metal catalysts fluidized bed reactor.spain

International Journal of Hydrogen Energy 32 (2007) 4821–4829www.elsevier.com/locate/ijhydene

Production of hydrogen and carbon nanofibers by thermal decompositionof methane using metal catalysts in a fluidized bed reactor

J.L. Pinillaa, R. Molinera, I. Suelvesa,∗, M.J. Lázaroa, Y. Echegoyena, J.M. Palaciosb

aInstituto de Carboquímica, CSIC, Miguel Luesma 4, 50018 Zaragoza, SpainbInstituto de Catálisis y Petroleoquímica, CSIC, Campus Universidad Autónoma, Cantoblanco, 28049 Madrid, Spain

Received 27 December 2006; received in revised form 6 July 2007; accepted 11 August 2007Available online 24 October 2007

Abstract

Thermo catalytic decomposition (TCD) of methane using Ni–Cu–Al catalyst in a pilot scale fluidized bed is studied. The conventionalmethod of catalyst preparation based on co-precipitation is compared versus an easier preparation method based on the fusing of the metallicnitrates. Catalysts prepared by both methods shown similar behavior. Fluidodynamic studies have shown that TCD can be carried out in afluidized bed reactor without reactor clogging provided that a methane velocity of two times the minimum fluidization velocity is used. Thishigh spacial velocity resulted in a reduction of the fraction of methane converted, but not in the quantity of carbon deposited per gram ofcatalyst. The optimum gas velocity should be chosen in terms of hydrogen production rates and fluidization quality. A semi-continuous fluidizedbed installation has been started up and operated to produce at the scale of 60 l h−1 of hydrogen and 15 g h−1 of carbon nanofibers.� 2007 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved.

Keywords: Hydrogen production; Carbon nanofibers; Thermal decomposition of methane; Fluidized bed reactor

1. Introduction

Hydrogen is an emerging alternative to conventional fuelsto reduce CO2 emissions. It is generally accepted, that in thenear-to-medium term hydrogen production will rely on fossilfuels, primarily natural gas. Thermo catalytic decomposition(TCD) of natural gas, with carbon being captured as a solid ofadded value product appears an interesting alternative to steamreforming [1] for hydrogen production.

Feasibility of TCD in economical terms is very sensible tothe carbon selling price which depends on the properties of thecarbon obtained. The quality of carbon produced from TCDlargely depends on the operation conditions and the type of cat-alyst used. Using metal-based catalysts lead to the productionof carbon forms of high quality whose high selling price wouldcompensate the high cost of the catalyst. TCD of methane us-ing Ni and Ni–Cu catalysts to produce hydrogen and novel car-bonaceous materials has been reported by many authors [2–8].The use of solar heating has also been reported to produce

∗ Corresponding authors. Tel.: +34 976733977; fax: +34 976733318.E-mail address: [email protected] (I. Suelves).

0360-3199/$ - see front matter � 2007 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights reserved.doi:10.1016/j.ijhydene.2007.08.013

hydrogen and carbonaceous materials by methane decomposi-tion [9,10].

TCD of methane using nickel catalysts has been extensivelystudied by our group using a fixed bed reactor. In the firstapproach [11], a commercial Ni catalyst (Ni content 65% wt),was used in TCD of methane to gain knowledge on the influenceof the operating conditions on the catalyst deactivation and onthe properties of the deposited carbon. It was shown, that thehigher the methane flow rate used the shorter the catalyst life.At temperature of 700 ◦C, the concentration of hydrogen in theoutlet gas was close to the thermodynamic value. SEM andTEM examinations showed that the deposited carbon appearedeither as large filaments a few nanometres in diameter emergingfrom Ni particles or as uniform coatings. XRD, FT-Raman andXPS revealed that in all cases the deposited carbon was highlyordered graphite whose structure apparently did not depend onthe operating conditions.

In the second stage [12,13], Ni and Ni–Cu catalysts wereprepared using different methods to evaluate the influence ofthe amount of nickel as catalyst, the role of copper as a possiblealloying element with Ni and the catalyst preparation procedureon methane conversion and on the properties of the produced

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4822 J.L. Pinilla et al. / International Journal of Hydrogen Energy 32 (2007) 4821–4829

carbon. The prepared catalysts showed a similar behaviour tothe commercial one, even better when a small concentrationof copper was introduced in its composition. At 700 ◦C thecatalyst’s activity did not decay after 8 h on-stream (exceptfor the catalyst with a percentage of nickel below 30%) and aweight ratio of carbon to nickel between eight and eleven wasobtained without catalyst deactivation.

Among the different methods of catalyst preparation de-scribed in the literature, co-precipitation and impregnation arethe most used. These methods include washing and filteringsteps which are time and energy consumed and involve the useof huge amounts of water. These could be important drawbacksfor their use at an industrial scale. For that reason an easierpreparation method based on the fusing of the metallic nitrateswas also studied at the same operating conditions [12,13]. Per-formance tests showed that the hydrogen production was nothighly dependent on the preparation method used. The cata-lyst deactivation was never reached at the fixed times used inthe reactivity tests for any of the preparation methods studied.Ni:Cu:Al catalysts promoted the formation of a well-orderedgraphitic carbon with relative large crystal domain sizes. Ni:Alcatalysts, however, promoted the formation of turbostraticcarbon.

Different aspects of the catalyst preparation procedure havebeen studied in more detail [14]. First of all the effects of cat-alyst calcination temperature on the hydrogen yield and on thecarbon obtained have been analyzed. A wide range of calci-nation temperatures have been studied in order to simulate theregeneration step of an industrial installation where catalyst issubmitted to high temperatures to burn the carbon layer coat-ing the catalyst particle. The higher hydrogen yields were ob-tained using catalysts calcined at 600 ◦C. This behaviour is at-tributed to the fact that at 450 ◦C the conversion of the nitricsalts into oxides is not complete, while at temperatures higherthan 600 ◦C there is an increment in the metal crystal size. Con-sistently, the efficiency for methane decomposition decreases.

The work presented above was carried out in a fixed-bedquartz reactor 2 cm i.d., 60 cm height, using pure CH4 as feed-ing gas and following the experimental procedure previouslydescribed [11–14]. One of the main problems associated to thefixed bed operation is that the reactor was completely filledwith produced carbon and the flow path for the reactant gaswas blocked. This behavior has been reported by Lee et al. [15]where carbonaceous catalysts were used and by [16–18] wheremetal catalyst were used. Therefore, in a continuous processof catalytic methane decomposition, the carbon produced mustbe removed from the reactor in order to maintain the reactionactivity and to avoid plugging by the carbon produced. Theissues related to the development of a reactor suitable for cat-alytic methane decomposition with continuous withdrawal ofcarbon product were treated by Muradov [19], using carbona-ceous catalyst. The fluidized bed reactor (FBR) was selected asthe most promising reactor for a large-scale operation. FBRsare used successfully in a multitude of processes both catalyticand non-catalytic. A FBR system provides constant flow ofsolids through the reaction zone, which makes it particularlysuitable for the continuous addition and withdrawal of carbon

particles from the reactor. In an FBR, the bed of catalyst parti-cles behaves as a well mixed body of liquid giving rise to highparticle-to-gas heat and mass transfer. The FBR was proposedin order to overcome the reactor plugging problem due to car-bon deposition, which resulted in the shut down of the FBR.

More recently different works have been published on thebehavior of the active carbons [15] and carbon blacks [20] ascatalysts and the modelling and scaling up [21] of a FBR forTCD of natural gas.

The FBR has also been proposed as an alternative for thelarge-scale nanotubes production from high hydrogen contentgases and not methane by CVD [22–25]. Morançais et al. haverecently reported the large-scale nanotubes production fromethylene using a FBR [26]. Methane decomposition with si-multaneous hydrogen and nanotubes production has been pro-posed by Quian et al. [27] using a two staged FBR. Despite this,scarce information is available in literature in which hydrogenand carbon nanofibers are simultaneously produced in an FBR[28] and so that, one of the main objectives of the present paperis to demonstrate the feasibility of the simultaneous productionof carbon nanofibers and hydrogen in an FBR, avoiding theclogging problems associated to the fixed bed operation whenusing the Ni:Cu:Al catalysts previously tested.

2. Experimental

2.1. Catalysts

Two types of catalyst have been prepared following theprocedures previously described for the FBR experiments[12–14]:

(1) Co-precipitation: a catalyst with Ni:Cu:Al of 78/6/16 ratiowas prepared by co-precipitation from an aqueous so-lution of the respective nitrates with sodium carbonate.The precipitates were then washed, dried and calcinedat 450 ◦C.

(2) Fusion: a catalyst with Ni:Cu:Al of 78/6/16 ratio wasprepared by fusing nitric salt of nickel and copper nitratewith nitric salt of aluminium followed by decomposi-tion of the mixtures at 350 ◦C and subsequent calcinationat 450 ◦C.

All samples were ground to fine powder (particle size lowerthan 100 �m).

2.2. Experimental apparatus

Experiments were carried out in the experimental apparatusschematically shown in Fig. 1. The FBR (with i.d. of 0.065and 0.8 m in height) was made of kanthal. Kanthal is a metalalloy treated to form an oxides layer which is very stable athigh temperature, avoiding material drop during the reaction. Ahorizontal perforated plate with 3 mm holes is used to divide thereactor in two stages. Quartz glass is used to prevent materialfrom plugging the holes.

J.L. Pinilla et al. / International Journal of Hydrogen Energy 32 (2007) 4821–4829 4823

Fig. 1. Scheme of the experimental apparatus employed.

All process variables (pressures, temperatures and flow rates)are recorded in continuous mode by a personal computer. Thegas enters the reaction zone using a stainless steel tube heated at550 ◦C by an electric furnace Watlow with a maximum powerof 888 W. The reactor is heated to the desired reaction tem-perature using an electric furnace Watlow, (maximum power1775 W). Thermocouples type K (Thermocoax) were used formonitoring pre-heater (by one thermocouple placed into thestainless steel tube) and reactor temperatures (by two thermo-couples, one fixed on the outer reactor wall and the other placedin the FBR, 3 cm above the perforated plate). Hydrogen andmethane flow rates were controlled by mass flow controllers(Bronkhorst).

The combined pressure drop across the distributor and flu-idized bed was measured with a differential pressure transducer.The pressure drop through the distributor and quartz glass wasseparately measured, so that the pressure drop across the flu-idized bed by itself could be calculated from Eq. (1):

�Pbed = �Pbed+perforate plate − �Pperforate plate. (1)

2.3. Operating procedure

All experiments were conducted at atmospheric pressure.Prior to activity tests, all catalysts were subjected to a reductiontreatment using a flow rate of pure hydrogen of 80 l/h for 3 h at550 ◦C. Then, pure methane (99.99%) was fed into the reactorand decomposed by the catalyst. The catalyst mass used wasfixed to a value of 20 g. The composition of the outlet gaswas determined by gas chromatography: two packed columns,Molecular Sieve 13× and Porapack, and TCD detector wereused. Methane conversion was calculated from the followingexpression:

�CH4= %H2

200 − %H2. (2)

The carbon deposited during each run (Cdep) was determined bydirect weight. Moreover, an estimation of the evolution of thecarbon product generated can be accomplished by the following


correlation:

Cdep.calc = MC

�

∫ t

0Finlet · �CH4

· dt , (3)

where � is the standard molar volume, MC is the atomic weightof carbon, Finlet is the methane flow rate fed to the reactor(l(SPT)/min), �CH4

is the methane conversion and t is the totalrun time.

The minimum fluidization velocity (umf) has been deter-mined experimentally using nitrogen gas at the reaction tem-perature (700 ◦C), by measuring the pressure drop caused by aknown mass of carbonaceous product (200 g) generated duringthe TCD of methane. The method used the extrapolation of thelinear part of the pressure drop curve (�p) versus gas velocityplot up to the value corresponding to the maximum theoreticalpressure drop (�pmax = W/S, where W is the mass of carbonproduct (g) and S is the cross-sectional area in m2). An inertgas was used to avoid changes and agglomeration problems.

The experimental umf value obtained with N2 has been cor-rected to account the effect of using CH4, using the ratio be-tween the theoretical umf at the reaction conditions calculatedfor CH4 and N2 by the method proposed by Wen and You [29],using as constant those proposed by Adánez and Abadanes [32]for carbonaceous material (umf-CH4/umf-N2 = 1.4).

The mean particle diameter of the carbon product obtainedby TCD of methane used for the determination of the minimumfluidization velocity was determined by hand sieving a repre-sentative sample and recording the weight fraction retained oneach sieve. The particle mean diameter is then calculated bythe following equation:

dp = 1∑all(xi/dpi)

, (4)

where xi is the weight fraction in the interval, dpi is the meanparticle diameter in the interval.

Bulk density, �b, was determinate by measuring the mass ofa known volume of particles placed inside a graduated cylinder,without packing down the particles.

3. Results and discussion

3.1. Fluidodynamic study: determination of the minimumfluidization velocity

As previously mentioned, in TCD, carbon is depositedmainly on the catalyst as filaments emerging from the nickelparticles. As a consequence, density and shape of the particlesin the reactor, and so, their fluidodynamic behaviour dramat-ically change as growing of carbon filaments progresses. Inaddition, agglomerations of particles resulting in reactor plug-ging have been observed at bench scale [15–18]. All thesephenomena demonstrated that fluidization of this system isnot an easy work. In order to gain information about thefluidodynamic behaviour of the system, a set of fluidizationexperiments were conducted.

First, the minimum fluidization velocity umf of the carbonproduct generated during TCD of methane was determined ex-

Table 1Properties of the carbon product generated during long TCD run (16 h)

Catalysts preparation Apparent density Mean particle diameter Catalystmethod (g/cm3) (�m) (%)a

Co-precipitation 0.646 209 8.30Fusion 0.604 206 8.84

T = 700 ◦C; catalyst: 20 g; methane flow rate = 70 l/h.acalculated by ash analysis.

0.0 0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 1.8 2.00

1

2

3

4

5

6

7

2 x umf-ebumf-pb umf-eb

ΔPmax=W/S

ΔP (

mbar)

uo (cm(STP)/s)

fusionco-precipitation

Fig. 2. Pressure drop �P across the bed vs. superficial fluidization velocityu. Fluidization gas: nitrogen; temperature: 700 ◦C.

perimentally. Carbon product generated in long run experiments(16 h) carried out using the same spactial velocity as in the FBRexperiments [11–14], and with the catalysts prepared by thetwo methods previously mentioned was selected conservativelyas model compound to the determination of the umf , thoughthe conditions for the fluidization at the early stage of the reac-tion are lower. During TCD reaction, particles got bigger andbigger, increasing mean particle diameter. It was also observedthat the apparent density diminished after the experiment, dueto the lower density of the carbon materials deposited. Carbonproduct was sieved and the mean particle diameter was deter-minate from Eq. (4). A wide particle size distribution, between50 and 1000 �m was observed. The mean particle diametersfrom the particle size distribution obtained from sieve analysiswere found to be 208 �m. Important features of the particlesare given in Table 1. We can conclude that the carbon productparticles belong to the Geldart Group A. These solids fluidizeeasily, with smooth fluidization at low gas velocities and con-trolled bubbling with small bubbles at high gas velocities. FCC(fuel catalytic cracking) catalyst is representative of these typesof solids [30].

The experimental values of �p measured for increasing val-ues of u are shown in Fig. 2. A characteristic curve is ob-tained, typical of the solids with wide particle size distribution,which indicates that a partial fluidization phenomena occurs[31]. When the gas velocity u0 is increased through these beds


Table 2Gas velocity, flow rate and space velocity for the different fluidization velocities tested

u0N2 (cm(STP)/s) u0CH4 (cm(STP)/s) CH4 flow rate (l(STP)/h) Space velocity (l(STP)/h/gcat)

umf-pb 0.47 0.66 78 3.9umf-eb 0.72 1.01 120 62xumf-eb 1.44 2.02 240 12

of solids, the smaller particles are able to slip into the void spacebetween the larger particles and be fluidized while the largerparticles remain stationary. The partial fluidization occurs, giv-ing an intermediate pressure drop �p. When increasing gasvelocity, �p approaches to W/S, showing that all the solidseventually are fluidized. The fluidizing gas velocity which re-sults of the extrapolation of the linear part of the �p versusgas velocity plot up to the value corresponding to the maxi-mum theoretical pressure drop is referred to as the minimumfluidization velocity for the partial bed (umf-pb). This value cor-responds to a nitrogen gas velocity of 0.47 cm (STP)/s. At this�p, the bed was partially fluidized. As the fluidizing gas ve-locity was further increased, the pressure drop again increasedand a greater fraction of the bed was fluidized. From a valueof 0.72 cm (STP)/s and beyond, the pressure drop becomes al-most constant and did not change with further increase in thegas velocity. At this stage the entire bed was fluidized. The ni-trogen gas velocity at which �P = W/A is referred here asthe minimum fluidization velocity for the entire bed (umf-eb).The gas velocity was further increased up to a value of 1.44 cm(STP)/s, which corresponds to two times the minimum fluidiza-tion velocity for the entire bed, 2 × umf-eb. Table 2 shows theexperimental umf with N2, the umf corrected to account the ef-fect of using CH4, the methane flow rate necessary to operateat the three fluidization velocities selected as well as the spacevelocity used.

3.2. Effect of the catalyst preparation method

Two types of catalyst were used in the pilot plant describedabove in order to study its behaviour in the TCD of methanein an FBR. Fig. 3 shows the hydrogen evolution in the outletgas for Ni:Cu:Al catalysts prepared by co-precipitation and byfusion using a methane flow rate corresponding to the mini-mum fluidization velocity for the partial bed umf-pb at 700 ◦C.In bench scale experiments [11–14], the hydrogen productionwas closed to the maximum thermodynamic value, finding nodifference between both catalysts. In pilot plant scale reactor,just small differences were found, observing the catalyst pre-pared by co-precipitation showed a slight better behaviour. Forthe co-precipitation catalysts, the average hydrogen concentra-tion in the outlet gas is around 65%, which corresponds to aCH4 conversion of 48%. For the fusion catalyst, the averagehydrogen concentration in the outlet gas is around 60%, whichcorresponds to a CH4 conversion of 43%. For both samples,catalysts deactivation did not occur after a 16 h run. In fact,catalyst activity is only slightly decreasing in spite of the highincrease of the amount of deposited carbon with time. At the

0 100 200 300 400 500 600 700 800 900 10000

10

20

30

40

50

60

70

80

%H

2 (

vol.)

Time (min)

Co-precipitation

Fusion

Fig. 3. Hydrogen production (%vol) for co-precipitation and fusion catalystin pilot plant scale. T = 700 ◦C, catalyst: 20 g, methane flow rate = 70 l/h.

end of the test, more than 200 g of carbon product were ob-tained. Table 3 shows the main parameters of umf-pb TCD testin the long run experiments featured for co-precipitation andfusion catalysts: the reaction time, the amount of carbon de-posited calculated from Eq. (3), the carbon formation rate thevolumetric hydrogen production as well as the average methaneconversion. Similar results were obtained with both catalystpreparation method used. In the TCD of methane, the catalystis a consumable expensive material contributing substantiallyto the final cost of the hydrogen produced. Co-precipitationmethod for catalyst preparation involves filtration and washingprocesses that could increase the overall production costs anindustrial scale. Consequently, a method for catalyst prepara-tion as simple as possible, like salt fusing, is highly desirable[12,13]. The catalyst prepared by the fusion method was se-lected for rest of the fluidization experiments.

It is worth to mention that the pressure drop measured dur-ing the experiments showed an exponential increase after a fewhours of reaction (figures not shown). This is attributed to thepoor fluidization quality of the particle bed as the particle in-ventory in the bed increased.

3.3. Effect of the fluidization conditions

3.3.1. Kinetics measurementsIt is well known that gas velocity is the most important pa-

rameter in the FBR operation because the fluidization qualityof the gas–solid contacting pattern is strongly dependent on gas


Table 3Hydrogen and carbon production and methane conversion for the FBR experiments

Fluidization velocity Reaction time Space velocity Carbon deposited Cdep/h H2 production Conversion(h) (l/gcath) (l/h)

umf-pba 16 4 241 15.06 76.8 0.48

umf-pbb 15 4 226 15.07 68.8 0.43

umf-ebb 8 6 118 14.75 50.4 0.21

2xumf-ebb 8 12 131 16.38 57.6 0.12

aCo-precipitation catalyst.bFusion catalyst.

0 100 200 300 400 5000

10

20

30

40

50

60

70

80

% H

2 (

vol.)

Time (min)

umf-pb

umf-eb

2 x umf-eb

Fig. 4. Effect of gas velocity on the hydrogen production over fusion Ni:Cu:Alcatalyst at the temperature of 700 ◦C.

velocity [15]. Fig. 4 shows the effect of the gas velocity on thehydrogen production over Ni:Cu:Al fusion catalyst at the oper-ating temperature of 700 ◦C. As can be observed in Fig. 4, theeffect of gas velocity on the hydrogen production was signifi-cant in the FBR. As we have mentioned before, for the umf-pb,which corresponds to a space velocity of 4 l/gcat/h, the hydro-gen production was 60%, and not deactivation was observedduring the 16 h run. For the umf-eb and 2 × umf-eb, which cor-respond respectively to a space velocity of 6 and 12 l/gcat/h,a different pattern was observed: the hydrogen concentrationremains constant during 150 min, obtaining a hydrogen con-centration of 40% and 30%, respectively, and from this pointonwards a progressive deactivation begins. At the end of therun, a hydrogen production of 25% and 15%, respectively, wasobserved for the umf-eb and 2 × umf-eb experiments.

Unlike the fixed bed operation, in the fluidized bed opera-tion it is assumed that all the gas in excess of that requiredfor the minimum fluidization passed through the bed as bub-bles. Therefore, the higher gas velocity increased the numberof bubbles and the size of them and these bubbles may exit thereactor without effective contacting the catalyst. Thus, the in-creasing gas velocity reduced the residence time in the reactorand lowered the contacting efficiency between gas and metal

0 100 200 300 400 5000.0

0.5

1.0

1.5

2.0

2.5

3.0

reaction r

ate

(m

mol C

H4/m

in/g

cat)

Time (min)

umf-pb

umf-pb

2 x umf-pb

Fig. 5. Effect of gas velocity on the methane reaction rate over fusionNi.Cu:Al catalyst at the temperature of 700 ◦C.

catalyst due to bubble formation [15]. The above mentionedhydrodynamic characteristic of fluidized bed explains why thegas velocity effect on the methane conversion was significant.It is worth to mention that a higher methane conversion can beachieved by an increase in the bed height. Fluidization of thebed can be carried out by the hydrogen produced. The fact thatthe theoretical umf for hydrogen calculated by the method pre-viously mentioned is 40% higher than the umf for methane isbalanced by the increase of the volume during the methane de-composition (2 mol of H2 are produced for each mol of methaneconverted). Fig. 5 shows the methane reaction rate for the runsfeatured at the three fluidization velocities. The same patternas in Fig. 4 is observed. However, the three curves fall in anarrow range meaning that negligible mass resistance occurs,and the gasin excess does not have an effective contact with themetal catalyst. As we have mentioned before, this gas in ex-cess improves the mixing of the bed and prevents the cloggingproblems observed when low fluidization velocities are used.

Table 3 shows the carbon deposited after each run, the car-bon deposition per hour, the average methane conversion aswell as the average hydrogen production. As can be observedin Table 3, close values of carbon deposition ratio and hydro-gen production ratio are observed. It can be observed that the


0 50 100 150 200 250 300 350 400 450 5000

10

20

30

40

50

umf-pb

umf-pb

2 x umf-pb

ΔPcalc

Δ P (

mb

ar)

Time (min)

0 50 100 150 2000

1

2

3

4

5

6

7

8

9

10

umf-pb

umf-pb

2 x umf-pb

ΔPcalc

Δ P (

mb

ar)

Time (min)

Fig. 6. (a) Variation in the pressure drop of the particle bed for different gasvelocities and (b) magnification of (a).

fluidized bed installation has been operated to produce at thescale of 60 l/h of hydrogen and 15 g/h of carbon nanofibers.

3.3.2. Fluidization qualityScarce information about the hydrodynamic behavior of the

bed of particles during the TCD is found in literature. In orderto confirm the correct fluidization of the bed of particles, thepressure drop across the particles bed was continuously mon-itorized, �Pmeasured. As we have seen before, the evolution ofthe amount of carbon deposited Cdep-calc as a function of thereaction time can be calculated from Eq. (4). From pressuredrop definition, we can calculate the evolution with time of thepressure drop of the fluidized bed �Pcalc as the carbon is de-posited in the catalyst and thus, increasing the mass particle in-ventory. By comparing �Pmeasured and �Pcalc we can determi-nate the quality of the fluidization: a good correlation betweenboth curves means that the bed fluidizes correctly.

Fig. 6a plots the �Pmeasured across the bed for different flu-idization regimes: umf-pb, umf-eb, 2 × umf-eb. Fig. 6b shows a

magnification of the Fig. 6a. �Pcalc was also plotted for a bet-ter comparison of the fluidization quality (to see if the reactorclogging occurs). Due to the same ratio of carbon deposited forthe three conditions studied, the same curve has been plottedfor all of them.

Figs. 6a and b show that for the three fluidization velocitiestested, the �Pmeasured was in good agreement with the �Pcalcduring the first stage of the reaction (until 150 min). For theumf-pb and umf-eb an exponential increase in the �Pmeasured wasobserved. However, when 2 × umf-eb is used, the �Pmeasuredcorrelates fairly well with the �Pcalc. That clearly indicates thatit is necessary to operate with a vigorous fluidization regime,at least two times the minimum fluidization for the entire bed.Otherwise, defluidization occurs leading to clogging problemduring the runs.

It is known that for binary mixtures where the particlesdiffer in size only, larger particles segregate preferentially to-ward the bottom of the bed (near the distributor plate), whilethe smaller particles accumulate near the free surface. Forcontinuous particle size distribution, the mean particle diam-eter decreases from the bottom of the bed to the free sur-face [33]. The fluidization experiments were carried out witha carbon product with a wide particle size distribution, as wehave seen from hand sieved analysis. Thus, we propose as ex-planation that, as the particle diameter gets higher, the big-ger particles tend to occupy the low region of the bed, andthe gas velocity is not sufficient to maintain the fluidizationof the bigger particles. Then defluidization occurs, leading tothe agglomeration of the particles that finally provokes theshut down of the reactor. In addition, the fluidization regimeleads to the attrition of the particles, avoiding the formation ofrelative big sized carbon particles. The agglomeration of thesolid is a problem in catalytic fluidized bed reactors. In mostcases high relative velocities (u/umf) are employed to avoiddefluidization [34].

At an industrial scale, the carbon product can be withdrawfrom the bottom of the fluidized bed, grinded to a desired par-ticle size, and recirculated, assuring a good fluidization of theparticle bed.

Thus, we can conclude that the lowering in methane con-version is balanced with good fluidization behaviour. So theoptimum gas velocity should be chosen in terms of hydrogenproduction rates and fluidization quality.

3.4. Characterization of the carbon product

Fig. 7a shows the SEM images for the carbon deposited fromthe co-precipitation catalyst, Fig. 7b shows the SEM imagesfor the carbon deposited from the fusion catalyst. They showthat nanofibers are present, appearing as long filaments emerg-ing from Ni particles coexisting with uniform coatings on theNi particles. Although the relative concentration of these twocarbons forms cannot be achieved, a simple examination of therespective images evidences that long filaments are more abun-dant than uniform coatings as it had been observed in the workcarried out using a small fixed bed reactor [11–14].


Fig. 7. SEM micrograph of the Ni:Cu:Al-78:6:16 catalysts after an 16 h run. Reaction T : 700 ◦C; (a) co-precipitation and (b) fusion. left, 8000×, right, 20 200×.

4. Conclusions

Performance tests showed that the hydrogen production wasnot highly dependent on catalyst preparation method and sothat an easier and cheaper one as the fusion of salts can be usedat an industrial scale.

Fluidodynamic studies have shown that TCD can be carriedout in a fluidized bed reactor without reactor clogging providedthat a methane velocity of two times the minimum fluidizationvelocity is used.

A semi-continuous fluidized bed installation has been startedup and operated to produce at the scale of 60 l h−1 of hydrogenand 15 g h−1 of carbon nanofibers.

Acknowledgments

The authors would like to thank CDTI for financial supportthrough the contract with Gas Natural to carry out the SPHERAProject included in the CENIT Program 2006. Y.E. the SpanishMEC for the FPI Grant.

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