optimization of the capri process

Optimization of the CAPRI Process for Heavy Oil Upgrading: Effect ofHydrogen and Guard BedAbarasi 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 recoveryand upgrade of heavy oil and bitumen. The technologies combine horizontal production well, in situ combustion, and catalyticcracking to convert heavy feedstock into light oil down-hole. The deposition of asphaltenes, coke, and metals can drasticallydeactivate the catalyst in the CAPRI process. A fixed bed microreactor was used to experimentally simulate the conditions in thecatalyst zone of the oil well of CAPRI. In this study, oil upgrading and catalyst deactivation in the CAPRI process wereinvestigated in the temperature range of 350−425 °C, pressure of 20 barg and residence time of 9.2 min. Additionally, a guardbed consisting of activated carbon particles prior to the active catalyst in a microreactor and/or the addition of hydrogen to thegas feed were used to minimize coke formation and catalyst deactivation through respectively removing and hydrocracking thecoke precursors. It was found that depending on the upgrading temperature, the viscosity of the produced oil reducedsignificantly by 42−82% and (American Petroleum Institute) API gravity increased by ∼2 to 7 °API relative to the feedstock of0.49 Pa·s and 13 °API, respectively. Conversely, the use of hydrogen further increased the API gravity by 2 °API and the viscosityby 5.3%. Notably, the coke content of the catalyst reduced from 57.3 wt % in nitrogen to 34.8 wt % in hydrogen atmosphere. Theuse 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 energyglobally. As conventional light crude oil production hasarguably reached its peak and its production begins to decline,attention has shifted to vast deposits of unconventional oilresources, that is, heavy oil, bitumen, and shale oil, etc. to meetthe ever rising energy demand.1 Heavy oil and bitumen aredense, viscous, and difficult and costly to extract, produce, andrefine. Conversely, heavy oil and bitumen account for about70% of the world’s total 9−13 trillion barrel oil resource.2,3

Therefore, the viability of these resources is dependent onrecovery and upgrading technology that will convert them tolight 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, andenergy demanding, usually being accomplished by pipelineheating and solvent dilution. However, heavy oil and bitumencannot be refined by present refineries without upgradingprocesses to convert them to synthetic light crude oil first so asto meet refinery feedstock specification.4 In the view of this, themajor cost associated with heavy oil and bitumen exploitation isthe additional cost incurred for the upgrading facility. This isone reason for their lesser exploitation in the past. Since 2005,rising oil prices have made the recovery and upgrading of heavyoils more economic compared to conventional fuels. This hasgreatly increased investment in the production of extra heavyoil and natural bitumen to supplement conventional oilsupplies, raising the production levels to more than 1.6mb·d−1 or just under 2% of world crude oil production.5 In2009 this increased to 2.3 mb·d−1 or less than 3% of the world

demand and is projected to meet about 10% of the world crudeoil demand in 2035.6 To achieve these projections innovation isneeded in technology to increase recovery levels and increasequality of the recovered oil by upgrading.Steam based technologies such as cyclic steam stimulation

(CSS) and steam assisted gravity drainage (SAGD) have beenthe most successful and commercialized techniques for heavyoil and bitumen recovery and upgrading. These processes relyon reducing the heavy crude oil and bitumen viscosity byheating the oil to improve oil flow from the reservoir to theproduction well.7,8 However, limitations imposed due togeology of the oil reservoirs1 and heavy environmentalfootprint, where 2−10 barrel of water has to be injected assteam for every barrel of oil produced (SAGD),9 are thepotential disadvantages. Moreover, the produced oil from theaforementioned technologies needs further upgrading andexpensive diluent in large quantities for transportation torefineries. To reduce costs and environmental footprint, a highpercentage recovery of oil in place and upgrading at the wellhead or even in situ within the well are desirable. This need isfurther augmented by the fact that surface upgrading processesare expensive to set up and energy intensive to run withextensive emission and negative environmental impact.10

Special Issue: NASCRE 3

Received: March 1, 2013Revised: April 24, 2013Accepted: April 24, 2013Published: April 24, 2013

Article

pubs.acs.org/IECR

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

pubs.acs.org/IECR

THAI-CAPRI is relatively a new technology simultaneouslyincorporating down-hole in situ catalytic upgrading withthermally enhanced oil recovery. THAI integrates in situcombustion with advanced horizontal well concepts, whereby asmall fraction of the reservoir oil is burnt to mobilize the heavyoil.11 Thermal cracked heavy oil produced during the processnot only aids oil recovery but also has the added benefit ofupgrading the oil.10 CAPRI is the catalytic extension to THAIdeveloped in collaboration with the Petroleum RecoveryInstitute (PRI), where the objective is to achieve furtherheavy oil upgrading in situ by placing an annular layer of catalystaround the perforated horizontal producer well to create adownhole catalytic reactor.12,13 The thermal-cracking reactionsof THAI, taking place ahead of the combustion zone, that is,the coke zone, creates the precursor for CAPRI. Oil upgradingis thought to occur by a combination of carbon-rejection(thermal cracking) and hydrogen addition reactions at thesurface of a hydroconversion (HCT) or hydrotreating (HDT)catalyst.12 From laboratory scale experiments of the THAI-CAPRI process Xia et al.10 reported recovery levels of 79% ofthe original oil in place (OOIP), API upgrading of 23 degreesand viscosity reduction of as low as 20−30 mPa.s usingLloydminster heavy crude oil as the feed and CoMo/alumina asthe catalyst. Detailed studies of the technology have beenreported elsewhere.12,14,15

Subsequently the CAPRI process was further investigatedusing a set of microreactors to replicate underground upgradingconditions and to optimize catalyst type, oil and gas flow rate,temperature, and pressure.14 Shah et al.16 pointed out a numberof shortcomings of the technology, such as asphaltenes, coke,and metal deposition drastically deactivates typical refineryHDS and HDM catalysts when extra heavy oil/bitumen fromAthabasca is used. In the current study, as an investigation toovercome these limitations of the THAI-CAPRI technique, anHDS CoMo/alumina was used along with a number ofmodifications to the catalytic bed and reaction media, such asthe use of a guard bed and hydrogen atmosphere. A guard bedinvolves packing inert porous particles above the catalyst inorder to adsorb or filter coke precursors from the feed beforethey reach the catalyst. The addition of hydrogen to the feedwas intended to augment catalytic hydroconversion andhydrocracking reactions in order to achieve a higher level ofupgrading than would be achieved with inert gasesalbeit atlow partial pressure. The source of hydrogen is expected to begasification and/or water gas shift reactions.17 A number oftemperature regions have been reported in a THAI-CAPRIreservoir18 and to investigate whether catalytic activity can bemaintained at lower temperatures compared to the previouslyoptimized temperature of 425 °C,14 experiments wereconducted at lower temperatures in the current study.

2. EXPERIMENTAL DETAILS2.1. Feedstock and Catalysts. The heavy crude oil was

supplied by Petrobank Energy and Resources Ltd. from itsWhitesands 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. Thecomposition and microstructural properties of the usedcatalysts such as specific surface area, pore volume, and porediameter were determined by Brunauer−Emmett−Teller(BET) technique and are presented in the SupportingInformation, Table 1S.

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

the thermal cracking is believed to be happening. Hence theprocess moves from the “Toe” position to the “Heel”. The oilin the MOZ remains hot upon reaching the horizontalproduction well due to heat gained from the combustionzone. Moreover the horizontal production well is always kepthot to keep the oil flowing. In the CAPRI process, theproduction well is packed with an annular catalyst layer. As thehot oil passes through the annular catalytic layer, catalyticupgrading is believed to happen. From Figure 1 a down-flowmicroreactor was used in the laboratory representing acylindrical core of 1 cm taken in a radial direction throughthe annular layer of the catalyst surrounding the producer wellas represented in Figure 2.

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

Table 1. Properties of the THAI Feed Oil

property value

density at 20 °C (g·cm−3) 0.9801 ± 0.002viscosity at 20 °C (mPa·s) 490API gravity (deg) ∼13

Figure 1. Schematic of the THAI-CAPRI process11.

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

Industrial & Engineering Chemistry Research Article

dx.doi.org/10.1021/ie400661x | Ind. Eng. Chem. Res. 2013, 52, 15394−1540615395

the catalytic bed with the help of gravity. The experiment wasinitiated by turning on the furnace and the trace heating andsetting the controls to the desired temperatures and pressures.Once the operating conditions of temperatures and pressureswere achieved in the reactor, the oil flow metering valve wasopened manually to initiate the oil flow from a feed tankpressurized by nitrogen gas. The flow metering valve wasadjusted 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 oilflow through the reactor. The THAI feed oil (i.e., heavy crudeoil) was delivered to the packed catalyst bed in down-flowmode to ensure flow by gravity and complete wetting of thecatalytic bed.The THAI feed oil was fed into traced heated lines having a

set temperature of 280 °C and into the furnace after achievingthe desired experimental temperature. The furnace providesisothermal 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 Figure3 for reference purposes only), were used as the reaction mediato simulate the combustion gases expected in a real THAI-CAPRI reservoir. These gases were mixed with the THAI feedoil in the mixing chamber and the gas−oil mixture passedthrough the reactor in concurrent flow. The gas−oil mixture ofthe partially vaporized THAI feed oil flowed downwardthrough the voids of the packed catalyst bed where itunderwent cracking reactions aided by heat consumption.The product stream coming out of the reactor passed

through a back pressure regulator (Swagelok Co. UK), whichregulated and maintained a constant pressure of 20 barg in thereactor. The upgraded products which include light oil andgases were passed to the gas−liquid separator, where the lightoil was collected, while the gaseous products were flashed off,and either vented or sent to a refinery gas analyzer (RGA) forconcentration and compositional analysis by gas chromatog-raphy. The analyses were performed periodically during eachexperimental run. The light oil sample was drained from thegas−oil separator initially every 20 min for the first hour, every30 min for 4 h, and thereafter every 40 min, and the collectedoil 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

shaped extrudate (Akzo) was used in this study. Thecomposition and microstructural properties of the usedcatalysts such as specific surface area, pore volume, and porediameter 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 producedoils were measured using an Anton Parr DMA 35 portabledensity meter (Anton Paar GmbH, Austria) at 15 °C andreported in kg·m−3. API gravity was calculated using eq 1,

= −API gravity141.5

SG131.5

(1)

Where SG represents specific gravityThe Bohlin CVO 50 NF rheometer (Malvern Instruments

Ltd., United Kingdom) was used to measure the viscosity of theTHAI feed oil and produced oils. All viscosity measurementswere performed at 20 ± 0.1 °C. Aluminum parallel plategeometry was used. The diameter of the used plate is 40 mmand it was made of aluminum material with a polished surface.The parallel plate gap size was set at 150 μm and a shear rate of100 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 anAgilent 6850N GC and calibrated in accordance with theASTM-D2887 was used to characterize the feed and producedoils. Agilent 6850N Network gas chromatograph system J&W125-10 is fitted with a DB-1 10 m length, 530 μm ID, and 2.65

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

Table 2. Operating Conditions in the Experiments

feed flow rate (mL·min−1) 1catalyst inventory (g) 6pressure (barg) 20reaction temperature (°C) 350−425WHSV (h−1) 9gas-to-oil ratio (mL·mL −1) 200−500



μm film thickness capillary column. Prior to the injection of thesample, the feed and produced oil samples were diluted withcarbon 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 volumepercentage of H2, CO, CO2, and C1−C5 hydrocarbons. TheRGA three channels include flame ionization detector (FID)and two thermal conductivity detectors (TCDs). The lighthydrocarbon components in the gas stream are determined bythe FID channel column HP-PLOT Al2O3S capable ofseparating C1 to nC5 including their 22 isomers based on thecalibrated table, while components heavier than nC6 are backflushed through the precolumn. One of the TCD with heliumcarrier gas is used for permanent gases analysis such as N2, CO,CO2, O2, and hydrocarbons to nC5, C6+. The other TCD withnitrogen as carrier gas determines gases like hydrogen andhelium in the gas stream. The oven column dimension is 27 m× 320 μm × 8 μm at a temperature of 200 °C. The RGA takes15 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, TG209 F1 Iris. A 10 mg sample of the spent catalyst was recoveredfrom the reactor and placed on an alumina crucible above themicrobalance. The microfurnace is programmed as follows:linear increase in temperature in range of 25 °C to 1000 °C anda heating rate of 20 °C·min−1. At 1000 °C an isothermalcondition was maintained for 20 min to enable total burnoff ofthe materials deposited on the spent catalysts. The total timefor each run is 49 min and the air flow rate during the TGanalysis was set at 50 mL·min−1.

3. RESULTS AND DISCUSSION3.1. Effect of Temperature. 3.1.1. Effect of Temperature:

Mass Balance. The effect of temperature upon the yield ofliquid, gas, and coke deposited upon the used catalyst wasstudied, together with the extent of upgrading of the producedoil as measured by viscosity, API gravity and simulateddistillation analysis. The mass of gas evolved during theupgrading reactions was calculated as the mass of the oilremaining after subtracting the masses of produced liquid andsolid deposits in the reactor. The mass balances of the threepseudo-products, that is, liquid, gas, and coke, were calculatedas percentage of the mass of THAI feed oil into the systemusing eqs 2 and 3:

=w

wyield (wt %) 100i

Feed (2)

= −

−

gas (wt %) 100 liquid yield (wt %)

coke yield (wt %) (3)

Where wi is the weight of component i and wFeed is the overallweight 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 lowertemperature favors lesser production of gases and coke andmore liquid products, suggesting a low degree of thermal andcatalytic upgrading. The amount of coke produced at 350 °Cwas 0.64 wt % compared to 1.86 wt % at 425 °C, and thecorresponding measured liquid yields were 97.4 and 93.8 wt %,respectively. The yield of gases was 1.96 wt % at 350 °C and4.34 wt % at 425 °C. The increased production of gas and coke

at the higher temperature can be attributed to increasedcatalytic cracking reactions with temperature rise as Krumm etal.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 showsthe API gravity of the produced oils at different processing

temperatures. It is clear that the API gravity of the produced oilincreased as the reaction temperature increases from 350 to 425°C. During the first 100 min, significant API gravity increasecan be observed at all of the three reaction temperatures andreaches a maximum value of 6.5° at 425 °C. This most likelyoccurs because during the earlier part of the reaction, thecatalyst is still active and has not deactivated. At lowertemperatures of 350 and 400 °C, an initial API increase of 5−6°can be observed compared to the THAI feed oil. At highertemperatures in the plateau region the level of upgrading alsoseems to depend on temperature, with about 1.7° increase at350 °C and increasing up to ∼4−5 °API upgrading at 425 °C.These trends in API gravity were expected as reported earlierby Shah et al.16 that 425 °C was the optimum upgradingtemperature for CAPRI, and further increase in temperatureleads to increased coke deposition which required the reactorto be shut down due to blockages of the reactor bed with coke.From the mass balance in Table 4, it can be observed that moreliquid production occurs at lower temperature; however, thelevel of upgrading at the lowest temperature of 350 °C is not assignificant as at the previously optimized temperature of 425°C, suggesting a trade-off between liquid production andquality. These results however, establish the fact that even atlower temperatures of 350 °C the API gravity increase of 1.7°

Table 3. Mass Balances; Catalyst, CoMo/Alumina; ReactionMedia, N2; Pressure, 20 barg; Oil Flow Rate, 1 mL·min−1;Gas Flow Rate, 500 mL·mL−1

temperature (°C) gas (wt %) liquid (wt %) coke (wt %)

350 1.96 97.4 0.64400 2.96 95.87 1.17425 4.34 93.8 1.86mean standard deviation 0.24 1.20 0.91

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



shows some measurable improvement for the catalyst packingaround the horizontal production well.In industry generally sulfides of Co, Ni, W, and Mo

supported catalysts, having a variety of pore structures andactive metal dispersion (active sites), are used for petroleumresidue hydroprocessing. The most important property for aresidue hydroconversion catalyst is pore diameter, becausefeedstock contains large molecules of asphaltene, metalchelates.20 Broadly heavy crude is typically processed intrickle-bed reactors at temperatures of 350−450 °C andpressures of 5−15 MPa of hydrogen.20 However, since nohydrogen was used in the results presented, the predominantreactions are assumed to be carbon rejection or limitedhydroconversion. In this case hydrogen is assumed to begenerated from side chain breaking and/or as a result of anumber of complex polymerization reactions resulting in theeventual formation of condensed polyaromatic species in ahydrogen deficient environment and releasing hydrogen.21 Theincrease in API gravity as observed from Figure 4 leads toimproved quality of the produced oil as Greaves and Xia13

noted that with a 7.9 °API increase only about 15% of diluentwas needed to meet pipeline specifications, compared to 30−50% required for nonupgraded bitumen produced from SAGDand 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 reactiontemperatures is presented in Figure 5. From Figure 5

substantial viscosity reduction of the upgraded oil can beobserved compared to the THAI feed oil, with larger viscosityreduction at higher temperatures. At reaction temperatures of350 and 400 °C the average viscosity of the produced oilreduced to 0.28 and 0.17 Pa.s, respectively, from the base valueof 0.49 Pa·s for the THAI feed oil. These are about 1.7 and 3times lower than the THAI feed oil viscosity. However higheraverage viscosity reduction from 0.49 Pa.s to 0.09 Pa.s wasobtained at 425 °C, which is approximately 5 times lower thanthat of the THAI feed oil. The level of viscosity reduction,which was up to 81% for the case of 425 °C in the presence ofCoMo/alumina catalyst in the presence of nitrogen, representsa significant step forward for the recovery and upgrading ofheavy oil/bitumen downhole. This would reduce the difficultiesoccurring during heavy oil production, transportation, andsurface processing as a result of upgrading of the oil can beachieved. Moreover, capital and operating costs of a surfaceupgrader can reach around hundreds of millions dollars, but adownhole or in situ oil upgrading process may reach the solecost of buying and installing the catalyst, which may deliver arelatively short payback time.10 Commercial catalysts are fairlycheap. Shah et al.1 estimated that 20 tonnes of new HDScatalyst 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 lowerAPI gravity of CAPRI produced crude oil (say 22°), then at$50/barrel, the catalyst cost represents only about 11/2 to 21/2days production. Early estimates suggest that the THAI/CAPRIprocess costs about one-third per producing barrel of theequivalent SAGD process.22

Saturates, aromatics, resins, and asphaltenes (SARA) analyseswere carried out and reported by Shah et al.16 for experimentsperformed in the same reactor and under the same conditionsas 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 theproduced oil at 425 °C were 16.37, 67.62, 9.11, and 6.9%,respectively. The amount of aromatics largely occurred at theexpense of naphthenes and asphaltenes. Shah et al.16 concludedthat the rise in API and viscosity was due largely tohydroconversion rather than HDS, HDM, or HDA activity.

3.1.4. Effect of Temperature on Boiling Point Distributionof Oils. Table 4 provides the results of the cumulative productpercentage yield of feed and produced oils at selected simulateddistillation temperatures. It can be observed that temperaturerise in the upgrading experiment favors a significant shifttoward lighter distillate fractions in comparison to the partiallyupgraded THAI feed oil. A dramatic 38 °C shift in the boilingpoint range at 75 wt % yield can be observed for 350 °C incomparison to THAI feed oil. This rises to 56 and 48 °C at 75wt % yield for the higher process temperatures of 400 and 425°C, respectively. An important observation can be made that at400 and 425 °C the shift toward lower distillable temperaturesis almost identical. The level of upgrading in the presence ofCoMo/alumina catalyst is also superior to thermal upgrading inthe presence of glass beads by 34 °C at 75 wt % cumulativeproduct percentage yield.

3.1.5. Effect of Temperature on Thermal DecompositionBehavior of the Spent Catalyst. Figure 6 provides thethermogravimetric analysis (TGA) thermograms or weightloss curves as a function of ramp temperature increase for theCoMo/alumina recovered catalyst after the upgrading experi-ments. 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 FlowRate, 500 mL·mL−1

cumulative product percent yield

15% 30% 45% 60% 75% 90%

temperature °CTHAI feed oil 218 296 360 402 438 475glass beads at 425 °C 190 258 321 379 424 471CoMo/alumina at 350 °C 207 265 313 358 400 455CoMo/alumina at 400 °C 183 233 284 334 382 442CoMo/alumina at 425 °C 183 233 285 339 390 442

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



of the experiment. The ramp temperature increase enables theinterpretation of the different chemical changes occurringduring the burnoff when compared to the oil TGA anddifferential thermogravimetric (DTG) curve. The isothermaltemperature of 1000 °C toward the end ensures that all carbonspecies are completely burnt off during the heating period.However, above 620 °C (see Figure 5), the deposits on thespent catalysts are defined as coke.23 Subsequently, to interpretthe different stages during the heating period, the derivative ofthe weight loss curve (DTG) obtained in the thermal burnoff inair 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 to208 °C represents loss due to devolatilisation of light oil andbeyond 620 °C is coke. Beside catalyst and reactor fouling andclogging, coke build up on the catalyst bed is one of the mainroutes for deactivation and shortening of lifespan.24 The cokecontent of the spent CoMo/alumina catalyst increased in theorder 48.4, 53.5, and 57.3 wt % at 350, 400, and 425 °Crespectively. The large quantity of coke 48−57 wt % of thespent catalyst, may have led to a large loss of surface area. Thisobserved trend is in line with the findings of Meng et al.25 oncoking behavior and catalyst deactivation for catalytic pyrolysisof heavy oil.Sanford26 has pointed out that during the earlier stages of

coking and hydrocracking reactions in residue conversion, C−Cbonds are broken. These reactions reject carbon at the surfaceof the catalyst pellets similar to the reactions occurring in theTHAI-CAPRI process. At higher temperatures of 425 °C andabove the coking reactions become more predominant, whichaccelerate catalyst deactivation and change in selectivity towardundesirable products,27,28 that is, the coke and gaseous species.This is evident from Table 3 where 4.34% of products formedwere gases at 425 °C compared to 1.96% at 350 °C, and cokewas 1.86 at 425 °C compared to only 0.64% at 350 °C. Similartrends can be observed from the spent catalyst TGA and DTGanalysis (Figure 6) where higher temperature favored highercoke formation and lesser liquids. This can be explained by thefact that at higher temperatures bond scission reactionsinvolving the side chains of compounds present in crudessuch as alkylaromatics increases and results in the formation ofmore coke, gases, and lower liquid products.29 This may also

explain that with an increase in temperature from 350 to 425°C the activity of the catalyst increased and hence the level ofhydroconversion. This resulted in higher upgraded oil in termsof API gravity and viscosity as evident from Figures 4 and 5 andalso the boiling point distribution shift toward lower boilingpoint hydrocarbons (Table 4) as reaction temperatureincreased. Zhao30 and Meng et al.25 reported that at reactiontemperatures of 400 °C and above, the cleavage of C−C bondsand the rate of cracking of the larger molecular weightcomponents becomes dominant, and subsequently promotetheir condensation reactions leading to increase coke contentand 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 cokeformation.

3.1.6. Effect of Temperature on Gas Products Distribution.Supporting Information, Table 2S provides quantitative analysisof the gas composition produced during the course of theexperiment. Average values of the gases have been reported inTable 2S for the course of the whole experiment as gascomposition varied during different stages of the experiments.From Table 2S it can be observed that trace gases from C1 toC5 (varying in structural composition) are produced for all thethree experiments namely glass beads at 425 °C and catalyticexperiments with CoMo/alumina conducted at 400 and 425°C. From the table it is clear that thermal cracking (over glassbeads) produces more of certain gas components such asmethane, propane, n-butane, trans-2-butane, and cis-2-butene,and n-pentane compared with catalytic cracking; however, someother components such as propene, i-butane, and 1-butenewere not produced over glass beads that were present over thecatalyst. The amount of gases also rises with temperature in thecatalytic runs. The combined value of the hydrocarbonsproduced was 1.2 for glass beads at 425 °C, 0.84 and 0.94vol % for catalytic runs at respective temperatures of 400 and425 °C. The levels of CO and CO2 were also identical in thetwo catalytic runs. The only exception is the producedhydrogen in the three different experiments, which rises withtemperature and in the presence of catalyst. These trends are inconformity with trends reported elsewhere, where gasificationhas largely been attributed to thermal cracking.19

3.1.7. Effect of Temperature: Thermal vs CatalyticCracking. To investigate the effect of temperature a controlledexperiment 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 400and 350 °C. This compares less favorably with the catalyticupgrading (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·sover glass beads compared to the catalytic runs viscosities of0.09, 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 higherviscosities reduction for the catalytic experiments compared tothe glass beads, establishing the superiority of catalytic crackingcompared to thermal cracking alone. SIMDIS values for glassbeads from Table 4 further confirm the low level of upgradingby the boiling points being 34 °C lower than the comparablecatalytic experiment for 75 wt % product yield. The level ofupgrading seen with glass beads may also have beencomplemented with the hydrogen produced during the process

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



as observed from the hydrogen in the outlet gases in the gasstream and reported in Supporting Information, Table 2S.The assumed hydroconversion (in the absence of any

externally added hydrogen) in the presence of CoMo/aluminacatalyst is further evidenced by the fact that a significantly lowerlevel of upgrading was achieved in the absence of the catalyst interms of API gravity or viscosity (Figures 4 and 5) or fromboiling point distribution of oil upgraded over only glass beadscompared 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 presenceof hydrogen is a process by which the hydrocarbon moleculesof 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 thepresence of a catalyst. These broken species are subsequentlyhydrogenated resulting in their conversion to lower molecularweight hydrocarbons. It has been reported that hydrogenpromotes saturation of olefins and aromatics, facilitates thetermination of free radical coke precursors formed during thecracking reactions, and increases the hydrogen-to-carbon (H/C) ratio of the produced oil. These reactions can subsequentlyremove a heteroatom, thereby suppressing coke yield andcatalyst deactivation.31

To investigate this effect in the THAI-CAPRI reactor,experiments were performed using CoMo/alumina catalysts ata temperature of 425 °C, pressure of 20 bar, H2-to-oil ratio of200 mL·mL−1, and WHSV of 9 h−1. The H2-to-oil ratio usedwith hydrogen was lower than the experiments presented in theearlier sections, but for the purposes of comparison anadditional experiment was carried out in a nitrogen atmospherewith an N2-to-oil ratio of 200 mL·mL−1. The effect ofhydrogen-addition on the API gravity upgrading of theproduced oil is presented in Figure 7. It is clear thathydrogen-addition results in higher API gravity of the producedoil 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 amaximum of 7 °API, compared to the THAI feed oil (13 °API).The fluctuations in the data from Figure 7 are largely due to

hydrogen gas flow control issues. During the course of theexperiment hydrogen consumption varied significantly andhenceforth 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 nitrogenatmosphere. The additional upgrading of 2 °API with hydrogenaddition is significant as it may represent a premium of $2−3for each barrel of oil produced. It also suggests that hydrogenaddition to the molecule is able to effect further upgrading thancatalytic carbon rejection which is thought to occur with onlynitrogen.CoMo/alumina supported on Al2O3 and/or SiO2 is a

bifunctional catalyst (i.e., metal and acidic), which is favorablefor hydrocracking and hydrogenation.32 The metal sitespromote hydrogen addition and the acidic sites of the supportspromote cracking reactions through a carbocationic mecha-nism.21,33 Catalytic upgrading in the presence of an HDS/HDTcatalyst, that is, CoMo/alumina, occurs by a number of steps.First breaking the larger molecules to give fragmented freeradicals, which then react with hydrogen radicals in order tostabilize the hydrocarbon chains and terminate the reaction. Itwould appear that in the absence of hydrogen, the active chainskeep reacting with each other resulting in the formation ofhigher molecular weight compounds by polymerization,increased coke formation, and adverse impact of viscosity andAPI upgrading of the produced oil. This has been illustrated ineqs 4 and 5.34 Moreover, since the molecular size of hydrogenis smaller than the polymer chains, the mass transfer rate isfaster leading to rapid termination of active chains.34 Duringhydrocracking reactions the possible splitting of C−C, C−S,C−N bonds results in free radical hydrocarbon chains. Thetermination of active chains of these free radicals is animportant step which is facilitated by the attachment ofhydrogen and results in the reduction of the viscosity of theproduced oils during the CAPRI process.

+ →active H active chains active chain termination (

low molecular weight)2

(4)

+→

active chain active chainactive chain termination (high molecular weight) (5)

The free radicals formed during thermal or catalytic crackingare very unstable and are usually stabilized by hydrogen.However, in a nitrogen atmosphere (as was the case in section3.1) the scarcity of free hydrogen results in radical linking toform larger molecules and/or double bonded olefins due tocollisions of carbon radicals.35 This may explain the relativelylower API upgrading when nitrogen atmosphere was used asopposed to hydrogen (Figure 7). Hydrogen may have beeninitially adsorbed and dissociated on the metal sites to formreactive hydrogen in this particular experiment via acidic sites(upon which hydrogen gives rise to a proton) and subsequentlyreacted with the cracking intermediates, free radicals, andolefins that require hydrogenation.36 API upgrading haspreviously been linked to increased hydrocracking andhydrogenation reactions,37 in agreement with the results anddiscussion presented here.

3.2.2. Effect of Hydrogen on Viscosity. Figure 8 shows acomparison of the viscosity reduction in the presence of eitherhydrogen or nitrogen as the gas atmosphere. From Figure 8average 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 andN2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 200mL·mL−1.



upgraded produced oil in hydrogen and nitrogen, respectively,compared to the 0.49 Pa·s for the THAI feed oil. Thisrepresents a 6.3% further reduction in the average viscosityreduction on top of the 80.9% reduction for nitrogen only. Thismakes 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 hydrogencan be attributed to the fast termination of free radicals in theflowing hydrogen atmosphere. The effect is to terminatepolymerization of the partially upgraded oil that couldotherwise lead to the formation of polyaromatic cokeprecursors.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 formationon the surface of CoMo/alumina catalyst is because reactionintermediates have been hydrogenated to form small stablemolecules. This agrees with the studies of Liu and Fan34 andGalarrage 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 recoverylevels and upgrade the quality of the recovered oil if useddownhole. Up to 20 mol % hydrogen was observed during BP’sISC oxygen field pilot at the Marguerite Lake, Alberta.17,39 Thesource was believed to be via gasification and/or water gas shiftreactions. During a catalytic upgrading test on heavy oil in acombustion tube at the University of Calgary, 3% of hydrogenwas measured in the produced gas.40 Although only about 2barg of hydrogen pressure was observed in these previousstudies compared to 20 barg of hydrogen pressure used in thisstudy, it is expected that the use of low pressures of hydrogenwill have a more prominent effect on the level of upgradingwith feeds having less aromatic character. The THAI partiallyupgraded oil used in this work is relatively rich in aromatics,57%,16 and aromatic crudes are widely known not to be suitablefeeds for catalytic hydroconversion.41 For this reason, CAPRImay be more suitable for less aromatic crudes.3.2.3. Effect of Hydrogen on Boiling Point Distribution of

Oils. Simulated distillation temperature ranges for hydrogenand nitrogen as reaction atmosphere are provided in Table 5. It

can be observed that with the addition of hydrogen a further 9°C shift toward lower temperature distillates can be observedfor hydrogen compared to nitrogen at 75 wt % product yield.

3.2.4. Effect of Hydrogen on BET and ThermalDecomposition Behavior of the Spent Catalyst. The nitrogensorption isotherms of the fresh and spent catalysts wereanalyzed using BET theory. The analyzed spent catalysts werecollected from the center of the fixed-bed reactor. Theadsorption−desorption isotherm of the fresh and spent catalystare compared for the nitrogen and hydrogen atmospheres inthe upgrading experiment, as shown in Figure 9a,b. Theadsorption−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 andN2; pressure, 20 barg; oil flow rate, 1 mL·min−1; gas flow rate, 200mL·mL−1.

Table 5. SIMDIS for the Feed and Upgraded Oils in N2 andH2: Catalyst, CoMo/Alumina; Temperature, 425°C;Reaction Media, H2 and N2; Pressure, 20 barg; Oil FlowRate, 1 mL·min−1; Gas Flow Rate, 200 mL·mL−1


15% 30% 45% 60% 75% 90%

temperature (°C)

THAI feed oil 218 296 360 402 438 475CoMo/alumina in N2 183 233 284 334 382 442CoMo/alumina in H2 171 220 269 320 373 433

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



the isotherm indicative of the type IV which is mesopores. Alarge hysteresis loop in the isotherm of the spent catalyst canalso be observed indicating deposited materials/coke in thepores and blockage of active sites.42 The significant decrease insurface area and the volume of adsorbed nitrogen for the spentcatalyst also indicates that deposits (e.g., carbon, metals, etc.)may have altered the pore structure, textural properties, andactivity.A further observation for the spent CoMo/alumina catalyst

was that the volume of adsorbed−desorbed N2 at the samerelative pressure is significantly lower compared to the freshCoMo/alumina. This loss in surface area may have resultedfrom the highly aromatic and coke-forming molecules such asresins and asphaltenes residing within the small pores of thecatalysts, due to limited diffusion of reactants and products.43

As a result further cracking of these molecules occurred yieldingcoke, causing preferential coke deposition, and hencenarrowing, of the larger pores and leading to total blockageof the smaller pores. On the other hand, when hydrogen wasintroduced the loss of surface area by the spent CoMo/aluminacatalyst was 78.6% compared to 99.8% in nitrogen atmosphereat the same operating conditions (i.e., 425 °C, 20 bar, and 200mL·mL−1). This indicates lesser deactivation due to theblockage of catalytic site by coke deposition in the presenceof hydrogen (see adsorption−desorption isotherm of fresh andspent CoMo/alumina obtained in the use of H2 in Figure 9a,b).This is because hydrogen helps to terminate some of the cokeprecursors 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

CAPRI reactor either in the presence of hydrogen or nitrogen.From the TGA curves the coke content was calculated to be57.3 wt % for catalysts operated under nitrogen decreasing to35 wt % in the presence of hydrogen. This indicates a 22 wt %reduction in coke content of the spent catalyst in the presenceof hydrogen. Suppression of coke formation due to hydrogenaddition is widely thought of as a capping of free radical cokeprecursors formed when carbon-to-carbon bonds split leadingto lower molecular weight compounds than the original THAIfeed oil molecules.44,45

A significant effect of hydrogen in this study was its ability tosuppress the formation of coke compared to nitrogen as theflowing gas atmosphere (Figures 9a,b and 10), and as iscommonly practiced in the refining industry the phenomenonis 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 ofthe produced gases when hydrogen was injected as the reactionatmosphere. The results can be compared with the gascomposition for nitrogen atmosphere at the same temperaturein Supporting Information, Table 2S. Although the gas flow rateof the experiment in Table 2S for nitrogen was higher than theresults for hydrogen in Table 3S, it can be seen that a similaramount of hydrocarbons were produced with the onlyexception being the level of methane produced. Twice ashigh methane concentration can be observed with hydrogencompared to the nitrogen atmosphere, making the combinedvalue of hydrocarbon gases produced with hydrogen 1.52 vol %compared to 0.94 vol % with nitrogen. The level of CO2 is 0.14vol % in the presence of hydrogen compared to 0.05 vol % inthe presence of nitrogen, suggesting relative completion of thethermal or catalytic reactions. For the run in hydrogenatmosphere 12.2 ± 3.9 vol % less hydrogen was observed atthe outlet compared to the inlet once analyzed with the RGA.The lower hydrogen at the reactor outlet compared with theinlet is due to a degree of dissolution in the liquid oil, and alsoto a chemical reaction along with some possible system losses.48

Equations of state for the calculation of hydrogen loss havebeen proposed by Maipur et al.,48 but in this work the outletgas flow was not measured, so overall loss of hydrogen in moleswas not calculated. However in a hydrogen environment thetotal volume percentage of hydrocarbon gases increasedcompared with its value in a nitrogen environment, as notedabove. This provides evidence that reaction of the oil feed withhydrogen did occur, with hydrogen acting to terminate freeradicals and produce light hydrocarbons that were cracked fromthe oil molecules. This is in conformity with the resultsreported by Kim et al.49 who measured hydrogen consumptionin the presence of an HDN catalyst. This may also explain therelatively higher upgrading in the presence of hydrogen interms of API gravity upgrading, viscosity reduction, andSIMDIS 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 thatbetter upgrading was obtained at high reaction temperature(i.e., 425 °C) at the expense of high coke formation leading torapid catalyst deactivation. It is well-known that largermolecular weight compounds such as resins and asphaltenescontribute largely to coke formation. Therefore, a guard bedwas introduced upstream of the catalyst bed to preventpremature catalyst deactivation due to coking by the adsorptionof macro-molecules from the THAI feed oil flowing throughthe active catalyst bed. For this reason activated carbon (AC)with a surface area of 819.92 m2·g−1 and pore diameter 412 nmwas used as the guard bed in the microreactor and was placedon top of the CoMo/alumina catalyst. Activated carbon wasused because of its affinity for macro-hydrocarbon moleculesand adsorptive selectivity for asphaltenes, resins, and cokeprecursors.50,51

3.3.1. Effect of Guard Bed on API Gravity. In Figure 11, theAPI gravity of the upgraded oil over activated carbon (i.e.,guard bed) only, CoMo/alumina only, and CoMo/aluminacatalyst with guard bed is presented. Notably, the use ofactivated 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; oilflow rate, 1 mL·min−1; gas flow rate, 200 mL·mL−1.



of the produced oil in the range of 3 to 9 °API points relative to13 °API of THAI feed oil. Conversely, 20 °API is the maximumobtainable API gravity using catalyst only, while 22 °API isachievable by adding activated carbon as guard-bed. Thisindicates an additional 2 °API gravity increment. Without theguard bed, the API gravity for the catalyst significantly fellshortly after the reaction started and reached its lowest value ofabout 16 °API at around 200 min into the experiment andremained at a plateau of lower API upgrading until the end ofthe experiment. However, the introduction of activated carbonas guard bed to the catalytic upgrading reactions sustained thecatalytic activities and API gravity of the produced oil above 16°API for almost 840 min before a noticeable decrease of APIgravity was observed in the produced oil. This adds additional10 h of catalytic activity compared to the use of CoMo catalystonly.3.3.2. Effect of Guard Bed on Viscosity. The plot of the

viscosity of the produced oil versus time-on-stream for theCoMo/alumina catalyst only, activated carbon only, andactivated carbon guard bed with CoMo/alumina catalyst isshown in Figure 12. It is clear that CoMo/alumina catalyst withguard bed system produced oil with lower viscosity than withCoMo/alumina catalyst and activated carbon only. This resultis consistent with that reported in Figure 11 for API gravity ofthe produced oil for each system. Thus, for CoMo/aluminacatalyst only the average degree of viscosity reduction is 81%for 20 h of operation, whereas with activated carbon upstreamthe CoMo/alumina catalyst bed an average degree of viscosityreduction is 89.5% in the same time of operation. This impliesan 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) observedwith the combination of guard bed, that is, AC and CoMo/alumina is largely because macro-molecules and cokeprecursors in the THAI feed oil may have been adsorbedonto the activated carbon.48 Thus relatively smaller or lesscontaminated molecules followed on to the CoMo/aluminacatalyst and cracked efficiently. This statement is furthersupported by the fact that neither the guard bed nor theCoMo/alumina catalyst achieved the higher upgrading on theirown. The guard bed acted as a sieve and allowed the diffusion

of asphaltenes into its catalytic pellet on its active sites. In thisrange of porosity, the catalyst has better metal retentioncapacity and asphaltenes elimination than HDS. The catalysthas high dispersion of metal and large pore diameter, whichprevents plugging of the pore network and a high metalretention up to 100% based on the fresh catalyst.52

3.3.3. Effect of Guard Bed on Boiling Point Distribution ofOils. From the SIMDIS boiling point ranges as presented inTable 6, it can be observed that the guard bed combination

with the CoMo/alumina catalyst does not differ in the boilingpoint ranges from the experiment where only CoMo/aluminacatalyst was used alone. In fact they are both almost identical.This suggests that the guard bed does not change the producedoils chemical properties.The sieve-like character of the guard bed can further be

evidenced from the fact that the AC bed is neutral in acidityand therefore does not perform cracking functions, instead theupgrading observed is only due to adsorbed macro-moleculesand thermal cracking reactions. SIMDIS provides a conclusiveproof to this statement where both the CoMo/alumina and thecombination of guard bed and CoMo/alumina boiling pointdistribution was almost identical, suggesting no chemical role ofthe guard bed.

3.3.4. Effect of Guard Bed on Thermal DecompositionBehavior of the Spent Catalyst. The TGA profile (i.e., weightloss as a function of temperature plot) and derivative weightloss curve for the spent activated carbon used as guard bed isshown in Figure 13. The estimated deposits of larger molecular

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

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

Table 6. SIMDIS of Produced Oil for Guard Bed Integratedwith Catalyst, Catalyst CoMo/Alumina: Temperature,425°C; Reaction Media, N2; Pressure, 20 barg; Oil FlowRate, 1 mL·min−1; Gas Flow Rate, 500 mL·mL−1


15% 30% 45% 60% 75% 90%

temperature (°C)

THAI feed oil 218 296 360 402 438 475CoMo/alumina 183 233 284 334 382 442CoMo/alumina + AC 193 247 292 336 381 438



weight compounds (e.g., resins and asphaltenes) on theactivated carbon guard bed was 31 wt %, which burns off inthe temperature range of 460−600 °C. This indicates that someof the macro-molecules have been adsorbed by activated carbonprior to catalytic process. Consequently, the TG and DTG ofspent CoMo/alumina catalyst recovered from AC guard bedinstalled above the CoMo/alumina catalyst and that withoutguard bed are presented in Figure 14. It is clear that the coke

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

4. CONCLUSIONSThe degree of upgrading in terms of API gravity of partiallyupgraded 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 3−5times compared to the THAI feed oil, in the temperature rangeinvestigated.

The improvement in API gravity and viscosity of theproduced oil when H2 was the injected gas compared tonitrogen alone was very significant. Also, the addition ofhydrogen to the CAPRI process suppressed coke formation bycapping the free radical coke precursors formed whenmacromolecules in the heavy oil split, thereby reducingdeactivation due to coke deposition on the catalyst. In thepresence of hydrogen, the catalysts not only promotehydrogenation of the hydrocarbon radicals, but also suppressexcessive cracking and polymerization reactions, wherebyextending catalyst activity by suppressing coke formation.The integration of activated carbon as guard bed achieved

further upgrading of the produced oil, providing two additionaldegrees 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, asmacro-molecules in the oil were adsorbed onto the activatedcarbon. Also, the coke content of the catalyst after upgradingwas reduced by 21.2% when activated carbon was used as aguard bed upstream of the CoMo/alumina catalyst bed.

■ ASSOCIATED CONTENT*S Supporting InformationComposition and properties of used catalysts, RGA analysis;gas composition of experiment conducted in nitrogen as thefeed gas, RGA analysis; gas composition of experimentconducted in hydrogen as the feed gas. This material isavailable free of charge via the Internet at http://pubs.acs.org.

■ AUTHOR INFORMATIONCorresponding Author*Tel.:+44 (0) 121 414 5295. Fax: +44 (0)121 414 5324. E-mail: [email protected] authors declare no competing financial interest.

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

■ ABBREVIATIONSDTG = differential thermal analysisHDS = hydrodesulfurizationHDT = hydrotreatingHDA = hydrodeasphaltizationHDS = hydrodesulfurizationHDM = hydrodemetlationOOIP = original oil in placeSARA = saturates, aromatics, resins and asphaltenesTGA = thermogravimetric analysisWHSV = weight hourly space velocity

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

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



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http://dx.doi.org/10.1016/j.fuproc.2012.09.059

http://dx.doi.org/10.1016/j.fuproc.2012.09.059

http://nextbigfuture.com/2009/01/petrobank-caprithai-processes-for.html

http://nextbigfuture.com/2009/01/petrobank-caprithai-processes-for.html

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