Fuel Processing Technology 92 (2011) 600–608

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Fuel Processing Technology

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Liquid fuel production from syngas using bifunctional CuO–CoO–Cr2O3 catalystmixed with MFI zeolite

Pravakar Mohanty a,c, K.K. Pant a,⁎, J. Parikh c, D.K. Sharma b

a Department of Chemical Engineering IIT Delhi India, 110016b Centre for Energy Studies IIT Delhi Indiac Departmnt of Chemical Engineering SVNIT Surat India, 395007

⁎ Corresponding author. Tel.: +91 1126596172; fax:E-mail address: kkpant@chemical.iitd.ac.in (K.K. Pan

0378-3820/$ – see front matter © 2010 Elsevier B.V. Adoi:10.1016/j.fuproc.2010.11.017

a b s t r a c t

a r t i c l e i n f o

Article history:Received 17 August 2010Received in revised form 29 October 2010Accepted 4 November 2010Available online 31 December 2010

Keywords:CuO–CoO–Cr2O3

MFI zeoliteCO hydrogenationLiquid fuelGTL

CuO–CoO–Cr2O3 mixed with MFI Zeolite (Si/Al=35) prepared by co-precipitation was used for synthesis gasconversion to long chain hydrocarbon fuel. CuO–CoO–Cr2O3 catalyst was prepared by co-precipitation methodusing citric acid as complexantwith physicochemical characterization byBET, TPR, TGA, XRD,H2-chemisorptions,SEM and TEM techniques. The conversion experiments were carried out in a fixed bed reactor, with differenttemperatures (225–325 °C), gas hourly space velocity (457 to 850 h−1) and pressure (28–38 atm). The keyproducts of the reaction were analyzed by gas chromatography mass spectroscopy (GC-MS). Significantly highyields of liquid aromatic hydrocarbon products were obtained over this catalyst. Higher temperatureand pressure favored the CO conversion and formation of these liquid (C5–C15) hydrocarbons. Higher selectivityof C5+hydrocarbons observed at lower H2/CO ratio and GHSV of the feed gas. On the other hand high yields ofmethane resulted,with a decrease in C5+ to C11+ fractions at lowerGHSV. Addition ofMFI Zeolite (Si/Al=35) tocatalyst CuO–CoO–Cr2O3 resulted a high conversion of CO-hydrogenation, which may be due to its large surfaceareaand small particle size creatingmore active sites. Thehomogeneity of various componentswas alsohelpful toenhance the synergistic effect of Co promoters.

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1. Introduction

The catalytic conversion of syngas to liquid hydrocarbons fuel andhigher alcohols is a potential route for providing clean fuel andpetrochemical feedstock [1,2]. Production of ultra clean fuel from coaland biomass derived synthesis gas (CO+H2) has received significantattention due to limited oil reserves, with increasing populations andeconomic developments. Renewable energy could be a sustainable andsignificant route to meet the future energy targets. The major part ofgas-to-liquids (GTL) technology is Fischer Tropsch synthesis whichconverts syngas into liquid fuel with awide range of liquid hydrocarbonfuel and high-value of added chemicals. Several types of catalysts havebeen used for the reaction [2]. However the process provides a widerange of hydrocarbons (C1–C60). Iron and cobalt based catalysts alongwith MFI Zeolites (ZSM-5) has been reported for potential commercialaspects because of their high catalytic activity and coke built tolerance[3–5]. A combination of an iron-based FT catalyst has shown highselectivity to olefins and oxygenates in presence of MFI zeolites or HYzeolites [2,4]. This activity enhanced the gasoline selectivity and anincreased concentration of high-octane branched and aromatic hydro-carbons by promoting oligomerization, cracking, isomerization, and

aromatization reactions on the MFI zeolites (Si/ Al=35) [2–6]. FTproduct distributions are generally controlled by the Anderson–Schulz–Flory (ASF) polymerization kinetics, resulting in a nonselectiveformation of any hydrocarbons [3–6]. In other dimension first step isto convert syngas to methanol over a methanol synthesis catalyst andsubsequently polymerize methanol to hydrocarbons over MFI Zeolite[7–9]. Many investigations have demonstrated the advantages of theone-stage processes by using bi-functional catalysts comparedwith thetwo-stage and three stage processes of synthesis gas conversion togasoline. Bifunctional catalysts have more potential to deliver betterperformance for syngas to higher hydrocarbons [10–14]. Zeolite has theacidic function most commonly used, due to its properties like activity,selectivity, limited deactivation, and thermal stability [15–21]. Howev-er, Cu–Co-based catalysts have been found effective for synthesis ofhigher alcohols produced from syngas [22–24]. In this study applicationand characterization of the CuO–CoO–Cr2O3 (CAT-A) mixed with MFIZeolite catalyst was studied for conversion of synthesis gas to higherhydrocarbon fuel. The catalyst was characterized by using thermogravimetric analyser (TGA), Surface area and pore volume analyser,X-ray powder diffraction (XRD), Temperature-programmed reduction(TPR), H2-chemisorptions, Scanning electron microscope (SEM), andTransmission electron microscopy (TEM). The effects of the operatingvariables on the catalytic performance of the (CAT-A+MFI zeolite)bifunctional catalysts were examined in a high-pressure fixed bedreactor.

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2. Material and methods

2.1. Experimental preparation of CuO–CoO–Cr2O3 (CAT-A)

All the chemicals used for catalyst preparation were of high purityand procured from Merck (Germany). Catalyst CuO–CoO–Cr2O3

(25:50:25 wt.%, CAT-A); was prepared by mixing a known amountof chromium nitrate (Cr (NO3)3 ·6H2O), copper nitrate, (Cu(NO3)2·6H2O), and cobalt nitrate, (Co (NO3)2·6H2O), in KOHsolution. Mixing of these three different metal precursors was carriedout in mechanical shaker; with deionized water. The solution waskept under constant stirring, at 65 °C for about 30 min. To thisprepared solution, citric acid was slowly added as a chelating agentand heated at 80 °C for 1 h followed by drying at 105±5 °C for 12 h.The addition of citric acid helps in dissolution of precipitated metalhydroxides and thereby resulting a homogeneous solution, becausethe organic ligand has a strong affinity towards Cu and Co ions duringthe solution preparation [11,12,17]. The foam-like solid thus obtainedwas calcined at 420±5 °C for 3 h in a muffle furnace. Blackish finepowder of CuO–CoO–Cr2O3 (1: 2: 1 mass ratio) was finally mixedwith MFI zeolite powder. The MFI zeolite (Si/Al=35) procured fromSüd-Chemie India (P) Ltd, was dried overnight at 105±2 °C andfollowed by crushing and sieving to 0.01±0.005 mm particle size.This MFI zeolite powder was physically mixedwith the CAT-A powder(dry mixing). The catalyst mixture thus obtained was a mixture ofCAT-A and MFI zeolite (Si/Al=35) in equal ratio and is designated asCAT-B. Finally the powder was pelletized followed by crushing todesired particle size for kinetic experiments.

Fig. 1. Schematic diagram of

2.2. Experimental set up for syngas conversion

The fixed bed reactor set up used was supplied by AutoclaveEngineers (BTRS). A schematic diagram of the experimental set up isshown in Fig. 1. The reactor is a stainless steel chamber of 12.6 mminternal diameter and 292.4 mm length, with a volume of catalyst upto 5 cm3 and is electrically heated. 3–5 g of the catalyst along withinert material silicon carbide, (SiC) was loaded in the reactor. Prior tothe reaction the reactor was kept under N2 pressure for the leak test,then the system was purged by flowing H2 for 30 min, followed bypre-reducing catalyst in a stream of H2–N2 for 4 h. Thereafter syngaswas introduced into the reactor. The flow rates of H2 and CO wereadjusted in the range of 40–100 (ml/min) using mass flow controllers.No temperature gradient was observed in the catalyst bed throughoutthe reduction process, but during the syngas conversion process, therewas some temperature rise observed due to the exothermic reaction.The analysis of the hydrocarbon was carried out with a Nucon gaschromatograph (AIMIL) provided with a thermal conductivitydetector (TCD) and a flame ionization detector (FID). The flow ratesof liquid and gaseous products were measured at regular intervals.The analysis of gaseous products like H2, CO, N2, CH4 and CO2, wasdone by Carbo-sphere column using Ar as carrier gas. All hydrocarbongases were analyzed by using a Porapak-Q column (i.d.: 2 mm, length:1.8 m). The liquid hydrocarbons were determined by using a capillarycolumn (petrocol fused silica column, L=60 m, i.d.=0.25 mm fusedsilica), on a FID mode. The oven, injector and detector temperaturewere maintained at 60, 120, and 220 °C respectively. Overall massbalance and carbon balance were performed based on inlet/outlet

the experimental setup.

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flows of gas composition analysis, and liquid product weight and runswith more than 95% carbon balance were considered for the analysisof the data.

2.3. Characterization of the catalysts

Final metal loading of the catalyst was determined by SEM–EDXanalysis using new ZEISS EVO Series Scanning ElectronMicroscope EVO50 and X-ray analysis. SEM EDX results are quantitatively presented inFig. 2 (A1, A2, B1) in which metals are identified by different peaksindicating the final weight % of metal loading on the catalyst. Theamount (wt.%) ofmetal loading of theCAT-Awas: Cu: Co: Cr (22:47:17).The metal content of the MFI mixed (1:1 w/w) catalyst was alsodetermined by SEM-EDX and the metal ratio was Cu–Co–Cr 13:24:10(rest Si and Al) due to the presence of Al and Si in MFI mixture. Thesurface areas and pore volumes of the fresh and used catalyst (CAT-Aand CAT-B) were analyzed using N2 adsorption. The BET surface area,total pore volume and pore size distribution of the catalyst wasdetermined from nitrogen adsorption/desorption isotherms measuredat −196 °C using Micromeritics ASAP 2010 apparatus [23,25]. Prior togas adsorptionmeasurements the catalystwasdegassed at180 °Cunderhigh vacuum for a period of 6 h. The total pore volumewas calculated ata relative pressure of approximately 0.99. TGA was used to analyzeweight loss during heat treatment and to estimate the amount of cokedeposited on the catalyst surface after reaction using a SDT Q600thermal analyzer of the precalcined catalysts. The TPR spectra wererecordedusing theMicromeritics PulseChemisorptionChemisorb 2720.TPR experimentswere carriedout at rampof 10 °C/min. The reactive gascompositions were H2 (10 vol.%; balance Ar). For the temperature-programmed reduction experiments, the catalyst sample was (0.5 g)placed in a quartz reactor and heated up to 750 °C. The metal surfacearea and metal dispersion was measured by H2 chemisorption usingMicromeritics Pulse Chemisorption Chemisorb 2720. For pulse chemi-sorptions 0.5 g of sample was taken in a quartz reactor, placed in a

Fig. 2. SEM-EDX analysis of catalysts: (A1)-SEM, (A2)-

furnace. Then the temperature of the furnacewas raised for reduction atramp of 10 °C/min under flow of 30 ml/min, using 9.9% H2 and rest Ar.After the reduction temperature was reached, under the flow of gasmixture, heating was stopped and the temperature was further cooleddown to 40 °C, and reduced catalyst sample was flushed with Ar, at aflow of 30 ml/min to remove specifically adsorbed H2. The reactor wascooled to 30 °C. At this temperature, 0.4 ml pulses of H2 injectedperiodically in the stream of nitrogen in order to find out the total H2

uptake. The cobalt surface area was calculated by assuming theadsorption stoichiometry of one H2 molecule per each fresh Co–Cr–Cusurface atom which can be shown in Eqs. (1) and (2) [26].

Percentage of Dispersion ¼ H2 uptake ⁎ atomic weight ⁎ stoichiometryð ÞPercentage of Metalð Þ

ð1Þ

¼number of Co0 atoms on surface ⁎ 100

� �

total number of Co atoms in sample ⁎ fraction reducedð Þ ð2Þ

XRD pattern was obtained using a Bruker-model operated at 40 kVand 20 mA using Cu Kα radiation with a wavelength of 1.54 A°. Thediffraction angle was varied from 20° to 80° at a scan rate of 0.02°/s.The average particle size of CoO in the calcined catalyst (CAT-A) wasestimated from Scherrer's equation using the most intense reflectionof CoO crystallites at different value of 2θ. The CoO, CuO and Cr2O3

particle size in the calcined sample were then converted to individualmetal diameter in the reduced catalyst by considering the relativemolar volumes of Coo and Co3O4 using the Eqs. (3–5) and followingthe concept for other metals [16–18,23].

d Co0� �

¼ 0:75 ⁎ d Co3O4ð Þ; : ð3Þ

EDX of CAT-A and (B1)-SEM, (B2)-EDX of CAT-B.

Table 1Physical properties of the catalysts.

Catalyst BET surfacearea (m2/g)

Pore vol.(cm3/g)

Avg. porediameter (Ao)

MFI zeolite (Si/Al-35) 276.6 0.29±0.05 6.6CAT-A 29.9 0.04±0.01 59CAT-B 138.5 0.22±0.04 65CAT-B — used 124.2 0.20±0.05 64

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The average size of Co3O4 particle was calculated according to theScherrer's equation [22].

d ¼ 0:89 ⁎ λ = B cosθð Þ ⁎ 1800= π

� �h ið4Þ

where d is the average crystallite diameter, λ is the wavelength ofradiation, and B is the full width half in degrees. Then the Coo metaldispersion was estimated from diameter of (Coo) by assuming aspherical geometry of the metal particles with uniform site density of14.6 atoms per nm2, using the following equation [10,16,26].

D = 96 = d½ � ð5Þ

where D is the % Coo dispersion and d is the mean particle size of Coo

in nm. The morphology of the catalyst sample was investigated usingSEM equipped with a new ZEISS EVO Series Scanning ElectronMicroscope EVO 50. The ZEISS EVO 50 is a versatile analyticalmicroscope with a large specimen chamber. Scanning ElectronMicroscope EVO 50 having the resolution 2.0 nm at 30 kV (SE withLaB6 option) with magnification range up to 1,000,000×. TEM wasperformed using PHILIPS CM12 microscope operated at an accelerat-ing voltage of 100 kV [17,18].

3. Results and discussion

3.1. Catalytic characterization

The surface area, pore volume, and average pore diameter of MFIZeolite (Si/Al=35), CuO–CoO–Cr2O3, (CAT-A) and CuO–CoO–Cr2O3+MFI Zeolite, (CAT-B) are reported in Table 1. The BET surface areas ofMFI Zeolite, CAT-A, fresh CAT-B, and used CAT-B are 276.6, 29.9, 138.5,and 124.2 m2/g, respectively. The surface area of the catalyst, CAT-Bdecreased after the reaction indicating that metal crystallite size

Fig. 3. TGA/DTA curves of pre-

increased slightly during the reaction [23,25]. The dried fine powderof pre-calcined CAT-A was analyzed using a SDT Q600 thermalanalyzer to determine the weight loss as a function of temperature.These TGA results are shown in Figs. 3 and 4. The reduction patternindicates that between 310 and 520 °C variation in wt.% wasapproximately 0.95%. Up to 460 °C calcinations it was sufficientto get oxidized and usage of CAT-A compound with MFI Zeolite(Si / Al=35) mixture. A total of 12% weight loss obtained duringcalcination indicates the decomposition of Cr(NO3)3 at 170±5 °C,and Cu(NO3)2 at 225±5 °C, and also due to the formation of CuCrO4

phase [27]. During calcinations it is easy to get different alloy shelland Co/ Cr spinel which is in agreement with the published literature[28–30]. Baker et al. [28] suggested that both copper and cobalt arerequired for higher alcohol synthesis and the presence ofmetallic cobaltis necessary either pure or alloyed with copper and chromium.

The reducibility of catalyst was examined by temperatureprogrammed reduction. Prior to temperature-programmed reductionexperiments, each catalyst sample was activated under flowing O2

(10 vol.%; balance He) at 600 °C for 60 min. Fig. 5 illustrates the TPRprofiles of two catalysts where a large peak centered at 334 °C in CAT-A, and at 336 °C in CAT-B was observed. Such a large peak can beattributed mainly to the reduction of Co3O4 species. Anotherreduction peaks was observed for CAT-B could be attributed to thereduction of Co3O4 to CoO and CoO to metallic CoO [16,23].Incorporation of Cu metal on the catalyst increased the reductiontemperature from 250 °C to 334±2 °C. Tien-Thao et al. [23] suggestedthat the copper provides an active site for higher alcohol synthesis, toactivate cobalt reduction, and possibly to modify the cobalt byalloying. After reduction of highly dispersed CuO and CoO aggregates,formation of small Cu–Co clusters occurs. Xiaoding et al. reported thatCu–Co alloy is responsible for higher alcohol synthesis [11]. Baker etal. proposed that the active centre is comprised of metallic cobaltatoms on copper clusters and the active catalyst consisted of copperparticles supported on Co/Cr spinel [28–31]. The coppermetal in thesecatalysts is less sensitive to atmospheric oxidation, as there is noseparate peak at lower temperatures for the reduction of copperhowever the copper metal remains in ionic states during reductionstep and reaction step in presence of CO and H2 [2]. The adsorptioncapacity of fresh CAT-A was examined during temperatureprogrammed reduction. Prior to temperature programmed reductionexperiments and before chemisorptions, argon was passed for 1 hfollowed by cooling from 750 °C to 40 °C at 30 ml/min. After reaching40 °C, H2 gas circulated in the analysis port for dosing purpose for TCD

calcined catalyst (CAT-A).

Fig. 4. TGA/DTA analysis of calcined catalyst (CAT-A).

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graphs [10,16,26]. The total metal dispersion as determined from thechemisorptions study was 3.6%. The average particle diameter of themetal crystallite as determined by H2 chemisorptions and XRD wasapproximately 20±3 nm. The intrinsic activities (TOF) for the FTsynthesis at 275 °C for the two catalysts (CAT-A and CAT-B) were7.4×10−3 s−1 and 18.2×10−3 s−1 respectively. The values are ingood agreement with (TOF=28.5×10−3 s−1) reported by Martinezet al. [16]. It must be stated that the TOF of the catalyst mostlydepends on the preparation method, different particle sizes and theoverall dispersion to achieve active metal sites.

Fig. 6 (A and B) shows the XRD pattern of catalyst samples. TheXRD peaks at (A) corresponding to Cu (JCPDS-04-0836) at 2θ of 43.44°and 50.44°. Peaks at 38.79° and 74.15° corresponds to Co3O4, whichwere observed (JCPDS-74-1657) in the fresh catalyst sample [2,13].Cobalt and chromium were found as Co3O4 and Cr2O3 crystallites inthe fresh and used sample as per (JCPDS-82-1484), Cu2O at (2θ)73.96° and 59.8° (JCPDS-03-0892)were clearly detected by XRD in theused catalyst samples in Fig. 6B XRD of used samples indicates thatCo3O4 was converted to CoO after reaction, whereas the reflectionlines of Cr2O3 are unchanged after reaction. Fig. 6B indicates thatCr2O3 was not reduced under different reaction conditions [13,18,23].Some phases like CrO3, CuCrO4 and CoCr2O4 [29,30] were also formedin fresh and used catalyst getting conformation from (JCPDS-77-0125,77-2474 and 78-0711) respectively.

Fig. 5. TPR analysis of CAT-A and CAT-B after calcination.

Fig. 7 (A–B) shows a typical SEM photograph of the calcinedcatalyst before and after reaction respectively. It is found that thecatalyst surface is covered with small crystallites of cobalt oxide(Co3O4), copper and chromium particles, which is in agreement withthe reported literature [17]. It is apparent that both CAT-A and zeolitein the fresh catalyst sample are uniformly distributed, i.e., zeoliteparticles are evenly distributed and covered by porous CAT-A oxides.In addition, the size of these small crystallites grew larger due tocrystallite agglomeration with MFI zeolite (Si/Al=35) during pellet-ization and heat temperature during reaction. Therefore the agglom-eration of individual Cu–Co–Cr metal oxide particles with MFI zeoliteand the formation of larger Cu particles may be a factor of decreasingthe catalytic performance, which is in agreement of the published data[17]. Thus high temperature causes agglomeration of small crystal-lites, which leads to catalyst deactivation under high temperaturereaction [31,32].

Cu–Co–Cr and MFI zeolite were well mixed in the fresh conditionsand resulting particles were both porous and nano-sized. TEM imagesof fresh CAT-A after calcination and the used catalyst CAT-B (after120 h of reaction under 225 to 325 °C) are depicted in Fig. 8 (A and B)respectively. In thesemicrographs, themorphology of the particle wasobserved spherical. The particles of fresh CAT-B are fairly uniformwith an average diameter of 18±3 nm and whereas for the used CAT-B it was 25±3 nm in size. The TPR results indicates only CuO hasreduced at temperatures below 423 °C; thus it is hypothesized thatthese distinctive particles are reduced metal copper which haspartially separated from the Co–Cr oxides after 120 h of reaction [17].Fig. 8 (A–B) shows both metal oxide phases are well fitted into eachother and metal grains are in the range of 15–30 nm. TEM Figuresclearly illustrates the crystalline nature of the sample. The samples arespastic structures, not single-crystalline structures, respectively.These are located close to the surface, their outer surface beingcoated by an amorphous or a poorly crystallized oxide which protectsthem from rapid oxidation in the presence of hydrogen at highertemperature. This tendencywith catalyst may be due to the compositematerial combination with MFI zeolite.

3.2. Syngas conversion to liquid hydrocarbons

The performance of the catalyst was investigated for syngasconversion to liquid fuel. Effect of temperature, reaction pressure, gashourly space velocity (GHSV) and CO/H2 ratio was studied in the fixedbed reactor set up [25]. The experiments were carried out at different

Fig. 6. XRD analysis of (A) fresh CAT-B and (B) used CAT-B.

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temperatures in the range 225 °C to 325 °C. The total pressure of thereaction was varied from 28 to 38 atm pressure with GHSV keptwithin 457 to 850 h−1 and the mole ratio of CO/H2 at 1 to 2. COconversion and product selectivity (%) on carbon basis was calculatedusing Eqs. (6) and (7) respectively [24,33].

CO conversion %ð Þ ¼ Moles COð Þin� Moles COð ÞoutMoles COð Þin

× 100 ð6Þ

Selectivity of product %ð Þ ¼ Moles of productð ÞMoles CO Reactedð Þ × 100 ð7Þ

CO conversion increased from 28% at 225 °C to 72% at 275 °Cand decreased to 68% at 325 °C. Although temperature favors thekinetics of hydrocarbon formation reactions, according to thethermodynamics only the light hydrocarbons, like methane, showshigher stability in these conditions. It shows the higher temperature

Fig. 7. SEM images of (A) CAT-B after calcination and (B) used CAT-B.

up to 275 °C is favorable for the CO conversion. The decrease inconversion, as temperature increased above 275 °C is probably due tothe thermodynamic restrictions of the exothermal reaction andcopper metal surface sintering, which leads to partial loss of catalystactivity [9]. The hydrocarbon concentration in the effluent increasedwith temperature. Hydrocarbon distribution toward gasoline (C10+to C15+) first increased as the reaction temperature increased from225 to 275 °C, and then decreased after 275 °C [24]. The amount ofcoke deposited on used catalyst was also determined after 120 h ofrun time. The deposited coke on the catalyst was estimated to be 2.8%by TGA analysis.

3.2.1. Effect of CO/H2 molar ratio on conversion and hydrocarbon yieldThe CO / H2, molar ratio plays a vital role in higher hydrocarbon

synthesis. Conventional syngas has high CO to H2 ratio as compared tonatural gas synthesized syngas. It can affect both reaction rates andactivity as lower H2/CO ratio tends to decrease CO conversion whileincreasing higher hydrocarbon selectivity. The feed H2/CO molar ratioof 2:1 was kept and its effect on CO conversion and hydrocarbondistribution has investigated. In Fig. 9 the CO conversion andhydrocarbon formation are compared with a H2/CO molar ratio of 1where CO conversion was 72% at 275 °C and then it decreased whenH2/CO molar ratio increased to ~2. Also significant changes in productwere obtained. A change in the ratio also affects the partial pressure ofH2 and CO under constant total pressure, thus the stoichiometry willbe different for each reaction. Higher hydrogen partial pressurereduces the coke formation, thus better activity of the catalyst wasobtained. Also at high carbon monoxide partial pressure, theformation of high molecular weight compounds increased. At lowerH2/CO ratio, CO insertion and C–C chain growth are more favorable,which enhanced the production of higher hydrocarbons [4,33].Therefore an optimum H2/CO ratio is desired for the liquid fuelhydrocarbon production.

Fig. 9 compares (C5+ range hydrocarbon) product distributions attwo different molar ratios of H2/CO. Higher H2/CO ratio favored theformation of straight chain paraffin products where as at lower H2/COratio of ~1, the formation of higher C/H ratio products (aromatics)favored. At higher H2/CO molar ratio the C5+ fraction increased at theexpense of the lighter fractions in the product stream. This is probablydue to the fact that a decrease in the H2/CO molar ratio favors thecondensation steps over the hydrogenation steps [9,20]. Consequentlyproducts of higher range of liquid hydrocarbons were formed andselectivity, toward hydrocarbon increased. Also olefin formation

Fig. 8. TEM images of (A) fresh CAT-A after calcination and (B) used CAT-B.

Fig. 10. Effect of temperature on CO conversion.

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suppressed producing higher paraffin and iso-paraffin. At (H2/CO=1–1.5) aromatic formation gradually increases, which is in accor-dance with the results reported in literature by Li. et al [7,20].

3.2.2. Effect of reaction temperature on conversion and hydrocarbonyield

The effect of temperature on CO conversion and product (C1+ toC15+) distribution are presented in Figs. 10 and 11 respectively. Theexperimental conditions were as follows: P (28–38 atm), GHSV (457to 850 h−1) at molar ratio of H2/CO of 1: 1. As shown in Fig. 11 thehydrocarbon distribution was affected by the change in temperature,higher hydrocarbons (mainly C5+) being favored at relativelymoderate temperatures. At low temperature formation of lower

Fig. 9. Effect of H2/CO molar ratio on higher hydrocarbon distribution at P=38 atm,T=275 °C and GHSV=585 h−1.

hydrocarbons favored mainly methane. The higher hydrocarbons aremore sensitive to temperature, and therefore undergo cracking andde-alkylation reactions. The identification of different key compo-nents synthesized during reaction was analyzed by GC-MS (Model:Perkin–Elmer clarus-600). These distinct identified peaks are sum-marized in Table 2 and programmed GC-MS graph has shown inFig. 12. Selectivity toward gasoline (C5+ fraction) first increased up to275 °C and decreases thereafter. It is reported that, the activationenergy values for condensation steps are smaller than thosefor hydrogenation of light hydrocarbons [20]. With respect to thehigher hydrocarbons and aromatic products (toluene, xylene, 1, 2,3-trimethylbenzene, and 1, 2, 4, 5-tetramethylbenzene), the C5+fractions increased and the C8+ fraction reach a maximum value at275 °C and then decreased with increase of temperature [34–36].

3.2.3. Effect of GHSV on conversion and hydrocarbon yieldExperimentswere carried out at different GHSV ranging from 457 to

840 h−1 with other operating conditions were kept constant at 275 °C,38 atm. and molar ratio of CO/H2~1. The effect of the space velocity onCO conversion and hydrocarbon distribution indicates that as spacevelocity increases CO conversion decreases, which has shown in Fig. 13.It is observed that the hydrocarbon selectivity increased up to a spacevelocity of 585 h−1 and decreased thereafter with further increase inspace velocity. The liquid hydrocarbons fraction (C8+) progressivelydecreased with increasing space velocity. Lower GHSV favoredhigher C/H ratio products, especially the aromatics like ethylbenzene,o-xylene, 1-methyl-2-ethylbenzene, naphthalene, styrene, n-decane,n-dodecane etc.

Fig. 11. Effect of temperature on liquid hydrocarbon distributions at P=38 atm,(H2/CO=1), and GHSV=585 h−1.

Table 2GC-MS analysis data.

GC-MS spectra Possible compounds

Peak-3.10 Aromatic hydrocarbons (Benzene, Ethylbenzene)Peak-5.64 Aromatic hydrocarbons (Toluene, O and P-xylene)Peak-8.20 Methyl Phenols, Benzyl alcoholsPeak-10.13 n-propylbenzenePeak-12.08 Cyclobenzene, CyclohexanePeak-13.74 Decane, n-undecane, n-tridecanePeak-15.17 Methyl benzenamines, AlcoholsPeak-16.46 Ploy aromatic hydrocarbonsPeak-17.91 Undecane, dodecanePeak-19.23 n-pentadecane, Decane, DodecanePeak-20.41 Eicosane, n-pentadecanePeak-21.48 Eicosane, n-pentadecane

Fig. 13. Effect of pressure and GHSV on CO conversion.

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3.2.4. Effect of reactor pressure on the conversion and product yieldAccording to the thermodynamics an increase in total pressure

increases the equilibrium towards the product side, and thusincreases the conversion. Pressure is always favorable for Fischer–Tropsch bi-functional catalysts [1,2]. The change in Gibbs free energyis always negative when pressure increases. Fig. 13 depicts the effectof change in total pressure and GHSV on CO conversion. Experimentsunder pressure of 28, 32, 35 and 38 atm., has been carried out in thepresent study. The other operating conditions were as follows:temperature=275 °C; CO / H2 molar ratio of 1; and gas spacevelocity=585 h−1. The increase in pressure favored the formation ofC8+ to C11+, while methane and C1–C3 gaseous hydrocarbonsdecreased with an increase in pressure. The primary reaction duringFT synthesis is the formation of α olefin which hydrogenates inconsecutive step to the corresponding paraffin, so pressure hassignificant effect on gaseous olefin to paraffin ratio. Furthermore thechange in total pressure did not affect the water gas shift reaction butchanged the equilibrium of product distribution and CO conversion inCO-hydrogenation.

Thus bifunctional transition metal oxide with MFI Zeolite catalystshave been found capable of converting syngas to liquid fuels in asingle reaction step as they contain both CO hydrogenation function(Cu–Co–Cr oxides) and hydrocarbon synthesis function. The liquidproduct obtained also include some groups of alcohols, ketones, andaldehydes as identified by GC-MS.

The basic reaction for the formation of these hydrocarbon productsis –CH2-group. Furthermore, a change in total pressure does not affectthe water gas shift reaction but changes the equilibrium productdistribution [1,4,22]. In general there are two steps involved in theconversion of syngas to hydrocarbons over a bifunctional catalyst.First, alcohols are synthesized from the syngas over an oxide catalyst.In the second step zeolite catalyst dehydrates these alcohols into

Fig. 12. GC-MS analysis of liqu

olefins and rearranges the remaining hydrogen and carbon atoms intoa concentrated high energy fuel [1–4]. Effect of temperature revealedthat aromatic fractions increased with temperature. Thus there is atradeoff between two trends: the CO hydrogenation reaction yieldingalcohols, mainly due to the Cu–Co–Cr metal oxides, with an optimumtemperature of 250–300 °C and the acid function-based reaction i.e.,conversion of oxygenates to hydrocarbons or aromatization, mainlydue to the MFI Zeolite catalyst [21]. The synthesis rates for bothalcohols and hydrocarbon are dependent on the syngas pressure overCu–Co–Cr metal oxide catalyst. Therefore, increasing syngas partialpressure results in an increased productivity of almost all productsexcept of CO2, which is not very sensitive to the change in syngaspressure. In the present study, a stoichiometric H2/CO ratio of oneseems to be an appropriate syngas composition for the synthesis ofhydrocarbons [25,26]. Cu–Co–Cr based catalysts are utilized for theproduction of higher alcohols. The chain growth of higher alcoholformation is favored at high cobalt concentration or low coppercontent. Mainly the cobalt ions in combination with metallic Cu or theinteraction phase are important to oxygenate synthesis, as it isimportant to carry out reduction with H2 at 250 °C, before syngaspassing inside Cu–Co interaction, which are the active site for thereaction. XRD results revealed that metallic copper and CoO and Cr2O3

are active components for CO hydrogenation. In higher alcoholsynthesis, copper sites are responsible for hydrogen dissociativechemisorption and CO associative adsorption since these have beenreported as the main elements in methanol synthesis [1,2,4,13].

id hydrocarbon product.

608 P. Mohanty et al. / Fuel Processing Technology 92 (2011) 600–608

4. Conclusions

CuO–CoO–Cr2O3 (CAT-A) mixed with molecular shape selectiveacid function zeolite (Si/Al-35) has been found effective for theconversion of synthesis gas to liquid hydrocarbons. Use of citric acid ascomplexant during catalyst preparation, has partly reduces thesamples and decrease the particle sizes. High yield of liquidhydrocarbons was observed over (CAT-A) catalyst mixed with MFIzeolite during syngas conversion. Temperature and pressure are themost significant variables affecting the C5+ yield and selectivity. COconversion increases from 28% at 225 °C to 72% at 275 °C; butdecreases thereafter. Selectivity towards gasoline (C5+ fraction)range of hydrocarbons also increased in the temperature range from225 to 275 °C. Higher the H2/CO ratio is favorable for formation ofproducts of lower C/H ratio (straight chain of paraffin), while lowerH2/CO ratio is favorable for higher C/H ratio products (aromatics).

Acknowledgments

The financial support provided by Department of Science andTechnology, Government. of India for sponsoring this research projectis gratefully acknowledged.

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