Fuel Processing Technology 132 (2015) 39–48

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Fuel Processing Technology

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Hydrogen from the high temperature water gas shift reaction with anindustrial Fe/Cr catalyst using biomass gasification tar rich synthesis gas

Simeone Chianese a,⁎, Jurgen Loipersböck b, Markus Malits b, Reinhard Rauch b,c, Hermann Hofbauer c,Antonio Molino d, Dino Musmarra a

a Department of Civil and Building Engineering, Design and Environment (DICDEA), Seconda Università di Napoli, Real Casa dell'Annunziata, Via Roma 29, 81031 Aversa (CE), Italyb BIOENERGY 2020+ GmbH, Wienerstraße 49, 7540 Güssing, Austriac Vienna University of Technology, Institute of Chemical Engineering, Getreidemarkt 9/166, 1060 Vienna, Austriad ENEA, National Agency for New Technologies, Energy and Sustainable Economic Development, UTTRI S.S. 106 Ionica, km 419+500, 75026 Matera, Italy

⁎ Corresponding author. Tel.: +39 081 5010 387; fax: +E-mail address: simeone.chianese@unina2.it (S. Chian

http://dx.doi.org/10.1016/j.fuproc.2014.12.0340378-3820/© 2015 Elsevier B.V. All rights reserved.

a b s t r a c t

a r t i c l e i n f o

Article history:Received 14 October 2014Received in revised form 10 December 2014Accepted 22 December 2014Available online 8 January 2015

Keywords:Water gas shift reactionHydrogen productionBiomass tar rich synthesis gas applicationFe/Cr based catalystSulfur poisoning deactivation

The high temperaturewater gas shift reaction (HTS) over an iron/chromium (Fe/Cr) industrial catalystwas inves-tigated in a pilot scale plant consisting of twofixed-bed reactors arranged in series and a biomass-derived tar-richsynthesis gas was used as a feed-stream. CO conversion and selectivity for the water gas shift reactionwere eval-uated through parameter variation. Four dry gas hourly space velocities (GHSVd) and two steam to dry synthesisgas ratios (H2O/SGd) equal to 52% v/v and 60% v/v were investigated at temperatures (T) of 350–450 °C. COconversion was investigated by varying H2S concentration 180–540 ppmv (dry basis) at a temperature of425 °C, considering two GHSVd. The highest CO conversion (~83%) was observed in the basis case at 60% v/vH2O/SGd, and 450 °C. The catalyst appeared to be resistant to sulfur poisoning deactivation, and achieved 48%CO conversion at the maximum H2S concentration used.

© 2015 Elsevier B.V. All rights reserved.

1. Introduction

Greater and greater attention is being paid to pollutants and green-house gas emissions, while the depletion of fossil fuel reserves has led toan increased interest in clean energy technologies and renewable ener-gy sources. Recent progress in industrial fuel cell technologies [1–4] haspromoted the exploration of hydrogen production technology [5–7].Hydrogen is a clean energy carrier and an environment-friendly fuelthat can be used as feedstock in low temperature fuel cells such as phos-phoric acid and proton-exchangemembrane fuel cells [8,9]without cre-ating pollutant and greenhouse gas emissions [10,11]. Hydrogen is alsoessential for the chemical refining of ammonia and methanol [12–16].

Nearly 96% of the global H2 production involves fossil fuels (48%hightemperature steam reforming of natural gas, 30% partial oxidation of re-finery oil, 18% coal gasification), resulting in high carbon dioxide emis-sions [17,18]. Biomass-based hydrogen production technologies, as acarbon-neutral and a renewable energy resource, could overcomeboth the problems of greenhouse gas emissions and fossil fuel depletion[19–21].

This studywas part of theHydrogen from biomass for industry projectperformed in the Güssing CHP biomass plant (Austria) by Bioenergy2020+GmbH. The project showed that high-purity hydrogen for direct

39 081 5010 436.ese).

integration into a refinery could be produced (Fig. 1), which involvedsteam gasification, a water gas shift reaction step and a pressureswing adsorption system (PSA).

A steam reforming step can be used to increase the overall biomassto hydrogen efficiency, converting the hydrocarbons in the raw gasfrom the CHP plant into hydrogen and carbon monoxide, and sendingthem to the water gas shift unit [22]. Using steam gasification, thesolid biomass can be thermochemically transformed into a gaseousmix-ture called synthesis gas (syngas) with a significant H2 (35–45% v/v)and CO (20–30% v/v) content [6,23]. The increased H2 content and de-creased CO content can be reached through the water gas shift reaction(WGSR) [24,25], where CO reacts with steam to produce H2 and CO2 inthe presence of a catalyst whose performance strongly depends on itscharacteristics and properties.

CO gð Þ þH2O gð Þ↔CO2

catalgð Þ þ H2 gð Þ ΔH0

R 298 Kð Þ ¼ −40:6 kJ=mol:

ð1Þ

TheWGSR is reversible and slightly exothermic. As a function of thetemperature range required for the catalyst to be effective, the WGSRcan be divided into high temperature (HTS) and low temperaturewater gas shift reactions (LTS). On an industrial scale, the HTS is usuallycarried out at 350–500 °Cwith a Fe/Cr catalyst, chosen for its selectivityand stability, while LTS is carried out at 150–300 °C with a Cu/Zn

Fig. 1. Flow chart of the hydrogen from biomass for industry project in the Güssing CHP biomass plant (Austria).

40 S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

catalyst [26–30]. In both cases, the CO conversion increases with reac-tion temperature [31–34]. At high temperatures, according to the LeChatelier's principle, the reaction shifts to the reagents and reduces COconversion, while the temperature dependence of the reaction ratecan be expressed by the Arrhenius' law, which indicates an increasingCO conversion as the reaction temperature increases [9].

This study focused on CO conversion by the high temperature watergas shift reaction over a Fe/Cr industrial catalyst. The selectivity of thewater gas shift reaction was assessed in a pilot scale plant consistingof two fixed-bed reactors arranged in series. The pilot plant was sup-plied with synthesis gas produced in the Güssing CHP dual-fluidizedbed steam gasification plant fuelled with biomass (wood chips). Thedry gas hourly space velocity (GHSVd), steam to dry synthesis gasratio (H2O/SGd), reaction temperature (T) and the influence of hydro-gen sulfide loading in the synthesis gas on CO conversion wereinvestigated.

2. Materials and methods

2.1. Catalyst

An industrial HTS Fe/Cr catalyst, ShiftMax®120 (Clariant), suppliedas Fe2O3/Cr2O3, was tested. It required the activation [26] of the Fe2O3

to Fe3O4 by reduction. Fe3O4 is susceptible to thermal sinteringdeactivation, which is reduced because chromium stabilizes the matrixaround the grains [29,35–37] while the Fe/Cr catalyst is resistant to thehydrogen sulfide poisoning deactivation [26]. Before feeding the cata-lyst with a reduction gaseous flow, the temperature of the catalyst bedwas increased to 250 °C using nitrogen as a thermal carrier to ensurethe catalyst was dry. Catalyst reduction was performed through a gas-eous mixture of N2, H2 and steam with a gas hourly space velocity of300 Sl L−1 h−1, a N2 flow of 75% v/v and a H2O/H2 ratio of 0.3 to avoidthe reduction of Fe3O4 to FeO [26].

2.2. Experimental apparatus

The WGS investigations were carried out in a pilot scale plantconsisting of two fixed-bed reactors arranged in series. A schematic di-agram of the experimental apparatus is reported in Fig. 2. Reactor 1 hadan internal diameter of 80.8 mm and Reactor 2 was 105.3 mm, bothwith a height of 850 mm, and the catalyst bed volumes were of 2.5 L,

for a total volume of 5 L. Additional steam was generated by a waterevaporator and a water flow regulated by a peristaltic dosing pump(ISMATEC). N2 was used as the thermal carrier to heat the plant in thestart-up phase before feeding synthesis gas to the experimental appara-tus, and nitrogen and synthesis gas flow rates were controlled by an or-ifice disc system (ENDRESS + HAUSER).

Four gas sampling points were installed: before Reactor 1, afterReactor 1, after Reactor 2 and after gas cleaning at the test rig. Electro-mechanical valves (AMVi) were operated to control gaseous streamflow rates. The pressure and temperature of the systemweremonitoredby three pressure sensors (APi) and 24 temperature sensors (ATi), re-spectively (Fig. 2). The temperature sensors were at a relative distanceof 75 mm. Eight heating systems (ABHi) were used to provide heat tothe apparatus. The experimental apparatus was completely automatedand controlled through a computerised system. The pilot plantwas sup-plied with a stream of synthesis gas produced in the Güssing CHP bio-mass plant and a fan located downstream of the experimental plantwas used to take syngas from the gasification plant. The typical compo-sition, on a dry basis, of the gas produced is reported in Table 1.

The syngas stream fed to the pilot plant was drawn off before therapeseed oil methyl ester (RME) scrubber of the gasification plant(Fig. 3) to use the tar-rich synthesis gas (Table 1). Tars indicate the un-desirable by-products of biomass gasification [38]—hydrocarbons withmolecular weights higher than benzene [39]. Tars are responsible forthe main operational problems in gasification plants and limit syngasapplications [40]. The tar sampling and analysiswere performed accord-ing to the tar measurement standard protocol by Neeft et al. [41], basedon a gravimetric method (i.e. molecular weight higher than naphtha-lene), asmodified by the Institute of Chemical Engineering (ViennaUni-versity of Technology) [42,43]. Steam and tar condensation wereavoided by heating the pipe between the evaporator and the pilotplant and between the CHP plant and the pilot plant, respectively.

Polytetrafluoroethylene sampling bags (RESTEK, 10 L) were used tosample the gas produced. The sampling line consisted of three flasks inan ice-bath, where the first two flasks were filled with glycol and thethird with a glass wool filter, in the gas stream flow rate direction. A di-aphragmpump (KNF)was used to draw the gas products from the sam-pling points into the sampling bags. The gas productswere analyzed andquantified by two gas chromatographs (GCs, Clarus 500 and Clarus 580,Perkin Elmer). The Clarus 500was equippedwith a Thermal Conductiv-ity Detector (TCD) and three packed columns: a 10 ft. Molecular Sieve

Fig. 2. Schematic representation of the pilot scale plant configured for WGSR.

41S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

5 Å for O2, N2, CO and CH4; a 6 ft. Porapak® Q 50/80 for CO2, C2H4 andC2H6 and a 10 ft. Porapak®N 50/80 for C3H6 and C3H8. H2 concentrationwas quantified as the complement to 100% v/v of the sum of the afore-said components. The Clarus 580was equippedwith an Agilent 355 Sul-fur Chemical Detector (SCD) using hydrogen and compressed air, and a30 m capillary column (Metal-Capillary, ID 0.32 mm, DF 4.00) for H2S,COS and organic sulfur components (like C4H4S). Helium was used asthe carrier for both GCs.

2.3. HTS investigations

The reaction temperature and gas hourly space velocity effects werestudied through parameter variation tests. Experiments were carriedout at 350–450 °C at a four dry gas hourly space velocities (GHSVd), per-formed under a slight underpressure, and CO conversion andWGSR se-lectivitywere assessed. Catalyst performancewas evaluated through COconversion (XCO) calculated by:

XCO %ð Þ ¼ COjin−COjout� �

mol=hð ÞCOjin mol=hð Þ � 100 ð2Þ

and hydrogen production was evaluated in terms of WGSR selectivity.Selectivity is one of themost important catalyst properties and indicates

Table 1Composition (dry basis) of the synthesis gas produced in the GüssingCHP plant.

Component Concentration [% v/v]

H2 35–45CO 20–30CO2 15–25CH4 8–12N2 1–3Tars [g/Nm3] 2–5H2S [ppmv] 150–300

its capacity to increase the rate of the desired reaction.WGSR selectivity[40] was calculated by:

WGSR selectivity %ð Þ ¼ CO2jout−CO2jin½ � mol=hð ÞCO2jout−CO2jin½ � þ CH4jout−CH4jin½ � mol=hð Þ�100: ð3Þ

Methanation COþ 3H2 ↔catal CH4 þH2O� �

was considered as theonly possible side reaction, and was added to the WGSR selectivityEq. (3) since methane was detected in the product gases. H2 and CH4

production flow rates were evaluated by using Eq. (4):

Production Flow Rate molNm‐3� �

¼ Componentjout‐Componentjin½ � mol=hð ÞVolumetric Syngas Flow Stream Nm3 =h

� � : ð4Þ

Reaction temperature was measured using the temperature sensorsin Reactor 1 (seven sensors) and Reactor 2 (four sensors), whichwere incontact with the catalyst bed. Before supplying product gas, the catalysttemperature was increased to 350 °C at a heating rate of 100 °C h−1

using N2 as the thermal carrier.The steam-to-dry synthesis gas ratio (H2O/SGd) investigations were

divided into two experimental sets. One,with aGHSVd equal to the basiscase (Case A) and the doubling of GHSVd basis case (Case B), was per-formed by adding steam to the syngas to achieve an average 60%v/vH2O/SGd. The other, with a GHSVd equal to 1.34*GHSVd basis case(Case C) and 2.68*GHSVd basis case (Case D), was performed usingthe syngas steamcontent and an average 52%v/vH2O/SGdwas achieved.

Syngas steam content was evaluated through the water balance ofthe CHP plant RME scrubber data, and pressure, temperature and flowrate of the synthesis gas and condensate flow rate were recorded. Theeffect of H2O/CO on the water gas shift reaction was evaluated. A com-parison between the H2O/CO and the reaction temperature effects onCO conversion was also carried out above 400 °C for both experimentalsets. Details of parameter variation experiments, including the order of

Fig. 3. Draw off point in the biomass Güssing CHP Plant (Austria).

42 S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

implementation and run-times, are reported in Tables 2 and 3. An over-all run time of 645 h was achieved.

The hydrogen sulfide loading effect was investigated at an averagereaction temperature of 429 °C and Case C and Case D GHSVd, whereinlet H2S ranged between 180 and 540 ppmv on a dry basis. The testswere performed without additional steam, achieving an average H2O/SGd = 51% v/v. The details of the hydrogen sulfide loading study are re-ported in Table 4.

The tar effect was investigated at the end of H2S loading test, after atotal time of exposure of the catalyst to tar rich syngas higher than800 h. Themeasurementswere performed in triplicate andwere carriedout at the inlet of the pilot plant (Inlet), at the outlet of Reactor 1 (OutletR1) and at the outlet of Reactor 2 (Outlet R2). The average values of tarconcentration for each measurement are reported in Table 5, together

Table 2Operating conditions for the 1st experimental set: CaseA; Case BGHSVd–60% v/vH2O/SGd.

Experiment[#]

Average reaction T[°C]

Run time[h]

H2O/SGd

[% v/v]H2O/CO[mol mol−1]

I.1 352 18 58 6.1I.2 379 24 61 7.9I.3 399 23 65 8.8I.4 423 22 58 7.9I.5 450 21 60 8.3

Table 3Operating conditions for the 2nd experimental set: Case C; Case D GHSVd–52% v/vH2O/SGd.

Experiment[#]

Average reaction T[°C]

Run time[h]

H2O/SGd

[% v/v]H2O/CO[mol mol−1]

II.1 451 21 52 5.5II.2 428 21 54 6.0II.3 402 20 53 5.9II.4 377 20 50 4.5II.5 348 20 50 4.2

with the details of the tar operative conditions. These results showthat inlet tar concentration is higher than 10 g/Nm3 on a dry basis(db) and a light increase can be observed in the outlet streams.

3. Results and discussion

3.1. Influence of reaction temperature and gas hourly space velocity on thewater gas shift reaction

The experimental measurements of CO conversion as a function ofthe reaction temperature (350–450 °C) are shown in Fig. 4, where theinfluence of the gas hourly space velocity and thermodynamic equilibri-um conversion line are included.

The catalyst performance was evaluated using four GHSVd. The re-sults in Fig. 4 show that the catalytic activity is strongly influenced bythe parameters studied. For both average H2O/SGd ratios, XCO increases

Table 4Operating conditions for the H2S loading test: Case C; Case D GHSVd–51% v/v H2O/SGd.

Experiment[#]

Average reaction T[°C]

Run time[h]

H2O/SGd

[% v/v]H2O/CO[mol mol−1]

H2S[ppmv,db]

1 421 20 49 5.3 1822 428 21 49 4.9 2993 431 39 53 5.9 2824 430 33 50 5.1 4405 433 27 51 5.5 537

Table 5Operating conditions for the tar effect test and tar measurement results.

Measurement[–]

Average T[°C]

GHSVd

[h−1]H2O/SGd

[% v/v]H2O/CO[molmol−1]

H2S[ppmv,db]

Total tar[mg/Nm3

db]

Inlet 351 52 7.5 357 10.8 ± 0.1Outlet R1 419 Case D 49 7.9 275 12.0 ± 0.2Outlet R2 429 Case C 45 16.1 332 11.8 ± 0.2

Fig. 4. GHSVd influence on CO conversion. (a) 60% v/v H2O/SGd, 6.1–8.8 mol mol−1 H2O/CO; (b) 52% v/v H2O/SGd, 4.2–6.0 mol mol−1 H2O/CO.

Fig. 5. Influence of reaction temperature on selectivity for WGSR. Case D GHSVd–52% v/v H2O/SGd, 4.2–6.0 mol mol−1 H2O/CO.

43S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

Table 6Influence of reaction temperature on CH4 production flow rate andWGSR selectivity. CaseD GHSVd–52% v/v H2O/SGd, 4.2–6.0 mol mol−1 H2O/CO.

Average reaction T CH4 production flow rate WGSR selectivity[°C] [mol Nm−3] [–]

348 0.04 0.95377 −0.02 1.0402 −0.01 1.0428 −0.03 1.0451 0.00 1.0

44 S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

monotonically with temperature, and the highest and the lowest XCO

are achieved at reaction temperatures of 450 °C and 350 °C, respectively.Moreover the lower the space velocity the higher the CO conversion.

In agreement with the literature, CO conversion increased as reac-tion temperature increased and space velocity decreased [27,44,45]and a minimum reaction temperature was required for catalytic activi-ty [26,46].

At an average H2O/SGd ratio of 60% v/v (Fig. 4a), the maximumXCO (~83%) was achieved in the basis case GHSVd (Case A), while inthe doubling of basis case GHSVd (Case B) the highest XCO achievedwas ~61%. The minimum XCO (~23%) was reached at the Case BGHSVd, while in the basis case GHSVd (Case A), the lowest XCO achievedwas close to 32%.

At an average H2O/SGd ratio of 52 %v/v (Fig. 4b), the maximum XCO

(~74%) was achieved at the Case C GHSVd, and at the Case D GHSVd,

Fig. 6. GHSVd influence on H2 purity and production flow

the highest XCO achieved was ~48%. The minimum XCO (~7%) wasreached at the Case D GHSVd, and at the Case C GHSVd the lowest XCO

achieved was ~15%. These results are in agreement with those reportedby Reddy et al. [47,48] and Maroño et al. [27,34]. As the space velocitydecreased, the contact time between the gaseous stream and the cata-lyst bed increased, indicating an increasing CO conversion as the spacevelocity was reduced.

At a reaction temperature of 350 °C and at the Case D GHSVd, with a2.2 mol mol−1 H2O/(CO + CH4 + CO2), the selectivity for WGSR was95–100% (Fig. 5). The results are not totally in line with the literature,where a selectivity of 100% was reported for an industrial HTS catalystindependently of the operating conditions [31,44,48–50]. Jeong et al.[32] investigated the Fe catalyst activity at different compositions.They evaluated selectivity considering the methanation reaction andobserved a selectivity lower than 100% at a reaction temperature of350 °C and a H2O/(CO+CH4+CO2) of 2mol mol−1, showingmethaneproduction. A strong correlation between catalyst composition, metha-nation reaction and selectivity was highlighted.

CH4 production flow rate as a function of reaction temperature is re-ported in Table 6. A positive CH4 production flow rate was assessed at348 °C, resulting in a production of methane while a negative CH4 pro-duction flow rate was evaluated in the temperature range 377–451 °C,resulting in a consumption of methane.

The influences of reaction temperature on H2 purity and productionflow rate are shown in Fig. 6 (H2O/SGd = 60% v/v) and Fig. 7(H2O/SGd = 52% v/v).

rate. 60% v/v H2O/SGd, 6.1–8.8 mol mol−1 H2O/CO.

Fig. 7. GHSVd influence on H2 purity and production flow rate. 52% v/v H2O/SGd, 4.2–6.0 mol mol−1 H2O/CO.

45S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

In agreement with CO conversion trends, H2 purity and productionflow rates increased as reaction temperature increased while seem notto be influenced by space velocity.

At an average H2O/SGd ratio of 60% v/v, the maximum H2 purity(~48% v/vdb) and H2 production flow rate (~13.0 mol Nm−3

db) wereachieved at the reaction temperature of 450 °C and Case A GHSVd.

At an average H2O/SGd ratio of 52% v/v, the maximum H2 purity(~46% v/vdb) and H2 production flow rate (~10.0 mol Nm−3

db) wereachieved at the reaction temperature of 450 °C and the Case C GHSVd

(1.34*GHSVd basis case).

3.2. Comparison between the reaction temperature and the H2O/CO effectson CO conversion

Since a real biomass-derived syngas was used as a feed-stream itwas not possible to define a precise and constant experimental steam-to-CO ratio. Nevertheless some remarks about the influence of H2O/COon the water gas shift reaction can be put forward.

Froma stoichiometric point of view, a H2O/CO of 1molmol−1wouldbe required to carry out the WGSR. In practice, in order to prevent anyside reaction that might decrease the catalytic activity—such as theBoudouard reaction (2CO ↔ C + CO2) by carbon deposition—theWGSR is performed at an excess of steam, operating at a H2O/COequal to 2 mol mol−1 (or higher) [26,44,49]. The steam-to-CO ratiosreached in the experiments are reported in Tables 2–4. At the Case Aand Case B (Table 2), the tests were performed by varying the H2O/CO

6.1–8.8 mol mol−1, while the H2O/CO was 4.2–6.0 mol mol−1 at theCase C and Case D (Table 3). Thanks to the H2O/CO used in theexperiments, the undesired reactions were avoided. The effects of theH2O/CO ratio and reaction temperature on CO conversion are shownin Fig. 8.

The results indicated that the higher the reaction temperature, thehigher the XCO value, even if the highest H2O/CO was not achieved atthe highest reaction temperature. At an average H2O/SGd ratio of 60%v/v (Fig. 6a), the maximum H2O/CO (~8.8 mol mol−1) was achievedat 399 °C, while a H2O/CO of ~6.0 mol mol−1 was reached at 402 °Cwith an average H2O/SGd ratio of 52% v/v (Fig. 6b). As it is shown inthe literature [44], far from thermodynamic equilibrium conditions,XCO wasmore influenced by reaction temperature than H2O/CO. No dif-ference in terms of GHSVd was observed.

3.3. Influence of H2S concentration on catalytic activity

The effect of H2S loading on CO conversion and catalyst stability isshown in Figs. 9 and 10, respectively. Activity and stability wereassessed at inlet H2S concentrations in the range 180–540 ppmv on adry basis (db), Case C and Case D GHSVd, with a 51% v/v H2O/SGd, andreaction temperature was kept constant, achieving average value of429 °C.

As H2S concentration increased, XCO decreased, reaching aminimumCO conversion of 48% and 4% at the Case C and Case D GHSVd, respec-tively (Fig. 9). As reported in the literature [26,51,52] a loss of activity

Fig. 8. Influence of H2O/CO and Reaction Temperature on CO Conversion. (a) 60% v/v H2O/SGd; (b) 52% v/v H2O/SGd.

Fig. 9. Influence of H2S loading on CO conversion. 429 °C average reaction temperature 51% v/v H2O/SGd, 4.9–5.9 mol mol−1 H2O/CO.

46 S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

Fig. 10. Influence of H2S loading on catalyst stability. 429 °C average reaction temperature 51% v/v H2O/SGd, 4.9–5.9 mol mol−1 H2O/CO.

47S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

was observed, although the catalyst had good resistance to hydrogensulfide poisoning deactivation at the highest space velocity. The de-creased activitywas due to the sulfiding reaction,where Fe3O4was con-verted to FeS [26,51].

Fe3O4 þ 3H2S↔þ 4H2O ΔH0R 298 Kð Þ ¼ −75:0 kJ=mol ð5Þ

H2S concentration in the inlet and outlet gaseous stream flows forReactor 1 and Reactor 2 were measured and a cumulative time on-stream was considered to evaluate the effect of H2S on the catalyst sta-bility (Fig. 10). Since real biomass-derived syngas was used, it was notpossible to control H2S concentration and a variation of H2S load againsttime on-stream was found, however no significant variation betweeninlet and outlet H2S concentration for each catalyst bed was observed,resulting an equilibrium regarding hydrogen sulfide adsorption/desorption.

Up to a time on-stream of 268 h, a decrease of activity was foundwith inlet H2S concentration increasing and a minimum CO conversionof 48% and 4%was reached at the Case C and Case DGHSVd, respectively.At a time on-stream of 310 h, a decrease of inlet H2S concentration wasobserved: for the Reactor 1 catalytic bed, no changes in activity wasfound, while for the Reactor 2 catalytic bed, activity increased, achievinga CO conversion close to 55%. The not-increase of XCO for the catalystbed of Reactor 1 may be related to a starting poisoning deactivationmechanism.

Over time, results show a good stability, highlighting an increase ofactivity for the catalyst of the Reactor 2 after a time on-stream higherthan 300 h.

3.4. Influence of tar concentration on catalytic activity

The effect of tar content on CO conversion is reported in Table 7. Thecatalytic activity was assessed at inlet total tar concentration equal to10.8 ± 0.1 g/Nm3 on a dry basis (db), an average reaction temperatureof 424 °C and Case C GHSVd, with an inlet 52% v/v H2O/SGd.

Table 7Influence of tar concentration on COconversion. 424 °C average reaction temperature, 52%v/v H2O/SGd, 7.5 mol mol−1 H2O/CO.

Average reaction T GHSVd XCO Total tar[°C] [h−1] [% v/v] [mg/Nm3

db]

424 Case C 53.1 10.8 ± 0.1

The catalyst showed a relevant activity since a CO conversion closeto 53% was achieved with a total tar concentration in the syngas ofabout 10.8 ± 0.1 g/Nm3

db, at Case C GHSVd. Fouling deactivation mech-anism originated by tar deposition onto the catalytic surfacewas not ob-served and tar reforming via catalytic cracking was not found.

4. Conclusions

The high temperaturewater gas shift reaction over a Fe/Cr industrialcatalyst was investigated in a pilot scale plant consisting of dual fixedbed reactors arranged in series, and a tar-rich synthesis gas derivedfrom a biomass gasification CHP plant was used as the feed-stream.The reaction parameters were varied and the reaction temperatureand gas hourly space velocity effects were studied.

Increases in CO conversion, H2 purity and production flow rate wereobserved with increasing temperature while decreasing space velocityseems to influence CO conversion only. The steam to CO molar ratiodid not influence the catalyst performance.

The hydrogen sulfide loading effect was investigated and cata-lyst stability was studied. A decreased catalytic activity was ob-served with increasing H2S concentration, although the catalystmanifested a good resistance to hydrogen sulfide poisoning deacti-vation. The catalyst showed a good stability, an equilibrium regard-ing hydrogen sulfide adsorbed and desorbed was observed and anincrease of activity for the catalyst of the Reactor 2 was highlightedafter a time on-stream higher than 300 h with decreasing inlet H2Sconcentration.

The tar effect was studied at the end of H2S loading test, after a totaltime of exposure of the catalyst to tar rich syngas higher than 800 h. Thecatalyst showed a good activity since a significant tar concentrationwasmeasured. Fouling deactivationmechanism originated by tar depositiononto the catalytic surface was not observed and tar reforming via cata-lytic cracking was not found.

Acknowledgments

The authors wish to thank the industrial partners involved in theHydrogen from biomass for industry project, the Güssing biomass CHPplant for supplying the synthesis gas and Clariant International Ltd. forsupplying the catalyst. The investigations were carried out within theCOMET funding competence centre program and the Austrian Climateand Energy Fund.

48 S. Chianese et al. / Fuel Processing Technology 132 (2015) 39–48

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