1-s2.0-S0009250999004443-main

q This paper is dedicated to Prof. Manfred Baerns on the occasion of his65th birthday.

1Present address: Lurgi OG l Gas Chemie GmbH, Frankfurt/Main,Germany.

2Present address: University of New Brunswick, Fredericton, Cana-da.

3Present address: Bayer AG, ZT-TE 4.4, Geb.E41, D-51368 Lever-kusen, Germany.

*Corresponding author. Tel.: #49-234-700-4102; fax: #49-234-709-4115.

E-mail address: [email protected] (L.Mleczko).

Chemical Engineering Science 55 (2000) 3955}3966

Reaction engineering investigations of CO2

reformingin a #uidized-bed reactorq

T. Wurzel1, S. Malcus2, L. Mleczko*,3Lehrstuhl fu( r Technische Chemie, Ruhr-Universita( t Bochum, D-44780 Bochum, Germany

Received 16 November 1998; accepted 21 July 1999

Abstract

CO2

reforming of methane to synthesis gas over an Ni (1 wt%)/a-Al2O

3catalyst was studied in lab-scale #uidized-bed reactors

(ID"3,5 cm). In the whole range of reaction conditions (pCH4

"pCO2

"25}45 kPa, pN2

"10}50 kPa, ¹R"700}8003C,

Hmf

"3,4,5 cm, m#!5

/<Q "2.8}7.3 g s ml~1, u/umf

"6.5}11.8) a stable isothermal operation was achieved. The catalytic performancestrongly depended on the oxidation state of the catalyst. When applying a reduced catalyst initial yields of carbon monoxide andhydrogen near the thermodynamic equilibrium were obtained. However, a slow decrease of methane conversion and syngas yieldcaused by carbon deposition was observed. The fresh unreduced catalyst exhibited signi"cantly lower activity. The in situ reducedcatalyst was more active but yielded CH

4and CO

2conversions lower than predicted by the thermodynamic equilibrium. The

reaction was not in#uenced by interphase gas exchange. Based on these results, reaction engineering modeling and simulation yieldeda global kinetic model which described the experimental data with an error of less than 10% was developed. ( 2000 Elsevier ScienceLtd. All rights reserved.

Keywords: CO2-reforming; synthesis gas; #uidized-bed modeling

1. Introduction

In recent years, considerable attention has been paid touse methane which is the main component of natural gasas feedstock for the chemical and petrochemical industry.The indirect route via synthesis gas (syngas) remainsthe main way to convert natural gas into liquid fuels,methanol, ammonia or oxygenates. Currently, syngas ismainly produced by steam reforming of methane (Adris,Pruden, Lim & Grace, 1996; Qin, Lapszewicz & Jiang,1994 ). However, the H

2: C

Oratio of 3 : 1 obtained by

steam reforming is higher than that needed for Fischer-Tropsch or methanol synthesis. In turn, further gas makeup is necessary in order to obtain syngas which meets therequirements of the down-stream technologies.

CH4#CO

2"2CO#2H2,

*RH0

298 K"#247.9 kJ/mol. (1)

Lately carbon dioxide reforming of methane (see Eq. (1))has gained increasing interest as an alternative process toproduce syngas (Edwards, 1995). This process yields thelow H

2: CO ratio of 1 : 1 which is desirable for synthesis

of oxygenated chemicals. CO2

reforming can also beused for producing high-purity CO (Teuner, 1985; Kurz& Teuner, 1990). Furthermore, this reaction is of envir-onmental signi"cance since it consumes two moleculeswhich contribute signi"cantly to the greenhouse e!ect(Rostrup-Nielsen, 1994). Finally due to the high en-dothermicity of the reaction it can be used for the chem-ical energy storage (Wang, Lu & Millar, 1996; Gadalla& Sommer, 1989). Carbon dioxide reforming still has nocommercial application by itself (Edwards & Maitra,1994) but it is used in mixed reforming processes toreduce the H

2: CO ratio (Teuner, 1987).

0009-2509/00/$ - see front matter ( 2000 Elsevier Science Ltd. All rights reserved.PII: S 0 0 0 9 - 2 5 0 9 ( 9 9 ) 0 0 4 4 4 - 3

Nomenclature

AT

cross-sectional area of reactor, m2

cB,j,i

concentration of component i in the j5) seg-ment of the bubble phase, mol m~3

cE,j,i

concentration of component i in the j5) seg-ment of the emulsion phase, mol m~3

dB,0

initial bubble diameter, mdB,j

local bubble diameter in j5) segment, mdP

particle diameter, lm, m*RH0

298Kstandard reaction enthalpy, kJ mol~1

emf

porosity at minimum #uidization condition,dimensionless

eB,j

local bubble hold-up in the j5) segment, di-mensionless

g earth acceleration, m s~2

ID inner diameter, cmh measuring height, cmH height, cmi component, dimensionlessj segment number in the #uidized bed, dimen-

sionlessK

iadsorption constant of component i, atm~1

KP,l

thermodynamic equilibrium constant of re-action l, var.

kl

kinetic constant of reaction l, var.kBE,i

mass transfer coe$cient in the j5) segment,s~1

l reaction number, dimensionlessjB

average life time of a bubble, sm

#!5mass of catalyst, g

mf at minimum #uidization condition, dimen-sionless

mN*

mass of nickel, gnR#0/7,BE,i,j

convective #ow of component betweenbubble and emulsion phase in the j5) seg-ment of the #uidized bed due to increase oftotal numbers of moles, mol s~1

N03*&

number of distributor holes, dimensionlessNR number of reactions dimensionlessm

#!5mass of catalyst, g

gg

gas viscosity, Pa s~1

p partial pressure, kPaQ

p,lreaction quotient of reaction l (Q

P,-"Ppvi,l

i,j,

var.r reaction rate, mol s~1 kg~1

#!5rB,j,l

rate of reaction l in the j5) segment of bubblephase, mol s~1 kg~1

#!5rE,j,l

rate of reaction l in j5) segment of emulsionphase, mol s~1 kg~1

#!5o#!5

catalyst density, kg m~3

og

gas density, kg m~3

oP

particle density, kg m~3

STP standard conditions, dimensionlesst time, min, h¹ temperature, 3C¹

Rreaction temperature, 3C

TOS time on stream, hu super"cial gas velocity, cm s~1

uj

super"cial gas velocity in the j5) segment,cm s~1

u/umf

#uidization number, dimensionlessuB,j

bubble rise velocity in the j5) segment, m s~1

li,l

stechiometric coe$cient of component i inreaction l, dimensionless

<Q volume #ow, m3 s~1

<B,j

volume of bubble phase in the j5) segment,m3

<CP,j

volume of cloud phase in j5) segment, m3

<EP,j

volume of emulsion phase in j5) segment, m3

Xi

conversion of component i, %X0

iinitial conversion of component i, %

>i

yield of component i, %

Against this background numerous works dealingwith the investigation of CO

2reforming in "xed-

bed reactors over supported metal catalysts were re-ported (for a review see Wang et al., 1996). The reactiontemperatures applied in these investigations ranged from5003C (Masai, Kado, Miyako, Nishiyama & Tsuruya,1988; ErdoK helyi, Cserenyi & Solymosi, 1993) up to12203C (Yu, Choi, Rosynek & Lunsford, 1993). Amongseveral transition metals which were proposed as activecomponents, nickel (e.g. Vernon, Green Cheetham& Ashcroft, 1992; Gadalla & Sommer, 1989), platinum(e.g. Masai et al., 1988; Vernon et al., 1992; Solymosi,Kutsan & ErdoK helyi, 1991) and rhodium (e.g. ErdoK helyi

et al., 1993; Richardson & Paripatyadar, 1990) havefound widest research interest. For several of these cata-lysts yields of CO and H

2near the thermodynamic

equilibrium were reported (e.g. Vernon, Ashcroft& Cheetham, 1991; Richardson & Paripatyadar, 1990).However, catalyst deactivation due to carbon depositionoccurred. This deposition limited application of a "xed-bed reactor for this reaction (Rostrup-Nielsen & BakHansen, 1993; Seshan, ten Barge, Hally, van Keulen& Ross, 1994; Swaan, Kroll, Martin & Mirodatos, 1994;Au, Hu & Wan, 1994). Furthermore, in this reactor typesteep temperature gradients caused by the high endother-micity of the reaction occurred. These, in turn, promoted

3956 T. Wurzel et al. / Chemical Engineering Science 55 (2000) 3955}3966

Fig. 1. Schematic sketch of the #uidized-bed reactor applied.

carbon deposition (Rostrup-Nielsen & Bak Hansen,1993; Zhang & Verykios, 1994).

In order to eliminate these obstacles which impede thecommercialization of the title reaction the application ofa #uidized-bed reactor was proposed (e.g. Olsbye, Wurzel& Mleczko, 1997). Excellent heat dissipation originatedfrom the solids mixing in this reactor type can be utilizedto obtain good temperature control isothermal opera-tion. Application of small particles should allow to over-come the intra-particle mass transfer limitation whichseriously reduces catalyst e$ciency in "xed beds (Adriset al., 1996). Until now, only few experimental studieshave dealt with the performance of a #uidized-bed reac-tor for CO

2reforming (Mirzabekova, Mamedov, Aliev

& Krylov, 1992; Blom, Dahl, Salgtern, Sortland, Spjel-kavik & Tangstad, 1994; Olsbye et al., 1997). The re-ported results show that conversions near thermo-dynamic equilibrium can be attained although catalystdeactivation has not been prevented (Blom et al., 1994).These studies aimed mainly at catalyst testing with re-spect to their catalytic and mechanical stability.

Against this background reaction engineering invest-igations of CO

2reforming over an Ni/a-Al

2O

3in a la-

boratory-scale #uidized-bed reactor were carried out.The work aimed at the optimization of reaction condi-tions for obtaining high syngas yields as a base for theestimation of kinetic parameters. Special attention waspaid to the in#uence of hydrodynamic and reaction con-ditions on the deactivation of the catalyst.

2. Method of investigation

2.1. Experimental

2.1.1. CatalystThe Ni (1 wt%)/a-Al

2O

3catalyst was prepared by

incipient wetness technique from a-Al2O

3(Janssen) with

an aqueous solution of Ni(NO3)2) 6H

2O (Merck). After

drying for 12 h at 1003C the catalyst precursor was cal-cined in air at 4703C for 10 h. Particle diameters in therange 71}160 lm were used. The experimentally deter-mined minimum #uidization velocity using nitrogen asthe #uidizing gas amounted to u

.&,8003C"0.006 m/s.

2.1.2. Catalyst characterizationThe fresh unreduced as well as samples of Ni/Al

2O

3catalyst used in di!erent series of experiments were char-acterized by applying several surface and bulk tech-niques. Carbon deposits were determined by means oftemperature programmed oxidation (TPO). X-ray dif-fraction (XRD) analysis was performed in order to deter-mine the phase composition of the catalyst samples andto correlate the catalytic activity to speci"c phases.Transmission electron microscopy (TEM) analysis ofdi!erent samples aimed at the visualization of nickelparticles and carbon deposits.

2.1.3. ApparatusExperiments were carried out at atmospheric pressure

in two #uidized-bed reactors made of quartz with innerdiameters of ID"3 cm and 5 cm (see Fig. 1). In thepreheating section (¸"80 cm, ID"2 cm) the feed gaswas warmed up to 4003C before distributing througha porous quartz plate (d

103%"40}90 lm). The bed tem-

perature was controlled by heating the catalytic sectionfrom outside. In order to reduce particle entrainmenta disengaging section and an internal cyclone werelocated at the top of the reaction zone.

For measuring the bed temperature, thermocoupleswithin a quartz tube were located in three axial positionsof the #uidized bed. Reactants (CH

4, CO

2and N

2serv-

ing as internal standard) and products (H2

and CO) wereanalyzed by on-line gas chromatography. The amount ofwater which was condensed down-stream from the reac-tor was calculated from the oxygen balance. Concentra-tion pro"les were measured applying an axial movablesample tube (OD"0.8 cm). A detailed description ofthe experimental equipment is given elsewhere (Mleczko,Pannek, Rothaemel & Baerns, 1996a).

T. Wurzel et al. / Chemical Engineering Science 55 (2000) 3955}3966 3957

Table 1Correlations used for the calculation of hydrodynamic parameters

Parameter Correlation Reference

umf u

mf"g

g

(33.72#0.0408d3pog(o

p!o

g)g~2

g)1@2!33.7

(ogdp)

Wen and Yu (1966)

kBE,j

kBE,j

"0.11/dB,j

Kobayashi and Arai (1967)

dB,0

dB,0

"1.3[(ATu0(N

03*&A

T)~1)2g~1]0,2 Werther, 1992

dB,j

d

dhj

dB,j

"A2e

B,j9p B

1@3!

dB,j

3juB,j

Werther (1992)

eB,j

eB,j

"0.8(uj!u

mf)/u

B,jWerther (1992)

jB

jB"280u

mf/g Werther (1992)

dB,j u

B,j"0.8(u

j!u

mf)#0.71]3.2Jgd

B,jWerther (1992)

2.1.4. Experimental conditionsIn all experiments a CH

4: CO

2ratio of 1 : 1 was

applied. Nitrogen-dilution was varied between 10 and50 kPa. The reaction was investigated between 700 and8003C. The reactor was operated in the bubbling regime(u/u

mf"6.5}11.8). Catalyst mass was varied between 106

and 177 g. The other experimental conditions are givenin the following part.

2.2. Modeling

2.2.1. Reactor modelThe hydrodynamics of a #uidized bed was described

by a model that originated from the bubble assemblagemodel of Kato and Wen (1969). In this model the#uidized bed is divided into segments. The height of thesegments corresponds to the local bubble diameter. Themodel is based on the two-phase theory of #uidization.According to this theory, the bed is divided into twophases, i.e. the emulsion phase and the bubble phase.Between the phases exchange of gas takes place; theinterphase gas exchange coe$cient was calculated fromthe correlation of Kobayashi and Arai (1967). Throughthe emulsion gas #ows with the minimum #uidizationvelocity which was calculated from the correlation ofWen and Yu (1966). The excess gas #ows in the form ofbubbles. The bubble phase consists of particle-free bub-bles that are surrounded by cloud and wake. The cloudand wake have the same porosity as the emulsion phaseand travel together with bubbles. For describing thedynamics of bubbles, i.e. bubble diameter and velocitythe correlation of Werther (1992) was used. The correla-tion of Werther describes growth of bubbles as a result ofcoalescence and splitting. However, during the reformingreaction local bubble diameter, super"cial gas velocityand the interphase gas exchange are also in#uenced bythe increase of gas volume due to chemical reaction. The

model equations are given below (Eqs. (2) and (3)). Thecorrelations used for hydrodynamic modeling are sum-marized in Table 1 (refer also Mleczko, Ostrowski& Wurzel, 1996b; Wurzel & Mleczko, 1998).

Model equation for emulsion phase:

ATumf

[cE,i,j~1

!cE,i,j

]#kBE,j<B,j

[cB,i,j

!cE,i,j

]

#

NR+l/1

li,l

rE,j,l

(1!emf

)o#!5<

EP,j!n

#0/7,BE,i,j"0. (2)

Model equation for bubble phase

AT(u

j!u

mf)[c

B,i,j~1!c

B,i,j]!k

BE,j<B,j

[cB,i,j

!cE,i,j

]

#

NR+l/1

li,l

rB,j,l

(1!emf

)o#!5<CP,j

#n#0/7,BE,i,j

"0 (3)

2.2.2. Rate equation and estimation of kinetic parametersReaction rate of CO

2reforming was calculated using

a Langmuir}Hinshelwood rate expression. Becausewater}gas-shift reaction takes place simultaneously dur-ing CO

2reforming (Olsbye et al., 1997) this reaction was

also taken into account to describe the experimentalresults. Reaction rate of the water}gas-shift reaction wascalculated by applying "rst-order rate equation. Forboth reactions vanishing reaction rates when approach-ing equilibrium were taken into account (see Eqs. (4) and(5)). These rate equations were already applied for de-scribing kinetics of dry reforming of methane over theNi/La/Al

2O

3catalyst (Olsbye et al., 1997).

rCO2~R%&

"

kCO2~R%&

pCH4

pCO2

(1#K1pCH4

#K2pCO

)(1#K3pCO2

)

]A1!Q

P,CO2~R%&K

P,CO2~R%&B, (4)


Table 2Frequency coe$cients estimated for CO

2-reforming and water}gas-shift reaction and adsorption constants adapted from Olsbye et al. (1997)

Reaction k0

K1

K2

K3

(mol s~1g#!5

atm~2) (atm~1) (atm~1) (atm~1)

CO2

reforming 0.0150$0.0039 0.52 10 27Water}gas-shift reaction 0.0120$0.0041

Fig. 2. Axial temperature as well as axial concentration pro"les ofreactant (CH

4, CO

2) and product (CO, H

2, H

2O) components concen-

tration pro"le (CH4

: CO2

: N2"1 : 1 : 0.86, u/u

mf"8.6, H

mf"7 cm,

¹R"8003C).

rWGS

"kWGS

PCO2

PH2A1!

QP,WGS

KP,WGS

B. (5)

Kinetic parameters were determined by comparingmodel predictions (using the aforementioned reactormodel) with the experimental data reported in this paperwhereas adsorption constants were adapted from Olsbyeet al. (1997) (see Table 2). For the non-linear parameteroptimization a random search method was applied. Inthis routine the goal function was built as the sum ofsquares of di!erences between measured and predictedconversions of methane and carbon dioxide. The detailsof the optimization routine are presented elsewhere(Stansch, Mleczko & Baerns, 1997).

3. Results and discussion

3.1. Reactor operation and catalytic performance

3.1.1. Reactor operationThe Ni/a-Al

2O

3catalyst was mechanically stable and

well #uidizable in the whole range of the investigatedreaction conditions. From the bed with the height of5.0 cm (m

#!5"177.2 g) during 129 h on stream 5.3 g cata-

lyst was lost; the average rate of entrainment amountedto 0.041 g h~1. The good #uidizability of the catalystresulted in a good temperature control (see Fig. 2). Thebubbling bed zone showed isothermal behavior. How-ever, in the grid zone (0}1 cm) a temperature gradient upto 80 K occurred. The temperature drop in the distribu-tor zone was signi"cantly higher than the one measuredin the same reactor for the oxidative coupling of methane(Mleczko et al., 1996a) and for the catalytic partial oxida-tion of methane to syngas (POx) (Mleczko & Wurzel,1997). This indicates that the temperature drop is notonly caused by the cold inlet gases but also by the highendothermicity of the reforming reaction. In the free-board the temperature increased up to 8253C since thispart of the reactor was still heated.

3.1.2. Catalytic performanceHighest conversions and syngas yields were achieved

at the following reaction conditions: ¹R"8003C,

Hmf

"5 cm, u/umf

"11.8, m#!5

/<QSTP

"4.7 g s ml~1,pCH4

: pCO2

: pN2

"1 : 1 : 2. Methane and carbon dioxide

conversion amounted to 90 and 93%, respectively. Hy-drogen and carbon monoxide yield amounted to 81 and90%, respectively; this corresponds to a H

2: CO ratio of

0.9.These results are in line with the published data. In

almost all studies conversions near the thermodynamicequilibrium were achieved. No di!erence between#uidized and "xed beds could be recognized. A compari-son between conversions and yields obtained in this workand those reported in the literature is di$cult due thedi!erences between reaction conditions and loading ofthe catalyst with an active component. In order to over-come this obstacle the activity of the investigated catalystwas compared with other systems on the basis of themass of the active component. This approach is justi"edby the linear dependency of the catalyst activity on the Niloading (Goetsch, Say, Vargas & Eberly, 1989). The con-tact times reported for "xed beds were shorter than theone applied in this work (m

N*/<Q

STP"0.047 g s ml~1). Al-

though the contact times for achieving thermodynamicequilibrium are comparable with those reported in theliterature, the activity of the catalyst applied in this workwas signi"cantly lower than the that of the catalyst pre-pared by SINTEF (Blom et al., 1994). Finally it should betaken into account that the validity of the linear depend-ence of the catalytic activity on the Ni loading is limited,e.g. for Ni/NaY}zeolithe catalysts the activity passed


Fig. 3. Dependence of methane and carbon dioxide conversions ontime on stream applying a fresh catalyst (CH

4: CO

2: N

2"1 : 1 : 0.22,

u/umf

"8.6, ¹R"8003C, H

mf"5 cm).

Fig. 4. XRD pattern of the fresh (unreduced) catalyst (a) and (b) of thecatalyst used for CO

2reforming after 120 h on stream

(CH4

: CO2

: N2"1 : 1 : 2, u/u

mf"11.8, H

mf"5 cm, ¹

R"8003C).

through a maximum with increasing nickel loading(Wang et al., 1996).

Since it is generally accepted that reforming reactionsare the rate-determining steps during partial oxidation ofmethane to syngas a comparison between CO

2reform-

ing and partial oxidation in #uidized beds can be made.When performing POx over various supported nickelcatalysts methane conversions close to the thermodyn-amic equilibrium were achieved at signi"cantly shortercontact times than in this work (Bharadwaj & Schmidt,1994; Olsbye, Tangstad & Dahl, 1994; Mleczko & Wur-zel, 1997; Santos, Menendez, Monzon, Santamaria, Miro& Lombardo, 1996). The fact that CO

2reforming itself is

slower than when it is performed during partial oxidationconditions was already stated in other works (e.g. Peters,Rudolf & Voetter, 1955; Lapszewicz & Xiang, 1992). Thissuggests that the reforming reactions are promoted whensome oxygen is present. In order to explain this e!ect itcan be postulated that carbon deposits which limit theoverall catalytic activity are burnt o! easier with oxygenthan with carbon dioxide. Furthermore, it is supposedthat the methane pyrolysis which is assumed to be therate-determining step of the reforming reactions can beaccelerated since hydrogen is withdrawn from adsorbedmethane by adsorbed oxygen.

3.2. Parameters inyuencing catalytic activity and stability

3.2.1. Activity of the fresh catalystIn the fresh (not-pretreated) catalyst nickel is com-

pletely oxidized. Therefore, when measuring the initialactivity of the catalyst the activity of the oxidized NiOsites for CO

2reforming can be determined. Furthermore,

e!ect of in situ reduction on catalytic activity duringreaction can be analyzed. Moreover, the in#uence of thereaction conditions on the catalyst state can be observedby determining the course of conversions over time onstream. In the experiment the methane-to-carbon dioxideratio amounted to 1 : 1. The feed was diluted with nitro-gen (p

N2"30 kPa). The height of the settled bed and

#uidization number (u/umf

) amounted to 5 cm (m#!5

"

177.2 g) and 8.6, respectively.The fresh catalyst showed a very low activity for CO

2reforming, e.g. the initial conversions of methane andcarbon dioxide amounted to 9 and 13%, respectively (seeFig. 3). Under reaction conditions activity of the catalystincreased. The characterization of the catalyst surfaceindicated that these changes in the catalytic activity canbe correlated with the oxidation state of Ni sites. XRDanalysis showed that the fresh catalyst contained com-pletely oxidized nickel (see Fig. 4a). The low initial con-versions indicate that oxidized nickel sites exhibit lowcatalytic activity for CO

2reforming. Therefore, slow

continuous reduction of nickel sites by non convertedmethane resulted in the increase of conversion. Thisexplanation is con"rmed by XRD (see Fig. 4b) and TEM

analysis of the used catalysts (after 120 h on stream)which showed that during CO

2reforming metallic Ni

sites were formed. The reduction process is promotedby increasing conversion since the partial pressure of


Fig. 5. Dependence of methane and carbon dioxide conversion (a) and(b) yield on time on stream for a regenerated (H

mf"5 cm) and a non-

regenerated (Hmf

"3 cm) catalyst (CH4

: CO2

: N2"1 : 1 : 2,

u/umf

"11.8, ¹R"8003C).

Fig. 6. Intensity of the CO2

signal during TPO for two catalyst samplesdi!ering in time on stream (CH

4: CO

2: N

2"1 : 1 : 2, u/u

mf"11.8,

Hmf

"5 cm, ¹R"8003C).

hydrogen which is a more e$cient reducing agent thanmethane increases during the activation period. The pos-tulated explanation for the increase of the activity of freshNi/Al

2O

3catalyst agrees well with published informa-

tion; also other investigators reported that an increase ofthe reduction degree of supported metal catalysts (Ni,Co) resulted in higher conversions when performing dryand steam reforming (Au et al., 1994; Dissanayake,Rosynek, Kharas & Lunsford, 1991; Santos et al., 1996).

After a period of 13 h the catalyst attained a stationarystate; no further substantial changes in conversions ofmethane and carbon dioxide occurred. Final conver-sions, that amounted to 46% for methane and 58% forcarbon dioxide were lower than those calculated for thethermodynamic equilibrium. In contrast to the resultsreported for the catalytic partial oxidation of methane tosynthesis gas for which methane conversions near thethermodynamic equilibrium after an in situ activation ofthe catalyst were achieved (e. g. Dissanayake et al., 1991;Santos et al., 1996), "nal conversions during CO

2reform-

ing were at comparable contact times signi"cantly lowerthan those predicted by thermodynamics.

3.2.2. Catalyst deactivationCatalytic stability was investigated by studying two

catalysts applied for dry reforming over a period of 60 hthat had exhibited di!erent initial activities. Di!erentcatalytic activity was established by treating the "rstcatalyst with O

2/N

2(1 : 1) for 1 h at 6003C after 60 h on

stream (CO2

reforming). The second catalyst was notpretreated prior to reaction. The bed with the stationaryheight of 5 cm for the "rst catalyst and 3 cm for thesecond one was #uidized with a gas velocityu"7.1 cm s~1 (u/u

mf"11.8). In both experiments the

reaction temperature and the methane-to-carbon dioxideratio amounted to 8003C and 1 : 1, respectively. The feedwas diluted with nitrogen (p

N2"50 kPa). The experi-

ments covered a period of approx. 100 h.For both catalysts two periods can be recognized (see

Fig. 5). In the "rst period the activity of the catalystvaried with TOS. In the second period both catalystsattained constant catalytic activity. The catalyst that hadexhibited high initial activity (X0

CH4"90%, X0

CO2"

93%) deactivated; after 50 h on stream conversions drop-ped to 78% for methane and 86% for carbon dioxide.Simultaneously, yields dropped from 82 to 75% for hy-drogen and from 90 to 83% for carbon monoxide. TPOmeasurements performed for the Ni/Al

2O

3catalyst (see

Fig. 6) indicate that during CO2

reforming continuouscarbon deposition occurred, which is assumed to beresponsible for the loss of activity. After 8 h on streamonly small amounts of carbon were deposited on thecatalyst which could be burnt o! at a temperature ofaround 6803C. In contrast, the specimen which was takenat the end of the reaction period (approx. 100 h) showeda substantial larger amount of carbon deposits which

could be burnt o! continuously over the whole range oftemperature. Furthermore, a second catalyst phase ofblack color was built up. The formation of the carbondeposits on the catalytic surface was also con"rmed by


Fig. 7. TEM photography of the catalyst after 100 h of CO2

reforming(CH

4: CO

2: N

2"1 : 1 : 2, u/u

mf"11.8, H

mf"5 cm, ¹

R"8003C).

Fig. 8. Dependence of methane conversion on time on stream with andwithout water addition (CH

4: CO

2: N

2"1 : 1 : 0.86, u/u

mf"8.6,

Hmf

"5 cm, ¹R"8003C).

XRD and TEM measurements. Furthermore, TEM anal-ysis indicates that a shell of carbon was built up aroundnickel crystallites (see Fig. 7). The blocking of active sitesby carbon deposits as the reason for the loss of catalyticactivity during CO

2reforming was also postulated by

other investigators (e.g. Rostrup-Nielsen & Bak Hansen,1993; Bitter, Hally, Seshan, van Ommen & Lercher, 1996).

The initial conversions and yields measured over thecatalyst that exhibited low initial activity amounted to66% for methane, 76% for carbon dioxide, 62% forhydrogen and 73% for carbon monoxide. For this cata-lyst conversions and yields rose in the initial period withtime on stream. Conversions increased to 73% for meth-ane and 82% for carbon dioxide. Simultaneously, yieldsrose to 72% for hydrogen and 78% for carbon monox-ide. The low initial activity of this catalyst charge isassumed to result from the passivation of deposited car-bon during the o!-stream period. The explanation thatthe aging of carbon deposits can be responsible for thedecay of catalytic activity is supported by the workperformed by Buyevskaya, Wolf & Baerns (1994) in theTAP reactor over an Rh catalyst. The initial increase ofconversions and yields observed during the activationperiod can be attributed to the substitution of the pas-sivated carbon by fresh, active carbon deposits during thereversed Boudouard reaction. When the passivated car-bon was completely substituted by fresh carbon depositsa stationary state of constant catalytic activity wasachieved.

The stationary state that was obtained after approx.50 h independent of the initial catalyst state is attributedto a dynamic equilibrium between reactions that con-sume deposited carbon (reversed Boudouard reaction)and produce carbon deposits (methane pyrolysis). Due tocontinuous catalyst deactivation when performing CO

2

reforming in "xed-bed reactors (e.g. Swaan et al., 1994)and the steady-state activity in the #uidized-bed theconclusion is drawn that catalyst deactivation is in-#uenced by the hydrodynamic conditions in the reactortype applied. In a #uidized-bed reactor deactivated cata-lyst particles are transported by means of the solid mix-ing into the carbon-dioxide-rich distributor zone wherethe carbon deposits can be consumed by the reversedBoudouard reaction. Therefore, a #uidized-bed reactorseems to be more suitable for performing CO

2-reforming

reaction than a "xed bed not only with respect to temper-ature control but also due to the slower deactivation ofthe catalyst. This is underlined by the di!erent deactiva-tion rates. For an Ni (6.2 wt%)/Al

2O

3}SiO

2catalyst the

deactivation rate in a "xed bed amounted to 3.7% h~1

(Swaan et al., 1994) whereas it amounted to 1.5% h~1

(Blom et al., 1994) and 0.6% h~1 (this work) in the#uidized bed.

3.2.3. Ewect of water on catalyst regenerationSince an excess of steam is known to prevent carbon

deposition during the steam-reforming process experi-ments were conducted in which the catalytic activity withand without steam addition (p

H2O"15 kPa) was investi-

gated. In these experiments the deactivated catalyst thatexhibited low initial activity was used. The methane-to-carbon dioxide ratio was held constant at 1 (p

CH4"

pCO2

"35 kPa). Nitrogen dilution amounted to 15 and30 kPa, respectively. The reforming reaction was carriedout at 8003C in a catalytic bed (H

mf"5 cm, m

#!5"64 g)

#uidized with a gas velocity u"5.2 cm s~1 (u/umf

"8.6).For both investigated feed-gas compositions a similar

dependence of catalytic activity on TOS was observed; inthe "rst 15 h methane and carbon dioxide conversionincreased followed by constant activity (see Fig. 8). Waterin the feed had no e!ect on the period of activation.However, slightly lower conversions were measuredwhen applying the dry feed due to thermodynamic


Fig. 9. Dependence of experimental (exp.) methane and carbon dioxideconversions and those predicted by thermodynamic equilibrium (eq.)on reaction temperature (CH

4: CO

2:N

2"1 : 1 : 0.86, u/u

mf"7.5,

Hmf

"3 cm).

reasons. The addition of water to the feedstock also hadno signi"cant in#uence on the catalytic stability of theNi/a-Al

2O

3catalyst. This behavior can be explained

by the stronger oxidation potential of carbon dioxidecompared to water. Since the catalyst is continuouslytransported in the carbon-dioxide-rich distributor zonecarbon gasi"cation by carbon dioxide is dominantcompared to water.

3.3. Concentration proxles

In order to elucidate the reaction pathway of CO2

reforming in a #uidized bed axial concentration pro"leswere measured. In these experiments the minimum bedheight H

mfand the p

CH4: p

CO2ratio amounted to 7 cm

and 1 : 1, respectively. Measurements were performed ina bed #uidized with a gas velocity of u"5.2 cm s~1

(u/umf

"8.6) carrying out the reaction at 8003C.The measured concentration pro"les are presented in

Fig. 2. The most signi"cant changes of the concentrationstook place within the "rst two centimeters of the #uidizedbed. In this region a decrease of the concentrations of thefeed gases methane and carbon dioxide from 35% at theinlet to 24.3% for methane and 17.6% for carbon dioxidewas observed. Simultaneously, the concentrations of theproducts hydrogen and carbon monoxide increased upto 11.4 and 14.8%, respectively. Also in the upper part ofthe bed, the concentrations of methane and carbon diox-ide decreased continuously, whereas the concentrationsof the selective products hydrogen and carbon monoxideincreased. However, with increasing distance from thegas distributor the concentration gradients decreased.Water yielded a maximum of 4.4% at a bed height of3 cm followed by a #at decline of 4% at the end of thebed. At the end of the bed a decay of carbon monoxideconcentration and an increase of the concentrations ofthe feed gases methane and carbon dioxide were mea-sured. In the freeboard (8}12 cm) no further changes inconcentrations occurred although the temperatureruns through a maximum in this region. This indicatesthat the reaction is heteregeneously catalyzed whereasgas-phase reactions have no in#uence on the extent ofreaction.

In the whole bed, concentration of carbon dioxide waslower than the methane concentration. Also the hydro-gen concentration was always lower than that of carbonmonoxide. This indicates that beside the CO

2-reforming

reaction the water}gas shift reaction also takes placewhich results in a H

2/CO ratio below 1. With respect to

the reaction rates it can be expected that the one forwater}gas reaction is faster than the one for CO

2-reform-

ing reaction since even in the distributor zone where thehighest reaction rates for CO

2reforming occur the di!er-

ence between the methane and carbon dioxide concentra-tion is signi"cant. Simultaneously, formation of watertakes place. The product distribution in the bed is also

in#uenced by the steam-reforming reaction. This wouldexplain the axial increase of the H

2/CO ratio (0.42}0.68)

since steam reforming produces more hydrogen thancarbon monoxide. However, steam reforming has onlyminor e!ect on the methane conversion.

The concentration pro"les presented above di!er sig-ni"cantly from those measured during partial oxidationof methane for which almost constant gas compositionwas attained in the distributor zone (Olsbye et al., 1994;Mleczko & Wurzel, 1997). It is therefore assumed thatthermodynamic equilibrium for partial oxidation isreached signi"cantly faster compared to CO

2reforming.

As mentioned earlier it is concluded that CO2

reformingitself is slower compared to its reaction rate during par-tial oxidation.

3.4. Ewect of reaction conditions on catalytic performance

3.4.1. Ewect of temperatureTemperature dependence of conversion was studied

between 700 and 8003C. The methane : carbon dioxideratio in the feed gas and the height of the bed amountedto 1 : 1 and H

mf"3 cm (m

#!5"106 g). The partial pres-

sure of nitrogen and the #uidization number amountedto 30 kPa and 7.5, respectively.

Upon increasing temperature from 700 to 8003C theconversions of methane and carbon dioxide rose from 29to 77% and from 43 to 84% (see Fig. 9). Simultaneously,the yields of the selective products hydrogen and carbonmonoxide increased from 25 to 73% for hydrogen andfrom 40 to 81% for carbon monoxide. In the investigatedtemperature range conversions and yields predicted bythe thermodynamic equilibrium were not achieved. How-ever, the deviations between experimental values and theones predicted by thermodynamic equilibrium decreasedwith increasing temperature. At a temperature of 8003Cthese deviations amounted to 11% for X

CO2and 17% for

>H2

. Hydrogen selectivity also rose with increasing


Fig. 10. E!ect of hydrodynamic conditions on methane and carbondioxide conversions (CH

4: CO

2: N

2"1 : 1 : 0.86, ¹

R"8003C,

u/umf

"6.5}9.2, Hmf

"3}5 cm).

temperature. These results indicate that the reaction iskinetically controlled in the investigated range of reac-tion conditions.

3.4.2. Ewect of hydrodynamic conditionsIn order to investigate the in#uence of hydrodynamic

conditions on conversions and syngas yields contact timewas varied between 4.4 and 7.3 g s ml~1 by changing theheight of the bed and the gas velocity. Fluidization num-ber was varied between 6.5 and 9.2 and the bed heightbetween 3 and 5 cm (106.2}177.2 g cat.). Bed temperatureand methane-to-carbon dioxide ratio at the reactor inletamounted to 8003C and 1, respectively.

As expected the conversions of methane and carbondioxide strongly depended on contact time (see Fig. 10).A monotonous growth of conversions and yields wasachieved with increasing contact time. No signi"cantdi!erence between conversions was observed when in-creasing contact time by using higher beds or lower gasvelocities. However, conversion was slightly less sensitiveto changes of the contact time when this parameter wasin#uenced by varying the mass of catalyst than whenchanging gas velocity. An increase of the contact time byreducing the #uidization number e.g. by changing gasvelocity results in smaller bubbles. This, in turn, reducesthe back-mixing of gas and improves mass transportbetween bubbles and the emulsion phase. Since the di!er-ences in the sensitivity of conversions to the changes ofthe contact time did not di!er signi"cantly it can beconcluded that in the investigated range of reaction con-ditions CO

2reforming was not controlled by the inter-

phase gas exchange.The "nding that the reactor performance is not con-

trolled by mass transfer limitation is another di!erence tothe partial oxidation of methane which con"rms that therates of both reactions are signi"cantly di!erent. How-

ever, with respect to large-scale application mass transferlimitation may become important. Simulations per-formed for catalytic partial oxidation (Wurzel &Mleczko, 1998) indicate that in an industrial-scale reac-tor the rate of reforming reactions was limited by themass exchange between bubbles and the emulsion phase.In turn, longer contact times are required in order toachieve the same conversions as in the lab-scale reactor.This di!erence was explained by a signi"cant change ofthe hydrodynamic conditions due to the formation oflarger bubbles; in the lab-scale reactor bubble sizeamounted to approx. 0.5}1 cm (visually observed) where-as bubble diameters up to 20 cm were calculated for anindustrial unit.

3.5. Simulation results

The conversions of methane measured in the station-ary state were used for estimation of the frequency coe$-cients in the rate equations for the CO

2reforming and

water}gas-shift reaction. These conversions were com-pared with the ones by the previously presented model ofthe #uidized bed. In this approach the estimated kineticparameters are biased by the uncertainty in the descrip-tion of the bed hydrodynamics, i.e. the back-mixing andinter-phase gas exchange. Because studies of the e!ect ofthe contact time showed that the measured conversionswere not limited by the inter-phase gas exchange thisapproach is justi"ed. Additionally, the correlations usedfor prediction of the bubble dynamics were veri"ed incold-#ow studies with the Ni/Al

2O

3catalyst (Wurzel

& Mleczko, 1998).The estimated frequency coe$cients and their con"-

dence regions (based on a con"dence number of 0.95) arepresented in Table 2. With the developed model themeasured conversions of methane and carbon dioxidecould be reproduced with an error lower than 10% (seeFig. 11a). It has to be pointed out that all experimentaldata points were below those predicted by thermo-dynamics, i.e. the good accuracy of the model is not dueto the thermodynamic limitations. The validity of themodel was con"rmed by good prediction not only of theintegral data but also of the concentration pro"les (seeFig. 11b). The concentrations calculated for the bubbleand emulsion phase follow the experimentally measuredconcentrations. In line with the experimental "ndingsthat the inter-phase gas exchange in a lab-scale#uidized-bed reactor has only small in#uence on theconversions the calculated gas composition in the emul-sion phase di!ered only slightly from the one in thebubble phase. Nevertheless, the better agreement of thevalues predicted for the bubble phase with the concentra-tions measured experimentally indicates that the pre-dicted reaction rates are slightly underestimated. Thisconclusion is also in agreement with the results presentedin the parity plot. Thus, this model can be used as


Fig. 11. Parity plot of experimental and simulated methane and carbondioxide conversion (a) and (b) measured and calculated concentrationpro"les (conditions see Fig. 2) during CO

2reforming in a #uidized bed.

a valuable tool for scaling-up of a #uidized-bed reactorfor dry reforming of methane.

4. Conclusions

Carbon dioxide reforming of methane was investigatedin a laboratory-scale (ID"3, 5 cm) #uidized-bed reac-tor. The catalyst was well #uidizable and almost isother-mal conditions were obtained over the whole range ofreaction conditions studied. Highest conversions andsyngas yields were achieved applying the followingreaction conditions: T

R"8003C, H

mf"5 cm, u/u

mf"

11.8, m#!5

/<QSTP

"4.7 g s ml~1, pCH4

: pCO2

:pN2

"1 : 1 : 2.Methane and carbon dioxide conversion amounted to 90and 93%, respectively. Hydrogen and carbon monoxideyield amounted to 81 and 90%, respectively; this corres-ponds to a H

2: CO ratio of 0.9. On the basis of experi-

mental results a model of the #uidized-bed reactor for dryreforming of methane was developed. This model, thatwas able to reproduce experimental conversions with anerror of less than 10% can be used as a valuable tool forthe reaction engineering studies of di!erent reactor con-"gurations proposed for the title reaction (see e.g.Mleczko, Malcus & Wurzel, 1997).

It has been shown that the performance of the Ni/a-Al

2O

3catalyst depended strongly on the oxidation state

of the active component. Oxidized nickel sites showeda very low catalytic activity compared to the reduced,metallic ones; applying a reduced catalyst initial yields ofcarbon monoxide and hydrogen near the thermodyn-amic equilibrium were obtained, whereas the freshunreduced catalyst exhibited signi"cantly lower activity.With increasing time on stream a slow decrease of meth-ane conversion and syngas yield was observed. Carbondeposition was found to be responsible for the deactiva-tion of a highly active catalyst. After several hours onstream constant level of conversions and yields wasachieved. This makes the signi"cant di!erence in thedeactivation behavior of the Ni/Al

2O

3catalyst in the

"xed-bed and the #uidized-bed reactor. Addition ofwater to the feedstock showed no signi"cant in#uenceeither on catalyst regeneration or on conversions.

The in situ reduced catalyst yielded CH4

and CO2

conversions lower than predicted by the thermodynamicequilibrium. CO

2reforming in the laboratory reactor

was not in#uenced by inter-phase gas exchange. Thisconclusion was also drawn from simulations of the lab-scale unit which resulted in a good reproduction of theexperimental data. With respect to the process commer-cialization the #uidized-bed reactor seems to be a suit-able reactor type. However, several serious problems likee!ect of pressure on the catalyst deactivation has to beinvestigated. Therefore, in the industrial-scale externalcatalyst regeneration might be necessary (Mleczko et al.,1997). Furthermore, the problem of metal dusting due tothe high CO content in the reformed gas has to be takeninto account with respect to the construction of the re-formed gas boiler (Grabke, Krajak, MuK ller-Lorenz& Strau{, 1996; Hohmann, 1996).

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1-s2.0-S0009250999004443-main

Documents

reaction number

j5 segment of emulsionphase

j5 segment of bubblephase

lrate of reaction

j5 segment

iconcentration of component

dimensionlessj segment

cient of component