7.New Membrane for Use as Hydrogen Distributor

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Microporous and Mesoporous Materials 109 (2008) 28–37

New membrane for use as hydrogen distributor forhydrocarbon selective hydrogenation

B. Chommeloux a,b, S. Cimaomo a, E. Jolimaitre a,*, D. Uzio a, P. Magnoux c, J. Sanchez b

a IFP Lyon, Catalysis and Separation Division, BP3, 69390 Vernaison, Franceb Institut Europeen des Membranes-UM2, CC 047 Pl. Eugene Bataillon, 34095 Montpellier Cedex 5, France

c LACCO-UMR CNRS 6503 , Universite de Poitiers, 40 avenue du Recteur Pineau, 86022 Poitiers Cedex, France

Received 18 September 2006; received in revised form 30 January 2007; accepted 17 April 2007Available online 29 April 2007

Abstract

Catalytic selective hydrogenation is a reaction widely used in petrochemical industry, particularly to produce intermediate olefins forpolymer production. The main drawback of these industrial conversion processes, carried out in fixed beds, is their low selectivity at highconversion levels: injection of the whole hydrogen stream along with the hydrocarbon feed at the inlet of the reactor does not allow anefficient control of the reaction selectivity. The use of a packed bed membrane reactor (PBMR) in order to improve the distribution ofhydrogen throughout the length of the reactor has been suggested as a possible solution to this problem. Consequently a new method toimprove the hydrogen distribution properties of MFI zeolite membranes, which are potential candidates for a PBMR, is presented in thispaper. The chemical composition of these new membranes were analysed and their performances were assessed before and after a chem-ical modification. After the modification treatment, the membranes were able to distribute hydrogen with a very low loss of hydrocarbonby counter diffusion.� 2007 Elsevier Inc. All rights reserved.

Keywords: Membrane distributor; Hydrogen; Coke formation; Modification

1. Introduction

Selective hydrogenation is a key reaction for the produc-tion of olefins, which are primary intermediate products forthe polymer industry. As an example, but-1-ene – an inter-mediate for the production of low density linear polyethyl-ene – is produced by selective hydrogenation of butadienefrom the C4-olefinic-cut. Industrial processes, based onmultiphase fixed beds packed with a palladium catalyst,usually operate with two beds. In the first bed, 98% ofthe butadiene is converted to but-1-ene, with a selectivityaround 40%. The second ‘‘finishing’’ reactor removes thelast butadiene traces [1,2]. In this second stage, the but-1-

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doi:10.1016/j.micromeso.2007.04.028

* Corresponding author. Tel.: +33 4 78 02 25 13; fax: +33 4 78 02 20 20.E-mail address: [email protected] (E. Jolimaitre).

ene selectivity drops significantly, leading to the formationof the total hydrogenation product (butane) and unwantedbutene isomers. This selectivity decrease can be explainedby the difficulty to mix homogenously the small hydrogengaseous flow with a large hydrocarbon liquid flow. Indeed,under these conditions, phase segregation may appear,leading to an over hydrogenation in the hydrogen rich zoneand insufficient conversion in the depleted hydrogen zone.

To maintain a high selectivity towards but-1-ene, severalresearch groups suggested working in the vapour phasewith a membrane reactor (MR) (Fig. 1) [3–5]. In all thesestudies, the palladium catalyst is located inside the mem-brane pores: contact time between the reagents (butadieneand hydrogen) and the catalyst can be controlled and thenminimised to avoid overhydrogenation. The use of theseMR leads to an improvement of butenes selectivity: Gaoet al. [5] obtained 95% of butadiene converted to butenes

mailto:[email protected]

MR

Hydrogen

stcudorP tnegaeR

Hydrogen

Hydrogen

stcudorP tnegaeR

Hydrogen

PBMR

Fig. 1. Differences between membrane reactor (MR) and packed bedmembrane reactor (PBMR).

B. Chommeloux et al. / Microporous and Mesoporous Materials 109 (2008) 28–37 29

with a selectivity of 95%, Lambert et al. [3] a butadieneconversion of 99% with a selectivity of 65% in a MRagainst 57% in a classical fixed bed reactor. In this config-uration, hydrocarbons and hydrogen have to diffusethrough the membrane to react in contact with the catalystinside the pores. Reactants and products are consequentlynot separated at the reactor output, which is characterizedby the loss of hydrocarbons (up to 50% [4]) in the hydrogenflow.

Another possible solution to improve butenes selectivityis to control the hydrogen quantity in contact with the cat-alyst and with butadiene through an inert membrane. Inthis configuration, the catalyst is packed in one of the reac-tor compartments, which are defined by the membrane [6].This concept of PBMR (Fig. 1.) has shown its efficiency todistribute oxygen for hydrocarbon oxidation [7,8]. One ofthe advantages of the PBMR is that oxygen and hydrocar-bons are not mixed at the output of the reactor. With thisdesign, the reaction efficiency is optimal and oxygen flowcan be reused.

In this study, we choose to work with an inert mem-brane PBMR to carry out the butadiene hydrogenation.

Efficiency of the MR is directly linked to the propertiesof the membrane material, which has to obey to somecharacteristics:

• A minimum permeation flux of the reactant moleculesthrough the membrane. Optimally, the membrane mustbe as selective as possible for hydrogen diffusion (no lossof hydrocarbon, reactant or product). Moreover, a goodcontrol of the hydrogen flux permeating through themembrane must be possible: the optimal hydrogen fluxis strongly dependent on the inlet butadiene concentra-tion and must therefore be adjusted by changing thepressure in the hydrogen compartment. Consequently,very small quantities of hydrogen must permeate whenthe transmembrane pressure is very low.

• The membrane permeation properties must be stable intime. First of all, the membrane material should not bemodified by contact with the hydrocarbon molecules(reactants and products). The membrane should alsoresist to the periodic thermal catalyst reactivation (3 hunder hydrogen flux at 200 �C).

• All these properties have to be satisfied under operatingconditions of the selective hydrogenation reaction(30 �C, atmospheric pressure).

Very few existing membrane materials fulfil all the pre-viously described prerequisites.

Generally, polymers which compose organic membranesinteract with hydrocarbons causing swelling and damage tothe material properties. In spite of an important improve-ment of their thermal stability, organic membranes aretherefore not suitable for our application because of theirsensibility to hydrocarbons [9].

Among all existing inorganic membranes, dense mem-branes and specially palladium membranes are good candi-dates. Hydrogen permeates with an infinite selectivity via asolution-diffusion mechanism through palladium mem-brane [10]. But this mechanism leads to acceptable hydro-gen permeability only at high temperature (above 300 �C)[11]. Moreover, the use of palladium membranes is not rec-ommended below 300 �C because of palladium embrittle-ment [12]. Consequently, their high cost and relativelylow permeability make them not interesting for our appli-cation [6].

Many porous inorganic membranes could be tested forthe PBMR application. VycorTM glass has a poor thermalstability [13] and carbon membrane synthesis is difficultto control because of membrane shrinkage [14]. c-Aluminamembranes are not suitable because their mean pore size istoo large and consequently, important loss of hydrocar-bons by counter diffusion has been observed [3]. Ulti-mately, zeolite membranes seem to be good candidatesbecause of their molecular sieving properties and thermalstability. Synthesis of supported LTA and MFI zeolitemembranes has been reported in literature [15–22]. Porediameter of these zeolites is ranged between 0.40 and0.55 nm and it is well known that linear hydrocarbonsadsorb and diffuse through their microporous structure.

Decreasing the membrane pores size appears to be apossible solution to enhance the hydrogen selectivity ofthe membrane without losing the thermal and chemicalresistance of the material. The Chemical Vapor Deposition(CVD) of tetraethyl orthosilicate-TEOS (kinetic diameterof 0.90 nm) [23], a post-coking treatment of 1,3,5-triisopro-pylbenzene (kinetic diameter of 0.84 nm) [24] have beeneffectively used to reduce the inter-crystalline porosityand defaults on MFI membranes. Indeed, these moleculesare bigger than MFI pore opening and consequently thetreatments do not modify the mean pore diameter. The cat-alytic cracking of silanes [25,26] decreases the pore size andimproves the hydrogen separation performances (selectiv-ity hydrogen/nitrogen improves by a factor 3), but thecomplexity of this method makes it quite difficult to beapplied.

An easy to implement treatment based on the oligomer-ization and cyclization of butadiene in the MFI zeolitepores has been reported by our research group [27]. Yanet al. [24] have reported a zeolite membrane treatment

30 B. Chommeloux et al. / Microporous and Mesoporous Materials 109 (2008) 28–37

based on the same principle: contacting a hydrocarbonmolecule with the zeolite at high temperatures (500 �C),in order to induce a reaction on the zeolite surface. How-ever, the molecule used by Yan et al. is 1,3,5-triisopropyl-benzene (TiPB), whose molecular diameter (8.4 A) ismuch larger than the MFI zeolite pore entrance. Conse-quently, coke formation occurs only around the MFI crys-tals, and the treatment is only useful for correcting theinitial membrane defects (non homogeneity of the crystallayer). On the contrary, butadiene is used in this study.This hydrocarbon molecule is small enough (4.9 A) to enterand react inside the zeolite porous network, and then mod-ify the zeolite membrane permeation properties. In thispaper, a pore size modification of a zeolite membrane isdescribed, along with the influence of the synthesis operat-ing conditions. Details about the composition and thelocalisation of the coke are given. Finally the performancesof these new membranes are evaluated by experimentalpermeation measurements.

2. Experimental

2.1. Experimental set-up

All the experiments were performed in a laboratory scalereactor (Fig. 2). The tubular membrane (length of 150 mm,inner and outer diameters of respectively 7 and 10 mm) wassealed in a module (diameter of 14.5 mm) and placed insidea heating system (cylindrical heating shells).

The purity of gases used (Air Liquide) is ranged between99.950 and 99.999%. Inlet flows were controlled with digi-tal mass-flowmeters provided by Bronkhorst. Inner andouter pressures were fixed by a regulator (Kammer valve)at the output of the reactor.

P

Chromato

Flow m

Ven

Butadiene

Butane

Hydrogen

Nitrogen

Gases for inner compartment

T

Mass flow controller

Three way valve

Pressure regulator andindicator

Temperature indicator

P

T

Fig. 2. Unit d

The flows in the inner and outer compartments wereanalysed at the both reactor’s input and output. Flowswere measured with a digital flowmeter (Agilent, range 0–42 L h�1 with a precision of 2%). Temperature was checkedat reactor’s output as shown on Fig. 2. The compositionwas analysed by gas chromatography on an aluplot column(Varian).

2.2. Original MFI membranes

Original MFI membranes were prepared at IFP [28,29]by in situ crystallization on multilayer tubular a-Al2O3

(Pall-EXEKIA) substrates, as following: a zeolite precursorsolution was prepared by first dissolving Aerosil 380 in anaqueous solution of TPAOH (tetrapropyl ammoniumhydroxide). Hydrothermal synthesis was then carried out.After cooling, membranes were washed, dried and firedin order to eliminate the template. As shown on Fig. 3,the gel was soaked into the pore system of the inner sideof the support, forming zeolite plugs in the top support’slayer. The resulting membranes were mechanically morestable than conventional zeolite top-layers grown on alu-mina supports [22].

Pure silicalite membrane should be obtained by this syn-thesis but it has been demonstrated that support dissolu-tion takes place during hydrothermal synthesis, leading tothe presence of aluminium species in the zeolite[21,22,30]. The charge created by aluminium insertion iscompensated by cation H+ because it is the only cationpresent in the synthesis medium. Pan and Lin [30] indicatea Si/Al ratio between 5 and 13 on a-alumina support withan Al-free precursor solution. Miachon et al. [22] synthes-ised MFI membranes using the same support used in thiswork and by following a very close experimental proce-

graphe

eter

t

P

Hydrogen

Nitrogen

Membrane or stainless steel tube

Temperature regulator

Gases for outer compartment

escription.

Fig. 3. Cross section SEM micrograph of membrane M-1: zeolite crystalsare deposited inside the support layers.

Table 2Selectivation operating conditions of different membranes

Membrane Hydrogen Temperature(�C)

Plateau duration(min)

M-1 Without 200 680M-2 Without 250 270M-3 Without 300 120M-4 5% volume then

without200 400


dure. These authors indicate an average Si/Al ratio ofabout 10 for all over the MFI crystals filling the pores.Therefore, we considered that our original membranespresent similar acidity (Si/Al ratio around 10).

Four membranes were used for this work; the depositedzeolite amount was around 2.4 wt.% except for membraneM-1 (Table 1).

2.3. Membrane selectivation

Prior to the modification, the membranes were heated at350 �C (1 �C min�1 ramp) under a nitrogen flow (innercompartment) during 4 h to remove adsorbed water intothe zeolite. The outer compartment was maintained atatmospheric pressure whereas a transmembrane pressureof 0.5 bar was applied.

For the modification of the membranes by hydrocarboncoking (also named ‘‘selectivation’’), the outer and innercompartments were maintained at atmospheric pressurewhereas 20% in volume of butadiene in nitrogen was flownin the inner compartment and pure nitrogen in the outercompartment (1.6 L h�1 for both fluxes). The reactor washeated at 1 �C min�1 until it reached a plateau.

Four membranes were treated at different experimentalconditions which are summarized on Table 2. In order toevaluate the effects of hydrogen on the efficiency of thetreatment, some experiments were performed addinghydrogen (5%) into the inner compartment. This additionshould decrease the coking rate or even inhibit the reactionbecause hydrogen is a well known inhibitor of cokingreactions.

Table 1Zeolite amount on each membrane

Membrane M-1 M-2 M-3 M-4

Zeolite (wt.%) 6.4 2.2 2.5 2.4

The membrane modification was followed on line bymeasuring the variation of the butadiene flows at the outletof both compartments. Pore size reduction leads to adecrease of butadiene permeability and therefore to adecrease of the butadiene flow at the outlet of the outercompartment and an increase at the outlet of the innercompartment. When both flows were stabilized, the selecti-vation process was considered to be completed and thenthe treatment was stopped. In spite of a similar synthesisprocedure, each one of the membranes presented slightlydifferent characteristics and therefore the selectivation timeobserved (plateau duration in Table 2) was different foreach membrane.

2.4. Membrane performances

Single gas permeances of hydrogen and nitrogen weredetermined for initial and selectivated membranes. Puregas was fed in the inner compartment at 30 �C and nogas was circulated in the outer compartment. Permeanceswere measured for several transmembrane pressures at dif-ferent temperatures (30, 200 and 250 �C).

Fig. 4 shows a schematic representation of the experi-mental methodology used to characterize the hydrogen dis-tribution along the membrane tube. Membrane wascharacterized by two criteria: the percentage of nitrogenretained in the inner compartment a ¼ F inner

N2=f inner

N2and

the percentage of hydrogen permeating through the mem-brane b ¼ F inner

H2=f outer

H2where F and f are the corresponding

outlet and inlet flows.For the evaluation of the membranes efficiency, pure

hydrogen was fed in the outer compartment and pure nitro-gen in the inner compartment. This last gas was used as atest-molecule to evaluate the membrane efficiency. Sincenitrogen is a smaller molecule than hydrocarbons reactantsand products (butadiene, butenes, butane) considered here,it was assumed that membrane performance with a hydro-carbon charge will be even greater than with a nitrogencharge. Inlet flows were set at 10.9 L h�1 in both compart-ments, the membrane module was maintained at 30 �C.

Outer compartment

Inner compartment

outerH2

f

innerN2

f

outerF

innerF

Fig. 4. Membrane characterization: flow comparison at the inlet (f) andthe exit (F) of the reactor.

0.000

0.050

0.100

0.150

0.200

0.250

0.300

0.350

0.400

0 200 400 600 800 1000

Time (min)

But

adie

ne f

low

(L

/h)

0

50

100

150

200

250

300

Tem

pera

ture

(˚C

)

Fig. 5. Butadiene flow evolution at the reactor exit in the inner (h) andouter (j) compartments during the selectivation of M-1 at 200 �C.

0.000

0.050

0.100

0.150

0.200

0.250

0.300

0.350

0.400

0 100 200 300 400 500 600

Time (min)

But

adie

ne f

low

(L

/h)

0

50

100

150

200

250

300

Tem

pera

ture

(˚C

)


0.000

0.050

0.100

0.150

0.200

0.250

0.300

0.350

0.400

0 100 200 300 400 500 600

Time (min)

But

adie

ne f

low

(L

/h)

0

50

100

150

200

250

300

Tem

pera

ture

(˚C

)



The transmembrane pressure difference – the differencebetween the total pressures in the outer compartment andin the inner compartment (atmospheric pressure) – wasvaried between 0.00 and 0.50 bar.

2.5. Crystals synthesis and modification

To gain some more insight on the nature of the coke,MFI zeolite crystals were subjected to the same cokingtreatment as the membranes. The temperatures were iden-tical (between 200 and 300 �C) and the butadiene flowswere adjusted to maintain the same butadiene flow/MFIcrystals weight ratio. As far as the coking conditions arevery similar for the crystals and the membranes, we consid-ered that the coke composition and localization in bothmaterials will be very close. Three MFI samples with differ-ent Si/Al ratio (Table 3) were chosen for this study. Thesamples, provided by Zeolyst, were chosen in a large rangeof Si/Al ratio in order to evaluate the influence of the zeo-lite acidity on the amount and composition of the coke.

Since the treatment takes place in a fixed bed, zeolitepowders had to be shaped in the form of extrudates withan alumina binder. Pural-SB3, zeolite powder and acidicwater were mixed for 30 min at 50 turn min�1 in a mixerN50EHT (Bravender). After extrusion, samples were driedat 80 �C for 12 h and calcinated at 550 �C (5 �C min�1

ramp) for 4 h under an air flow of 1 L h�1 g�1.Extrudates were packed inside a stainless steel tube of

the same dimension as the membrane. The tube was placedin the previously described laboratory unit (Fig. 2), wherethe treatment was performed flowing butadiene at atmo-spheric pressure (1.4 L h�1). The reactor was heated at1 �C min�1 to reach a plateau of 200 �C or at 2 �C min�1

to reach 350 �C. Temperature was then kept constant dur-ing 37 h. At the end of the treatment, heating was stoppedand the reactor cooled down until the reactor temperaturereaches 30 �C. The reactor was then stripped with 3.3 L h�1

of nitrogen for 2 h.

3. Results and discussion

3.1. Membrane selectivation

Figs. 5–7 show the evolution of the butadiene flows dur-ing the thermal treatment of membranes M-1, M-2 andM-3 at three different temperatures. The treatment was

Table 3Identification of the three zeolite samples and their main characteristics

Zeolitesamplename

Compensationcation

Si/Al moleratio

Sample name for

Freshextrudate

Extrudateafter treatment

At 200 �C At 350 �C

Z-15 H+ 15 ± 2 E-15 E-15_200 E-15_350Z-38 H+ 38 ± 2 E-38 E-38_200 –Z-73 H+ 73 ± 2 E-73 E-73_200 E-73_350

stopped when both flows were stabilized, that is why theduration was different for each selectivation. It is observedthat for all the membranes the butadiene quantity flowingto the outer compartment decreases dramatically duringthe first 200 min of reaction. For the treatment at 200 �C(Fig. 5), the butadiene flow continues to increase in theinner compartment meaning that the coking reaction stilltakes place even after the temperature stabilization. Atthe same time, the butadiene flow seems to be stable inthe outer compartment. This result is either due to the lack

Fig. 8. Cross section SEM micrograph of modified membrane M-1.


of precision on the flow measurement at very low flow val-ues or to the consumption of butadiene by coke formationduring the permeation through the membrane. At 250 and300 �C, the flows stabilize earlier: it is very probable thatselectivation is quicker and is achieved before the end ofthe temperature ramp.

The butadiene permeability is very low for each of themodified membranes, meaning the treatment was success-ful in all selectivation conditions. However, it is difficultto compare the permeability values because measurementswere not carried out at the same temperature and perme-ability is strongly linked to this parameter. As explainedabove, each initial zeolite membrane presents differentcharacteristics (effective thickness, inter-granular porosity,number of defects, acidity, etc). Evaluation of their perfor-mances as distributor is therefore necessary to evaluate theeffect of the selectivation temperature.

Influence of the addition of hydrogen in the inner com-partment was studied on membrane M-4. When the treat-ment is performed with hydrogen, no significant variationof the membrane permeability properties is observed. Thisresult is in accordance with the assumption that membranemodification is the result of oligomerization and cyclizationof butadiene inside the pores: hydrogen being an inhibitor ofthese reactions, coke formation should logically be reducedwhen hydrogen is present inside the zeolite porosity.

In order to confirm this assumption, further work hasbeen carried out in order to analyse the chemical composi-tion and the exact location of the coke.

Though the effect of coke formation on membrane per-meability was drastic, the relative quantity of coke presentin the membrane structure is almost negligible (the mem-brane weight increase after selectivation is only around0.4%). The very small amount of coke present in the mem-brane makes the precise analysis of its composition verydifficult: only elementary analysis and microscopic obser-vations were therefore performed on the membranesamples.

3.2. Coke characterization in the membrane

Coke characterization was carried out on the modifiedmembranes M-1 and M-2. Each membrane was cut in slicesto form three samples taken at different membrane’s length(input, middle and output). The samples were observed bySEM (scanning electron microscopy). The carbon–hydro-gen weight ratio (C/H) was then calculated for each sampleusing elementary analysis from results for carbon andhydrogen.

The samples were first observed by SEM to localise thecoke. For this purpose, the samples were coated with a5 nm thick palladium–platinum layer and analysis werecarried out on a JEOL JSM 6340 F apparatus using anelectron acceleration voltage of 1 kV. The observationswere performed in a cross section of the membrane tube,from the inner membrane surface to the outer membranesurface. SEM micrographs of original (Fig. 3) and modified

(Fig. 8) membranes were very close and no coke depositwas observed.

Elementary identification was then performed by EDX(energy dispersive spectroscopy) of X-ray. No carbon wasidentified, neither on alumina and zeolite crystals norbetween zeolite and alumina crystals, in the detection limiti.e. 0.5 wt.% in a volume of 1 lm3 (Fig. 9). These resultstend to prove that carbon is certainly mainly present insidethe cavities of the zeolite crystals.

The carbon weight percent per zeolite weight and thecarbon–hydrogen weight ratio (C/H) – calculated using ele-mentary analysis results for carbon and hydrogen – arereported in Table 4.

The amount of coke deposited in the membrane is verylow but the carbon amount is directly proportional to thezeolite quantity. This can be also consider as evidence thatcoking takes place preferentially inside the zeolite crystals,which represent only a very small percent of the membranetotal mass.

Selectivation temperature seems to be a key parameterto control the coke composition. From the C/H weightratio (Table 4), it is shown that the coke deposited on themembrane modified at 250 �C (M-2) is more aromatic thanthe one on membrane treated at 200 �C (M-1). Selectiva-tion temperature controls the coke composition, whereasit has no influence on the coke concentration. However,the coke amount reported to the zeolite quantity is veryclose for both membranes. In the membranes, the coke rep-resents around 10% weight of the zeolite crystals. It is dif-ficult to compare this amount of coke with themicroporous volume of the zeolite because the chemicalcomposition of the coke is not precisely known.

For membrane M-1, the carbon concentration seems todecrease between the input and the output of the reactor,but the variation is within the uncertainty of the measure.Furthermore, this tendency is not confirmed for membraneM-2.

Fig. 9. EDS spectrum of the inner surface of the membrane M-1.

Table 4Elementary analysis on three samples of membranes treated at 200 and250 �C

Membrane Position in thereactor

Carbon/zeolite(weight ratio)

C/H (weightratio)

M-1(200 �C)

Input (2 cm) 9 ± 2 8 ± 1Middle (6 cm) 8 ± 2 7 ± 1Output (10 cm) 7 ± 2 11 ± 1

M-2(250 �C)

Input (2 cm) 9 ± 2 11 ± 1Middle (6 cm) 9 ± 2 12 ± 1Output (10 cm) 9 ± 2 11 ± 1


3.3. Coke characterization on zeolite crystals

MFI zeolite crystals were subjected to the same cokingtreatment as the membranes. Using the proceduredescribed before, more than 2 g of coked zeolite were pro-duced for each temperature. These samples were then char-acterized by elementary analysis and by an identification ofcoke molecules. As mentioned in Section 2, and since thecoking conditions are very close for both materials (mem-branes and pellets) results concerning coke localizationand composition for the pellets have been extrapolated tothe membrane. Analysis of the coke composition and loca-tion was carried out by the method detailed by Guisnet andMagnoux [31]. Coked samples were first analysed byelementary analysis to determine the total percentage ofcarbon (in the pores of the zeolite and on the crystal

Table 5Amount and composition of coke

Samples Carbon content (wt.%)

Total Inside the porosity

E-15_200 19.6 ± 0.2 14.0 ± 0.2E-38_200 18.1 ± 0.2 12.9 ± 0.2E-73_200 8.3 ± 0.2 7.1 ± 0.2

E-15_350 26.2 ± 0.2 25.5 ± 0.2E-73_350 23.1 ± 0.2 14.3 ± 0.2

a C24–C52 unsaturated compounds of molar mass between 300 and 600 g mo

surface). The samples were then washed with dichloro-methane to eliminate all carbonaceous compounds of thezeolite surface. After drying, a new elementary analysiswas carried out to determine the percentage of carboninside the pores of the zeolite. The washed samples werethen dissolved in hydrofluoric solution in order to solubi-lize mineral compounds and liberate carbonaceous com-pounds of the zeolite porosity. An extraction withdichloromethane was finally carried out and the solublecomponents were analysed by GC/MS experiments. Theresults are summarized on Table 5.

Total carbon content varies between 8% and 26% of theextrudates mass i.e. between 11% and 43% of the crystalsmass. This result is in good agreement with carbon con-tents found on the membrane zeolite crystals (an averageof 10%). This slightly lower value for the membranes canbe explained by two different factors. First of all, the massof zeolite present in the alumina membrane is not knownprecisely. It is determined by measuring the mass of themembrane before and after the zeolite synthesis step. Sincethe zeolite represents only around 2% of the total mem-brane weight, this value cannot be determined with a greatprecision. Furthermore, it is not possible to know if all themass gain is due to zeolite formation or if there is also someamorphous phase formation in the membrane aluminalayer. In that case, the mass of zeolite in the membranewould be over-evaluated and the carbon concentration inthe zeolite crystals underestimated. Another possible expla-nation is that the zeolite crystals in the membrane are sub-

Soluble coke wt.% Composition soluble coke

100 Naphtalene, Indene100 Oligomersa, Naphtalene100 Oligomersa, Indene

0 –50 Naphthalene, Phenanthrene

l�1.

Table 6Comparison of C/H weight ratio trands for crystals and membranesamples

Samples C/H (weight ratio)

EA Experimental (soluble coke)

E-15_200 7 ± 1 13.5–15.0E-38_200 7 ± 1 8.7–15.0E-73_200 7 ± 1 8.0–13.5


jected to a shorter coking treatment compared to the zeo-lite crystals in the pellets. The coke molecules may not havethe time to form/diffuse into the whole network. Moreover,studies of butane adsorption on silicalite crystals show thatsaturation quantity at room temperature is around 10% ofthe mass crystal [33,34]. This gives us an indication on thecoked crystals degree of pore filling: the pore network isprobably nearly completely packed with the coke mole-cules. After the treatment at 200 �C, at least 80% of cokeis located inside the zeolite porosity. At 350 �C, coke alsoseems to be predominantly situated inside the porosity,even though results are more difficult to analyse becauseof the presence of insoluble coke in the porosity. This resultis very important and in opposition with the previousresults reported by Yan et al. [24]. It is now clear that itis possible to form large coke molecules inside the MFIpore network and not only on the crystals surface. Indeed,the zeolite structure had to be completely destroyed byhydrofluoric acid in order to retrieve 80% of the totaldeposited coke. When extrapolated to membrane selectiva-tion, this confirms that coke molecules should be locatedmainly in the zeolite crystals and not on the membranesupport.

As can be seen on Table 5, the total carbon amountincreases with the coking temperature and to the acidiccharacter of the zeolite. Moreover, the amount of solublecoke is also a function of the coking temperature: at200 �C almost all the coke is soluble whereas at 350 �C atleast 50% of the coke is insoluble. The aromaticity of thecoke is therefore more pronounced at 350 �C. For the sam-ples treated at 200 �C, a clear influence of the zeolite acidityon the soluble coke composition can be noticed. For theless acidic zeolite (sample E-73_200), soluble coke is consti-tuted of a mixture of oligomers and aromatics, whereas forthe more acidic zeolite (sample E-15_200), only aromaticmolecules are detected. Coke aromaticity increases withzeolite acidity and coking temperature. It can be underlinedthat, at 200 �C, oligomers are constituted of C24–C52 unsat-urated compounds of molar mass between 300 and600 g mol�1. The carbon number of these molecules isnot always a multiple of 4, which suggests rearrangementand cracking of oligomers followed by the condensationof cracking products with unsaturated compounds (butadi-ene and olefinic products). These results are similar withthe trends observed for the propene coking on MFI zeolitecrystals [32].

Coke concentration increases directly with zeolite acid-ity and coking temperature, which is quite logical sinceboth parameters enhance the coking reactions kinetics.

Total weight percent of hydrogen was also determinedby elementary analysis (EA) in order to calculate theC/H weight ratio. Results obtained are compared withC/H ratio calculated with the formula of soluble cokecompound (Table 6). There are significant discrepanciesbetween the results obtained by EA and the correspondingcalculated values: EA C/H ratio is lower because of wateradsorption on the samples during analysis. Nevertheless,

since the same problem occurred with the membrane sam-ples, the EA C/H weight ratios can be qualitatively com-pared. The C/H weight ratios on the membrane treatedat 200 �C (Table 5) and of the crystals treated at 200 �C(Table 6) are in the same range (around 7). This resultprobably means that the coke has the same composition(oligomers with a molar mass around 300 g mol�1 andsome aromatics). For membrane treated at higher temper-ature, coke is more aromatic.

3.4. Single gas permeances

To gain more insight on the diffusion mechanisms insidethe treated membranes, the membrane selectivated at300 �C was characterized by single gases (nitrogen andhydrogen) permeation measurements at three different tem-peratures. Initial performances were determined at 30 �Cfor a transmembrane pressure of 1 bar. After the coking,permeation flows were too low to obtain a representativevalue so permeances were measured for higher transmem-brane pressure. The experimental results are show onFig. 10.

It is clear that both hydrogen and nitrogen permeancesdrop after selectivation. This effect can be attributed to apartial blockage of the zeolite pores. It is also possible thatthe pore network and/or the crystals surface are completelyplugged with the coke molecules, and that the remainingpermeation flow takes place mainly through the membranedefects. However, we can observe on Fig. 10. that hydrogenand nitrogen permeances increase with temperature. Thisbehaviour cannot be explained by a Knudsen or Poiseuilleflow mechanism, as it would be the case if permeationoccurred mainly through the membrane defects. Indeed,the temperature effect seems to indicate that most of thepermeation flow takes place by diffusion through the cokedmicropores.

A lot of MFI membranes have been synthesised anddescribed in the literature. Single gas permeances are diffi-cult to compare because the membrane synthesis methodare not always the same (number of zeolite layers, acidity,synthesis reactor design) and because the permeance testsare not always carried out in same experimental conditions(temperature, presence of a sweep gas, transmembranepressure applied, . . .). Nevertheless, the hydrogen perme-ances of our membranes (between 2.10�7 and10.10�7 mol m�2 s�1 Pa�1 at 30 �C) are in the same

0.0E+00

1.5E-07

3.0E-07

4.5E-07

0 3 4 5 6 7 8

Transmembrane pressure (105. Pa)

Init

ial p

erm

eanc

e (m

ol.m

2 .s-1

.Pa-1

)

0.0E+00

5.0E-09

1.0E-08

1.5E-08

2.0E-08

Per

mea

nces

aft

er s

elec

tiva

tion

(mol

.m2 .s

-1.P

a-1)

21

Fig. 10. Permeances of initial (h) and selectivated membrane forhydrogen (j) and nitrogen ( ) at 30 (j), 200 (�) and 250 �C (m).


range than values found in the literature (2 · 10�7–30 ·10�7 mol m�2 s�1 Pa�1 [35–37] between 25 and 30 �C).

3.5. Membrane performances as hydrogen distributor

In the test conditions described in Section 2, hydrogenand nitrogen fluxes through the membrane were comparedbefore and after selectivation. As shown on Fig. 11, thetreatment leads to a modification of mass transfer throughthe membrane.

As expected, membrane selectivation leads to a simulta-neous decrease of the nitrogen and hydrogen flows throughthe membranes. An important experimental observation isthat hydrogen flow through the membrane is controlled bythe pressure applied in the outer compartment. By adjust-ing this pressure, it is possible to regulate the hydrogenquantity flowing to the inner compartment, so as to bring

Before

Before

Before

After

After

After

Before

Before

Before

After After After

0.0

1.0

2.0

3.0

4.0

5.0

6.0

0.00 0.25 0.50

Trans membrane pressure (bar)

Flo

w t

hrou

gh t

he m

embr

ane

(L/h

)

Fig. 11. Evolution of hydrogen (j) and nitrogen ( ) flows throughoriginal (before) and modified (after) membrane M-1.

Table 7Effect of the coking and of the test under harder conditions on the performancpressure in both compartments

Membrane Before selectivation After selectivati

a b a

M-1 0.67 ± 0.08 0.32 ± 0.05 0.95 ± 0.12M-2 0.76 ± 0.09 0.17 ± 0.03 0.98 ± 0.12M-3 0.60 ± 0.07 0.28 ± 0.03 1.00 ± 0.12

to the catalyst bed a stoichiometric hydrogen quantity forselective hydrogenation. This increase of hydrogen flowleads to a decrease of nitrogen permeation by counter dif-fusion and consequently leads also to a higher retention ofthe reagent inside the inner compartment. Table 7 showsthe values of the criteria a and b for all the original andmodified membranes.

The presence of coked molecule inside zeolite porosityresults in an increase of criterion a and a decrease of crite-rion b for all three membranes. The butadiene treatment isefficient to improve the membrane performances as hydro-gen distributor.

The quantity of reagent lost by counter diffusion is acritical parameter for the present application. One of themain objectives of this work is to find a membrane ableto retain at least 95% of the charge in the inner compart-ment (a > 0.95). Results from Table 7 show that originalMFI membranes are not acceptable (0.60 < a < 0.76). Onthe other hand, the three modified membranes present verygood performances as hydrogen distributor: at least 95% ofnitrogen is retained in inner compartment and less than 6%of hydrogen is permeating. The treatment is therefore effi-cient: the objective is achieved since leaks by counter diffu-sion are very limited for modified membranes.

Influence of the treatment temperature on the mem-branes performances is not easy to be explained becauseof the experimental errors and the initial difference betweenthe original three zeolite membranes. Even if the selectiva-tion temperature is different, all three modified membranespresent very close efficiencies. However, criterion b seemsto decrease with treatment temperature (b = 0.09 at200 �C and 0.00 at 300 �C), which means that the transferkinetics is lower for higher selectivation temperature. Thisresult can be explained by the coke composition: at highertemperature, the coke molecules are more voluminous (ascan be seen on Table 5) and the volume accessible for dif-fusion in the zeolite network is reduced.

Membranes’ performances were stable for 80 h. More-over modified membranes appear to be resistant after threehours at 200 �C under hydrogen flow: no modification ofthe performances was observed (Table 7).

4. Conclusions

In this paper, the hydrogen distributor properties ofMFI membranes were enhanced by using an easy to imple-ment post-treatment in order to decrease the pores size.

es of the three membranes as hydrogen distributor at 30 �C and the same

on After 3 h at 200 �C under hydrogen flow

b a b

0.09 ± 0.01 0.98 ± 0.12 0.09 ± 0.010.01 ± 0.01 0.98 ± 0.12 0.01 ± 0.010.00 ± 0.01 1.00 ± 0.12 0.00 ± 0.01


This method is very efficient as each modified membranesdistributes an appropriate flux of hydrogen while hydrocar-bons are maintained in the reaction compartment with neg-ligible counter diffusion. The permeability properties werenot affected by the selectivation temperature.

The membranes modification is based on coke synthesisby butadiene oligomerization and cyclization. MEB analy-sis, EDX spectrum and total coke concentration highlightthat coke is only deposited in the zeolite porous structure.

Coke composition is controlled by coking temperature:at 200 �C the coke is composed of polyolefins and somearomatic compounds and at 350 �C of aromatics and poly-aromatics. Aromatics and polyaromatics are expected to bemore stable than oligomers. It will be therefore preferableto use high temperature selectivated membrane to be sureof their long time stability.

The membranes described in this paper were used in amembrane reactor to perform butadiene hydrogenation inthe vapour phase. As expected, butene selectivity for agiven butadiene conversion was greatly improved in themembrane reactor compared to the fixed bed reactor.These results will be discussed in a later article.

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7.New Membrane for Use as Hydrogen Distributor