DOI: 10.1002/cctc.201402434

Catalytic Steam Gasification of Cellulose Using ReactiveFlash VolatilizationFan Liang Chan and Akshat Tanksale*[a]

Introduction

Lignocellulosic biomass is a renewable energy resource thatcan be derived from organic sources such as energy crops, ag-ricultural residues, forestry residues and recycled cardboardand paper.[1] Biomass utilisation for fuels and energy produc-tion is recognised as one of the most promising solutions forthe energy crisis and anthropological CO2 emission problems.[2]

Thermochemical processes for biomass conversion, such ascombustion, gasification and pyrolysis, can be used for powergeneration and biofuels production.[2] Among the biomass-conversion technologies, gasification is one of the most inter-esting technologies from both an industrial and academic re-search point of view because of its high conversion efficiency.[3]

Biomass gasification can be achieved at temperatures in excessof 700 8C in the presence of oxygen or air with or without ad-ditional steam. However, at this temperature, a significantamount of condensable oxygenated hydrocarbons, commonlyreferred to as tar, is produced. In the absence of catalysts, tar-free gasification requires higher temperatures (�1000 8C).[4]

Tar removal is a major hurdle that hinders the commerciali-sation of biomass gasification.[5] Many factors can affect theamount and type of tar formed during gasification. These fac-tors include gasifier type and design, operating parameters(temperature, pressure, heating rate and residence time), typeof feedstock and the type of catalyst used. The optimisation ofthese factors may maximise the efficiency of gasification withminimum tar formation. Dolomite- and CeO2/SiO2-supportedNi, Pt, Pd, Ru and alkali metal oxides have been used in thepast to catalyse gasification reactions, reduce tar formation, im-prove conversion efficiency and improve the product gas

purity.[6] As Ni-based catalysts are used industrially for thesteam reforming of methane and naphtha,[7] they are expectedto catalyse the steam reforming of tars and the subsequentwater–gas shift reaction to produce H2. However, monometallicNi catalysts suffer from rapid deactivation because of coke for-mation if they are used as primary catalysts in fluidised-bedgasifiers.[6c, 8] The doping of Ni catalysts with a small amount ofnoble metal increases its reforming activity, reducibility andcoke resistance.[9]

Conventionally, biomass gasification is performed in one ofthe following reactor setups:

* Fluidised-bed gasifier with a downstream catalytic tar-clean-ing reactor

* Fast pyrolysis reactor with a downstream catalytic steam re-former

* Catalytic fluidised-bed gasifier* Entrained-flow gasifier

Fluidised-bed reactors are capital-intensive and hence large-scale reactors are required to make them economical. The cap-ital cost of the gasifier and gas-cleaning system can accountfor 66.7 to 85.5 % of the total capital cost.[10] Entrained-flowgasifiers are more expensive because of their higher flow ve-locities and temperatures in excess of 1000 8C, which requireNi-based alloys (e.g. , Hastelloy) to provide oxidation resistanceand high-temperature strength. In comparison, the capital costof fixed-bed reactors for biomass gasification is significantlylower. Therefore, a new approach called reactive flash volatili-sation (RFV) was proposed recently for cellulose gasification.[11]

RFV uses a fixed-bed gasifier with a carbon space velocityand carbon mass flow rate 10–100 times higher than that offluidised-bed reactors.[11a] As a result, RFV reactors require sig-nificantly less catalyst and a smaller reactor volume to process

Biomass gasification is considered to be one of the most prom-ising technologies to deliver renewable energy. However, tarformation in the gasifier is one of the main challenges. Ni-based catalysts are one of the most effective transition-metalcatalysts in biomass gasification for tar cracking and reforming.Alumina-supported Ni, Pt-Ni, Ru-Ni, Re-Ni and Rh-Ni catalystswere tested for their activity in the reactive flash volatilisation(RFV) of cellulose to produce synthesis gas in 50 ms reactiontime. RFV is a catalytic gasification process that utilises a high

carbon space velocity and mass flow rate with oxygen andsteam as gasification agents. Re-Ni, Rh-Ni and Ru-Ni supportedcatalysts showed a higher gasification efficiency than the othercatalysts, which was because of their higher metal surfacearea, high reducibility and lower CO desorption temperature.The highest gasification efficiency was achieved at 750 8C witha carbon-to-oxygen ratio of 0.6 and a carbon-to-steam ratio of1.0 without any oxygen breakthrough.

[a] F. L. Chan, Dr. A. TanksaleDepartment of Chemical EngineeringMonash UniversityClayton, VIC 3800 (Australia)E-mail : akshat.tanksale@monash.edu

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a unit mass of carbon feedstock compared to fluidised-bedgasifiers. Therefore, RFV reactors can be economically viable insmall-scale operation, which is useful for distributed or decen-tralised biomass-processing facilities. The advantage of decen-tralised facilities is that the cost of transporting large volumesof low-energy-density biomass is minimised by placing thesefacilities closer to the source of biomass.[12]

Similar to conventional gasification, the chemistry of RFV iscomplex and yet to be understood fully. However, it is general-ly accepted that the major reactions include pyrolysis, oxida-tion, partial oxidation, reduction, steam reforming and water–gas shift reactions.[5, 13] Side reactions such as methanation mayalso occur but to a lesser extent.[14] These reactions are sum-marised in Table 1.

In RFV, the formation of char is avoided by oxidising the cel-lulose pyrolytic products rapidly into gases by steam reformingand partial oxidation reactions. The heat generated from theoxidation reaction is used in the steam reforming reaction,which makes the RFV process autothermal. A conversion of99 % of the feed at �70 % hydrogen selectivity can be ach-ieved without catalyst deactivation.[11b] However, so far the RFVstudies have been performed at temperatures >800 8C and byusing supported Rh-Ce catalysts.[11] There is a need to developa low-cost catalyst that is active and selective at low tempera-tures to make this process economically attractive. In thisregard, promoted Ni catalysts may play a vital role.

Results and Discussion

Catalyst characterisation

Nitrogen physisorption

The specific surface area, total pore volume and pore size ofthe catalysts prepared in this project are summarised inTable 2. The BET surface area of the commercial alumina sup-port used in this study was 101.62 m2g�1, which decreased byvarying amounts for the impregnated catalysts. This is expect-ed to be because of pore blockage caused by impregnatedmetals. Among the impregnated catalysts, the Ru-Ni catalysthad the highest specific surface area of 95.74 m2g�1, and theRe-Ni catalyst had the lowest surface area of 43.62 m2g�1. This

is because the Re-Ni catalyst was not calcined beforethe BET measurement was performed, therefore, themetal precursors are expected to remain on the alu-mina surface and block a large proportion of thepores. After reduction, the BET surface area of theRe-Ni catalyst increased to 74.16 m2g�1.

X-ray fluorescence spectroscopy

The results of the X-ray fluorescence (XRF) elementalanalyses of the catalysts are presented in Table 2.The NiO and Al2O3 contents were measured directlyin the analyses, and the promoter metals contentwas calculated based on mass balance. The XRF re-sults and calculated promoter contents for Ni, Pt-Ni,Ru-Ni and Rh-Ni are in agreement with the nominal

values, which indicates the effectiveness of the catalyst prepa-ration procedure.

H2 temperature-programmed reduction

The reducibility of Ni catalysts is an important factor to deter-mine the activity of the catalyst. It is known that Ni0 is activefor steam reforming and the water–gas shift reaction, whereasNi2+ and NiO are not active. H2 temperature-programmed re-duction (TPR) was conducted to test the reducibility of pre-pared catalysts (Figure 1). The monometallic Ni catalyst hadthe highest onset reduction temperature of �420 8C, and the

Table 1. Gasification reactions and their enthalpies for C6 compounds.

Reaction Stoichiometric equation DH0r

(T=27 8C, x = 6)

pyrolysisCxHyOz ! (1�x) CO+

y2 H2+C (1) 180

CxHyOz ! (1�x) CO+ðy�4Þ

2 H2+CH4 (2) 300

partial oxidationCxHyOz+

12 O2 ! x CO+

y2 H2 (3) 71

CxHyOz+O2 ! (1�x) CO+CO2+y2 H2 (4) �213

CxHyOz+2 O2 ! x2 CO+

x2 CO2+

y2 H2 (5) �778

steam reformingCxHyOz+H2O ! x CO+y H2 (6) 310CxHyOz+n H2O ! a CO+(x�a) CO2+y H2 (7) 230CxHyOz+(2 x�z) H2O ! x CO2+(2 n+

y2�z) H2 (8) 64

water–gas shift CO+H2O ! CO2+H2 (9) �41methanation CO+3 H2 ! CH4+H2O (10) �206

Table 2. Comparison of SBET, total pore volume, amount of active metal, metal dispersion, NiO, Al2O3 and promoter contents of the investigated catalysts.

Catalyst Nitrogen physisorption CO chemisorption XRF spectroscopySBET Vpore (BJH) Irreversible CO Metal disper- NiO [wt %] Al2O3 [wt %] Promoter [wt %][m2 g�1] (P/P0) [cm3 g�1] uptake [mmol g�1] sion [%] Theoretical Measured Theoretical Measured Calculated[a]

Al2O3 101.62 0.18 n.d. n.d. – – – – –Ni 91.99 0.21 12.32 0.66 13.59 15.13�0.02 86.41 84.09�0.07 –Pt-Ni 77.18 0.14 67.88 3.87 12.37 13.25�0.02 86.53 84.30�0.05 1.37Ru-Ni 95.74 0.23 41.96 2.33 12.36 11.96�0.07 86.44 86.95�0.05 1.09Re-Ni[b] 43.62 0.07 78.92 4.49 12.36 10.25�0.08 86.43 50.42�0.18 n.d.Rh-Ni 77.92 0.15 37.67 2.09 12.36 14.99�0.08 86.41 83.72�0.10 1.29

[a] Calculated based on mass balance. [b] Uncalcined catalyst.

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reduction was not complete at 800 8C. The low reducibility ofthe Ni catalyst is attributed to the formation of nickel alumi-nate (NiAl2O4) spinel[15] if Ni supported on alumina is calcinedat or above 500 8C.[16] All the promoted catalysts showeda lower onset reduction temperature. The onset reductiontemperature increased in the order: Ru-Ni (165 8C) <Pt-Ni(240 8C)�Re-Ni (240 8C)<Rh-Ni (362 8C)<Ni (420 8C). It isknown from the literature that noble metals promote the re-duction of NiO through the surface migration of a chemisorbedhydrogen atom, a phenomenon known as the hydrogen spill-over effect.[17] The most promising catalyst was Re-Ni, whichwas reduced almost fully at 800 8C.

The total hydrogen consumed by each catalyst in H2-TPRwas calculated by integrating the area under the curve usingthe trapezoidal method (Table 3). As the Re-Ni catalyst was not

calcined before the H2-TPR measurement, the precursors of Reand Ni present on the surface of the catalyst react with hydro-gen through the reactions shown in Equations (11) and (12)and lead to a high hydrogen consumption.[18]

2 NH4ReO4þ7 H2 ! 2 Reþ8 H2OþNH3 ð11Þ

NiðNO3Þ2þ6 H2 ! NiþN2þ6 H2O ð12Þ

CO chemisorption

The results obtained from CO chemisorption are listed inTable 2. The Re-Ni catalyst had the highest amount of irreversibleCO uptake, and therefore, the largest metal surface area and dis-persion. Overall, all the promoted Ni catalysts exhibited a highermetal surface area than the monometallic Ni catalyst. This is be-cause of the higher reducibility of the promoted catalyst andthe higher dispersion of high-atomic-weight noble metals.

CO temperature-programmed desorption

The CO temperature-programmed desorption (TPD) profiles ofall the catalysts used in this study are shown in Figure 2. COchemisorption is not activated on Ni and noble metal activesites,[19] therefore, at room temperature (�21 8C) full coveragecan be achieved easily. CO desorption at elevated tempera-tures is a complex process that is not well understood.[18]

Nonetheless, it is well known that the rate of desorption isa function of heat and activation energy of adsorption, andlow activation energy sites exhibit a higher rate of CO desorp-tion at lower temperatures. The profiles for CO desorption areshown in Figure 2 a, whereas Figure 2 b shows profiles for CO2

Figure 1. Comparison of the TPR profiles of the investigated catalysts: ~: Ni,*: Pt-Ni, !: Ru-Ni, &: Re-Ni, ^: Rh-Ni.

Table 3. Comparison of the amount of H2 consumed and amount of COand CO2 adsorbed on each catalyst.

Catalyst H2 consumed inTPR [mmol gcat

�1]CO desorbedin TPD

CO2 desorbedin TPD

CO+CO2 desor-bed in TPD

Ni 124 31.98 1637 1669Pt-Ni 163 33.45 1025 1058Ru-Ni 306 29.01 1789 1818Re-Ni 344 22.41 2124 2146Rh-Ni 107 31.49 2501 2533

Figure 2. a) CO-TPD profiles and b) CO2-TPD profiles of the investigated cata-lysts. ~: Ni, *: Pt-Ni, !: Ru-Ni, &: Re-Ni, ^: Rh-Ni.

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evolution, which is formed from the dehydroxylation of alumi-na and subsequent reaction with desorbed CO.[9a, 20]

The low-temperature CO and CO2 desorption peaks thatappear at 120 8C for the Ni, Ru-Ni and Re-Ni catalysts were at-tributed to the weak multilayer chemisorption of CO over thecatalyst. A low-temperature CO2 peak was also observed forthe Rh-Ni catalyst but the corresponding CO peak was not ob-served, which suggests that Rh-Ni promotes the rapid conver-sion of CO into CO2 at �120 8C.

The second CO desorption peak observed at a higher tem-perature of 275–300 8C can be attributed to the strong mono-layer adsorption of CO on the metal surface that requiresa higher temperature and energy to desorb. Broad CO and CO2

desorption peaks were observed for Ni, Pt-Ni and Rh-Ni cata-lysts in this temperature range (275–320 8C), whereas a should-er peak was observed for the Re-Ni catalyst in this temperaturerange. The CO uptake of each catalyst was quantified by inte-grating the area under the curves for CO and CO2 desorptionusing the trapezoidal method. Re-Ni and Rh-Ni showed thehighest amount of CO+CO2 desorption in that order (Table 3).

In summary, from the results of the characterisation of thecatalysts used in this project, it can be observed that com-pared to the monometallic Ni catalysts, the promoted Ni cata-lysts had a significantly lower reduction temperature, higherdispersion and higher amount of active metal sites. This is ex-tremely important, because only a small amount of noblemetal was used in this study, which resulted in a significant im-provement in the catalyst properties. Therefore, it is expectedthat the promoted Ni catalysts will be more active for the RFVrun. Overall, the Re-Ni catalyst showed the highest metal sur-face area from H2-TPR, CO-TPD and CO chemisorption.

Reactive flash volatilisation of cellulose

Effect of gasification temperature

The catalysts were tested for their activity and stability in RFVat three operating temperatures (700, 750 and 775 8C) by keep-ing the carbon-to-oxygen (C/O) and carbon-to-steam (C/S)

ratios constant at 0.5 and 1.0, respectively. The temperaturerange for gasification was selected based on the thermody-namic analysis of Colby et al. who used HSC chemistry soft-ware to demonstrate that the CO selectivity increased with theincreasing temperature, whereas the H2 selectivity peaked ataround 700 8C.[11a] A significant amount of char was producedat 700 8C with all the catalysts (Figure 3), except Re-Ni, whichled to the eventual clogging of the reactor (reduction in theproduct gas flow rate). The Re-Ni catalyst showed a char selec-tivity of only 1 %, whereas all other catalysts showed a char se-lectivity of 19–35 %. The gasification efficiency increased withthe increasing temperature, whereas the gas selectivity wasthe highest at 750 8C. Re-Ni, Ru-Ni and Rh-Ni catalysts per-formed best at this temperature in terms of a high gas yield(in that order) and no char. The composition of synthesis gason a dry basis produced in these runs is shown in Figure 4. Ingeneral, all the catalysts showed a high selectivity to CO overmethane and C2 compounds as a result of the high activity forsteam reforming reactions.

A low char selectivity and high gasification efficiency wereobserved at 750 and 775 8C because steam reforming and par-tial oxidation reactions are favourable at higher tempera-tures.[21] These reactions convert the char and tar into gaseousproducts. It is known from the literature that the cellulose par-ticles undergo oxidative pyrolysis on the hot surface of cata-lysts to form a film of bio-oils that undergo catalytic steam re-forming and partial oxidation.[11a] Therefore, high temperaturesled to the complete conversion of cellulose into gases withsmall amount of unconverted volatile organics. However, forthe Ni and Ru-Ni catalysts, temperatures higher than 750 8Cwere not favourable because of the sintering of metal particlesat this temperature.[22] The sintering of active metal reducesthe metal surface available for reactions and, therefore, re-duces the gasification efficiency. Ni catalysts supported on alu-mina also suffered from the transformation of NiO/Al2O3 toNiAl2O4 spinel at temperatures above 750 8C.[23] A noble-metalpromoter is able to suppress this phase transformation of Nicatalyst into spinel by keeping the Ni in a reduced state by thehydrogen spillover mechanism.[9c, 24]

Figure 3. Effect of RFV temperature and catalyst promoters on the selectivity of char (&), tar (&) and gas (&) based on carbon balance. The numbers on thetop of each bar represent the gasification efficiency, which is the percentage of carbon in the gas and tar combined. Reaction conditions: cellulose flowrate = 15 g h�1, C/O = 0.5, C/S = 1.0 and residence time = 50 ms. Here, tar is defined as water-soluble organics.

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The H2/CO and CO/CO2 ratios are shown in Figure 5 a and b,respectively. These ratios are important for the downstreamapplication in biofuels production. For example, for synthesisof dimethyl ether a H2/CO ratio close to 1 is desirable becauseof the overall stoichiometry of DME synthesis from CO and H2

[Eq. (16)] . Whereas, for methanol synthesis, a H2/CO ratio of 2.0is desirable [Eq. (13)] . Therefore, the ability to tune the H2/COratio by changing the reaction parameters is an advantage in

this process [Eqs. (13)–(16)] .

2 ðCOþ2 H2 Ð CH3OHÞ ð13Þ

2 CH3OHÐ CH3OCH3þH2O ð14Þ

COþH2OÐ H2þCO2 ð15Þ

Overall : 3 COþ3 H2 Ð CH3OCH3þCO2 ð16Þ

A H2/CO ratio of close to 1.0 is achieved at 750 8C with mostthe catalysts (Figure 5 a), whereas a wider distribution of thisratio is achieved at lower and higher temperatures. The resultsranged from a minimum of 0.66 with Re-Ni at 700 8C to a maxi-mum of 1.57 with Rh-Ni at 750 8C. In general, the CO/CO2 ratioshowed an inverse relationship to temperature. The CO/CO2

ratio decreased at high temperatures because of the completeoxidation of carbon and the higher water–gas shift activity onpromoted catalysts.[24] A low CO/CO2 is not desirable as it re-duces the calorific value of the synthesis gas. Therefore, furthertests were performed at 750 8C, and the effect of feed ratios ofC/O and C/S on the product H2/CO and CO/CO2 ratios weretested to achieve the desired target.

Effect of carbon-to-oxygen ratio in the feed

The effect of the C/O ratio was tested on all the promoted cat-alysts. The monometallic Ni catalyst was not considered forthese tests because of the high char selectivity observed withthis catalyst at all of the temperatures tested. The cellulosefeed rate was kept constant at 15 g h�1, similar to the testsabove, and the oxygen flow rate in these tests was controlledsuch that the C/O ratio was 0.5, 0.6 and 0.7. The reactor tem-perature and C/S ratio for these runs were kept constant at750 8C and 1.0, respectively. The results, illustrated in Figure 6,show that all the promoted Ni catalysts demonstrated a highgasification efficiency of 88–100 % at all the C/O ratios tested.In general, the gasification efficiency decreased with the in-creasing C/O ratio, which was expected because a high C/Oratio means a low oxygen flow rate and, therefore, a low par-tial oxidation. This result is consistent with previous observa-tions in which higher C/O ratios in air-steam gasification led toa lower carbon conversion.[25] Although a low C/O ratio may

Figure 5. Effect of temperature on a) H2/CO and b) CO/CO2 of the investigat-ed catalysts at a C/O ratio of 0.5 and C/S ratio of 1.0. ~: Ni, *: Pt-Ni, !: Ru-Ni,&: Re-Ni, ^: Rh-Ni.

Figure 4. Effect of RFV temperature and catalyst promoters on product gas composition. C2 compounds include ethane, ethylene and acetylene. Reactionconditions: cellulose flow rate = 15 g h�1, C/O = 0.5, C/S = 1.0 and residence time = 50 ms.

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lead to high gasification temperatures (hot spots in the cata-lyst bed), which is favourable for steam reforming, the higherrate of oxidation reactions may also result in the oxidation ofCO to CO2 and H2 to H2O. Therefore, low C/O ratios may de-grade the quality of synthesis by increasing the CO2 and H2Omole fraction. This is precisely what was observed in our ex-periments (Figure 7). A C/O ratio of 0.5 resulted in a CO2 molefraction in the dry synthesis gas in excess of 50 %. If the C/Oratio was increased to 0.6, the CO2 mole fraction decreased to�40 %. However, a further increase of the C/O ratio resulted inincomplete gasification in most cases, that is, the formation ofchar. The Ru-Ni catalyst performed the best amongst all thepromoted catalysts. The char selectivity with the Ru-Ni catalystwas zero at all the C/O ratios. Ru is known to exhibit a higheractivity for C�C bond cleavage than Re, Rh and Pt in thatorder.[26] Ru also shows a higher water–gas shift reaction activi-ty than Pt and Rh and it is nearly equal to that of Re.[27] There-fore, it is expected that Ru will perform well under gasificationconditions, which requires a combination of C�C bond cleav-age and water–gas shift reaction. The water–gas shift activitywas high because of the high H2/CO ratio and low CO/CO2

ratio observed for the promoted Ni catalysts (Figure 8).In summary, a higher C/O ratio, especially C/O = 0.6, was fa-

vourable as it improved the gasification efficiency, synthesis

gas quality and the mole fractions of H2 and CO in the productgas. In addition, the catalysts are less likely to be re-oxidisedunder an oxygen deficit environment (high C/O), which canimprove the catalyst service life significantly.

Effect of the carbon-to-steam ratio in the feed

The effect of the C/S ratio was tested on all the promoted cat-alysts. The monometallic Ni catalyst was not considered forthese tests because of the high char selectivity observed withthis catalyst at all the temperatures tested in the previous sec-tion. The cellulose feed rate was kept constant at 15 g h�1, sim-ilar to the tests above, and the steam flow rate in these testswere controlled such that the C/S ratio was 1.0, 1.5 and 2.0.The reactor temperature and C/O ratio for these runs werekept constant at 750 8C and 0.6, respectively. The results showthat all the promoted Ni catalysts showed a high gasificationefficiency of 82–100 % at all the C/S ratios tested (Figure 9). Ingeneral, the gasification efficiency reduced with the increasingC/S ratio (i.e. , a lower steam flow rate), which is expected be-cause of the reduction in steam reforming and water–gas shiftactivity at high C/S ratios. This result is consistent with the cat-alytic steam gasification[28] and air-steam gasification[29] of bio-mass in previous studies. Although lower C/S ratios increase

Figure 6. Effect of C/O feed ratio and catalyst promoters on the selectivity of char (&), tar (&) and gas (&) based on carbon balance. The numbers on thetop of each bar represent gasification efficiency, which is the percentage of carbon in the gas and tar combined. Reaction conditions: cellulose flowrate = 15 g h�1, 750 8C, C/S = 1.0 and residence time = 50 ms. Here, tar is defined as water-soluble organics.

Figure 7. Effect of C/O feed ratio and catalyst promoters on product gas composition. C2 compounds include ethane, ethylene and acetylene. Reaction condi-tions: cellulose flow rate = 15 g h�1, 750 8C, C/S = 1.0 and residence time = 50 ms.

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the rate of steam reforming, they can reduce the temperatureof the catalyst bed because steam reforming is highly endo-

thermic. The required heat for steam reforming is provided bypartial oxidation and water–gas shift reactions. In this sectiona high C/O ratio was used, therefore, high C/S ratios were notfavourable because the amounts of both the gasifying agents(O2 and steam) were low, which results in incomplete gasifica-tion. The molar composition of the product gases are shownin Figure 10, and it is clear that higher C/S ratios led toa higher fraction of CO2 in the product gas because the oxida-tion reactions dominated the gasification reaction if the steamflow rate was reduced (higher C/S ratios). The fraction of CO2

increased from �40 % to over 50 % if the C/S ratio was in-creased from 1.0 to 1.5. The H2/CO and CO/CO2 ratios againstthe changes in C/S ratio are shown in Figure 11. The H2/COratio recorded in this part of the study was the highest amongall the previous tests reported in the previous two sections. Onaverage the H2/CO ratio ranged from 1.0 to 1.5, which is be-lieved to be because a high C/S ratio leads to oxidation-domi-nated gasification and conversion of CO into CO2. This led toa reduction in the CO/CO2 ratio as seen in Figure 11 b.

Among the catalysts tested in this study, the Re-Ni and Ru-Ni catalysts performed the best in terms of gasification efficien-cy and hydrogen yield. As seen in the catalyst characterisationsection, the Re-Ni and Ru-Ni catalysts had a high metal surfacearea, low reduction temperature and high CO uptake in CO-TPD, which combined explains the superior performance ofthese two catalysts in the gasification experiments.

Overall, the results from the catalyst characterisation andRFV sections combined show that the Re-Ni, Rh-Ni and Ru-Nicatalysts were the best catalysts among those tested in thisproject. Among the conditions tested in this study, the bestconditions were determined to be a reaction temperature of750 8C, C/O = 0.6 and C/S = 1.0. This conclusion is based on theoverall gasification efficiency, in particular the low char selectiv-ity and the product gas composition, principally, the low CO2

mole fraction in the gas phase. These conditions resulted inthe sustained performance of the catalysts (>300 min) withoutnoticeable deactivation of the catalyst. The duration of the

Figure 8. Effect of C/O on a) H2/CO and b) CO/CO2 of the investigated cata-lysts at 750 8C and a C/S ratio of 1.0. *: Pt-Ni, !: Ru-Ni, ^: Rh-Ni, &: Re-Ni.

Figure 9. Effect of C/S feed ratio and catalyst promoters on the selectivity of char (&), tar (&) and gas (&) based on carbon balance. The numbers on thetop of each bar represent gasification efficiency, which is the percentage of carbon in the gas and tar combined. Reaction conditions: cellulose flowrate = 15 g h�1, 750 8C, C/O = 0.6 and residence time = 50 ms. Here, tar is defined as water-soluble organics.

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gasification study could not be extended because of healthand safety aspects. The gasification rig had to be operatedmanually and monitored at all times, and it was considereda potential hazard to leave the rig unattended overnight,given the toxic and highly flammable nature of the productgases.

Stability of the catalysts

The rapid deactivation of Ni catalysts because of coke forma-tion is a well-known problem in gasification and hydrocarbonreforming reactions.[6c, 8–9, 30] Commercial Ni catalysts used asprimary catalysts in biomass gasification are especially proneto deactivation because of fouling, sintering and coke deposi-tion.[6c, 8, 31] Methods to improve the activity and coke resistanceof supported Ni catalysts include doping with noble met-als,[17b,c] alkali metals,[21] rare earth metals[17b] or transition met-als.[17c, 21, 24a] In this project, the first strategy to promote the Nicatalyst with noble metals was adopted. Coke deposition overthe spent catalysts used in these investigations was studied byTEM, and the resulting images are shown in Figure 12. Thespent catalyst samples selected for imaging had a char selec-tivity of 2–12 %, which is believed to be in the range in whichthe catalysts started to deactivate. Coke deposition of varyingdegrees was found on all five catalysts. The highest coke depo-sition was found on the monometallic Ni catalyst, which ex-plains its rapid deactivation that led to a high char selectivity.Modification of the Ni catalyst by the addition of noble metalsreduced the extent of coke deposition to a great extent. Re-Niand Rh-Ni showed the highest coke resistance among the cata-lysts tested. This result concurs with previous reports in whichnoble-metal-promoted Ni was used in biofuel steam reform-ing[18] and the partial oxidation of propane.[21] Interestingly, themonometallic catalyst, which showed the highest coke deposi-tion, was used for the shortest time in the gasification run(180 min), whereas, Re-Ni and Rh-Ni, which showed negligiblecoke deposition, were used in the gasification run for over300 min. Therefore, the rate of coke deposition on the Ni cata-lyst was several times higher than that on the Re- and Rh-pro-moted catalysts. Carbon deposited on the catalysts can be cat-egorised into three types: amorphous, filamentous and graphi-tic.[24b] Amorphous carbon is formed at the lowest temperaturerange of T�570 8C, followed by filamentous carbon at570 8C<T<1000 8C and finally graphitic carbon at the highesttemperature range of T�1000 8C. In our study, only filamen-tous carbon deposits were found on the spent catalyst. This isbecause the reaction temperatures used in our experiments(700–775 8C) were favourable for the formation of filamentouscarbon.

Figure 10. Effect of C/S feed ratio and catalyst promoters on product gas composition. C2 compounds include ethane, ethylene and acetylene. Reaction con-ditions: cellulose flow rate = 15 g h�1, 750 8C, C/O = 0.6 and residence time = 50 ms.

Figure 11. Effect of C/S ratio on a) H2/CO and b) CO/CO2 of the investigatedcatalysts at 750 8C and a C/O ratio of 0.6. ~: Ni, *: Pt-Ni, !: Ru-Ni, &: Re-Ni,^: Rh-Ni.

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Catalyst deactivation may also be caused by the sintering ofmetal crystallites to lead to the loss of active metal surfacearea.[22, 25–26] High temperatures and the presence of steam, re-ferred to jointly as hydrothermal conditions, may lead to thesintering of Ni nanoparticles.[27] To test the hydrothermal stabil-ity of the catalysts, TEM images of the spent catalysts were re-corded. The spent catalyst samples selected for the imaginghad a char selectivity of 2–12 %. Images of the fresh (calcined)Re-Ni and Rh-Ni catalysts compared to the spent catalysts,

which had been used in gasifi-cation runs for over 300 min,are shown in Figure 13. Al-though it was difficult to quan-tify the sintering effect, it canbe said qualitatively that themetal particle size of the spentcatalysts was slightly larger thanthat of the fresh catalyst. How-ever, the effect of sintering ongasification is unclear becausethe catalyst was stable for theduration of the study. We couldnot perform longer gasificationruns to determine the maxi-mum duration for which thecatalyst can remain active.

In addition to testing thecoke deposition and metal sin-tering during the gasificationruns, the stability of the aluminasupport under hydrothermalconditions was also studied.Phase transformation from g-

alumina into d-alumina or q-alumina phases may take place athigh temperatures, which leads to the loss of support surfacearea.[28–29] The loss of support surface area may lead to lowercatalytic activity by increasing the diffusional and film masstransport resistances. The support surface area of the fresh(calcined) catalysts was compared against spent catalysts inthe gasification runs. Results from the BET measurement of thefresh and spent catalysts are listed in Table 4. The results indi-cate clearly that the surface area of the spent catalyst had de-creased, which was an indication of possible phase change.

TEM images (Figures 12 and 13) of the spent catalysts alsosuggest that all five catalysts had gone through some phasetransformation. The TEM images of spent catalysts showedthat alumina particles had agglomerated, whereas the particlesin the fresh catalyst were much smaller and had a moreporous texture. Phase transformation of g- to d-alumina mayoccur at temperatures as low as 780 8C, and d-alumina may fur-ther transform into q-alumina at around 950 8C.[28–29] To test thehydrothermal stability of the commercial g-alumina supportused in this study, we heated the alumina samples to 750 and

Figure 12. TEM images of spent catalysts a) Ni, b) Pt-Ni, c) Ru-Ni, d) Rh-Ni and e) Re-Ni. The spent catalyst samplesselected for the imaging had a char selectivity of 2–12 %. Scale bars = 100 nm.

Figure 13. TEM images of a) spent and b) fresh Rh-Ni, and c) spent andd) fresh Re-Ni. Scale bars = 100 nm.

Table 4. Comparison of SBET of fresh (calcined) and hydrothermally treatedcatalysts.

Catalysts Fresh catalyst Conditions Spent catalyst Change inSBET [m2 g�1] T [8C] C/O C/S SBET [m2 g�1] SBET [%]

Ni 91.99 750 0.5 1.0 25.42 �72.36Pt-Ni 77.18 750 0.5 1.0 31.25 �59.51Ru-Ni 95.74 775 0.5 1.0 34.78 �63.67Rh-Ni 77.92 750 0.6 1.5 40.64 �47.84Re-Ni 74.16[a] 750 0.6 1.5 31.83 �57.08Re-Ni 74.16[a] 750 0.7 1.0 28.56 �61.49

[a] Re-Ni was reduced. All other catalysts were calcined only.

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800 8C under air (calcination) and air+steam atmospheres andthen analysed their morphology using powder XRD. The resul-tant XRD patterns are shown in Figure 14. Pattern (a) corre-sponds to the as-received commercial g-alumina sample. Pat-

terns (b) and (c) correspond to the g-alumina support calcinedat 750 and 800 8C, and patterns (d) and (e) correspond to theg-alumina support treated under air+steam atmosphere at 750and 800 8C, respectively. XRD patterns (a) and (b) are nearlyidentical, which indicates that calcination at 750 8C does notchange the alumina phase. However, the increase of tempera-ture and addition of steam has an effect on the alumina phasechange. The peaks at 2 q= 21.3 and 23.78, which correspond tod-alumina, increase in intensity from patterns (b) to (e). If thetreatment temperature under air is increased from 750 to800 8C, a small but measurable impact in the phase transforma-tion of g- to d-alumina is observed. It is, however, underair+steam atmosphere that the phase transformation is morepronounced.

Therefore, from the results in Table 4 and Figure 14 it can beconcluded that the alumina support used in this study was notcompletely stable during RFV. Although all the catalysts usedin this project were calcined at 600 8C in air, except Re-Ni, thereaction temperatures of 700–775 8C in the presence of steamcaused the alumina pores to collapse, which led to a loss ofthe BET surface area (SBET) and phase transformation of g- to d-alumina. Although the phase transformation of g- to d-aluminabegins at �780 8C, which is higher than all the RFV experi-ments, the presence of steam in the reactor lowered the phasetransition temperature.

In summary, coke deposition, active metal sintering and alu-mina phase transformation all contributed, in that order, to-wards the deactivation of the catalysts used in this study. Themonometallic Ni catalyst had the worst performance becauseof high coke deposition and Ni sintering. The addition ofa small amount of noble metal prevented coke deposition toa great extent and reduced Ni sintering. If we combine the re-sults presented in the catalyst characterisation and RFV sec-

tions, it can be said that the noble-metal-promoted Ni catalystswere effective for RFV because of their superior metal surfacearea, high steam reforming and water–gas shift activity andgood coke resistance. In particular, Re-Ni, Rh-Ni and Ru-Ni, inthat order, provided a much improved performance over themonometallic Ni catalyst.

Conclusion

Various reaction conditions were tested for reactive flash volati-lisation, which include the effects of temperature, carbon-to-oxygen (C/O) feed ratio, carbon-to-steam (C/S) feed ratio andcatalyst promoter. The highest gasification efficiency was ach-ieved at 750 8C, C/O = 0.6 and C/S = 1.0. It was observed thatRe-, Rh- and Ru-promoted Ni catalysts performed the best inthat order. This conclusion is based on the catalyst properties,gasification efficiency and stability. Monometallic Ni was activefor reactive flash volatilisation; however, the catalyst sufferedrapid deactivation because of coke deposition, metal active-site sintering and alumina support pore collapse because ofphase change. Coke deposition and sintering were reducedsignificantly by using noble-metal promoters. The noble-metalpromoters increased the activity of the Ni catalysts by increas-ing the metal surface area, reducibility and reducing the COdesorption temperature. The noble-metal promoters also in-creased coke resistance and reduced sintering. Therefore, thepromoted Ni catalysts were active for longer with little to nodeactivation.

Experimental Section

Catalyst preparation

Five Ni-based catalysts were developed using the impregnationmethod. They were Ni/Al2O3, Pt-Ni/Al2O3, Ru-Ni/Al2O3, Rh-Ni/Al2O3

and Re-Ni/Al2O3. Nickel nitrate (Ni(NO3)2·6 H2O), alumina (Al2O3) andchemicals such as H2PtCl6, RuCl3, RhCl3 and NH4ReO4 obtained fromSigma–Aldrich were used as the precursors for the catalyst synthe-sis. First, the required amount of nickel nitrate was measured anddissolved in distilled water. Then, the corresponding amount ofalumina and metal promoter precursor were added into the solu-tion. The monometallic Ni/Al2O3 catalyst had a Ni content of11 wt %, and all bimetallic catalysts had Ni contents of 10 wt % anda metal promoter content of 1 wt %. To ensure a homogeneousmix of all precursors with alumina, the solution was heated up to65 8C and maintained for 5 h under constant stirring. The solutionswere then dried overnight in a 100 8C oven. Dry solid was recov-ered and calcined in a muffle furnace at 600 8C for 6 h under air at-mosphere. Only the Re-Ni/Al2O3 catalyst was not calcined at hightemperature because of the high volatility of rhenium(VI) and (VII)oxides. The resultant calcined catalysts were reduced in situ beforethe reaction studies under 60 mL min�1 of 5 % H2/N2 flow at 400 8Cfor 5 h.

Catalyst characterisation

Catalysts were characterised using techniques that include N2

physisorption, CO chemisorption, XRF spectroscopy, TEM, H2-TPRand CO-TPD.

Figure 14. Powder XRD patterns of g-alumina support: a) as received; cal-cined at b) 750 8C; c) 800 8C; and treated in air+steam atmosphere atd) 750 8C and e) 800 8C.

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Nitrogen physisorption : The catalyst specific surface area, poresize distribution and pore volume were measured using N2 physi-sorption by using a Belsorp mini-II instrument. The surface areawas characterised using the BET method, whereas the pore sizedistribution and pore volume were measured using the Barrett–Joyner–Halenda (BJH) method on the adsorption–desorption iso-therm curve. Catalyst (200 mg) was degassed under high vacuumat 120 8C for at least 5 h prior to the measurement.

XRF spectroscopy : The elemental composition of Ni and aluminaof the prepared catalysts were determined by using an AmetekSpectro iQ II XRF, and the promoter content was calculated basedon mass balance.

CO chemisorption : The amount of active metal sites and the dis-persion of the catalysts were measured using CO chemisorption byusing a Micromeritics ASAP 2020. Catalyst (200 mg) was loadedinto a flow-through quartz tube reactor in which an in situ reduc-tion was performed. The sample was reduced with 5.22 % H2/N2

gas mixture at 400 8C for 5 h and then cooled to RT under N2 flow.The chemisorption measurement was started once the samplereached 30 8C.

TPR : TPR measurements were performed to investigate the reduci-bility of the fresh catalyst and examine the interaction between Ni,metal promoter and the support. Measurements were performedin a custom-built instrument that consisted of a vacuum compart-ment fitted with an Agilent Technologies TPS Compact vacuumpump and a Stanford Research Systems Residual Gas Analyser(RGA) 300. A simplified schematic diagram of the setup is present-ed in Figure 15. Catalyst (500 mg) was loaded into a custom-madequartz reactor that was placed in a vertical tube furnace (Labec)and heated from RT to 800 8C at a heating rate of 10 8C min�1.During the heating process, a 5.22 % H2/N2 gas mixture was fedcontinuously into the reactor at a flow rate of 50 mL min�1 toreduce the metal oxide to its pure metal state. The resultant gaswas introduced into the vacuum chamber through a capillary tubeand a leak valve. The RGA, which is a quadrupole mass spectrome-ter, analysed the gases in the vacuum chamber and reported the

partial pressure of gases with respect to time on stream. To reportthe results, the partial pressures were converted into molar flows,based on the volumetric flow rate used and the time on streamwas converted to temperature, based on the ramp rate of the fur-nace.

CO-TPD : CO-TPD was performed to examine the interaction be-tween Ni and the metal promoter and to measure the strengthand number of metal active sites available on the catalyst surface.The same custom-built instrument described for the H2-TPR studywas used in this investigation. In this study, fresh catalyst samplewas loaded into the quartz reactor and reduced in situ under60 mL min�1 of 5.22 % H2/N2 gas mixture at temperature of 400 8Cfor 5 h followed by purging with 100 mL min�1 of He gas at 400 8Cfor 1 h. The catalyst was then cooled to RT under He gas. To dopethe catalyst surface with CO, 10 % CO/He gas mixture was fed intothe reactor for 15 min at a flow rate of 100 mL min�1 followed bypurging with He gas at 100 mL min�1 for 2 h to remove any excessCO such that only a monolayer CO adsorbed was left on the sur-face of the catalyst. Once the CO partial pressure, monitored byusing the RGA, was stable, the He flow was subsequently reducedto 50 mL min�1, and the TPD data were recorded as the samplewas heated from RT to 800 8C at 10 8C min�1. Typical CO-TPD prod-ucts include CO and CO2. It is believed that the formation of CO2 islargely because the water–gas shift [Eq. (17)] and Boudouard[Eq. (18)] reactions occurred if the adsorption layer becomesmobile at high temperature.[9a] However, CO2 formation is largelydominated by Equation (17). As a result, both CO and CO2 partialpressures were monitored and reported in the CO-TPD analysis.

2 COþ2 OHAl2 O2! 2 CO2þH2 ð17Þ

2 COÐ CO2þC ð18Þ

Powder XRD : Powder XRD measurements were performed byusing a Rigaku Miniflex powder diffractometer with monochroma-tised CuKa radiation (l = 0.154 nm) at 40 kV and 15 mA.

Figure 15. Schematic diagram of the custom-made TPR-TPD-TPO instrument.

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Catalytic activity evaluation

Reactor setup : The catalytic activity was evaluated by usinga bench-scale reactor setup that consisted of a 25 mm OD and700 mm long quartz tube reactor, a K-Tron twin screw powderfeeder (K-MV-KT20), an Alltech HPLC pump (model 426), three Tele-dyne Hasting mass flow controllers for nitrogen and oxygen,a Labec vertical split tube furnace, a Brooks Instrument DLI Seriesevaporator and a custom-built gas–liquid separator. A schematicdiagram of the reactor setup is shown in Figure 16.

Catalyst (1 g) was loaded into the quartz reactor prior to assem-bling it in the furnace. The catalyst was held in the centre bya porous quartz disk. Before the catalytic tests, in situ reductionwas performed with H2/N2 gas mixture at 400 8C for at least 4 h.RFV was performed at 700, 750 and 775 8C with various C/O and C/S ratios. During these experiments, 15 g h�1 of cellulose was gravityfed by the twin screw powder feeder and a mixture of O2, N2 andsteam was fed into the reactor. The amounts of steam and O2 fedinto the reactor were controlled by the HPLC pump and the massflow controller, respectively. N2 was used as a control parameter tovary the reactor space velocity. Product gases from the reactorwere fed into a custom-built gas–liquid separator in which the gas-eous products were sampled periodically for composition analysis.Gases were analysed by using a Shimadzu gas chromatograph GC-2014 equipped with a molecular sieve 5 � column (60/80 mesh, 1/8 inch diameter, 6 feet in length) using a thermal conductivity de-tector and a flame ionisation detector. The volumetric flow rate ofthe gaseous product was determined after each GC analysis byusing a 50 mL bubble flow meter.

Detailed operating parameters: temperature 700, 750, 775 8C, cellu-lose type Cellets 200, cellulose feed rate 15 g h�1, N2 flow rate

235 sccm, residence time 50 ms, C/O ratio 0.5, 0.6, 0.7, C/S ratio1.0, 1.5, 2.0.

The gas yield and composition results were computed by usingthe trapezoid rule, which integrates the molar flow rate of eachgas species over total run time. Carbon balance was performed bycalculating the moles of carbon atoms in the gas, liquid and solidproducts collected from the reactor. Carbon in the gas phase wascalculated by measuring the amount of carbonaceous molecules inthe gas phase using GC (described above). Carbon atoms in theliquid phase (tar) were calculated by measuring the total organiccarbon by using a Shimadzu TOC-L series analyser. A liquid sample(0.5 mL) was diluted with deionised water (200 mL) prior to theanalysis. The number of moles of carbon atoms in the solid prod-ucts (char) was calculated by measuring elementary carbon byusing a PerkinElmer 2400 Series II CHNS/O system. Based on thecarbon balance, gas selectivity, tar and char are reported accordingto Equations (19)–(21):

Sgas ½%� ¼moles of carbon atoms in the gas phasemoles of carbon atoms in the products

� 100 ð19Þ

Star ½%� ¼moles of carbon atoms in the liquid phase

moles of carbon atoms in the products� 100 ð20Þ

Schar ½%� ¼moles of carbon atoms in the solid phase

moles of carbon atoms in the products� 100 ð21Þ

Biomass feedstock : High-purity microcrystalline cellulose, Cellets200 (Pharmatrans Sanaq AG) was used in our studies. The particlesize of Cellets 200 was 200–355 mm.

Figure 16. Schematic diagram of the RFV reactor setup. MFC = Mass flow controller.

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Acknowledgements

The authors are grateful for the financial support from the RuralIndustries Research and Development Corporation (RIRDC) projectgrant PRJ-004758 and the Department of Chemical Engineering,Monash University.

Keywords: biomass · nickel · platinum · rhenium · supportedcatalysts

[1] M. Ni, D. Y. C. Leung, M. K. H. Leung, K. Sumathy, Fuel Process. Technol.2006, 87, 461 – 472.

[2] A. Kumar, D. D. Jones, M. A. Hanna, Energies 2009, 2, 556 – 581.[3] L. Devi, K. J. Ptasinski, F. J. J. G. Janssen, Biomass Bioenergy 2003, 24,

125 – 140.[4] a) J. J. Hern�ndez, G. Aranda-Almansa, A. Bula, Fuel Process. Technol.

2010, 91, 681 – 692; b) K. Qin, P. A. Jensen, W. Lin, A. D. Jensen, EnergyFuels 2012, 26, 5992 – 6002.

[5] G. W. Huber, S. Iborra, A. Corma, Chem. Rev. 2006, 106, 4044 – 4098.[6] a) S. Rapagn�, N. Jand, A. Kiennemann, P. U. Foscolo, Biomass Bioenergy

2000, 19, 187 – 197; b) K. Tomishige, M. Asadullah, K. Kunimori, Catal.Today 2004, 89, 389 – 403; c) E. G. Baker, L. K. Mudge, M. D. Brown, Ind.Eng. Chem. Res. 1987, 26, 1335 – 1339.

[7] a) R. J. Farrauto, C. Bartholomew, Introduction to Industrial Catalytic Pro-cesses, Chapman and Hall, London, 1997; b) J. R. Rostrup-Nielsen, SteamReforming Catalysts : An Investigation of Catalysts for Tubular Steam Re-forming of Hydrocarbons, Taknisk Forlag, Copenhagen, 1975.

[8] X. T. Li, J. R. Grace, C. J. Lim, A. P. Watkinson, H. P. Chen, J. R. Kim, Bio-mass Bioenergy 2004, 26, 171 – 193.

[9] a) A. Tanksale, J. N. Beltramini, J. A. Dumesic, G. Q. Lu, J. Catal. 2008,258, 366 – 377; b) A. Tanksale, Y. Wong, J. N. Beltramini, G. Q. Lu, Int. J.Hydrogen Energy 2007, 32, 717 – 724; c) A. Tanksale, C. H. Zhou, J. N. Bel-tramini, G. Q. Lu, J. Inclusion Phenom. Macrocyclic Chem. 2009, 65, 83 –88.

[10] D. Y. C. Leung, X. L. Yin, C. Z. Wu, Renewable Sustainable Energy Rev.2004, 8, 565 – 580.

[11] a) J. L. Colby, P. J. Dauenhauer, L. D. Schmidt, Green Chem. 2008, 10,773 – 783; b) J. R. Salge, Science 2006, 314, 801 – 804.

[12] a) Y. N. Shastri, L. F. Rodriguez, A. C. Hansen, K. C. Ting, Biofuels, Bioprod.Bioref. 2012, 6, 21 – 31; b) P. J. Dauenhauer, J. Colby, L. D. Schmidt, SaltLake City, 2007.

[13] a) D. Sutton, B. Kelleher, J. R. H. Ross, Fuel Process. Technol. 2001, 73,155 – 173; b) Y. Wang, C. M. Kinoshita, Solar Energy 1993, 51, 19 – 25.

[14] a) A. Kumar, K. Eskridge, D. D. Jones, M. A. Hanna, Bioresour. Technol.2009, 100, 2062 – 2068; b) A. Haryanto, S. D. Fernando, L. O. Pordesimo,S. Adhikari, Biomass Bioenergy 2009, 33, 882 – 889.

[15] S. Wang, Appl. Catal. B 1998, 16, 269 – 277.[16] J. M. Rynkowski, T. Paryjczak, M. Lenik, Appl. Catal. A 1993, 106, 73 – 82.[17] a) A. Tanksale, J. N. Beltramini, G. M. Lu, Renewable Sustainable Energy

Rev. 2010, 14, 166 – 182; b) M. El Doukkali, A. Iriondo, P. L. Arias, J. Re-quies, I. Gandar�as, L. Jalowiecki-Duhamel, F. Dumeignil, Appl. Catal. B2012, 125, 516 – 529; c) Z. Hou, P. Chen, H. Fang, X. Zheng, T. Yashima,Int. J. Hydrogen Energy 2006, 31, 555 – 561.

[18] L. P. R. Profeti, E. A. Ticianelli, E. M. Assaf, Int. J. Hydrogen Energy 2009,34, 5049 – 5060.

[19] D. O. Hayward, B. M. W. Trapnell, Chemisorption, 2ednd edButterworths,London, 1964.

[20] S. D. Jackson, B. M. Glanville, J. Willis, G. D. McLellan, G. Webb, R. B.Moyes, S. Simpson, P. B. Wells, R. Whyman, J. Catal. 1993, 139, 207 – 220.

[21] M. Shiraga, D. Li, I. Atake, T. Shishido, Y. Oumi, T. Sano, K. Takehira, Appl.Catal. A 2007, 318, 143 – 154.

[22] J. Sehested, J. A. P. Gelten, S. Helveg, Appl. Catal. A 2006, 309, 237 – 246.[23] A. Shrotri, L. K. Lambert, A. Tanksale, J. Beltramini, Green Chem. 2013,

15, 2761 – 2768.[24] a) K. Omata, Y. Endo, H. Ishii, A. Masuda, M. Yamada, Appl. Catal. A

2008, 351, 54 – 58; b) M. Asadullah, T. Miyazawa, S. I. Ito, K. Kunimori, S.Koyama, K. Tomishige, Biomass Bioenergy 2004, 26, 269 – 279.

[25] a) L. Wei, S. Xu, L. Zhang, C. Liu, H. Zhu, S. Liu, Int. J. Hydrogen Energy2007, 32, 24 – 31; b) P. M. Lv, Z. H. Xiong, J. Chang, C. Z. Wu, Y. Chen,J. X. Zhu, Bioresour. Technol. 2004, 95, 95 – 101.

[26] J. H. Sinfelt, Adv. Catal. 1973, 23, 91 – 119.[27] D. C. Grenoble, M. M. Estadt, D. F. Ollis, J. Catal. 1981, 67, 90 – 102.[28] B. Moghtaderi, Fuel 2007, 86, 2422 – 2430.[29] M. Campoy, A. G�mez-Barea, F. B. Vidal, P. Ollero, Fuel Process. Technol.

2009, 90, 677 – 685.[30] R. T. K. Baker, M. A. Barber, P. S. Harris, F. S. Feates, R. J. Waite, J. Catal.

1972, 26, 51 – 62.[31] A. Tanksale, J. N. Beltramini, G. Q. Lu, Devel. Chem. Eng. Mineral Process.

2006, 14, 9 – 18.

Received: June 11, 2014

Published online on August 21, 2014

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