H2, CO, CO2, and C1−C5 hydrocarbons. The RGA three channels include flame ionization detector (FID) and two thermal conductivity detectors (TCDs). The light hydrocarbon components in the gas stream are determined by the FID channel column HP-PLOT Al2O3S capable of separating C1 to nC5 including their 22 isomers based on the calibrated table, while components heavier than nC6 are back flushed through the precolumn. One of the TCD with helium carrier gas is used for permanent gases analysis such as N2, CO, CO2, O2, and hydrocarbons to nC5, C6+. The other TCD with nitrogen as carrier gas determines gases like hydrogen and helium in the gas stream. The oven column dimension is 27 m × 320 μm × 8 μm at a temperature of 200 °C. The RGA takes 15 min to complete the analysis of one gas sample. A thermogravimetric analyzer (TGA) was used to determine the amount of coke deposit on the spent catalysts. In this study, TGA was carried out with NETZSCH-Geratebau GmbH, TG 209 F1 Iris. A 10 mg sample of the spent catalyst was recovered from the reactor and placed on an alumina crucible above the microbalance. The microfurnace is programmed as follows: linear increase in temperature in range of 25 °C to 1000 °C and a heating rate of 20 °C·min−1. At 1000 °C an isothermal condition was maintained for 20 min to enable total burnoff of the materials deposited on the spent catalysts. The total time for each run is 49 min and the air flow rate during the TG analysis was set at 50 mL·min−1.

Table 3. Mass Balances; Catalyst, CoMo/Alumina; Reaction Media, N2; Pressure, 20 barg; Oil Flow Rate, 1 mL·min−1; Gas Flow Rate, 500 mL·mL−1 gas (wt %)

liquid (wt %)

coke (wt %)

350 400 425 mean standard deviation

1.96 2.96 4.34 0.24

97.4 95.87 93.8 1.20

0.64 1.17 1.86 0.91

at the higher temperature can be attributed to increased catalytic cracking reactions with temperature rise as Krumm et al.19 reported similar trends in the distribution of gas, liquid, and coke during the catalytic cracking of heavy oil fraction. 3.1.2. Effect of Temperature on API Gravity. Figure 4 shows the API gravity of the produced oils at different processing

3. RESULTS AND DISCUSSION 3.1. Effect of Temperature. 3.1.1. Effect of Temperature: Mass Balance. The effect of temperature upon the yield of liquid, gas, and coke deposited upon the used catalyst was studied, together with the extent of upgrading of the produced oil as measured by viscosity, API gravity and simulated distillation analysis. The mass of gas evolved during the upgrading reactions was calculated as the mass of the oil remaining after subtracting the masses of produced liquid and solid deposits in the reactor. The mass balances of the three pseudo-products, that is, liquid, gas, and coke, were calculated as percentage of the mass of THAI feed oil into the system using eqs 2 and 3: wi yield (wt %) = 100 wFeed (2)

Figure 4. API gravity of CAPRI produced oil at 350, 400, and 425 °C: catalyst, CoMo/alumina; reaction media, N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 500 mL·mL−1.

temperatures. It is clear that the API gravity of the produced oil increased as the reaction temperature increases from 350 to 425 °C. During the first 100 min, significant API gravity increase can be observed at all of the three reaction temperatures and reaches a maximum value of 6.5° at 425 °C. This most likely occurs because during the earlier part of the reaction, the catalyst is still active and has not deactivated. At lower temperatures of 350 and 400 °C, an initial API increase of 5−6° can be observed compared to the THAI feed oil. At higher temperatures in the plateau region the level of upgrading also seems to depend on temperature, with about 1.7° increase at 350 °C and increasing up to ∼4−5 °API upgrading at 425 °C. These trends in API gravity were expected as reported earlier by Shah et al.16 that 425 °C was the optimum upgrading temperature for CAPRI, and further increase in temperature leads to increased coke deposition which required the reactor to be shut down due to blockages of the reactor bed with coke. From the mass balance in Table 4, it can be observed that more liquid production occurs at lower temperature; however, the level of upgrading at the lowest temperature of 350 °C is not as significant as at the previously optimized temperature of 425 °C, suggesting a trade-off between liquid production and quality. These results however, establish the fact that even at lower temperatures of 350 °C the API gravity increase of 1.7°

gas (wt %) = 100 − liquid yield (wt %) − coke yield (wt %)

temperature (°C)

(3)

Where wi is the weight of component i and wFeed is the overall weight of the THAI feed oil. Table 3 displays the mass balance of gas, liquid (light oil), and coke, from which it can be observed that the lower temperature favors lesser production of gases and coke and more liquid products, suggesting a low degree of thermal and catalytic upgrading. The amount of coke produced at 350 °C was 0.64 wt % compared to 1.86 wt % at 425 °C, and the corresponding measured liquid yields were 97.4 and 93.8 wt %, respectively. The yield of gases was 1.96 wt % at 350 °C and 4.34 wt % at 425 °C. The increased production of gas and coke 15397



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substantial viscosity reduction of the upgraded oil can be observed compared to the THAI feed oil, with larger viscosity reduction at higher temperatures. At reaction temperatures of 350 and 400 °C the average viscosity of the produced oil reduced to 0.28 and 0.17 Pa.s, respectively, from the base value of 0.49 Pa·s for the THAI feed oil. These are about 1.7 and 3 times lower than the THAI feed oil viscosity. However higher average viscosity reduction from 0.49 Pa.s to 0.09 Pa.s was obtained at 425 °C, which is approximately 5 times lower than that of the THAI feed oil. The level of viscosity reduction, which was up to 81% for the case of 425 °C in the presence of CoMo/alumina catalyst in the presence of nitrogen, represents a significant step forward for the recovery and upgrading of heavy oil/bitumen downhole. This would reduce the difficulties occurring during heavy oil production, transportation, and surface processing as a result of upgrading of the oil can be achieved. Moreover, capital and operating costs of a surface upgrader can reach around hundreds of millions dollars, but a downhole or in situ oil upgrading process may reach the sole cost of buying and installing the catalyst, which may deliver a relatively short payback time.10 Commercial catalysts are fairly cheap. Shah et al.1 estimated that 20 tonnes of new HDS catalyst for 500 m horizontal producer well cost about $60− 100k. Petrobank predict a flow rate per horizontal well of 90 m3 oil·day−1, of about 800 barrels·day−1. Allowing for the lower API gravity of CAPRI produced crude oil (say 22°), then at $50/barrel, the catalyst cost represents only about 11/2 to 21/2 days production. Early estimates suggest that the THAI/CAPRI process costs about one-third per producing barrel of the equivalent SAGD process.22 Saturates, aromatics, resins, and asphaltenes (SARA) analyses were carried out and reported by Shah et al.16 for experiments performed in the same reactor and under the same conditions as reported in Table 2 of this paper. They found that saturates, aromatics, resins, and asphaltene contents were 15.38, 57.04, 20.18, and 7.4% for the THAI feed and typical values for the produced oil at 425 °C were 16.37, 67.62, 9.11, and 6.9%, respectively. The amount of aromatics largely occurred at the expense of naphthenes and asphaltenes. Shah et al.16 concluded that the rise in API and viscosity was due largely to hydroconversion rather than HDS, HDM, or HDA activity. 3.1.4. Effect of Temperature on Boiling Point Distribution of Oils. Table 4 provides the results of the cumulative product percentage yield of feed and produced oils at selected simulated distillation temperatures. It can be observed that temperature rise in the upgrading experiment favors a significant shift toward lighter distillate fractions in comparison to the partially upgraded THAI feed oil. A dramatic 38 °C shift in the boiling point range at 75 wt % yield can be observed for 350 °C in comparison to THAI feed oil. This rises to 56 and 48 °C at 75 wt % yield for the higher process temperatures of 400 and 425 °C, respectively. An important observation can be made that at 400 and 425 °C the shift toward lower distillable temperatures is almost identical. The level of upgrading in the presence of CoMo/alumina catalyst is also superior to thermal upgrading in the presence of glass beads by 34 °C at 75 wt % cumulative product percentage yield. 3.1.5. Effect of Temperature on Thermal Decomposition Behavior of the Spent Catalyst. Figure 6 provides the thermogravimetric analysis (TGA) thermograms or weight loss curves as a function of ramp temperature increase for the CoMo/alumina recovered catalyst after the upgrading experiments. Each catalyst sample was removed for analysis at the end

Table 4. SIMDIS for the Feed and Upgraded Oils at 350, 400, and 425°C: Catalyst, CoMo/Alumina; Reaction Media, N2; Pressure, 20 barg; Oil Flow Rate, 1 mL·min−1; Gas Flow Rate, 500 mL·mL−1 cumulative product percent yield 15% THAI feed oil glass beads at 425 CoMo/alumina at CoMo/alumina at CoMo/alumina at

°C 350 °C 400 °C 425 °C

218 190 207 183 183

30% 296 258 265 233 233

45%

60%

temperature °C 360 402 321 379 313 358 284 334 285 339

75%

90%

438 424 400 382 390

475 471 455 442 442

shows some measurable improvement for the catalyst packing around the horizontal production well. In industry generally sulfides of Co, Ni, W, and Mo supported catalysts, having a variety of pore structures and active metal dispersion (active sites), are used for petroleum residue hydroprocessing. The most important property for a residue hydroconversion catalyst is pore diameter, because feedstock contains large molecules of asphaltene, metal chelates.20 Broadly heavy crude is typically processed in trickle-bed reactors at temperatures of 350−450 °C and pressures of 5−15 MPa of hydrogen.20 However, since no hydrogen was used in the results presented, the predominant reactions are assumed to be carbon rejection or limited hydroconversion. In this case hydrogen is assumed to be generated from side chain breaking and/or as a result of a number of complex polymerization reactions resulting in the eventual formation of condensed polyaromatic species in a hydrogen deficient environment and releasing hydrogen.21 The increase in API gravity as observed from Figure 4 leads to improved quality of the produced oil as Greaves and Xia13 noted that with a 7.9 °API increase only about 15% of diluent was needed to meet pipeline specifications, compared to 30− 50% required for nonupgraded bitumen produced from SAGD and CSS operations. 3.1.3. Effect of Temperature on Viscosity. The viscosity of the produced oil as a function of time-on-stream at the reaction temperatures is presented in Figure 5. From Figure 5

Figure 5. Viscosity of CAPRI produced oil at 350, 400, and 425 °C: catalyst, CoMo/alumina; reaction media, N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 500 mL·mL−1. 15398



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explain that with an increase in temperature from 350 to 425 °C the activity of the catalyst increased and hence the level of hydroconversion. This resulted in higher upgraded oil in terms of API gravity and viscosity as evident from Figures 4 and 5 and also the boiling point distribution shift toward lower boiling point hydrocarbons (Table 4) as reaction temperature increased. Zhao30 and Meng et al.25 reported that at reaction temperatures of 400 °C and above, the cleavage of C−C bonds and the rate of cracking of the larger molecular weight components becomes dominant, and subsequently promote their condensation reactions leading to increase coke content and coke yield. Although, higher reaction temperature (i.e., 425 °C) led to improved viscosity (and API gravity; section 3.1.4) of the produced upgraded oil, it is at the expense of more coke formation. 3.1.6. Effect of Temperature on Gas Products Distribution. Supporting Information, Table 2S provides quantitative analysis of the gas composition produced during the course of the experiment. Average values of the gases have been reported in Table 2S for the course of the whole experiment as gas composition varied during different stages of the experiments. From Table 2S it can be observed that trace gases from C1 to C5 (varying in structural composition) are produced for all the three experiments namely glass beads at 425 °C and catalytic experiments with CoMo/alumina conducted at 400 and 425 °C. From the table it is clear that thermal cracking (over glass beads) produces more of certain gas components such as methane, propane, n-butane, trans-2-butane, and cis-2-butene, and n-pentane compared with catalytic cracking; however, some other components such as propene, i-butane, and 1-butene were not produced over glass beads that were present over the catalyst. The amount of gases also rises with temperature in the catalytic runs. The combined value of the hydrocarbons produced was 1.2 for glass beads at 425 °C, 0.84 and 0.94 vol % for catalytic runs at respective temperatures of 400 and 425 °C. The levels of CO and CO2 were also identical in the two catalytic runs. The only exception is the produced hydrogen in the three different experiments, which rises with temperature and in the presence of catalyst. These trends are in conformity with trends reported elsewhere, where gasification has largely been attributed to thermal cracking.19 3.1.7. Effect of Temperature: Thermal vs Catalytic Cracking. To investigate the effect of temperature a controlled experiment was conducted in the presence of glass beads. Average API gravities at 350, 400, and 425 °C were 13.8, 13.8, and 14.8 Pa·s, respectively. This represented an increase of 1.8° at 425 °C; however only 0.8° at the lower temperatures of 400 and 350 °C. This compares less favorably with the catalytic upgrading (Figure 4) where an average upgrading of 3.8, 2.5, and 1.7° occurred at 425, 400, and 350 °C, respectively. Average viscosity was measured as 0.22, 0.29, and 0.37 Pa·s over glass beads compared to the catalytic runs viscosities of 0.09, 0.18, and 0.24 Pa·s at 425, 400, and 350 °C temperatures, respectively. This represented a 2.4, 1.6, and 1.5 times higher viscosities reduction for the catalytic experiments compared to the glass beads, establishing the superiority of catalytic cracking compared to thermal cracking alone. SIMDIS values for glass beads from Table 4 further confirm the low level of upgrading by the boiling points being 34 °C lower than the comparable catalytic experiment for 75 wt % product yield. The level of upgrading seen with glass beads may also have been complemented with the hydrogen produced during the process

Figure 6. TGA and DTG of spent CoMo/alumina catalyst under reaction temperatures of 350, 400, and 425 °C: catalyst, CoMo/ alumina; reaction media, N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 500 mL·mL−1.

of the experiment. The ramp temperature increase enables the interpretation of the different chemical changes occurring during the burnoff when compared to the oil TGA and differential thermogravimetric (DTG) curve. The isothermal temperature of 1000 °C toward the end ensures that all carbon species are completely burnt off during the heating period. However, above 620 °C (see Figure 5), the deposits on the spent catalysts are defined as coke.23 Subsequently, to interpret the different stages during the heating period, the derivative of the weight loss curve (DTG) obtained in the thermal burnoff in air atmosphere is also presented in Figure 6. Figure 6 shows that the weight loss process of the spent catalyst can be divided into several steps; the region from 25 to 208 °C represents loss due to devolatilisation of light oil and beyond 620 °C is coke. Beside catalyst and reactor fouling and clogging, coke build up on the catalyst bed is one of the main routes for deactivation and shortening of lifespan.24 The coke content of the spent CoMo/alumina catalyst increased in the order 48.4, 53.5, and 57.3 wt % at 350, 400, and 425 °C respectively. The large quantity of coke 48−57 wt % of the spent catalyst, may have led to a large loss of surface area. This observed trend is in line with the findings of Meng et al.25 on coking behavior and catalyst deactivation for catalytic pyrolysis of heavy oil. Sanford26 has pointed out that during the earlier stages of coking and hydrocracking reactions in residue conversion, C−C bonds are broken. These reactions reject carbon at the surface of the catalyst pellets similar to the reactions occurring in the THAI-CAPRI process. At higher temperatures of 425 °C and above the coking reactions become more predominant, which accelerate catalyst deactivation and change in selectivity toward undesirable products,27,28 that is, the coke and gaseous species. This is evident from Table 3 where 4.34% of products formed were gases at 425 °C compared to 1.96% at 350 °C, and coke was 1.86 at 425 °C compared to only 0.64% at 350 °C. Similar trends can be observed from the spent catalyst TGA and DTG analysis (Figure 6) where higher temperature favored higher coke formation and lesser liquids. This can be explained by the fact that at higher temperatures bond scission reactions involving the side chains of compounds present in crudes such as alkylaromatics increases and results in the formation of more coke, gases, and lower liquid products.29 This may also 15399



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hydrogen gas flow control issues. During the course of the experiment hydrogen consumption varied significantly and henceforth resulted in variable API gravity upgrading. The increase in API gravity of 5 to 7 °API in the presence of hydrogen compares to only 3 to 6 °API obtained in nitrogen atmosphere. The additional upgrading of 2 °API with hydrogen addition is significant as it may represent a premium of $2−3 for each barrel of oil produced. It also suggests that hydrogen addition to the molecule is able to effect further upgrading than catalytic carbon rejection which is thought to occur with only nitrogen. CoMo/alumina supported on Al2O3 and/or SiO2 is a bifunctional catalyst (i.e., metal and acidic), which is favorable for hydrocracking and hydrogenation.32 The metal sites promote hydrogen addition and the acidic sites of the supports promote cracking reactions through a carbocationic mechanism.21,33 Catalytic upgrading in the presence of an HDS/HDT catalyst, that is, CoMo/alumina, occurs by a number of steps. First breaking the larger molecules to give fragmented free radicals, which then react with hydrogen radicals in order to stabilize the hydrocarbon chains and terminate the reaction. It would appear that in the absence of hydrogen, the active chains keep reacting with each other resulting in the formation of higher molecular weight compounds by polymerization, increased coke formation, and adverse impact of viscosity and API upgrading of the produced oil. This has been illustrated in eqs 4 and 5.34 Moreover, since the molecular size of hydrogen is smaller than the polymer chains, the mass transfer rate is faster leading to rapid termination of active chains.34 During hydrocracking reactions the possible splitting of C−C, C−S, C−N bonds results in free radical hydrocarbon chains. The termination of active chains of these free radicals is an important step which is facilitated by the attachment of hydrogen and results in the reduction of the viscosity of the produced oils during the CAPRI process.

as observed from the hydrogen in the outlet gases in the gas stream and reported in Supporting Information, Table 2S. The assumed hydroconversion (in the absence of any externally added hydrogen) in the presence of CoMo/alumina catalyst is further evidenced by the fact that a significantly lower level of upgrading was achieved in the absence of the catalyst in terms of API gravity or viscosity (Figures 4 and 5) or from boiling point distribution of oil upgraded over only glass beads compared to similar catalytic run (Table 4). 3.2. Effect of Injected Hydrogen. 3.2.1. Effect of Injected Hydrogen on API Gravity. Catalytic upgrading in the presence of hydrogen is a process by which the hydrocarbon molecules of oil are broken into simpler molecules, by splitting of the C− C, C−S, and C−N bonds of high molecular weight species, under the addition of hydrogen at high pressure and in the presence of a catalyst. These broken species are subsequently hydrogenated resulting in their conversion to lower molecular weight hydrocarbons. It has been reported that hydrogen promotes saturation of olefins and aromatics, facilitates the termination of free radical coke precursors formed during the cracking reactions, and increases the hydrogen-to-carbon (H/ C) ratio of the produced oil. These reactions can subsequently remove a heteroatom, thereby suppressing coke yield and catalyst deactivation.31 To investigate this effect in the THAI-CAPRI reactor, experiments were performed using CoMo/alumina catalysts at a temperature of 425 °C, pressure of 20 bar, H2-to-oil ratio of 200 mL·mL−1, and WHSV of 9 h−1. The H2-to-oil ratio used with hydrogen was lower than the experiments presented in the earlier sections, but for the purposes of comparison an additional experiment was carried out in a nitrogen atmosphere with an N2-to-oil ratio of 200 mL·mL−1. The effect of hydrogen-addition on the API gravity upgrading of the produced oil is presented in Figure 7. It is clear that hydrogen-addition results in higher API gravity of the produced oil compared to upgrading in nitrogen as reaction media. From Figure 7 the average increase in API gravity when hydrogen is added to the injected gas is 5 ± 1.1 °API, with a maximum of 7 °API, compared to the THAI feed oil (13 °API). The fluctuations in the data from Figure 7 are largely due to

active H 2 + active chains → active chain termination ( low molecular weight)

(4)

active chain + active chain → active chain termination (high molecular weight)

(5)

The free radicals formed during thermal or catalytic cracking are very unstable and are usually stabilized by hydrogen. However, in a nitrogen atmosphere (as was the case in section 3.1) the scarcity of free hydrogen results in radical linking to form larger molecules and/or double bonded olefins due to collisions of carbon radicals.35 This may explain the relatively lower API upgrading when nitrogen atmosphere was used as opposed to hydrogen (Figure 7). Hydrogen may have been initially adsorbed and dissociated on the metal sites to form reactive hydrogen in this particular experiment via acidic sites (upon which hydrogen gives rise to a proton) and subsequently reacted with the cracking intermediates, free radicals, and olefins that require hydrogenation.36 API upgrading has previously been linked to increased hydrocracking and hydrogenation reactions,37 in agreement with the results and discussion presented here. 3.2.2. Effect of Hydrogen on Viscosity. Figure 8 shows a comparison of the viscosity reduction in the presence of either hydrogen or nitrogen as the gas atmosphere. From Figure 8 average viscosities of 0.06 and 0.09 Pa·s can be observed for the

Figure 7. Effect of hydrogen addition on API gravity of produced oil: catalyst, CoMo/alumina; temperature, 425 °C; reaction media, H2 and N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 200 mL·mL−1. 15400



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Table 5. SIMDIS for the Feed and Upgraded Oils in N2 and H2: Catalyst, CoMo/Alumina; Temperature, 425°C; Reaction Media, H2 and N2; Pressure, 20 barg; Oil Flow Rate, 1 mL·min−1; Gas Flow Rate, 200 mL·mL−1 cumulative product percent yield 15%

30%

45%

60%

75%

90%

438 382 373

475 442 433

temperature (°C) THAI feed oil CoMo/alumina in N2 CoMo/alumina in H2

218 183 171

296 233 220

360 284 269

402 334 320

can be observed that with the addition of hydrogen a further 9 °C shift toward lower temperature distillates can be observed for hydrogen compared to nitrogen at 75 wt % product yield. 3.2.4. Effect of Hydrogen on BET and Thermal Decomposition Behavior of the Spent Catalyst. The nitrogen sorption isotherms of the fresh and spent catalysts were analyzed using BET theory. The analyzed spent catalysts were collected from the center of the fixed-bed reactor. The adsorption−desorption isotherm of the fresh and spent catalyst are compared for the nitrogen and hydrogen atmospheres in the upgrading experiment, as shown in Figure 9a,b. The adsorption−desorption curve reveals a large hysteresis loop in

Figure 8. Effect of hydrogen-addition on viscosity of produced oil: catalyst, CoMo/alumina; temperature, 425 °C; reaction media, H2 and N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 200 mL·mL−1.

upgraded produced oil in hydrogen and nitrogen, respectively, compared to the 0.49 Pa·s for the THAI feed oil. This represents a 6.3% further reduction in the average viscosity reduction on top of the 80.9% reduction for nitrogen only. This makes the combined viscosity reduction equal to 87.2% compared to the viscosity of the THAI partially upgraded oil. This additional degree of upgrading achieved with hydrogen can be attributed to the fast termination of free radicals in the flowing hydrogen atmosphere. The effect is to terminate polymerization of the partially upgraded oil that could otherwise lead to the formation of polyaromatic coke precursors. From the above results it can be deduced that the increased activity in the presence of hydrogen in terms of API increase, viscosity reduction, and decreased amount of coke formation on the surface of CoMo/alumina catalyst is because reaction intermediates have been hydrogenated to form small stable molecules. This agrees with the studies of Liu and Fan34 and Galarrage et al.38 Viscosity reduction (or API gravity upgrading, section 3.2.1) of the produced oils in the presence of hydrogen and an HDS/ HDT catalyst as observed in this work can increase recovery levels and upgrade the quality of the recovered oil if used downhole. Up to 20 mol % hydrogen was observed during BP’s ISC oxygen field pilot at the Marguerite Lake, Alberta.17,39 The source was believed to be via gasification and/or water gas shift reactions. During a catalytic upgrading test on heavy oil in a combustion tube at the University of Calgary, 3% of hydrogen was measured in the produced gas.40 Although only about 2 barg of hydrogen pressure was observed in these previous studies compared to 20 barg of hydrogen pressure used in this study, it is expected that the use of low pressures of hydrogen will have a more prominent effect on the level of upgrading with feeds having less aromatic character. The THAI partially upgraded oil used in this work is relatively rich in aromatics, 57%,16 and aromatic crudes are widely known not to be suitable feeds for catalytic hydroconversion.41 For this reason, CAPRI may be more suitable for less aromatic crudes. 3.2.3. Effect of Hydrogen on Boiling Point Distribution of Oils. Simulated distillation temperature ranges for hydrogen and nitrogen as reaction atmosphere are provided in Table 5. It

Figure 9. Fresh and spent (at 425 °C reaction temperature) CoMo catalyst adsorption−desorption isotherm in the presence of (a) nitrogen and (b) hydrogen: spent catalyst, CoMo/alumina; temperature, 425 °C; reaction media, H2 and N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 200 mL·mL−1. 15401



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the isotherm indicative of the type IV which is mesopores. A large hysteresis loop in the isotherm of the spent catalyst can also be observed indicating deposited materials/coke in the pores and blockage of active sites.42 The significant decrease in surface area and the volume of adsorbed nitrogen for the spent catalyst also indicates that deposits (e.g., carbon, metals, etc.) may have altered the pore structure, textural properties, and activity. A further observation for the spent CoMo/alumina catalyst was that the volume of adsorbed−desorbed N2 at the same relative pressure is significantly lower compared to the fresh CoMo/alumina. This loss in surface area may have resulted from the highly aromatic and coke-forming molecules such as resins and asphaltenes residing within the small pores of the catalysts, due to limited diffusion of reactants and products.43 As a result further cracking of these molecules occurred yielding coke, causing preferential coke deposition, and hence narrowing, of the larger pores and leading to total blockage of the smaller pores. On the other hand, when hydrogen was introduced the loss of surface area by the spent CoMo/alumina catalyst was 78.6% compared to 99.8% in nitrogen atmosphere at the same operating conditions (i.e., 425 °C, 20 bar, and 200 mL·mL−1). This indicates lesser deactivation due to the blockage of catalytic site by coke deposition in the presence of hydrogen (see adsorption−desorption isotherm of fresh and spent CoMo/alumina obtained in the use of H2 in Figure 9a,b). This is because hydrogen helps to terminate some of the coke precursors formed during the CAPRI process. Figure 10 presents the amount of coke formed on the CoMo/alumina catalyst after 1200 min of operation in the

A significant effect of hydrogen in this study was its ability to suppress the formation of coke compared to nitrogen as the flowing gas atmosphere (Figures 9a,b and 10), and as is commonly practiced in the refining industry the phenomenon is known to increase the yield of distillates and its quality.46−48 3.2.5. Effect of Hydrogen on Gas Products Distribution. Supporting Information, Table 3S shows the composition of the produced gases when hydrogen was injected as the reaction atmosphere. The results can be compared with the gas composition for nitrogen atmosphere at the same temperature in Supporting Information, Table 2S. Although the gas flow rate of the experiment in Table 2S for nitrogen was higher than the results for hydrogen in Table 3S, it can be seen that a similar amount of hydrocarbons were produced with the only exception being the level of methane produced. Twice as high methane concentration can be observed with hydrogen compared to the nitrogen atmosphere, making the combined value of hydrocarbon gases produced with hydrogen 1.52 vol % compared to 0.94 vol % with nitrogen. The level of CO2 is 0.14 vol % in the presence of hydrogen compared to 0.05 vol % in the presence of nitrogen, suggesting relative completion of the thermal or catalytic reactions. For the run in hydrogen atmosphere 12.2 ± 3.9 vol % less hydrogen was observed at the outlet compared to the inlet once analyzed with the RGA. The lower hydrogen at the reactor outlet compared with the inlet is due to a degree of dissolution in the liquid oil, and also to a chemical reaction along with some possible system losses.48 Equations of state for the calculation of hydrogen loss have been proposed by Maipur et al.,48 but in this work the outlet gas flow was not measured, so overall loss of hydrogen in moles was not calculated. However in a hydrogen environment the total volume percentage of hydrocarbon gases increased compared with its value in a nitrogen environment, as noted above. This provides evidence that reaction of the oil feed with hydrogen did occur, with hydrogen acting to terminate free radicals and produce light hydrocarbons that were cracked from the oil molecules. This is in conformity with the results reported by Kim et al.49 who measured hydrogen consumption in the presence of an HDN catalyst. This may also explain the relatively higher upgrading in the presence of hydrogen in terms of API gravity upgrading, viscosity reduction, and SIMDIS shift as reported in sections 3.2.1−3.2.3 of this paper. 3.3. Effect of Guard Bed. In section 3.1 it was found that better upgrading was obtained at high reaction temperature (i.e., 425 °C) at the expense of high coke formation leading to rapid catalyst deactivation. It is well-known that larger molecular weight compounds such as resins and asphaltenes contribute largely to coke formation. Therefore, a guard bed was introduced upstream of the catalyst bed to prevent premature catalyst deactivation due to coking by the adsorption of macro-molecules from the THAI feed oil flowing through the active catalyst bed. For this reason activated carbon (AC) with a surface area of 819.92 m2·g−1 and pore diameter 412 nm was used as the guard bed in the microreactor and was placed on top of the CoMo/alumina catalyst. Activated carbon was used because of its affinity for macro-hydrocarbon molecules and adsorptive selectivity for asphaltenes, resins, and coke precursors.50,51 3.3.1. Effect of Guard Bed on API Gravity. In Figure 11, the API gravity of the upgraded oil over activated carbon (i.e., guard bed) only, CoMo/alumina only, and CoMo/alumina catalyst with guard bed is presented. Notably, the use of activated carbon as guard bed further increased the API gravity

Figure 10. TGA and DTG of spent catalyst (CoMo/alumina) obtained from CAPRI reactor with and without hydrogen-addition: temperature, 425 °C; reaction media, H2 and N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 200 mL·mL−1.

CAPRI reactor either in the presence of hydrogen or nitrogen. From the TGA curves the coke content was calculated to be 57.3 wt % for catalysts operated under nitrogen decreasing to 35 wt % in the presence of hydrogen. This indicates a 22 wt % reduction in coke content of the spent catalyst in the presence of hydrogen. Suppression of coke formation due to hydrogen addition is widely thought of as a capping of free radical coke precursors formed when carbon-to-carbon bonds split leading to lower molecular weight compounds than the original THAI feed oil molecules.44,45 15402



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Figure 11. API gravity of produced oil for guard bed integrated with catalyst, catalyst CoMo/alumina, and activated carbon (AC): temperature, 425 °C; reaction media, N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 500 mL·mL−1.

Figure 12. Viscosity of produced oil for guard bed integrated with catalyst, catalyst CoMo/alumina, and activated carbon (AC): temperature, 425 °C; reaction media, N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 500 mL·mL−1.

of the produced oil in the range of 3 to 9 °API points relative to 13 °API of THAI feed oil. Conversely, 20 °API is the maximum obtainable API gravity using catalyst only, while 22 °API is achievable by adding activated carbon as guard-bed. This indicates an additional 2 °API gravity increment. Without the guard bed, the API gravity for the catalyst significantly fell shortly after the reaction started and reached its lowest value of about 16 °API at around 200 min into the experiment and remained at a plateau of lower API upgrading until the end of the experiment. However, the introduction of activated carbon as guard bed to the catalytic upgrading reactions sustained the catalytic activities and API gravity of the produced oil above 16 °API for almost 840 min before a noticeable decrease of API gravity was observed in the produced oil. This adds additional 10 h of catalytic activity compared to the use of CoMo catalyst only. 3.3.2. Effect of Guard Bed on Viscosity. The plot of the viscosity of the produced oil versus time-on-stream for the CoMo/alumina catalyst only, activated carbon only, and activated carbon guard bed with CoMo/alumina catalyst is shown in Figure 12. It is clear that CoMo/alumina catalyst with guard bed system produced oil with lower viscosity than with CoMo/alumina catalyst and activated carbon only. This result is consistent with that reported in Figure 11 for API gravity of the produced oil for each system. Thus, for CoMo/alumina catalyst only the average degree of viscosity reduction is 81% for 20 h of operation, whereas with activated carbon upstream the CoMo/alumina catalyst bed an average degree of viscosity reduction is 89.5% in the same time of operation. This implies an additional 8.5% decrease in the produced oil viscosity. The larger viscosity reduction as evidenced from Figure 11 (or API gravity increase from Figure 12, section 3.3.1) observed with the combination of guard bed, that is, AC and CoMo/ alumina is largely because macro-molecules and coke precursors in the THAI feed oil may have been adsorbed onto the activated carbon.48 Thus relatively smaller or less contaminated molecules followed on to the CoMo/alumina catalyst and cracked efficiently. This statement is further supported by the fact that neither the guard bed nor the CoMo/alumina catalyst achieved the higher upgrading on their own. The guard bed acted as a sieve and allowed the diffusion

of asphaltenes into its catalytic pellet on its active sites. In this range of porosity, the catalyst has better metal retention capacity and asphaltenes elimination than HDS. The catalyst has high dispersion of metal and large pore diameter, which prevents plugging of the pore network and a high metal retention up to 100% based on the fresh catalyst.52 3.3.3. Effect of Guard Bed on Boiling Point Distribution of Oils. From the SIMDIS boiling point ranges as presented in Table 6, it can be observed that the guard bed combination Table 6. SIMDIS of Produced Oil for Guard Bed Integrated with Catalyst, Catalyst CoMo/Alumina: Temperature, 425°C; Reaction Media, N2; Pressure, 20 barg; Oil Flow Rate, 1 mL·min−1; Gas Flow Rate, 500 mL·mL−1 cumulative product percent yield 15%

30%

45%

60%

75%

90%

438 382 381

475 442 438

temperature (°C) THAI feed oil CoMo/alumina CoMo/alumina + AC

218 183 193

296 233 247

360 284 292

402 334 336

with the CoMo/alumina catalyst does not differ in the boiling point ranges from the experiment where only CoMo/alumina catalyst was used alone. In fact they are both almost identical. This suggests that the guard bed does not change the produced oils chemical properties. The sieve-like character of the guard bed can further be evidenced from the fact that the AC bed is neutral in acidity and therefore does not perform cracking functions, instead the upgrading observed is only due to adsorbed macro-molecules and thermal cracking reactions. SIMDIS provides a conclusive proof to this statement where both the CoMo/alumina and the combination of guard bed and CoMo/alumina boiling point distribution was almost identical, suggesting no chemical role of the guard bed. 3.3.4. Effect of Guard Bed on Thermal Decomposition Behavior of the Spent Catalyst. The TGA profile (i.e., weight loss as a function of temperature plot) and derivative weight loss curve for the spent activated carbon used as guard bed is shown in Figure 13. The estimated deposits of larger molecular 15403



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The improvement in API gravity and viscosity of the produced oil when H2 was the injected gas compared to nitrogen alone was very significant. Also, the addition of hydrogen to the CAPRI process suppressed coke formation by capping the free radical coke precursors formed when macromolecules in the heavy oil split, thereby reducing deactivation due to coke deposition on the catalyst. In the presence of hydrogen, the catalysts not only promote hydrogenation of the hydrocarbon radicals, but also suppress excessive cracking and polymerization reactions, whereby extending catalyst activity by suppressing coke formation. The integration of activated carbon as guard bed achieved further upgrading of the produced oil, providing two additional degrees of API gravity and viscosity reduction of 8.5% compared to the use of HDS/HDT catalyst alone. In addition, the catalytic activity was sustained for a longer period, as macro-molecules in the oil were adsorbed onto the activated carbon. Also, the coke content of the catalyst after upgrading was reduced by 21.2% when activated carbon was used as a guard bed upstream of the CoMo/alumina catalyst bed.

Figure 13. TGA and DTG of activated carbon guard bed placed upstream of the CoMo/alumina catalyst: temperature 425 °C; reaction media, N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 500 mL·mL−1.

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weight compounds (e.g., resins and asphaltenes) on the activated carbon guard bed was 31 wt %, which burns off in the temperature range of 460−600 °C. This indicates that some of the macro-molecules have been adsorbed by activated carbon prior to catalytic process. Consequently, the TG and DTG of spent CoMo/alumina catalyst recovered from AC guard bed installed above the CoMo/alumina catalyst and that without guard bed are presented in Figure 14. It is clear that the coke

ASSOCIATED CONTENT

* Supporting Information S

Composition and properties of used catalysts, RGA analysis; gas composition of experiment conducted in nitrogen as the feed gas, RGA analysis; gas composition of experiment conducted in hydrogen as the feed gas. This material is available free of charge via the Internet at http://pubs.acs.org.

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AUTHOR INFORMATION

Corresponding Author

*Tel.:+44 (0) 121 414 5295. Fax: +44 (0)121 414 5324. Email: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors acknowledge the financial support of PTDF, Nigeria, EPSRC (Grant No. EP/E057977/1 and No. EP/ J008303/1), United Kingdom and Petrobank Energy and Resources Ltd., Canada, for supplying the heavy crude oil used in this study. The TGA used in this research was obtained through Birmingham Science City: Hydrogen with support from Advantage West Midlands (AWM) and partial funding by the European Regional Development Fund (ERDF).

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Figure 14. TGA and DTG of spent CoMo/alumina catalyst obtained from a guard bed reactor and without guard bed: temperature 425 °C; reaction media, N2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 500 mL·mL−1.

content in the experiment with guard bed placed ahead of the CoMo/alumina was 39.6 wt % compared to the run where only CoMo/alumina was used which is 57.4 wt %.

4. CONCLUSIONS The degree of upgrading in terms of API gravity of partially upgraded THAI feed oil was 1.7 and 3 °API at 350 and 400 °C, respectively. However there was a larger upgrading effect of 7 °API points at the highest temperature investigated of 425 °C. The viscosity of the produced oil was also reduced by 3−5 times compared to the THAI feed oil, in the temperature range investigated.

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ABBREVIATIONS DTG = differential thermal analysis HDS = hydrodesulfurization HDT = hydrotreating HDA = hydrodeasphaltization HDS = hydrodesulfurization HDM = hydrodemetlation OOIP = original oil in place SARA = saturates, aromatics, resins and asphaltenes TGA = thermogravimetric analysis WHSV = weight hourly space velocity REFERENCES

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