Direct route from ethanol to pure hydrogen through …Membrane reactor Hydrogen production...

Direct route from bio-ethanol to pure hydrogen throughautothermal reforming in a membrane reactor: experimentaldemonstration, reactor modelling and designCitation for published version (APA):Spallina, V., Matturro, G. E., Ruocco, C., Meloni, E., Palma, V., Fernandez, E., Melendez Rey, J., PachecoTanaka, D. A., Viviente, J. L., van Sint Annaland, M., & Gallucci, F. (2018). Direct route from bio-ethanol to purehydrogen through autothermal reforming in a membrane reactor: experimental demonstration, reactor modellingand design. Energy, 143, 666-681. https://doi.org/10.1016/j.energy.2017.11.031

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Energy 143 (2018) 666e681

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Energy

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Direct route from ethanol to pure hydrogen through autothermalreforming in a membrane reactor: Experimental demonstration,reactor modelling and design

V. Spallina a, c, *, G. Matturro a, C. Ruocco b, E. Meloni b, V. Palma b, E. Fernandez c,J. Melendez c, d, A.D. Pacheco Tanaka c, J.L. Viviente Sole c, M. van Sint Annaland a,F. Gallucci a, **

a Chemical Process Intensification, Department of Chemical Engineering and Chemistry, Eindhoven University of Technology, Eindhoven, The Netherlandsb Processes and Catalysis for Energy and Environment Depollution, Industrial Engineering Department, University of Salerno, Fisciano, Italyc Membrane Technology Group, Energy and Environment Division, Tecnalia, San Sebasti�an-Donostia, Spaind Chemical Engineering and Environmental Department, University of the Basque Country UPV/EHU, Bilbao, Spain

a r t i c l e i n f o

Article history:Received 11 April 2017Received in revised form19 September 2017Accepted 5 November 2017Available online 8 November 2017

Keywords:Ethanol reformingPalladium membranesMembrane reactorHydrogen productionExperimental demonstrationModelling

* Corresponding author. Chemical Process Intensifiical Engineering and Chemistry, Eindhoven UniversiThe Netherlands.** Corresponding author. Chemical Process Intensifiical Engineering and Chemistry, Eindhoven UniversiThe Netherlands.

E-mail addresses: [email protected] (V. Spallina), f.

https://doi.org/10.1016/j.energy.2017.11.0310360-5442/© 2017 The Authors. Published by Elsevie

a b s t r a c t

This work reports the integration of thin (~3e4 mm thick) Pd-based membranes for H2 separation in afluidized bed catalytic reactor for ethanol auto-thermal reforming. The performance of a fluidized bedmembrane reactor has been investigated from an experimental and numerical point of view. Thedemonstration of the technology has been carried out over 50 h under reactive conditions using 5 thinPd-based alumina-supported membranes and a 3 wt%Pt-10 wt%Ni catalyst deposited on a mixed CeO2/SiO2 support. The results have confirmed the feasibility of the concept, in particular the capacity to reacha hydrogen recovery factor up to 70%, while the operation at different fluidization regimes, oxygen-to-ethanol and steam-to-ethanol ratios, feed pressures and reactor temperatures have been studied. Themost critical part of the system is the sealing of the membranes, where most of the gas leakage wasdetected. A fluidized bed membrane reactor model for ethanol reforming has been developed andvalidated with the obtained experimental results. The model has been subsequently used to design asmall reactor unit for domestic use, showing that 0.45 m2 membrane area is needed to produce theamount of H2 required for a 5 kWe PEM fuel-cell based micro-CHP system.© 2017 The Authors. Published by Elsevier Ltd. This is an open access article under the CC BY license

(http://creativecommons.org/licenses/by/4.0/).

1. Introduction

H2 represents an important product for the chemical industry[1], and its demand is constantly increasing, also due to its potentialuse as automotive fuel and energy carrier [2]. Nowadays, fossil fuels(and mostly natural gas) are used for the production of more than95% of the global H2. Currently, H2 is almost entirely used asfeedstock within the refining and chemical industries to convertraw materials into higher value chemicals (e.g. NH3, CH3OH) or

cation, Department of Chem-ty of Technology, Eindhoven,

cation, Department of Chem-ty of Technology, Eindhoven,

[email protected] (F. Gallucci).

r Ltd. This is an open access article

refinery hydro-treating processes [3]. In addition, H2 represents themost important energy carrier for the future and can play animportant role in reducing the anthropogenic greenhouse gasemissions. Recently, fuel cells have been developed and launched tothe market and a lot of research is ongoing to investigate anddevelop the technologies for the transition towards H2 [4]. Amongthe various renewable feedstock alternatives, ethanol (EtOH) isconsidered as an attractive feedstock due to its relatively highhydrogen content, abundant availability, non-toxicity, storage orhandling ease and safety [5e7]. Moreover, the ethanol can beproduced renewably by biomass sources (agricultural wastes,forestry residuals or organic municipal wastes, etc.), also called bio-ethanol, and is mixed with water (about 15% on weight basis ofEtOH [8]). The ethanol-to-hydrogen conversion is normallyattained by ethanol reforming reactions that can be classified in: i)steam reforming (ESR) and ii) auto-thermal reforming (EAR) in the

under the CC BY license (http://creativecommons.org/licenses/by/4.0/).

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V. Spallina et al. / Energy 143 (2018) 666e681 667

presence of an oxidant (air or oxygen). Sun et al. [9] have conducteda thermodynamic study on ethanol reforming comparing steamreforming, auto-thermal reforming and partial oxidation and theyhave found the optimal conditions at 900 K, with steam-to-ethanolequal to 6 and oxygen-to-ethanol equal to 0.25. Similarly, de Avilaet al. [10] presented their study including also the dry reformingand they found a mismatch in the maximum H2 mol fraction(around 773e873 K) and H2 yield (when the temperature is higherthan 1123 K).

The reaction pathways for ethanol reforming include severalpossible reactions which are influenced by the catalyst and theoperating conditions used in system [5,11,12]. Different supported/unsupported catalysts have been proposed in the literature for theESR reaction which have been summarized in a recent review fromHou et al. [11], where both noble metal catalysts (Rh, Ru, Pt, Pd)operated in the range of 650e750 �C with a steam-to-ethanol ratio(H2O/EtOH) of 1e3 (and also O2 co-feeding) and transition metals(especially Ni and Co) operated at a lower temperature(350e650 �C) with a higher H2O/EtOH ratio (up to 10 on a molarbasis) have been considered. However, the addition of smallquantities of noble metals to transition metals-based catalysts wasshown to improve the ethanol conversion and H2 yield even at lowtemperatures and/or steam-to-carbon ratios [13]. As support ma-terial, alumina (Al2O3) has beenwidely studied because of its strongthermal and mechanical stability, despite the fact that it may allowthe formation of C2H4 which is a precursor for coke formation. Ingeneral, supports such as CeO2 and mixed oxides systems haveshown to be highly favourable in terms of EtOH conversion and H2selectivity reducing the sintering effect because of metal dispersionthrough the support [13,14]. As a result, the application of the highsurface area of a CeO2/SiO2 mixed system has shown improvedcatalytic stability with respect to SiO2 free samples [15].

H2 production from ethanol is normally carried out via multiplesteps of conversion and separation processes [12,16]. The reformingis carried out at 750e800 �C, where the ethanol is decomposed inother gaseous species (CH4, CO, H2) as indicated in Equation (1),making the conversion of ethanol into H2 thermodynamicallylimited. In the ethanol auto-thermal reforming (EAR), the presenceof oxygen reduces the H2 selectivity, however, no external furnaceis required to supply the required heat of reaction for the highlyendothermic ESR. Due to the thermodynamic equilibria, largequantities of CO are produced. The main reactions occurring in theEAR reactor are:

C2H5OH / H2 þ COþ CH4 DH298K0 ¼ 51:3kJ=mol (1)

CH4 þ H2O4 3H2 þ CO DH298K0 ¼ 206:2kJ=mol (2)

CH4 þ 0:5O24 2H2 þ CO DH298K0 ¼ �38kJ=mol (3)

COþ H2O4 H2 þ CO2 DH298K0 ¼ �41:2kJ=mol (4)In the complete process, the H2 yield can be enhanced by

employing additional water-gas-shift (WGS) reactors, producing H2and CO2 while reducing the CO content below 1e0.5 vol% In someprocesses, the presence of CO in the H2-rich stream acts as a poisonfor the materials of the downstream units (e.g. fuel cells, NH3synthesis, etc.). In these cases, a further purification step is requiredto obtain high purity H2, and usually a methanation reactor isconsidered to further decrease the CO content to below 10 ppmconverting CO into CH4 [17]. Alternatively, pressure swing absorp-tion (PSA), cryogenic distillation, or membrane technology can beused to produce high purity H2 (>99.99%).

The overall H2 yield of ethanol reforming is usually about 4.3

molH2/molEtOH [16]. However, in case pure H2 is required, theapplication of PSA (with an H2 separation factor of about 80%) leadsto an overall efficiency (based on the LHV) of 65% (

V. Spallina et al. / Energy 143 (2018) 666e681668

and the composition (H2O/EtOH from 3 to 4.2) resulting in anoptimal system performance (higher than 40% of electrical effi-ciency compared to 33.1% of the conventional case) at 12 bar and aH2O/EtOH ratio equal to 3.6 and 0.4m2 of membrane area. In case ofH2 separation without sweep gas, the optimal performance thatcould be achieved at 6 bar and 0.3 bar at the permeate side required0.3 m2 membrane area (40.3% based on Ethanol lower heatingvalue). Hedayati et al. [21] have compared the experiments andmodelling of a dynamic unit to produce H2 for fuel cell and theyfound that the changes in the fuel flow rate responds much fasterthan the electric part. Recently, Arratibel et al. [34] have summa-rized the operating conditions and performance of differentmethane and ethanol reforming membrane reactor for pure H2production that have been tested and presented.

In the present paper, a fluidized bed membrane reactor (FBMR)has been tested for ESR and EAR and evaluated from an experi-mental and numerical point of view. The catalyst material used is3 wt%Pt-10 wt%Ni deposited onto a mixed CeO2/SiO2 support toincrease the stability of the catalyst as presented in a previous work[14]. Five ultra-thin ceramic-supported Pd-Ag membranes (4 mmthickness) of 120e160 mm length and 10 mm of external diameterhave been sealed and integrated in a lab-scale fluidized bed (pre-viously used to demonstrate a WGS membrane reactor [30]) andtested under reactive conditions. The proof-of-concept has beenachieved for different operating conditions. A two-phasephenomenological model for the fluidized bed reactor has beenvalidated with the experimental results and subsequently used forthe design of a commercial scale unit.

2. Description of the technology

The schematic representation of a fluidized bed membranereactor is shown in Fig. 1. The inlet gases, EtOH/H2O/air, are fed tothe bottom part of the reactor through a 40 micron porous plate

feed (i.e. C2H5OH + H2O + Air)

Catalyst

Pd/Ag membrane

Support

retentateCO2+H2O (+H2,CO, CH4)

H2

Fig. 1. Schematic layout of a fluidized bed membrane reactor.

distributor. All the EtOH is partly oxidized with O2 and decomposedinto other gases (reaction (1)), while the produced CH4 and CO startto react with H2O (reactions (2e4)). The supported Pd-Ag mem-brane tubes are immersed inside the catalyst bed to extract thehydrogen from the reaction zone and enhance the methanereforming and water gas shift reactions. The pressure drop over thebed is negligible, while mass and heat transfer coefficients arehigher due to the vigorous solids mixing resulting in a virtuallyuniform temperature throughout the reactor.

The presence of the membranes can improve the bubble-to-emulsion mass transfer rate, since the membranes act as internalsin the granular suspension promoting bubble breakage [35]. Due tothe gas extraction through the membranes, the solids circulationpattern may be modified compared to the conventional fluidizedbed reactor, since densified zones may be formed close to themembranes [36,37].

The integration of catalyst and membranes has been carried outat two different scales: in the first part, a single tube membranereactor has been used to carry out permeation tests using differentreactor layouts (empty, fluidized and packed bed configurations) toverify the validity of the kinetic model proposed in Ruocco et al.[14]. Subsequently, 5 membranes have been sealed and used in alarger prototype, which has been operated close to industrialoperating conditions.

2.1. Catalyst preparation

The CeO2/SiO2 support material was prepared by addingcalcined (at 600 �C for 3 h, heating rate of 10 �C min�1) silica gel(particle size distribution 150e250 mm, supplied by Sigma-Aldrich)to an aqueous solution of Ce(NO3)3$6H2O (Strem Chemicals). TheCeO2/SiO2 weight ratio, previously optimized [15], was fixed at 30%.Impregnation was carried out at 80 �C for 2 h on a heating andstirring plate. Then, the support was filtered by means of a Buchnerfunnel, dried overnight at 120 �C and calcined as reported above forthe bare silica. Ni (10 wt%) and Pt (3 wt%)metal loadings refer to thetotal ceria mass were sequentially added to the support, startingfrom Ni(NO3)3$6H2O and PtCl4 (Strem Chemicals) as salt pre-cursors, respectively.

BET specific surface areas and porosity of the support as well asthe catalyst were determined from the absorption isotherms at77 K, acquired by means of a ThermoScientific Surfer. Structuralproperties were characterized by powder X-ray diffraction (XRD)using a D-8 Advance Bruker WAXRD diffractometer under a CuKaradiation of 1.5406 Å. XRD spectra allowed the calculation ofaverage crystallite size by means of Scherrer equation. H2 Tem-perature Programmed Reduction (TPR)measurements were carriedout in situ in the laboratory apparatus previously described [15].The temperature was increased to 600 �C (heating rate of 10 �C/min) and held for 1 h under 5%H2 in N2 (500 Ncm3∙min�1 of totalflow rate). The support and catalysts were prepared and charac-terized (XRD and TPR analysis) in the Proceed laboratories of theUniversity of Salerno.

2.2. Membrane preparation

Alumina asymmetric tubes with a 10mmoutside diameter wereused as membrane supports. Pd-Ag thin layers were deposited ontothe alumina tubes using the simultaneous (Pd and Ag) electrolessplating technique reported in Pacheco Tanaka et al. [38]. After theplating step, the membrane layers were annealed at 550 �C with aH2-N2 gas mixture. The final membranes are approximately140e230 mm in length and with a Pd-Ag thickness of around3e4 mm (Fig. 2).

Al2O3 100 nm

Pd-Ag

Fig. 2. Cross-section SEM image of a Pd-Ag membrane layer deposited onto an Al2O3asymmetric porous tube (100 nm pore size, 10/4 mm OD/ID).


2.3. Reactor design and test rig

Two different reactors have been used in this work. The firstreactor is a single tube membrane reactor (Fig. 3a) which has beenused for the permeation tests and the validation of the kineticmodelling by feeding a mixture of CH4, CO and H2 and H2O underfluidization conditions. The membrane (named E275) is 5.5 cm inlength and has an external diameter of 10 mm (Fig. 3b). The sealingmethod has been proposed by Chen et al. [39] in which graphiteferrules, instead of metal ferrules are used. The same method has

Single tube reactorLR = 29 cmDR = 3.2 cm

Fig. 3. a) Schematic layout of the single tube membrane reactor; b) Pd-Ag membraneused for the experiments.

been successfully used for a long-term test under reactive condi-tions [24,30,40]. The reactor is placed in an electrically heated ovenin order to have a constant temperature along the entire reactor. Asecond membrane (E273) of 10.4 mm (ID) and 109.8 cm length wasalso used for the permeation test using respectively the empty bed,packed bed and fluidized bed configurations.

The second reactor has an inner diameter of 10 cm and a lengthof 1 m is located inside an electric heated jacket and details areshown in Fig. 4 and Fig. 5. This experimental set-up allows per-forming reactive permeation experiments with a maximum of fivemembranes. There are three thermocouples, and two pressuresensors distributed over the reactor height. A vacuum pump islocated at the permeate side, in order to separate H2 at 10e30mbar.The system further includes feed connections to pure gas compo-nents as O2, H2, N2, CO, CO2, CH4 and pressurized air, as well asliquid tanks containing H2O and EtOH. Both EtOH and water areconnected to a CEM (Controlled Evaporator Mixer) where bothcomponents can be evaporated before feeding to the reactor. All thetubes connected to the reactor inlet are traced, as well as the tubesused for the retentate, to maintain the temperature well above theboiling point of water and ethanol. In order to measure the EtOHcontent in the retentate stream, an Agilent Technologies Cary 630FTIR with CaF2 windows was used together with a RED-SHIFT gassampling system. The FTIR was calibrated prior to the experimentsusing the classical Lambert-Beer law in typical adsorption spectrafor the gases CO2, CH4 and EtOH. After that, the retentate passesthrough a cooler where the water and ethanol (if present) arecondensed and separated in a flash column. The retentate is con-nected to a Sick® analyzer to measure the retentate composition(dry basis). A similar analyzer is also connected to the permeateside, which is able to measure hydrogen concentrations in therange of 0e100% and COX concentrations on a ppm level. The flowrate at the permeate side is measured with a Horiba film flowmeter. The membranes can also be assessed individually, by closingvalves manually which are placed on top of the reactor.

3. Reactor model

The phenomenological two-phase model for the membrane-assisted fluidized bed reactor is depicted in Fig. 6. This model hasbeen developed in the past years and used for different processesinvolving H2/O2 membrane reactors [18,41]. Both emulsion andbubble phases are considered as a cascade of a number of contin-uous stirred tank reactors (CSTRs) and the size of the CSTRs isdirectly related to the extent of gas back-mixing in each phase. Themodel assumptions are as follows: i) the reactor consists of twophases, viz. bubble and emulsion phases; ii) the gas through theemulsion phase is assumed completely mixed in each CSTR and atincipient fluidization conditions; iii) the bubble phase is consideredto be in plug flow (and therefore several CSTR are used to properlyaccount for it); iv) the reactions involved in the system occur only inthe emulsion phase (i.e. the bubbles are assumed devoid of parti-cles); v) the gas that permeates through the membrane is takenfrom both bubble and emulsion phases in a ratio according to theirlocal phase fractions; note that the gas extracted from the emulsionphase is replaced by part of the gas transferred from the bubblephase in order to respect the minimum fluidization velocity; vi) thetemperature is assumed constant along the entire bed.

The overall component mass balance equations have beenformulated accounting for the chemical reactions and the gaspermeation in the source terms. More details on the material andenergy balances and additional parameters for the membrane-assisted fluidized bed reactor model can be found in Gallucciet al. [32]. (see Table 1)

The kinetic model is based on the results presented by Ruocco


et al. [14], which have been validated against experimental resultsand compared with other models available in the literature:

⁃ C2H5OH þ 3.5a O2 / (1 e a)C2H5OH þ 2aCO2 þ 3aH2O: thereaction is completely shifted toward the products until theoxygen is completely consumed;

⁃ (1 e a)C2H5OH / (1 e a)CH4 þ (1 e a)H2 þ (1 e a)CO: fullconversion is assumed with an infinite reaction rate;

⁃ CH4 þ H2O4 COþ 3H2: the reaction rate of the steammethanereforming (SMR) reaction has been determined experimentally;

⁃ CO þ H2O4 CO2 þ H2: water gas shift reaction (WGS) has beenassumed at chemical equilibrium.

An empirical power-law expression is used to describe the ki-netics of SMR. The power-law expression generally used for theSMR is:

RSMR ¼ kSMR�paCH4p

bH2Op

cCOp

dH2

�ð1� bÞ

where b ¼

1

KSMReq

p3H2pCOpCH4pH2O

!

And KSMReq ¼ exp�� DG

SMR

RT

�

a ¼ 0:96b ¼ 0:28c ¼ d ¼ 0

(5)

kSMR ¼ k0; SMR exp�� Eact; SMR

RT

� Eact;SMR ¼ 72:8 kJ=molk0;SMR ¼ 8:02� 10�2mol=ðkgcats PaaþbþcþdÞ

(6)

The H2 permeation rate has been described (molH2/m2) withSieverts' law:

JH2 ¼ PH2;0 exp�� EH2 ;p

RT

��p0:5H2;r � p0:5H2;p

� EH2;p ¼ 4:57 kJ=molPH2;p ¼ 1:14� 10�4mol=ðs m2 Pa0:5Þ

(7)

The membrane permeation rate is discussed in more detail thenext section.

4. Results

4.1. Catalyst characterization

The specific surface area, average pore diameter (Dp) and thepore volume (Vp) of the support and the final catalyst are shown inTable 2. The N2 adsorption-desorption isotherm curves at 77 K(Fig. 7 (a) and (b)), according to the IUPAC classification, belong totype IV, typical of mesoporous structures [42]. The mesoporousstructure of the CeO2-SiO2 support as well as its specific area wasunaffected by the deposition of the active species. In both cases abroad hysteresis loop was observed, which is typical of mesoporoussolids; however, the presence of micropores (pores with sizessmaller than 2 nm) cannot be excluded [43]. The XRD patterns ofthe support and the catalysts are shown in Fig. 7 (c). The broad peak

around 23� can be attributed to amorphous SiO2 while thediffraction peak at 43.2�, observed in the spectrum of the Pt-Ni/CeO2-SiO2 sample, is related to the NiO crystalline phase [44].Diffraction lines of the CeO2 fluorite-type structure are also visiblein both patterns. As a result of the high surface area of the silicasupport (400 m2 g�1), very low dimensions for ceria and nickeloxides crystallites were recorded (equal to 82 and 85 Å respec-tively). Similarly to the results shown in Table 2, the deposition ofthe active species had no effect on the ceria crystallite sizes (78 vs82 Å).

The catalyst was also characterized by H2-TPR (Fig. 7 (d)) inorder to investigate the Ni as well as Pt interactions with the CeO2/SiO2 support. Two zones were observed, at 70e200 �C and200e600 �C, related to the noble and the non-noble metal reduc-tion respectively [45]. Both low and high temperature reductionzones can be deconvoluted into two peaks, which proves thepresence of oxides particles with different interactions with thesupport. In the case of NiO, for example, the peak observed at294 �C can be ascribed to the reduction of bulk particles, havingweak interaction with the support, whereas the peak at 384 �C isattributed to nickel oxide particles in intimate contact with CeO2/SiO2. The total hydrogen consumption (3802 mmolH2∙gcat�1) is almosttwice the expected one, which suggests that active species depo-sition on the catalyst may improve the bulk-phase oxygen reduc-tion of CeO2, commonly negligible below 600 �C [46]. This effect isenhanced by the better Ni reducibility observed after Pt addition,

related to spillover phenomena and commonly reported in theliterature, which also resulted in lower reduction temperatures fornickel oxide phase [47]. On the other hand, the low dimension of

ceria crystallites as well as the CeO2-SiO2 interactions may furtherimprove ceria surface oxygen reducibility [48].

4.2. Membrane permeation in a single tube reactor

The permeation tests were carried out with the samemembraneat the same operating conditions while the reactor configurationswere modified. After sealing the membrane (using a torquewrench with a momentum >10 Nm), the membrane leakages aredetected by immersing the membrane inside ethanol and feedingHe internally to the membrane (with a DP of 1 bar). The He leakageis totally concentrated in the sealing part and corresponds to0.036 ml min�1. After that, the membrane was located inside thereactor and heated up to 400 �C in a N2 environment. During theheating, air was shortly flowing inside the reactor in order to cleanthe membrane surface from any unwanted species. When thesystem is at 400 �C, a mixture of H2/N2 is used to activate themembrane. Membrane activation is reached when the H2

Multi-tubes reactorLR = 100 cmDR = 10.2 cm

a)

b)

c)

d)

Fig. 4. a) and b) Schematic layout of the multi-tubular membrane reactor; c) membrane tube bank arrangement prior to placement inside the reactor; d) zoomed in picture of themembrane arrangement.

FT-I

R

a)

b) c) d)

Fig. 5. a) P&ID of the setup used for the multi-tubular experiments; b) Vessels of EtOH and H2O used for the experiments, c) Electric oven and reactor; d) Gas analysers for thepermeate (left) and retentate (right) streams.

EE

B B B B

Permeated H2,e

Perm. H2,b

Bubble-to-emulsionmass and energy

exchange

Bubble-to-emulsionmass and energy

exchange

Feed gas

Perm. H2,b

Permeated H2,e

Perm. H2,b Perm. H2,b

retentate

Figure 6. Schematic of the two-phase phenomenological model [32].

1 The H2/N2 ideal perm-selectivity is the average ratio of the H2 and N2 per-meance values at 450 �C in the range of 2.5e4.5 bar.


permeation is no longer increasing after two consecutive fluxmeasurements.

The first tests have been carried out using the E273 membrane.The membrane permeation rate was measured at 400 �C withthree different configurations and the comparison is shown inFig. 8. In the empty configuration (represented by a circle in Fig. 8),no solids are placed inside the reactor and the permeation testsare carried out using different H2/N2 compositions (100/0, 80/20,60/40, 50/50), where 4 Nl∙min�1of H2 is fed to the system whilethe amount of N2 is properly adjusted. The total pressure at thefeed/retentate side is varied from 1.5 to 4 bar, while the permeateside pressure is kept at 1 bar. Due to limitations in the totalpressure, when decreasing the H2 content fewer measurementswere possible in order to guarantee a positive H2 permeationdriving force from the retentate to the permeate side. After thetest without solids inside the reactor, two more tests have beencarried out. Firstly the reactor has been filled with 95 g of CeO2/SiO2 with particle size in the range of 150e250 mm, so thatpermeation tests could be carried out under fluidization condi-tions (these results are marked with triangles in Fig. 8). Finally, thereactor has been completely filled with inert material with aparticle size of 1e1.5 mm to have the membrane operated in apacked bed configuration (square markers in Fig. 8).

Two main conclusions can be derived from the analysis of theseresults. The first conclusion is a confirmation that the extent ofconcentration polarization is considerable. The slope of the curve -which corresponds to the membrane permeance according toequation (7) - decreases by a 53% (respectively 2.18$10�3 mol m�2

s�1Pa�0.5 and 1.03$10�3 mol m�2 s�1Pa�0.5) when comparing thegas feed compositions of 100% H2 and 50/50H2/N2. These results arealso confirmed by several other works available in the literature[49e53], specifically in the presence of high flux membrane as isthe case of ultra-thin (~1 mm thickness) Pd-based membranes. Thesame trend is also shown in case of fluidized and packed bedconfigurations. Remarkably, only a negligible improvement(þ6e8%) is observed when using the fluidized bed configurationwith respect to the packed bed configuration. Helmi et al. [30] havedemonstrated that the H2 permeation in a membrane reactor isenhanced when employing a fluidized bed configuration due to theincreased radial dispersion; they have carried out an experimentalcampaign operating the system at 3 < u/umf

Table 1Hydrodynamic parameters and mass transfer coefficients used in the model.

Parameters Equation Ref.

Archimedes number

Ar ¼ dprg�rp � rg

�g.mg (8)

[36]

Minimum fluidization velocity

umf ¼�mg

.dprg

� � ffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffið27:2Þ2 þ 0:0408Ar

q� 27:2

�(9)

[37]

Bed voidage at minimum fluidization velocity

εmf ¼ 0:586Ar�0:029 rgrp

!0:021(10)

[37]

Velocity of rise of swarm of bubblesub ¼ uo � umf þ ubr (11)

[36]

Rising velocity of single bubble

ubr ¼ 0:711�g,db;avg

�1=2(12)

[36]

Emulsion velocity

ue ¼ uo � d$ub1� d (13)[36]

Average bubble diameter

db;avg ¼ db;max ��db;max � dbo

�exp

��0:3H

DT

�(14)

[38]

Initial bubble diameter

dbo ¼ 0:376�uo � umf

�2(15)

[35]

Bubble phase fraction

dbn ¼ubub

(16)

[35]

Emulsion phase fractionden ¼ 1� dbn (17)

[35]

Maximum superficial bubble gas velocityusb;max ¼ uo � umf (18)

[35]

Initial superficial bubble gas velocity

usb;o ¼ ubr;odbo where dbo ¼�1� Hmf

.Hf�

(19)

[35]

Height of bed at minimum fluidization velocity

Hmf ¼ Hs1� εs1� εmf

(20)

[38]

Height of bed expansion

Hf ¼ HmfC1

C1 � C2(21)

[35]

where,

C1 ¼ 1�ub;oub;avg

exp��0:275

DT

�(22)

C2 ¼ub

ub;avg

1� exp

��0:275

DT

�

(23)

Average bubble rise velocityub;avg ¼ uo � umf þ ubr (24)

[35]

Gas exchange coefficient

Kbc ¼ 4:5�umfdp

�þ 5:85

Dgg1=4

db

!(25)

[35]

Kce ¼ 6:77�Dgεmf ub

db

�1=2(26)

1Kbe

¼ 1Kbc

þ 1Kce

(27)


Table 2Physiochemical properties of the fresh CeO2-SiO2 and Pt-Ni/CeO2/SiO2.

Sample BET (m2∙g�1) Dp (nm) Vp (cm3∙g�1)

CeO2/SiO2 254 2.82 0.802Pt-Ni/CeO2/SiO2 255 2.86 0.807

Figure 7. N2 adsorption (square)-desorption (triangle) isotherms at 77 K for CeO2/SiO2 (a) aprofile after deconvolution for the Pt-Ni/CeO2/SiO2 sample (d).

0

0.15

0.3

0.45

0.6

0.75

0.0 40.0 80.0 120.0 1

H2

flux,

mol

m-2

s-1 ]

PH2,f0.5 - P

packed

empty

fluidized

50%H2-5

Fig. 8. H2 flux as a function of the permeation driving force for different H2-


range of 450e550 �C; the pressure at the feed/retentate side hasbeen varied from 2 to 4 bar; the gas flow rate has been adjusted inorder to always keep a u/umf equal to 3. As reference case, the O2/EtOH ratio has been set at 0.4 (on a molar basis) and the H2O/EtOHat 6. The amount of EtOH fed to the system has been varied be-tween 100 and 250 Nml∙min�1 to ensure a constant gas compo-sition and fluidization regime.

nd Pt-Ni/CeO2/SiO2 (b), XRD spectra of the support (1) and the final catalyst (2) (c), TPR

60.0 200.0 240.0 280.0 320.0

H2,p0.5 , Pa0.5

100%H2

80%H2-20%N2

60%H2-40%N2

0%N2

N2 compositions and configurations at 400 �C with the E273 membrane.

Table 3Membrane performance used for the prototype before reactive tests.

Membranes Name Length Area H2 Permeance (pure gas) H2/N2 SelectivityError! Bookmark not defined.

e e mm cm2 mol s�1 m�2 Pa�0.5 PH2/PN2

Membr_1 FL-4 146.6 47.85 1.46E-03 893Membr_2 FL-5 148.4 48.44 8.58E-04 1319Membr_3 FL-6 175 57.12 8.08E-04 590Membr_4 FL-7 187 61.04 7.36E-04 2836Membr_5 FL-8 153.5 50.10 8.96E-04 803

total area 264.56


For all the tests, the EtOH conversion at the reactor outlet wasabove 99.9%, demonstrating that the catalyst was highly activewithrespect to ESR and EAR reactions and the system is at its thermo-dynamic equilibrium, when accounting for the hydrogen that hasbeen extracted from the system.

From the analysis of the composition at the retentate side(shown in Fig. 9), the following conclusions can be derived:

0.0%

1.0%

2.0%

3.0%

4.0%

5.0%

6.0%

2 3 4

gas

com

posi

tion,

%vo

ldry

feed pressure, barH2 CO2 CH4 CO

0.0%

1.0%

2.0%

3.0%

4.0%

5.0%

6.0%

2 3 4

gas

com

posi

tion,

%m

ol,d

ry

feed pressure, bar

0.0%

1.0%

2.0%

3.0%

4.0%

5.0%

6.0%

2 3 4

gas

com

posi

tion,

%m

ol,d

ry

feed pressure, bar

500 °

550

450

Fig. 9. Results of permeation test at different temperature (450e550 �C) and feed pressurefeed flow rate has been varied according to the u/umf.

⁃ The CH4 conversion increases when increasing both the tem-perature or pressure, while the CO content increases at highertemperatures, which confirmes that i) the amount of catalystand the residence time in the reactor are sufficiently high toavoid any kinetic limitation in the conversion of the reactantsand ii) the higher the amount of permeated H2, the more the

0

100

200

300

400

500

2 3 4

P erm

e at e

dH

2,m

l/min

feed pressure, bar

0

100

200

300

400

500

2 3 4

Per m

eat e

dH

2,m

l/min

fee pressure, bar

0

100

200

300

400

500

2 3 4

Per m

eat e

dH

2,m

l/min

feed pressure, bar

C

°C

°C

(2e4 bar). The feed composition is EtOH 1.7%, H2O 10.3%, Air 3.25%, N2 to balance. The

0%

20%

40%

60%

80%

100%

470

480

490

500

510

520

530

540

550

0 0.1 0.2 0.3 0.4

HR

F(%

)

Hy d

roge

npe

rmea

tion

(Nm

l/min

)

O2/EtOH mol ratio

Hydrogen permeation HRF

Fig. 10. Effect of the O2/EtOH ratio on the amount of hydrogen permeated and the hydrogen recovery factor in the EAR membrane reactor.


equilibrium is shifted towards the products. Specifically, at550 �C the CH4 fraction is only 0.1e0.3 vol%.

⁃ At increased pressure the amount of H2 that permeates throughthe membrane increases due to the higher driving force be-tween the retentate and the permeate side (where 0.03 bar(a) isconsidered when using the vacuum pump).

⁃ The CO, CO2, and H2 concentrations in the permeate side werefrequently measured during experiments. The permeated COand CO2 concentrations were in the range of 100e300 ppm,whereas the hydrogen concentration was measured constantlyat values above 98%, where the remaining 2% is expected to beN2 for all the experiments.

⁃ The error in the carbon balance is always

0%

10%

20%

30%

40%

50%

60%

70%

2 3 4

gas

com

posi

tion

(%m

ol,d

ry,N

2-fre

eba

sis)

feed pressure, bar

0%

10%

20%

30%

40%

50%

60%

70%

2 3 4

gas

com

posi

t ion

(%m

old r

yan

dN

2- fr e

eba

sis)

pressure, barH2 - Experimental H2 - ModelCO2 - Experimental CO2 - ModelCH4 - Experimental CH4 - ModelCO - Experimental CO - Model

0

100

200

300

400

500

2 3 4

Perm

eate

dH

2,m

l /mi n

feed pressure, bar

H2 permeation - Experimental H2 permeation - Model

0

100

200

300

400

500

2 3 4Pe

rmea

ted

H2,

ml/m

infeed pressure, bar

0

100

200

300

400

500

2 3 4

Perm

eate

dH

2,m

l/min

feed pressure, bar

0%

10%

20%

30%

40%

50%

60%

70%

2 3 4

gas

com

posi

tion

(%m

oldr

yan

dN

2-fre

eba

sis)

pressure, bar 500 °C

550 °C

450 °C

Fig. 11. Comparison of experimental results and model predictions for three different temperatures. Left: retentate composition; right: permeated H2. The retentate composition isconsidered after normalization at dry conditions and after removing the amount of N2.

0%

1%

2%

3%

4%

5%

6%

7%

8%

0% 1% 2% 3% 4% 5% 6% 7% 8%

Pred

icte

d va

lue,

mol

frac

tion

%dr

y

Experimentally measured value, mol fraction %dry

H2 CO2 CH4 CO

Fig. 12. Parity plot of the results (modelling vs experiments).

0.0%0.5%1.0%1.5%2.0%2.5%3.0%3.5%4.0%4.5%5.0%

0 2 4 6 8 10 12 14 16 18

gas

frac

tion,

%m

ol

reactor axial length, cm

H2

COCH4

emulsion bubble

Fig. 13. Axial gas composition profiles (at T ¼ 500 �C, pfeed ¼ 4 bar) for the emulsion and bubble phases.

Fig. 14. Leakages from the sealing of the membrane after reactive test during a He leaktest.


higher than in the emulsion phase. This can be explainedby the higher methane conversion rate by the SMR in theemulsion phase than the CH4 mass transfer rate from the emul-sion to the bubble phase (viz. 0.11$10�5 vs. 0.78$10�6 mol s�1), inparticular at the end of the reactor because of the increasedbubble size.

After the experimental campaign (approximately 50 h withEtOH, three weeks under H2/N2 environment), the membraneswere taken out of the reactor and placed in an ethanol solution toperform helium leak test experiments (Fig. 14). One membranewas broken during the shutdown, two of the membranes wereleaking mostly from the bottom sealing, while the remaining two

JH2 ¼ PH2;0 exp�� EH2 ;p

RT

��p0:5H2;r � p0:5H2;p

� EH2;p ¼ 4:57 kJ=molPH2;0 ¼ 1:03g10�3mol=ðs m2 Pa0:5Þ

(28)

membranes were damaged at the upper sealing. This confirmsthat the sealing of the ceramicesupported membranes is one ofthe major concerns in this technology that needs to be further

developed. Specifically the following recommendations can bemade: 1) use of a Swagelok connection with a larger internaldiameter (e.g. 12 mm instead of 10 mm); 2) use of predefinedgraphite ferrules that could avoid the need of any thermal ormechanical pre-treatment, so that no parts of the membrane arein direct contact with the Swagelok where usually the membranebreaks. This also requires an even better accuracy control of theshape of the ceramic supports.

The catalyst particle size distribution was also measured afterthe experiments and the results indicate that the averaged particlediameter has not changed (dp,av z190 mm). However, smaller par-ticles (dp < 50 mm) present in the fresh catalyst were absent. Theseparticles have been blown out of the reactor during the experi-ments, since the filter to remove particles from the retentate sidehad to be unblocked a few times.

4.5. Design of FBMR for mCHP unit

The FluidCELL project will demonstrate the feasibility of aFBMR using EtOH as feedstock integrated with a PEM fuel cellup to 5 kWel. This unit is expected to be used for domestic ap-plications to supply electricity and heat. The FBMR is assumed tobe operated at 550 �C, 12 bar, 1 bar at the permeate side usingH2O as sweep gas (H2Osw/H2,p ¼ 3.7 kgH2O/kgH2), with atthe feed/retentate side a H2O/EtOH ratio of 3.8 and an (vol.), O2/EtOH ratio of 0.4 (%vol); The resulting feeding gas mass flow rateis 6.87 kg/h and the composition as calculated from previouswork [16] is EtOH 14.9%, H2O 56.7%, O2 5.6%, N2 22.5% (molarbasis).

As discussed before, the H2 content affects the H2 permeationrate because of the prevailing concentration polarization. Based onthe permeation tests a different pre-exponential factor (PH2,0) hasbeen adopted in this analysis (Eq. (28)).

Since both the kinetics andmass transfer rates do not give rise toany limitation, the membrane area required to separate 3.2 Nm3/hof H2 is estimated at 0.44 m2. Based on this reactor design, the

0

0.1

0.2

0.3

0.4

0.5

0.6

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4

gas

com

posi

tion

,vol

.

reactor axial length, m

N2

H2O

H2CH4

CO2

CO

Fig. 15. Axial gas composition profiles at the feed/retentate side.


minimum amount of catalyst needed to keep the retentate at itsthermodynamic equilibrium has been calculated. The modellingshows that for a total amount of 7 kg of solidmaterial required to fillthe reactor, about 250 g of catalyst is sufficient to achieve therequired conversion. In terms of fluidization variables, the mini-mum u/umf is ranging from 6 to 9: at the beginning the volumetricflow rate increases because H2 is produced while consuming CH4;in the second part of the reactor, the u/umf decreases because of thelarger H2 permeation rate in comparison with the CH4 conversion.The computed gas profiles at the feed/retentate are shown in Fig. 15and the amount of H2 permeating through the membranes is re-ported in Fig. 16.

5. Conclusions

The current work has provided an experimental demonstrationof a membrane reactor for ethanol auto-thermal reformingtogether with the validation of a phenomenological model to pre-dict the reactor performance. The integration of EtOH conversionand H2 separation via thin film Pd-based ceramic supportedmembranes has been studied and presented. The results of thepermeation tests have shown a considerable influence of

Fig. 16. Volumetric flow rate of H2 permeated along the reactor length.

concentration polarization for the empty, packed and fluidized bedconfigurations. The three different configurations did not showsignificant differences in the permeation rate, because the radialdiffusion/dispersion rates were comparable for the three consid-ered cases. The permeation tests under reactive conditions havebeen carried out for more than 50 h (at reactive conditions) and intotal 3 weeks under high temperature and fluidization conditionswith H2/N2, reaching a HRF ranging from 30 to 70% (close to in-dustrial application). The Pt-Ni based catalyst supported on amixedCeO2/SiO2 has demonstrated a high activity and stability for theEAR confirming previous results. The two-phase phenomenologicalmodel of a fluidized bed membrane reactor has been described andall its parts including hydrodynamics, mass transfer, kinetics andpermeation models were discussed. The validation of the modelhas been carried out and the model has been subsequently used forthe design of a membrane reactor for mCHP applications. For such aconfiguration 0.44 m2 of the thin-film supported Pd-based mem-branes are needed to separate 3.2 Nm3∙hH2�1 using H2O as sweep gasresulting in an overall HRF of 67% with a substantial reduction inthe amount of required catalyst in the fluidized bed (


FBMR Fluidized Bed Membrane ReactorHRF Hydrogen Recovery FactorPEM Proton Exchange MembranePFR Plug Flow ReactorSMR Steam Methane Reforming

SymbolsAr Archimedes Number, [�]Ci Concentration of component i, [mol m�3]C1, C2 Coefficient for the bed height expansion, [�]Di Diffusion coefficient of component i, [m2 s�1]db Bubble diameter, [m]dp Particle diameter, [m]dt Tube diameter, [m]Ea Activation energy, [kJ mol�1]g Gravity constant, [m s�2]H Bed height, [m]J Partial flux, [mol s�1]Keq Equilibrium constant, [�]Kbe, Kbc, Kce Gas exchange coefficients (bubble ecloud-emulsion

phase), [s�1]ki Forward reaction rate constant for reaction i, [mol kgcat�1

s�1 Pa�2]PH2,0 Permeability of Pd-Ag membrane to H2, [mol s�1 m�1

Pa�n]pi Partial pressure of component i, [Pa] or [bar]R Gas constant, [J mol�1 K�1]Re Reynolds number, [�]Remf Reynolds number at minimum fluidization, [�]rx Reaction rate of reaction x, [mol s�1]Mi Molecular weight component i, [g mol�1]S Selectivity, [�]T Temperature, [K]ug Gas velocity, [m s�1]umf Minimum fluidization velocity, [m s�1]yi Molar fraction of component i in gas phase, [�]4 Sphericity, [�]d Bubble/emulsion phase fraction, [�]ε Porosity, [�]εmf Porosity at minimum fluidization velocity, [�]mi Viscosity of component i, [kg s�1 m�1]ri Gas/particle Density, [kg m�3]Kij i-phase to j-phase gas exchange coefficient, [s�1]

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Direct route from ethanol to pure hydrogen through autothermal reforming in a membrane reactor: Experimental demonstration, ...1. Introduction2. Description of the technology2.1. Catalyst preparation2.2. Membrane preparation2.3. Reactor design and test rig

3. Reactor model4. Results4.1. Catalyst characterization4.2. Membrane permeation in a single tube reactor4.3. Prototype multi-tubular reactor test4.4. Validation of the model4.5. Design of FBMR for μCHP unit

5. ConclusionsAcknowledgementsNomenclatureReferences

Direct route from ethanol to pure hydrogen through …Membrane reactor Hydrogen production...

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