G Morra

A Desmartin-Chomel

C Daniel

U Ravon

D Farrusseng
email: david.farrusseng@ircelyon.univ-lyon1.fr

R Cowan

A Krusche

C Mirodatos

M Krusche

High-throughput gas phase transient reactor for catalytic material characterization and kinetic studies

Keywords: High throughput catalysis, Combinatorial catalysis, Parallel reactor, Transient experimentation, Xylene hydrogenation, Oxygen storage, Gas adsorption, Pressure swing adsorption

de niveau recherche, publiés ou non, émanant des établissements d'enseignement et de recherche français ou étrangers, des laboratoires publics ou privés.

Introduction

High-throughput (HT) screening relies on a fast testing of large libraries of compounds by using parallel reactors. This research methodology, massively applied for drug development, was extended to Material and Catalysis science in the late 1990s. Commonly, a HT research project can be carried out in two stages. The primary stage consists at screening large and as diverse as possible libraries to identify promising candidates called "hits." The latter are then optimized in a secondary screening stage with targeted libraries. The primary screening is usually characterized by a high degree of parallelization and a very fast or parallel analysis while the secondary screening is driven by higher quality data which are obtained under more realistic operating conditions and shall consider scale-up issues, fluid dynamic and process control.

Numbers of parallel experimental set-ups were developed for gas-phase reaction. An overview of current developments can be found in [START_REF] Karlsson | Microreactor Technology and Process Intensification[END_REF][START_REF] Schuth | [END_REF]. HT testing equipments for primary screening can be divided into two categories. The first ones are monolith or micro-channel type reactors in which each channel is coated with a catalyst layer having a specific composition [3][4][5][6][7]. The second ones are addressable arrays of thin films on wafers [8][9][10][11] or single beads [12]. In most cases, the gas phase is analysed by mass spectrometry using a sampling capillary mounted on a XY robotic actuator to test all the catalysts individually and sequentially. In contrast, HT testing equipments for secondary screening are usually fixed bed reactors with catalyst loading in the range of 20-200 mg. They consist in multi-tubular reactor modules made of steel, quartz or ceramics [13][14][15][16][17][18][19][START_REF] Bricker | Scaling up of catlysts discovered from small-scale experiments[END_REF][START_REF] Bergh | High-throughput workflow development: strategies and examples in heterogeneous catalysis[END_REF][START_REF] Schunk | Mastering the challenges of catalysts screennig in highthroughput experimentations for heterogeneously catalyzed gas-phase reactions[END_REF][START_REF] Serra | Two exemplified combinatorial approaches for liquid-solid and gas-solid processes in oil refining and fine chemicals[END_REF][START_REF] Paul | [END_REF].

Although the secondary screening provides a quantitative evaluation of catalytic performances under operating conditions close to the real applications, much more detailed investigations on catalytic performances must be carried out before scaling up the best selected systems towards mini-pilot scale reactors. This stage, also called tertiary screening, comprises catalysts activation/deactivation studies, kinetic studies, mass and heat transfer analysis, poisoning under real inlet gases and catalyst characterization, in and/or ex situ [25,[START_REF] Weinberg | Impact of high-throughput screening technologies on chemical catalysis[END_REF]. Secondary and tertiary screenings are operated on a reduced number of materials. A few numbers of parallel equipments were developed for these purposes. Perez-Ramirez et al have designed a six flow reactor setup was for HT kinetic studies [START_REF] Perez-Ramirez | [END_REF]. The key feature of the catalytic bench relies on the fluid feeding system which enables to feed each of the six reactors at a different flow rate or pressure while the proportion of reactant remains the same. This design is very well appropriate when one looks at adjusting the contact time to obtain many conversion data points in appropriate ranges for kinetic modelling. The "Spider TM " reactor, developed by Serra et al, consists of 16 parallel high-pressure reactors which can be fed independently for oil refining applications [18].

Whatever the degree of sophistication of these reactors, a key type of catalytic processes can hardly be caught correctly. They relate to catalytic instability phenomena like activation/deactivation processes occurring during steady-state operations, or catalyst behaviour under transient operating conditions which may correspond to permanently unsteady processes employing moving or fluidized bed, post-combustion, reverse flow reactors, domestic applications, etc. If one considers for example deactivation processes, their rates depend usually on catalyst formulation and operating conditions. In a sequential testing process, the first and the last catalyst of a loaded series of six or more catalysts are evaluated at very different time on stream, depending on the time required for analysing each sample. This generally results in bias activity values [14].

A way to tackle that issue is to perform HT kinetic studies on a truly parallel system. A 16-channel reactor system equipped with a spatially resolved FTIR was developed at the University of Delaware to perform kinetic measurements under truly parallel conditions [28]. The system was successfully applied for CO oxidation and NOx storages. In contrast to sequential analysis by gas chromatography and/or mass spectrometry, this system offers the opportunity to record catalytic measurements every 4-12 s for all catalysts which then enable to have access to kinetic data under controlled and similar time-on-stream conditions.

In this paper, a 16-parallel reactor designed for HT transient screening and HT kinetics is presented. The fluidic supply that enables parallel transient experimentation and the reactor modules are described. The capability, versatility and limits of the testing units are illustrated through a series of examples: (1) deactivation under steady-state conditions and kinetic study of o-xylene hydrogenation, (2) characterization of oxygen storage capacity (OSC) of ceria-doped catalysts by transient redox operations, and (3) HT gas adsorption isotherms on microporous materials by transient adsorption/desorption cycles. Another successful application dealing with transient temperature runaway during temperature programmed catalytic soot combustion will be reported in a forthcoming paper.

Experimental

Reactor system for fast material screening

A proprietary 16 channel-multitubular reactor (Fig. 1)which has now been commercialised as the SWITCH 16 reactor system by AMTEC GmbH [29] -was jointly developed within the frame of a joint EU program [30] by IRC-CNRS and AMTEC.

The fluidic supply consists of a combination of two separate feed delivery modules and two 16-port valves that are placed before and after the multi-channel reactor array. The reactors are made from Inconel 600 and are connected to the inlet and outlet lines using VCR fittings, and have internal diameters of 7 mm. They can be operated at up to 600 • C and 13 bars. Each feed delivery module is equipped with five mass flow controllers (MFC) and a controlled evaporator mixer (CEM-Bronkorst). All the tubing and valves can be heated in order to prevent vapour condensation. The upper 16-port valve enables transient experimentation to be carried out in all 16 channels. The two 16-port valves can be synchronized so that when the upper valve switches to a given channel the lower valve selects the corresponding outlet channel which is connected to analytical systems. This system thus offers the opportunity to feed one selected reactor with a selected feed composition (Feed 2) while the 15 others are fed with a second composition (Feed 1), which can be identical to or different from Feed 1 (Fig. 2). This design has several advantages regarding testing capabilities and the type of experiments that can be carried out, that is, it is possible to evaluate catalysts under exactly the same ageing period and more specifically to measure the initial activity of each catalyst. Testing conditions can be varied in the selected channel while the other channels are maintained under standard or inert conditions, and in addition, catalyst deactivation can be studied in the selected channel while the 15 other channels are being regenerated. Finally, the design ensures that very accurate and reproducible flows are guaranteed in the selected channel by the 16-port valve. In the cases of permanent gas, the standard deviation is 0.4%. It is usually higher in the case of vapours and it depends very much of pressure, flows and nature of vapours. In the case of xylene hydrogenation the partial pressure of xylene is measured with a standard deviation of 4%. Conversion calculations are based on the partial pressure of n-heptane used as internal reference. There is therefore no need to use flow restrictors, which is an asset when acquiring suitable transient signals. However, in case of steady state experimentations (not reported here), flow constrictor systems can be used to guarantee equal flow distribution in the 15 other channels. Downstream, both lines are equipped with mechanical back pressure regulators (BPR) which control pressure in the common and selected lines. The back pressure regulators can be bypassed if required.

In the next presented case studies, the reaction products were monitored by gas chromatography and/or on-line mass spectrometry. GC analysis was performed by using an Agilent 3000 version QUAD instrument equipped with a four-channel module. Channel A was used for H 2 ,O 2 ,N 2 ,CH 4 and CO measurements (MolSieve 5A, BF PPU Plot, with Ar as carrier gas), Channel B for CO 2 and H 2 O (PPU, BF PPQ, with He as carrier gas). Channel C for cis-and trans-dimethyl-cyclohexane (OV1) and Channel D for o-xylene quantification (Stabilwax). Mass spectrometer (SM) analysis was performed using an Inficom CPM instrument.

Case study 1: kinetic studies of o-xylene hydrogenation 2.2.1. Background and objective

Toluene and xylenes are typical model molecules used for fundamental studies on the hydrogenation of aromatic hydrocarbons which takes place in petro-chemical refinery units. In contrast with most of combinatorial studies where the catalyst optimization for higher activity or selectivity is targeted, the search of relationships between the catalyst descriptors (metal nature, dispersion, support interaction, doping effects) and the catalytic performances is the main objective when investigating model reactions. Although the model reaction of o-xylene hydrogenation is investigated for several decades, its mechanism is still a matter of debate. A number of conceptual reaction pathways were introduced such as the stepwise [31], rollover [32] and interfacial [33] mechanisms. Usually a kinetic model is supported by experimental data on a specific class of catalysts and the extrapolation of the derived model to other systems is always questionable. The investigation of a kinetic model which holds for diverse catalysts whatever be the active metallic sites, dopants and supports meets this above mentioned objective for searching quantitative relationships between catalyst description and catalytic performance.

Another major issue to be tackled for any kinetic study is deactivation phenomena (carbon deposits, particle growth, metal/support interaction, alloy formation or demixing) which may be specific for each tested catalyst and which cannot a priori be predicted. This issue may be even more critical during HT parallel investigation. When a catalyst is under testing, the deactivation of other catalysts is not monitored. Indeed, the choice of catalytic systems that do not deactivate with time on stream will overcome any aging issues during parallel screening [14]. However, this extremely restrictive solution cannot be envisaged in any industrial oriented project. The investigation of aging kinetics is therefore a prerequisite for handling properly reaction kinetics.

This first case study aims at acquiring a pool of kinetic data in view to further discriminate various kinetic models, while monitoring correctly the catalyst ageing all along the kinetic campaign. Only the initial part of this study will be presented here, to demonstrate the capability of the equipment.

A library of 100 bi-metallic catalysts supported on various supports was prepared and tested in the hydrogenation of oxylene. For sake of clarity, only a selection of six representative catalysts is presented here (Table 1), characteristic of diverse catalytic properties (high/low activity and deactivation shapes) and catalyst formulation (metals and supports). Results on the whole library will be reported elsewhere.

Methodology

The deactivation was first carried out by testing the catalysts one by one for about 1 h, that is, contacting each one with the reacting mixture while the others are maintained under H 2 /N 2 flow. The switching protocol described in the experimental section was used for this purpose. Several equations were investigated to model the observed deactivation profiles. Then, kinetic data were obtained by varying H 2 partial pressure and contact time (total flow) for a given catalyst at a given temperature. When the whole set of experimental conditions is completed, the catalyst is placed in the initial conditions applied for deactivation studies and analyses are performed. Next, the selection valve switches to the next catalyst. At the end of the cycle (16 catalysts), the whole heating block is set to the next temperature dwell and the next cycle is started as soon as the temperature is stabilized.

Testing conditions

For the initial deactivation study, catalysts are tested with a stream of 0.5%(mol) o-xylene, 5%(mol) heptane, 25%(mol) hydrogen and nitrogen as diluent (Feed 2). The other channels are maintained under H 2 /N 2 flow (Feed 1). The catalysts are previously reduced under H 2 at 400 • C for 2 h.

For the kinetic study, reactors are loaded with 50, 100, 200 or 400 mg of catalysts depending on their previously measured activity. The H 2 partial pressure was varied from 0.5%(mol) to 48% (mol), the temperature from 100 to 200 • C and the total flow from 20 mL CNTP /min to 50 mL CNTP /min. Each GC analysis was duplicated.

Generally speaking, it was observed that conversion can attain the thermodynamic equilibrium for one catalyst but is below detection limits for a second one although the two tested systems have "similar" compositions. As an example, the doping of Pd/alumina catalyst with Ge can multiply the activity by a factor 20 [34]. This parameter sensitivity causes experimental issues when one wants to investigate kinetics. As a consequence, the catalyst mass in the channel, that is, contact time, must be adjusted with respect to the initial activity in order to measure conversion and selectivity in appropriate ranges (outside diffusion regime but being large enough to be correctly measured) for further kinetic modelling.

Kinetic modelling

In order to test different kinetic models, a catalyst (catalyst #7) which does not deactivate and shows low conversion in the experimental conditions was selected as reference catalysts. A large dataset of more than 200 data points was collected and used as benchmark for model discrimination study and to simulate the impact of deactivation rates on the kinetic parameters.

About 10 different existing models and derivatives were investigated, from simple power-law to more advanced models based on diverse mechanism steps. For the present study, the selected model (referred to as Cin1) assumes that the adsorption of reactants is competitive and H 2 does not dissociate. It can be described by the following steps:

A + * K A ←→ A * H 2 + * K H ←→ H * 2 A * + 2H * 2 K AH 4 ←→ AH * 4          Equilibrated, (r 1 )AH * 4 + H * 2 k 1 -→ C + 2 * , (r 2 )AH * 4 + H * 2 k 2 -→ T + 2 *
where A stands for the aromatic, C and T for cis-DMCH and trans-DMCH, respectively [32,35,36].

A rate equation for o-xylene conversion was built accordingly and it was also checked that no diffusion limitations occurred under the testing conditions. For that, the criterion τ = -ln(1 -X)/k, where τ is contact time; X, conversion; k, rate constant, was found to be fulfilled for all tested catalysts.

Case study 2: oxygen storage capacity measurements on ceria-doped catalyts 2.3.1. Background and objective

Ceria-based mixed oxides (Ce x M 1 -x O y ) are versatile solid oxygen exchangers. At high temperatures (400-800 • C) the redox cycle Ce 3+ ⇔ Ce 4+ +e -facilitates oxygen storage and release from the bulk fluorite structure. In addition, the surface redox chemistry of ceria is sensitive even at low temperatures to crystal structure defects [37] which can be tuned by substituting some of the Ce cations with ions of different size and/or charge [38,39]. This makes them ideal candidates for catalytic oxidation applications where their oxygen storage capacity (OSC) can be used in a dynamic way, such as the three-ways catalysis, and/or the hydrogen production under non-stationary conditions [40].

This second case study aims at showing how the effects of doping the ceria support of platinum based catalysts on OSC can be investigated using the parallel testing reactor set-up above presented, aiming at discovering new materials [41].

Methodology

A set of 12 catalysts containing 2 wt% Pt on Ce 0.9 M 0.1 O x was prepared (where M = Pb, Bi, Zr, V, W, Mo, Y, La, In, Sn, and two combinations of Zr/Bi, respectively) by impregnating platinum on various doped ceria supports. In practice, one reactor channel was loaded with a dopant free high surface area Pt/CeO 2 catalyst for comparison. The temperature was varied stepwise from 100 to 200 • C every 20 • C. OSC measurements were quantified by H 2 consumption in transient mode using the parallel reactor under isothermal conditions: Catalyst samples were oxidised in air for 15 min, then flushed with N 2 and then reacted with a reducing mixture of 10% H 2 in He while measuring continuously the hydrogen consumption by the Mass spectrometer for 3 min. While one sample was analyzed under 10% H 2 flow, the other ones were maintained under N 2 flow. After a cycle was completed, the temperature was increased by 20 • Candane w cycle started following the previous experimental sequence.

Case study 3: adsorption isotherms of microporous materials 2.4.1. Background and objective

The selective capture of carbon dioxide is a major objective for any large-scale process, CO 2 being one of the main greenhouse effect gases. For instance its separation from methane is required in the processing of natural gas. Generally, pipeline specifications for natural gas require a carbon dioxide concentration below 3%. The technology most widely used for carbon dioxide removal is amine absorption, but amine plants are costly and subjected to strong environmental constraints. An alternative, which is more recently applied, is the removal of carbon dioxide by means of an adsorption process. Pressure swing adsorption (PSA) technology has gained interest due to low energy requirements and low capital investment costs. Its main issue is to find an appropriate adsorbent which shall adsorb selectively one component over a mixture while not being adsorbed too strongly, otherwise the regeneration step can negatively affect the economy of the process. Separation by PSA can be based either on the different adsorption affinity of the components (thermodynamic separation), or on the different diffusion rates in the adsorbent (kinetic separation) [START_REF] Ruthven | Pressure Swing Adsorption[END_REF][START_REF] Ackley | [END_REF]. For thermodynamic separation, the two key criteria of adsorbent effectiveness are the working capacity (WC) and the working selectivity (WS). The definition of working capacity is given in Eq. (1):

WC A = Q A (x, P, T ) ads -Q A (x, P, T ) des (1)
where Q A ,(x, P, T) ads and Q A ,(x, P, T) des are the equilibrium loadings of A (mol/g) corresponding to the gas phase mole fraction (x), total pressure (P) and temperature (T) at adsorption and desorption steps, respectively. The working capacity physically represents the net amount of adsorbate transacted on and off of the adsorbent for each complete cycle, providing a good indication of the amount of adsorbent required for the separation.

The working selectivity of component A relative to component B (defined in Eq. ( 2)) is a key indicator since it represents the separation efficiency. The operating conditions will dictate the selectivity and hence the performances of the adsorbents.

α = WC A WC B (2)
Here, the adsorption isotherms are assumed to follow Langmuir behaviour; that is, the sorbents exhibit monolayer coverage as the saturation pressure is approached and have an energetically homogeneous surface. However, this parameter can be easily extended to other adsorption models as well. Isotherms were modelled using the Langmuir equation for a single component which can be written in the following form:

Q = Q sat Kp (1 + Kp) (3)
where K = k exp(-H/RT) and p is the partial pressure of the single component in the diluted feed. This third case study aims at demonstrating the ability of the Switch 16 set-up for rapidly acquiring the above described key indicators for CO 2 adsorbent screening.

Methodology

The adsorption isotherms are calculated from breakthrough curves measurements taken at different total pressure [44]. Breakthrough curves consist at quantifying gas uptakes of an adsorbent bed in a transient operating mode. A gas mixture is sent to a selected channel for which the pressure is fixed at a given value. Partial pressures are measured on line by MS for a few minutes till the adsorbents get saturated. While adsorption measurements take place at the selected channel, the other channels are flushed under He flow to allow gas desorption for next measurements under different conditions of temperature and/or pressure.

In the reported study, channels are loaded with 200 mg of sieved adsorbents. The feed is a gas mixture of 12.5% CO 2 , 12.5% N 2 in He with a total flow rate of 40 ml min -1 . The breakthrough curves were recorded at different total pressure from 1.4 to 6 bars. The adsorbents are zeolites supplied by ENI Technology (NaX, NaY, NaZSM5, KL, NaBeta) and purchased from Flucka for 13X or made at the laboratory (NaZSM5 with Si/Al = 27, 240). Microporous metal organic frameworks Cu 2 (BTC 3 ) and a second undisclosed compound were prepared by precipitation.

Results and discussion

Kinetic studies of o-xylene hydrogenation

Deactivation processes

The deactivation decay curves acquired for six representative catalysts over a period of several hours are reported in Fig. 3. As can be seen, the more active the systems, the more it is deactivating. Two general aging trends were observed: an initial exponential deactivation for all catalysts in the very first hours on stream, followed/or not by a linear deactivation. These two types of deactivation profiles were described by a simple exponential model (Deac1) or by a combination of an exponential and a linear model such as (Deac2)

A = A ∞ + k e -(t/τ) (Deac1) A = A ∞ -νt + k e -(t/τ) (Deac2)
where A is the activity; A ∞ is the residual activity; ν is the rate of linear decrease; k is the pre-exponential factor and τ the relaxation time.

For both deactivation types, three catalysts exhibiting high (catalysts #1 and #4), medium (catalysts #2 and #5) and low activities (catalysts #3 and #6) are presented. As such, the range of initial conversions spans from 20% for catalysts #1 and #4 to 1% for catalyst #6.

Concerning the possible models for activity decay with time on stream, let us briefly recall that according to Levenspiel [45], the most general expressions are summarized in Table 2.

For the present case of o-xylene conversion, no concentration dependency was evidenced, (i.e., rate of decay was essentially independent of o-xylene, hydrogen or DMCH concentration). Additional TPO experiments performed on spent catalysts have shown that coke deposition occurred on deactivated samples. Therefore coking effects may be the cause of this activitydependant poisoning. A mechanism by coke deposition which encapsulate the metal particles and/or block the pores by deactivation progressively the catalyst bed from top to bottom of the reactor would qualitatively fit with the observed decay curves. In addition, better fits were observed in most cases when an overlaying linear trend is added to the initial exponential model. This linear and constant deactivation, independent from activity, is assumed to be limited in time (for sake of model validity). This underlying mechanism could be a non-selective poisoning of the catalytic bed coming from a toxic component contained in the feedstock (e.g. traces of oxygen or of another pollutant). It would be classified as a side-by-side activity.

Indeed, these speculations require further analyses, such as establishing the decay rates as a function of coke contents (via advanced coke deposition kinetics acquired on an oscillating thermobalance TEOM, for instance) which might ascertain the decay origin. However, within the scope of the present study, this advanced knowledge of the aging process is not required for further discrimination of the reaction kinetics.

Table 3 presents the results of the aging modelling. As mentioned above, catalyst #7 showed no deactivation and has been used as a benchmark to test kinetics models and to simulate the deactivation impact. As previously noted, the most active catalysts show the biggest global deactivation. Although catalysts #3 and #7 have similar compositions, they thus behave quite differently in terms of deactivation (catalyst #3 activity dropped by near 20% within 5 h). This may be due to the use of different precursors (palladium (II) acetate for catalyst #3 and palladium (II) dichlorine for catalyst #7) and of different synthesis protocols.

Depending on the model (Deac1 and 2), one or two times dependency parameters can be calculated. Since the exponential decay deactivation is activity dependant, the exponential relaxation time τ can be used as a catalyst deactivation descriptor. In contrast, the linear decreasing rate ν indicates a deactivation phenomenon which is independent of the activity. As a consequence, this parameter could be used to describe the interaction of a polluting environment on a specific catalyst.

Deactivation and kinetic model integration

The target to develop a methodology allowing a fast access to kinetic parameters for a large number of catalysts imposes that the number of experimental conditions tested shall be limited. In so far, kinetic parameters of the 100 catalysts were derived from 15 to 30 data points using the kinetic model presented in the section "Experimentals."

We first have investigated the ability of the model to describe catalyst performances while deactivating by using raw kinetic data, that is, not corrected from deactivation. Fig. 4 shows that the kinetic model Cin1 is able to adequately describe both deactivating (catalyst #3b) and non-deactivating (benchmark catalyst #7) systems from type Deact1. Similarly, the model predicts as well the activity for a system of type Deact2 (catalyst #6) despite a strong deactivation (about 50%) in the period where kinetic data are collected.

This means that one cannot use the standard deviation between experimental and calculated data to discriminate experiments strongly influenced by the deactivation from experiments little or not affected by deactivation.

Quantification of the impact of deactivation on kinetic parameters was studied. Kinetic parameters were calculated by applying diverse rates of deactivation on the basis of the above-described aging models (see extrapolated activity at t 0 (mmol/h/g cata )inTable 2). Complete release of the methodologies and results will be published in a forthcoming publication. Nevertheless it can be pointed out here the large impact of the deactivation on the computation of kinetic parameters, the deepest effect being on adsorption entropies and enthalpies, especially widening prediction bands.

Thus, it is utmost important to take into account deactivation profiles while processing data for reaction kinetic modelling. By reprocessing the data according using appropriate deactiva-tion models "intrinsic" kinetic parameters ( H, S, E a ) can be calculated as catalyst descriptors.

Oxygen storage capacity measurements on ceria-doped catalysts

The results on the OSC measurements by transient parallel experimentations on various ceria-doped catalysts are shown in Fig. 5.

The reference catalyst containing 4.75 wt%Pt on high surface area ceria shows a very high OSC whatever the temperature. In contrast, the 2 wt.% Pt doped samples prepared for this study show lower OSC under the prevailing conditions, demonstrating the large impact of Pt concentration on OSC extent. As a matter of fact, it was shown elsewhere [46] that part of the Pt loading can form oxide-like clusters inside a fluorite structure, ensuring both an accelerated oxygen diffusion from bulk to surface and along the ceria surface and the storage of overstoichiometric oxygen species (including hydroxyl groups).

Among the 13 dopants, four (Pb, W, Mo and Sn) allow to increase the OSC by factors of two to three, mostly independently of the temperature within the investigated range. In contrast, vanadium doped ceria shows a continuous increase of the OSC with the temperature to attain values of 120 mol/g.

In general, these results clearly demonstrate that solid state of ceria supports determines the level of OSC and of oxygen transfer dynamics. As a matter of fact, by doping the fluorite structure with exo-cations, various structural effects are induced, such as creating and/or stabilising oxygen vacancies [39], which affects directly the intrinsic oxygen diffusion and storage capacity. In the presence of a metal, the reducible ceria provides adsorbed oxygen species that can spill over from the support to the metal particles, the reduced Ce 3+ centres being subsequently reoxidized by gaseous oxygen (Eqs. ( 4) and ( 5)).

2Ce 4+ + O 2-+ * → 2Ce 3+ + O * (4) 2Ce 3+ + 1/2O 2 → 2Ce 4+ + O 2- (5)
In the presence of hydrogen, ceria can be reduced in the bulk (Eq. ( 6)), creating anionic vacancies and increasing the mobility of the remaining lattice oxygen atoms. The resulting hydroxyl anions can combine to give water and O 2-(Eq. ( 7)) with the Ce 3+ cations reoxidised by molecular oxygen to Ce 4+ as shown above.

2Ce 4+ + 2O 2-+ H 2 → 2Ce 3+ + 2OH - (6) 2OH -→ H 2 O + O 2- (7)
A common feature for the above mentioned doping cations favoring the OSC is that the most stable forms of the corresponding oxides are essentially amphoteric or acid except V which can be either acid or basic. By maintaining an optimized level of surface acidity, these dopants would somehow neutralize ceria basicity (hydroxyl groups) and therefore inhibit formation of stable adspecies such as formates and carbonates which would in turn inhibit oxygen mobility and spillover towards platinum particles [41]. Indeed more advanced catalyst characterization such as in situ FTIR acid-base measurements are required for supporting these assumptions, which also points out the limits of the present HT investigation.

Finally, as reported in Ref. [41], it was also observed for this catalyst library that strong relationships exist between OSC and CO preferential oxidation in a hydrogen flux (PrOx). As a general trend, it was found that the most performing CO oxidation samples exhibited the poorest OSC, indicating that excess of oxygen delivered by the system is quite detrimental for the selective oxidation of CO in the presence of hydrogen.

From an engineering point of view, the method of HT transient OSC measurements at iso-temperature described in the present study was proved to be efficient for assessing key features of oxygen storage and transfer dynamics in diverse materials libraries. This approach can therefore be used as an alternative guideline for optimizing primary screened formulations.

Adsorption isotherms of microporous materials

Adsorption isotherms of gas and vapours in porous materials are obtained by means of breakthrough curve measurements. The experimental procedure is very similar to the OSC measurements described above. Examples of breakthrough curves on seven different microporous materials are shown in Fig. 6. When the selected channel is fed by the gas mixture, some of the gases are preferentially adsorbed as indicated by a decrease of the partial pressure on the mass spectrometer. When the porous network is saturated, the steady-state is achieved and another breakthrough curve can be operated on the next sample. Thanks to the fluidic supply design, when one adsorbant is under testing, the other ones are desorbed under flowing sweep gas (He). As a consequence, desorption is allowed to proceed 15 times longer than adsorption. This operating mode fits well with the usual fact that desorption under sweep gas takes longer time than adsorption, especially when strongly activated adsorption takes place, which can be an issue for fast material screening.

Because of some dead volumes in the tubing and channels, a sharp decrease of the pressure is observed in the first seconds for all cases. Adsorption capacities are calculated by subtracting blank experiments performed with a non-porous SiC sample.

From breakthrough curves at different pressure, gas adsorption isotherms were established as shown in Fig. 7.

In good agreement with literature data, CO 2 is adsorbed preferentially over N 2 on zeolites with a high aluminum content such as 13X and NaX. In contrast, CO 2 adsorption is much less pronounced for NaZSM-5 with a low aluminium content and Cu 3 (BTC) 2 metal organic frameworks due to non polar frameworks [47].

Higher CO 2 adsorption capacities are observed for HZSM5, Na-ZSM5-27 and for our reference materials NaX and 13X. The reproducibility was validated using five replicates of NaZSM5-240.

Isotherms obtained at different temperatures were modeled using Langmuir equation in order to calculate adsorption parameters. As an example, an activation energy of 20 kJ/mol was obtained for the reference 13X zeolite in good agreement with literature data.

This experimental protocol enables to measure 16 isotherms per day for pure gases or for gas mixtures. From the isotherms, working capacity and selectivity can be calculated as well according Eqs. ( 1) and ( 2) at different adsorption and desorption pressure conditions. In turn, the device authorizes a rational screening of adsorbents because it enables to evaluate different materials while defining the best process conditions for each one.

Conclusions and perspectives

New HT technologies offer the opportunity to increase by several orders of magnitude the number of tested samples. However, instead of screening catalysts against a single criterium such as activity or selectivity, the double approach of HT kinetic and HT characterization was privileged to obtain relevant quantitative information for further Quantitative Structure-Activity Relationship studies, thus providing new guidelines for discov-ery. The case studies presented here showed how to access key parameters/descriptors using HT transient methods for catalysis (hydrogenation reaction), oxygen storage capacity (redox materials) and adsorption capacity (microporous adsorbents). This double approach should facilitate the simultaneous optimization of both material and process engineering.

Fig. 1 .

 1 Fig. 1. Operating principle of the multi-channel reactor (left), and flow scheme (right).

Fig. 2 .

 2 Fig. 2. Feed system scheme showing the two 16-port valve positions and flow through the selected reactor.

Fig. 3 .

 3 Fig. 3. Deactivation profiles during o-xylene hydrogenation for six selected catalysts. (a) Exponential (Deac1) and (b) exponential + linear (Deac2) decay models. Part I (from 0 to 60 min) corresponds to the collection of deactivation data; Part II (from 60 to 250 min), to the collection of kinetic data and Part III, to the acquisition of data in the initial deactivation conditions after 4 h on stream.

Fig. 4 .

 4 Fig. 4. Parity plots resulting from the kinetic modelling for benchmark catalyst #7 (a) and catalyst #3 (b), using non-corrected data. The dashed lines represent the 95% prediction bands.

Fig. 5 .

 5 Fig. 5. Oxygen storage capacity (OSC) measurements as a function of the temperature expressed in micromol per gram of catalysts for various Pt/ceria-doped materials.

Fig. 6 .

 6 Fig. 6. N 2 -CO 2 mixture breakthrough curves of seven microporous materials obtained at 100 • C and five bars total pressure: dashed lines CO 2 , plain lines N 2 .

Fig. 7 .

 7 Fig. 7. CO 2 adsorption isotherms at 120 • C from a mixture N 2 /CO 2 /He on zeolites. Five replicates of the sample NaZSM5-240 demonstrate the measurement reproducibility.

Table 1

 1 Selected catalysts which display diverse formulation, initial activity and deactivation profile in o-xylene dehydrogenation

	Catalyst ID	Metal	Support	Catalyst
		content (wt.%)		weight (mg)
	1	0.5 Rh	Low surface Al 2 O 3	30
	2	0.5 Ni, 0.16 Mo Low surface Al 2 O 3	50
	3	0.5 Pd	Low surface Al 2 O 3	50
	4	0.5 Rh, 0.19 In High surface Al 2 O 3	30
	5	0.5 Pt	Low surface Al 2 O 3	50
	6	4 Ni, 0.77 In	Low surface Al 2 O 3	100
	7 (benchmark)	0.3 Pd	Low surface Al 2 O 3	200
	Catalyst #7 was characterized and tested in details and is used as a reference
	benchmark.

Table 2

 2

	Summary of deactivation models
	Activity decay	Concentration-dependence	Aging rate equation
	Independent	None	-da/dt = k d a m
	Parallel Series Side-by-side	Reactant A Product B Feedstock impurity P	-da/dt = k d C n A a m -da/dt = k d C n B a m -da/dt = k d C n P a m

a is the activity relative to fresh catalyst at a fixed point in a catalyst bed; k d deactivation rate constant; C is the concentration of reactant A, product B or feedstock poison P; n and m are orders of decay on concentration and activity, respectively.

Table 3

 3

	Results of deactivation kinetics
	Catalyst ID	Deactivation	Coefficient of	Extrapolated activity	Global deactivation	Decreasing rate ν	Relaxation time τ (h -1 )
		model	determination (R 2 )	at t 0 (mmol/h/g cata )	[(A(t 0 ) -A(t end ))/A(t 0 ); %]	(h -1 ) (Deac2)	(Deac1 and 2)
	#1	Deac1	0.99	2.29416	77	-	2.12
	#2	Deac1	0.99	0.40329	43	-	0.57
	#3	Deac1	0.99	0.0269	18	-	0.86
	#4	Deac2	0.99	4.22971	80	278	0.26
	#5	Deac2	0.99	0.92671	67	57	0.13
	#6	Deac2	0.99	0.05593	54	1.58	0.32

Acknowledgements

This work was supported by the EU FP6 Integrated Project TOPCOMBI (Project No. NMP2-CT2005-515792) and IFP-Lyon.


E Odier

Y Schuurman

C Mirodatos
email: claude.mirodatos@ircelyon.univ-lyon1.fr

Non-stationary catalytic cracking of methane over ceria-based catalysts: Mechanistic approach and catalyst optimization

Keywords: Methane cracking to hydrogen, Noble metals supported on ceria, Pt, Ir, Rh, Ceria doping, Transient kinetic modeling, In situ DRIFT, Non stationary processes

The non-stationary cracking of methane over various noble metal/CeO 2 -doped catalysts at 400 and 600 8C was followed by DRIFT spectroscopy and on the basis of the identified elementary steps a simplified kinetic modeling is proposed. The production of H 2 by direct decomposition of CH 4 on the noble metal is improved by the capacity of ceria to store carbonaceous surface species thanks to: (i) the spillover of carbonyls from noble metal particles towards basic hydroxyls formed on partially reduced Ce sites and (ii) the reverse spillover of ceria oxygen towards metal to oxidize the carbon issued from methane cracking. The resulting formate adspecies are in turn oxidized into carbon dioxide during the regeneration step. Doping the ceria with basic lanthanide oxides and replacing Pt by more efficient and eventually better dispersed metals for methane decomposition like Rh and Ir lead to significant improvements in the hydrogen productivity.

Introduction

Within the perspective of sustainable hydrogen production for feeding fuel cells and/or generating synthetic fuels, new catalytic processes have to be designed for meeting non-stationary conditions imposed by the targeted application. For example, frequent start-up and shutdown operations have to be foreseen for any domestic stationary or on-board reformers. Non-steady-state processes can also be considered as an alternative to conventional continuous ones, being designed specifically to take advantages of transient operations. Recently, a novel process for hydrogen production with purity higher than 90% was proposed by sorption-enhanced steam methane reforming reactions in two parallel fixed bed reactors operated in a cyclic manner. First, through addition of a CO 2 sorbent into a reforming reactor, the reactions of reforming, water-gas shift and CO 2 sorption were combined, and more CH 4 was expected to convert H 2 in one reactor. Second, regeneration of the sorbent was carried out in the other reactor. The hydrogen production and sorbent regeneration processes were carried out simultaneously in the two fixed bed reactors, operated in a cyclic manner by switching a methane/ steam feed and an Ar-containing feed between the two reactors at a fixed feed switchover time [1].

On a similar basis, other promising non-stationary processes were proposed years ago, like the cyclic two-step reforming which combines a step of hydrocarbon cracking into hydrogen and carbon deposits followed by a step of catalyst regeneration, as proposed in Refs. [2][3][4] for the case of natural gas conversion:

step I CH 4 ! C deposited on catalyst + 2H 2 step II C deposited on catalyst + O 2 ! CO 2
For this case, both the oxygen and the carbon storage capacity of the catalytic systems have to be considered as key factors for ensuring proper operations control. As a matter of fact, oxygen stored on the catalyst may partly oxidize the hydrogen formed during step I, but also favors the reoxidation in step II. In turn, a high capacity of C storage associated with fast dynamics appears as a prerequisite for such a transient process. In turn, the capacity to store carbon deposits during the cracking step is essential for avoiding a fast deactivation of the catalyst but also for allowing a fast and controlled regeneration that is not detrimental for the catalyst structure.

At variance with Ni based systems that may store carbon as carbide and graphite at high temperature (700-800 8C), supported noble metals, which operate at lower temperature, require C storage assistance from the support. For this case, a low temperature process is expected to be more efficient since the rate of adspecies diffusion on the support has to be comparable with the rate of carbon formation on metal during the cracking step.

When associated to oxygen storing materials like CeO 2 and doped CeO 2 solid solutions, noble metals were already recognized as effective catalysts for the partial oxidation of methane (POM), water gas shift (WGS) equilibration and selective oxidation of CO (SELOX) [5][6][7]. Their potential for being operated under transient conditions has already been successfully demonstrated in our group by considering the model catalyst Pt/CeO 2 , investigated by means of in situ IR studies [8], mechanistic and kinetic modeling [4]. The exact and respective role of the metal and the support remains however a key question, which has to be answered in order to design improved materials able to meet the requirements of these relatively low temperature transient operating conditions.

As a continuation of our previous studies, this paper will first give an overview of the main findings on this model catalyst and then present new results on a series of noble metals supported on doped ceria, aiming at: (i) improving the catalysts performances and (ii) gaining a better understanding of the dynamic processes of this two-step process.

Experimental

Materials

High surface area ceria powder used either pure or modified by La and Pr oxides doping (210-180 m 2 /g) (provided by Rhodia) were first calcined at 600 8C, then impregnated with an aqueous solution of Me(OH) 2 (NH 3 ) 4 salts (Me = Pt, Rh and Ir) at various contents. Precursors were calcined at 500 8C for 8 h, reduced in a hydrogen flow at 300 8C for 15 h and then kept under inert gas (Ar) before reaction. A commercial 0.5% Pt/ gAl 2 O 3 catalyst was used as reference. The composition of all ceria-based catalysts after calcination is reported in Table 1.

Catalysts characterization

The specific surface area of the samples was calculated from the N 2 adsorption-desorption at À196 8C by using the BET method, after desorption under vacuum for 3 h at 300 8C.

Noble metal dispersion was evaluated for some cases by the method of temperature-programmed reduction (TPR) over used samples. Its principle, well adapted to the case of noble metals supported on redox supports like ceria, consists in a TPD measurement carried out in the presence of hydrogen in the sweep gas [9]. Typically, the sample is first re-reduced at 400 8CinH 2 flux for 4 h, then cooled down and heated under a flux of 4 vol% H 2 in Ar from 25 to 800 8C with a ramp of 10 8C/ min. By integrating the low temperature peak of hydrogen desorption (at ca. 120 8C), an evaluation of the noble metal dispersion was obtained.

Testing procedure

The reaction was carried out at 400 8C under atmospheric pressure either in fixed bed reactor (4 mm i.d.) or in an in situ DRIFT cell (SPECTRATECH, 5 mm i.d.) [4,8]. Experiments consisted in flowing alternatively 20-50% CH 4 in inert gas (during 1 min), then pure inert gas (as a flush) and finally 20% O 2 in inert gas (as long as necessary to eliminate all the reversible C stored), all mixtures with a total flow rate of 20 ml/ min (STP). The gas concentration at the reactor outlet was continuously monitored by on line mass spectrometry. For the case of DRIFT experiments, infrared spectra were recorded The catalysts are listed and ranked according to increasing hydrogen yield.

with a Nicolet Magna-IR 550 spectrometer, using a MCT detector with a resolution of 4 cm À1 at a frequency of one spectrum every 3 s. The methane conversion was calculated as the amount of methane converted along the whole cracking step over the amount admitted during that step. The selectivity was the amount of hydrogen collected at reactor outlet as dihydrogen over the total amount of hydrogen contained in the reacted methane. [CO max ] was the highest CO concentration observed during the methane cracking step.

Transient curve modeling

In order to describe adequately the transient experimental curves, the fixed bed reactor was described as a cascade of perfectly stirred tank reactors in series, allowing us to account for the axial dispersion. For a low degree of dispersion (Bo < 50) the number of stirred tanks, N, and the Bodenstein number, Bo, are related as: N % Bo/2 [4]. The fixed bed reactor was described by a system of ordinary differential equations (ode's). No radial gradients were taken into account and a onedimensional model was applied. Mass balances were developed for both the gas phase and the adsorbed phase. The reactor was considered to be isothermal.

The experimental reactor consists of three distinct zones: (1) the tubing from the switching valve to the catalyst bed, (2) the catalyst bed and (3) tubing to the mass spectrometer. Each zone was taken into account explicitly in the model. The resulting set of differential equations was solved numerically by means of an improved Euler method with variable step size control. Parameter optimization was accomplished by a Mead Nelder simplex algorithm [4].

Results and discussion

Catalysts characterization 3.1.1. BET surface area

As can be seen in Table 2, the initial high ceria surface area (184 m 2 /g after calcination at 600 8C) was found to be slightly decreased to about 180, 131 and 163 m 2 /g after 0.8% Pt, 5% Pt and 0.5% Ir addition, respectively. The decrease in surface area after metal addition (the higher the larger content of noble metal) is ascribed to the action of the acidic impregnation solution, activating the surface of Ce-Me-O particles (hydroxylated and chlorinated). This would favor a re-stacking of the primary particles and also the incorporation of part of the noble metal atoms as oxidic species within the nanodomains, forming new intercrystalline/domain boundaries, partly plugging the surface nanoporosity of the mixed oxide [10,11].

A significant decrease in specific area was noted after long term catalytic tests carried out at 400 and 600 8C (ca. 42 m 2 /g for 5% Pt/CeO 2 ), which indicates a severe sintering of the ceria material, possibly due over heating during the regeneration steps (similar surface area decrease was obtained after ceria calcination at 800 8C).

Noble metal dispersion

The evaluation of Me/ceria catalysts from TPR measurements gave values comprised between 12% (for 5 wt.% Pt/ CeO 2 ) to 30% for the 0.5 wt.% Ir/CeO 2 catalysts. A slight decrease in metal dispersion was noted after catalytic tests carried out from 400 to 600 8C (e.g. 10 and 15% for the Pt and Ir/CeO 2 catalyst, respectively), in line with the over heating phenomena mentioned above. Note that for reference samples, a reasonable agreement in metal dispersion was obtained with SEM measurements (e.g. particles of 2-3 nm were observed for the Ir/ceria materials).

Reaction mechanism over reference Pt/CeO 2 3.2.1. Catalytic results over Pt/CeO 2 sample

Under forced non-steady-state conditions (Fig. 1), a methane conversion of ca. 9 wt.% with an H 2 selectivity of ca. 40 wt.% was obtained at 400 8C for a gas hourly space velocity (GHSV) of 3400 h À1 during the methane pulse over the unpromoted 0.8 wt.% Pt/CeO 2 , leading to an average hydrogen yield of 3 wt.% (Table 1). The remaining hydrogen extracted from methane was released as H 2 O essentially during the oxidation step. Less than 1% CO was detected during methane cracking into H 2 or during the oxidative regeneration step, which was the prerequisite for the selection of this catalyst. Only a small CO 2 peak and a permanent trace release of water were detected during the methane pulse.

In situ DRIFT analysis

Three series of IR spectra have been recorded under reaction conditions with the following samples: 0.8 wt.% Pt/CeO 2 (A), CeO 2 (B), 0.5 wt.% Pt/Al 2 O 3 (C). Under CH 4 flow, the main kinds of observed species are listed in Table 3 and illustrated in Fig. 2: gaseous species (CH 4 ,C O 2 ), adspecies on platinum (gem dicarbonyl, linear carbonyl, bridge-bonded and multibonded carbonyl species, depending probably on the Pt dispersion and interaction with the support) and adspecies on cerium oxide (hydroxyls -OH, formates -CO 2 H, carbonates -CO 3 ), in agreement with Li et al. [12].

Gaseous carbon dioxide is present at the beginning of the methane pulse step I for samples A and C because of the previous regeneration of the sample by oxidation step II: the oxidized Pt surface favors initially the total oxidation of methane until Pt is reduced and the cracking starts.

C storage.

The three samples were compared to determine precisely which kinds of adspecies participate efficiently to C storage. The Pt-free sample B (CeO 2 ) shows that cerium oxide can activate methane and oxidize it into surface carbonates (see Fig. 2a). But in this case methane conversion is very low and this oxidation, which could be initiated at 200 8C [13], corresponds to the interaction of methane with surface lattice oxygen anions created by the previous regeneration step. These carbonates are essentially irreversible (only a minor part can be outgassed as CO 2 during the regeneration step) and therefore cannot be considered as efficient C storage agents. They are also present in sample A (0.8 wt.% Pt/CeO 2 ). However, at higher temperature, like under POM or methane combustion conditions, these carbonates may be decomposed easily into CO 2 , forming a pool of carbon but also of oxygen able to spillover towards the noble metal particles under transient conditions [14].

With sample C (0.5 wt.% Pt/Al 2 O 3 ), the decomposition of CH 4 on Pt is rapidly stopped after methane admission because the non-reducible alumina support is unable to store carbonaceous adspecies (CH 4 conversion < 1%). Only Pt carbonyls are formed, resulting from CH 4 partial oxidation by oxidized Pt sites.

On sample A (0.8 wt.% Pt/CeO 2 ), Pt carbonyls are also detected but with different spectral characteristics: a shift of the maximum absorption (from 2061 (C) to 2042 (A) cm À1 ) followed by a shoulder around 1968 cm À1 indicates that the nature and structure of carbonyls have changed (more reduced Pt acceptors and extension from single to multi-bonded carbonyls) (Fig. 2b).

The carbonyl species, which in turn can either desorb into the gas phase or be stored on the ceria surface by reverse spillover, can only be formed thanks to oxygen provided by ceria (surface lattice oxygen which can migrate towards platinum) reacting with carbonaceous species provided by the methane decomposition. This CO storage on ceria occurs through the formation of formates (CO + OH on reduced Ce 3+ ions) as can be seen in Fig. 2c. The presence of reduced ceria is confirmed during the regeneration step: at the beginning of the pulse, O 2 is firstly totally consumed before CO 2 and H 2 O grow a The observation of gem dicarbonyls on Pt/Al 2 O 3 might come from a higher Pt dispersion, favoring the relative concentration of coordinatively unsaturated sites (CUS). Another explanation like the existence of partially oxidized Pt atoms after catalyst regeneration under oxygen (which would also lead to high frequency carbonyl bands) cannot be excluded.

up simultaneously to a maximum after which a breakthrough of oxygen occurs. The mass balance calculations for the regeneration step show that a part of injected O 2 is used to the reoxidation of the ceria (Ce 3+ ! Ce 4+ ).

A more quantitative approach of carbon storage can be attempted on the basis of the ceria surface potential for storing formate species. The amount of carbon stored during the methane cracking step was 138 mg C/g cat for the 0.8 wt.% Pt/ CeO 2 catalyst (Table 1). OSC measurements (oxygen uptake after H 2 reduction at 400 8C) indicated that about 2870 mmol O/g cat are available during redox cycles. The commonly agreed OH concentration on a ceria surface can be estimated to ca. 13 OH/nm 2 which makes a potential capacity for storing formates of 2.2 mmol C/m 2 . For a BET surface of 130 m 2 /g cat , it makes about 286 mmol C/g cat , i.e., 3432 mg C/g cat . By comparing that figure to the effective storage of carbon (138 mg C/g cat ), it comes that only 4% of the capacity of carbon storage is effectively used during the cracking step (this percentage may reach 14% in the worse case after ceria severe sintering during long term runs at 600 8C). This important statement tends to indicate that the BET surface of the ceria support is not determining for the carbon storage. Accordingly, other parameters like the diffusion rate of CO along the surface to reach the basic OH groups able to fix the carbonyl groups produced by the Pt particles should be considered. That parameter would include metal dispersion and ceria acido-basic surface properties as sensitive factors as well.

H storage. OH groups are detected essentially in ceria containing samples (A and B

). During the cracking step, two types of hydroxyl groups can be distinguished (Table 3): at high frequency (around 3720 cm À1 ) the most basic type I, which is consumed after methane admission (negative peak), while a second one at a lower frequency (around 3600 cm À1 ), type II, is accumulating. Another observation is that the IR bands of the hydroxyls are larger in the Pt containing sample A than for the Pt-free sample B. This might indicate that the Pt/ ceria interaction improves the reducibility of the support by the fast decomposition of CH 4 ,H 2 or H 2 O on Pt that provides reactive H atoms able to form OH groups on ceria. In the presence of O 2 (regeneration step), the low frequency band at 3665 cm À1 disappears while the high frequency band at 3690 cm À1 is restored within a few seconds after exposure. This can be related to a reoxidation of the ceria surface and the generation of type I OH groups as a result of the formate oxidation/decomposition. In contrast, the more acidic type II hydroxyls are formed during the cracking step due to the reaction of hydrogen (as dihydrogen arising from methane cracking or as water arising from the initial combustion of methane) with basic ceria surface oxygen. In turn, these type II hydroxyls are decomposed during reoxidation, likely via water formation.

In order to ascertain the above mechanistic proposal, the series of IR spectra recorded after introducing methane was analyzed through the 2D correlation method described in Ref. [8]. The synchronous correlation analysis revealed a strong correlation between the following IR bands transients: formates, high frequency hydroxyls (linked to Ce 3+ ) and carbonyls on Pt. It comes therefore that these rather basic OH groups are required for formate formation, in agreement with Li et al. [16], which supports the following reaction steps: where the active sites are high frequency OH groups linked to Ce 3+ ions (or Ce 4+ with oxygen vacancies), CO being provided as Pt carbonyls migrating towards the support through interfacial sites.

Role of OH groups for C storage.

Among the three tested systems, only the Pt/CeO 2 catalyst displays both a significant C storage capacity and the ability to accumulate formates at the CeO 2 surface. It can therefore be concluded that the formation of formates contributes directly to an efficient C storage, though it requires hydrogen atoms (to form the OH groups), provided by the CH 4 decomposition, which in turn decreases the H 2 yield. The other forms of carbonaceous adspecies participating to C storage -carbonates -are more stable therefore less adapted to fast and reversible storage. When contacted to the reducing atmosphere characteristic of the cracking step (CH 4 +H 2 mixing), the ceria surface is partially reduced by providing oxygen atoms to platinum for oxidizing the carbon atoms arising from methane cracking. In the mean time, the oxygen migration from ceria bulk to surface, necessary to maintain a fully oxidized state, is not fast enough to locally avoid H storage as OH groups arising from the reaction of hydrogen (gas phase and/or Pt-H) with surface oxygen. It can therefore be concluded that formates formation needs partially reduced ceria containing oxygen vacancies and hydroxyl groups.

Overall scheme of the cyclic process

From the IR study, a detailed description of the elementary steps that control the overall process is now available. The carbon issued from methane cracking is stored mainly as carbonyl, formate and carbonate species, and the hydrogen species not released in the gas phase are stored both as formate and hydroxyl groups. In analogy with proposed models for POM and WGS on similar Pt/CeO 2 catalysts, elementary reactions can be proposed as follows, as discussed in Ref. [4]:

CH 4 ðgÞþPt ? CH x;Pt þð4 À xÞH Pt ? C Pt þ 4H Pt (1) CH 4 ðgÞþ4PtO ! CO 2 þ 2H 2 O þ 4Pt (2) C Pt þ O Pt=supp ! CO Pt (3) CO Pt þ OH supp ? HCOO supp (4)
2H Pt ? H 2 ðgÞ (5)

H Pt þ O supp ? OH supp (6)
The experimental data from Fig. 1 are compared with the model curves in Fig. 3. Adequate profiles have been calculated for both methane and hydrogen curves (other products are not represented here). In particular, the long tailing of hydrogen desorption after the methane pulse is finished can be directly related to the reversibility of step (6).

These main aspects of hydrogen production under step transient conditions over the reference catalyst Pt/CeO 2 , summarized in this first part, were used as a basis for improving the catalyst performance by changing the noble metal nature and/or concentration and by doping the ceria.

Catalytic performances over Me/doped ceria systems

Table 1 summarizes the performance (conversion, selectivity, hydrogen yield and carbon storage capacity) for the different Me/ doped ceria catalysts tested under cyclic conditions.

Influence of temperature

By increasing the cracking temperature from 400 to 600 8C, both conversion and hydrogen selectivity, and therefore the yield, increase significantly. In turn the CO concentration tends also to increase, as expected from a faster oxygen release from the ceria support, able to oxidize the carbon deposits during the cracking step. Higher temperatures are also in favor of an easier desorption of CO adsorbed on the noble metal particles. From these statements it can be concluded that the low temperature domain is interesting for having a low CO concentration in the hydrogen rich effluent (<1%) but will require methane recycling since the conversion of methane remains rather low ($22%). In turn, at 600 8C, much improved hydrogen yields are obtained (up to 34%) with the drawback of a larger CO concentration (9-10%). This precludes any direct feeding into a state-of-the art PEMFC without downstream H 2 purification, but is quite compatible for feeding a SOFC or MCFC.

Influence of contact time

By decreasing GSHV (increasing contact time) of a given system, the conversion, at 400 8C, was not proportionally Fig. 3. Outlet gas concentrations obtained during step I over 5 wt.% Pt/CeO 2 catalyst at 400 8C compared with curves calculated according to the model [10].

increased, as shown in Table 4 for the 0.8% Pt/CeO 2 sample. This feature indicates that the cracking does not proceed under an effective catalytic regime. In turn it would be essentially controlled by the diffusion of methane cracking fragments from noble metal to support, as deduced previously from the calculation of the effective occupancy of the ceria support surface by formate adspecies (only 8% at 400 8C).

Effect of ceria doping

Ceria-based mixed oxides (Ce x M 1Àx O y ) are versatile solid oxygen exchangers. At high temperatures (400-800 8C) the redox cycle Ce 3+ , Ce 4+ +e À facilitates oxygen storage and release from the bulk fluorite lattice. This makes them ideal candidates for catalytic oxidation applications [6,7,[9][10][11][12]. However, the surface redox chemistry of ceria is sensitive even at low temperatures to crystal structure defects, which can be tuned by substituting some of the Ce cations with ions of different size and/or charge [6,10,14].

By comparing the systems Pt/ceria and Pt/ceria-zirconia, detrimental effects of ZrO 2 doping on hydrogen productivity were observed (worst yield at 600 8Ci nTable 1). Zirconia addition, known to massively improve oxygen mobility, will therefore ensure a permanent flux of oxygen towards the Pt surface, which will promote the oxidation of hydrogen during the methane cracking step. This flux of oxygen also explains the largest CO concentration observed at 600 8C for the zirconiadoped system. Therefore, in contrast with POM reaction where the zirconia doping was recognized as beneficial for the hydrogen yield [10,14], the reverse effect is observed for the sequential cracking process.

At variance with zirconia, ceria doped with lanthanide oxides improved significantly the catalyst performances (best H 2 yields obtained at 400 and 600 8C, Table 1). In addition, for a given metal (e.g. Ir), ceria doped with praseodymium or lanthanum lead to an increase in C storage capacity and therefore a decrease in the CO 2 release during the cracking step. Since this positive effect cannot be ascribed to a surface area or metal dispersion effect (both parameters did not vary in a significant way for all catalysts), it is tentatively ascribed to changes in surface properties. As a matter, higher concentration of basic OH groups tends to develop after lanthanide oxides doping [10,14]. This effect is likely to favor carbon trapping via carbonates and formates as shown previously by IR spectroscopy. However, the mean CO concentration was not affected by the ceria doping with lanthanides since it is essentially controlled by the enthalpy of desorption on the metal phase.

Effect of metal nature and dispersion

Among the transition metals, the noble metals can be ranked according the required energy for methane dissociation, based on the theoretical calculation reported by Au et al. [START_REF] Au | Methane conversion vs. hydrogen selectivity data obtained with the sequential process at 400 and 600 8C and domains of performances for a conventional commercial POM process (dashed grey) and targeted for new process development[END_REF] (Fig. 4). This gives the following sequence: Rh > Ru > Ir > Pt > Pd.

In the present study, the best catalytic results in term of hydrogen yield were obtained over Rh and Ir with or without ceria doping (Table 1) in accordance with the predicted trends. This indicates that in addition to the diffusion from metal to support, the cracking process also depends on the methane dissociation over the metallic phase. It can also be noted that the best performances of Ir and Rh supported catalysts were obtained with very low metal content (0.5 wt.%, respectively) and the highest metal dispersion (ca. 30%). Therefore, at this stage of the study, it can be proposed that: (i) the nature of the metal is likely to monitor the dynamics of methane dissociation and (ii) its dispersion (likely related to its content) determines the concentration of interface sites (between metal and ceria) which allow the transfer of metal carbonyls to the basic hydroxyls of the support to form formate and the reverse spillover of ceria oxygen to metal sites for oxidizing the carbon issued from methane cracking.

Storage capacity

The carbon storage capacity, calculated from the transient regeneration curves, varies from 187 to 902 mg C/g cat . This indicates that a larger and larger part of the surface storage capacity (evaluated previously from the theoretical maximum concentration of surface OH groups) is used during the cracking step. A linear correlation is observed between the amount of converted methane and the sum of the amount of gas phase CO plus the amount of carbon stored on the support, thus confirming a closed carbon mass balance within the experimental error. Indeed, no straightforward relationship exists between this capacity and the hydrogen yield since the former depends on many factors such as the basicity of the support (e.g. to accommodate carbonate species), the availability of lattice oxygen and/or defects for ensuring the formation and spillover of the various carbon-containing adspecies. In turn, the hydrogen yield is essentially controlled by the methane decomposition over the metal phase.

Comparison between the best performances achieved for the sequential and the conventional oxidative steam reforming processes

In order to assess the efficiency of this sequential process, the obtained performances in terms of conversion versus hydrogen selectivity were reported for both 400 and 600 8Con a diagram where the domains corresponding to the existing conventional POM performances and the targeted domain for the development of new processes are reported (Fig. 5).

For both temperatures, as seen above, an effective improvement of the yield is achieved by changing the process and catalysts parameters. However, it is interesting to note that all the data align along two curves essentially determined by temperature, which might indicate that improvements are still possible.

As can be seen, the best performances obtained at 600 8C compare reasonably with a commercial steady-state oxidative steam reforming process, but with the advantage of lower CO x concentrations in the hydrogen rich feed, which might be beneficial for feeding directly a SOFC or a new generation high temperature PEMFC. In terms of hydrogen productivity, it was calculated that the sequential process could provide up to 1 m 3 / h/kg of hydrogen at 600 8C, while the targeted value would be around 5 m 3 /h/kg. Thus, improvements in both the process as well as the catalyst formulations are still necessary, in order that the sequential process might be an effective alternative for transient applications in the future.

Conclusion

By investigating a sequential two-step process for producing hydrogen rich effluents from natural gas at relatively low temperature (up to 600 8C), numerous elementary steps dealing with transfer/spillover phenomena of carbon, oxygen and hydrogen containing adspecies between the active metal and support phases were identified and quantified. In situ IR spectroscopy analysis made it possible to prove that during methane cracking over Pt/CeO 2 catalyst, formate species are formed via the direct reaction of CO carbonyls with basic OH groups of ceria surface vibrating at 3720 cm À1 . The reactivity of the latter, associated with a high rate of surface diffusion, originates the catalyst capacity for storing carbon quantitatively and reversibly. Significant improvements of this storage capacity and therefore of the hydrogen productivity were achieved both by changing the metal phase to Ir and Rh as predicted from theory and by doping the ceria support by basic lanthanide oxides to increase the concentration of these reactive OH groups.

By comparing the best performances obtained with the sequential process and with a conventional oxidative steam reforming processes it was concluded that the sequential process might be an alternative for transient applications in the future.

Fig. 2 .

 2 Fig. 2. DRIFT spectra at 400 8C: (a) under CH 4 flow for samples A (0.8% Pt/ CeO 2 ), B (CeO 2 ) and C (0.5% Pt/Al 2 O 3 ), (b) changes in normalized intensity of the three main bands of carbonyl species over sample A during the CH 4 pulse (Pt-CO(1): multi-bonded at ca. 1930 cm À1 , Pt-CO(2): bridge-bonded at ca. 1970 cm À1 , Pt-CO(3): linear at ca. 2040 cm À1 ) and (c) changes in sample A spectra (carbonate and formate wavenumbers range) during CH 4 pulse, striping under He and regeneration under O 2 .

Table 1

 1

	Catalysts composition and performance (GHSV = 3400 h À1 )
	Me	Catalysts		Temperature (8C)	Conversion%	S H 2 %	Y H 2 %	[CO] max (%)	H 2 prodced	C storage
		Loading (wt.%)	Ceria doping						(ml/min)	(mg C/g cat )
	Pt	0.8	Undoped	400	8.6	39.5	3.4	0.2	0.7	138
	Pt	5	Undoped	400	15.6	37	5.7	0.1	1.1	nd
	Pt	5	23.7% Pr 6 O 11	400	14.2	39.8	5.6	0.2	1.2	228
	Ir	0.5	10.0% La 2 O 3	400	12.5	47.0	5.9	0.2	1.3	464
	Ir	0.5	Undoped	400	18.9	39.2	7.4	0.2	1.6	187
	Ir	0.5	23.7% Pr 6 O 11	400	22.7	32.9	7.5	1.5	1.6	902
	Rh	0.5	Undoped	400	13.7	64.6	8.8	0.3	1.9	597
	Pt	5	14.5% ZrO 2	600	38.2	75.0	28.7	14.0	6.1	217
	Ir	0.5	Undoped	600	36.7	81.6	29.9	8.9	6.4	235
	Pt	5	Undoped	600	36.7	83.7	30.7	9.0	6.6	585
	Ir	0.5	23.7% Pr 6 O 11	600	33.6	92.9	31.2	8.7	6.7	377
	Rh	0.5	Undoped	600	34.1	92.9	31.7	10.8	6.8	361
	Pt	5	10.0% La 2 O 3	600	32.6	97.6	31.8	9.0	6.8	461
	Pt	5	23.7% Pr 6 O 11	600	38.5	84.2	32.4	9.5	6.9	228
	Ir	0.5	10.0% La 2 O 3	600	36.2	94.6	34.3	9.5	7.3	414

Table 2

 2 BET surface area for high surface area ceria material and after Pt and Ir addition and testing

	Support/catalyst	BET area (m 2 g À1 )
	HSA CeO 2 uncalcined	195 AE 2
	HSA CeO 2 calcined at 600 8C	184 AE 2
	HSA CeO 2 calcined at 800 8C5 7 AE 1
	0.8% Pt/CeO

2 calcined at 500 8C, reduced at 300 8C but unreacted 180 AE 2 5% Pt/CeO 2 calcined at 500 8C, reduced at 300 8C but unreacted 131 AE 2 5% Pt/CeO 2 after long term run at 400 and 600 8C 42 AE 1 0.5% Ir/CeO 2 calcined at 500 8C, reduced at 300 8C 163 AE 2 Fig. 1. Inlet and outlet gas concentrations obtained over 0.5 wt.% Pt/CeO 2 tested at 400 8C by sequentially feeding a fixed bed reactor (4 mm i.d.) with methane, then helium and then oxygen.

Table 3

 3 Main IR bands observed during the CH 4 reaction step over A, B, C samples at 400 8C

	Gas phase
	Gaseous CH 4	3200-2800, 1360-1250 cm À1
	Gaseous CO 2	2390-2280 cm Basic Ce-OH
	Type II hydroxyl groups	3640 cm À1 (A, B)	Acidic Ce-OH
	Unidentate carbonate	1066, 1456 cm À1 (A, B)
	Bidentate carbonate	1028, 1560 cm À1 (B)
	Bridged carbonate	1015, 1398 cm

À1 Solid phase (Pt) Carbonyls on Pt a 2061 cm À1 (C): gem dicarbonyl 2042 cm À1 (A): linear carbonyl, 1968 cm À1 (A): bridge-bonded carbonyl 1928 cm À1 (A): multi-bonded carbonyl Solid phase (cerium oxide for samples A and B) Type I hydroxyl groups 3720 cm À1 (A, B) À1 (A) Formate 1378, 1580 cm À1 (A)

Table 4

 4 Changes in catalytic performance with GHSV for the 0.8% Pt/CeO 2 sample at 400 8C

	GHSV (h À1 )	Conversion (%)	H 2 selectivity (%)	H 2 yield (%)	[CO] max (ppm)
	1700	11.6	30.8	3.6	170
	3400	8.6	39.5	3.4	300
	8500	4.8	38.3	1.9	380

Fig.

4

. Methane dissociation energy over various metals. Reproduced from Ref.

[15]

.

Acknowledgements

Air Liquide is fully acknowledged for partial support of that work and providing reference data for the commercial POM process. Rhodia is acknowledged for kindly providing the undoped and doped ceria samples.


Jérôme Canivet

Sonia Aguado

Cécile Daniel

David Farrusseng

Engineering the Environment of a Catalytic Metal-Organic Framework by Postsynthetic Hydrophobization

niveau recherche, publiés ou non, émanant des établissements d'enseignement et de recherche français ou étrangers, des laboratoires publics ou privés.

Engineering the Environment of a Catalytic Metal-Organic Framework by Postsynthetic Hydrophobization

Jerome Canivet, Sonia Aguado, Cecile Daniel, and David Farrusseng* [a] Heterogeneous catalysis is of paramount importance in many areas of the chemical and energy industries. Often, however, reactions can be hindered or reaction rates limited by poisoning effects originating from moisture in the air or from the water formed during the organic transformation. Water can be adsorbed and block catalytic sites, leading to their deactivation. [1][2][3] This drawback has motivated the design and engineering of catalytic materials with hydrophobic features, such as the hydrophobic outer shell of enzymes, to prevent water-induced catalyst poisoning.

Metal-organic frameworks (MOFs) represent an extensive class of porous organic-inorganic crystalline materials. [4][5][6][7][8][9][10][11] Due to their calibrated pore size, they are regarded as new shapeselective catalysts analogous to zeolites. MOFs have already been reported to catalyze a broad range of organic transformations involving their Lewis acid nodes as well as their Brønsted acid-base properties. [12,13] Many reports have dealt with carbon-carbon bond formation catalyzed by unmodified MOFs through reactions such as Suzuki-Miyaura cross-coupling, [14] the Mukaiyama aldol reaction, [15] alkylation [16,17] or polymerization. [18,19] To obtain more sophisticated MOF catalysts, many research groups have examined the functionalization of these materials. The post-synthetic modification (PSM) of MOFs, an appealing route toward functionalized frameworks, [20][21][22][23][24] involves chemical modification of the solid after formation of the crystalline structure, assuming that the primitive MOFs employed are sufficiently porous and robust. Post-synthetic modification can provide a wide range of isotopological structures from a single MOF by treating it as a substrate with a variety of organic reagents. The insertion of pendant groups onto or into the MOF makes it possible to add chemical functionality while retaining the MOF's overall framework. Cohen and co-workers extensively studied the covalent organic PSM of a variety of amino-functionalized MOFs. [25][26][27][28][29] This strategy was used by his group and others to generate metal complexes on MOFs, creating a new class of Lewis acid catalysts. [22,23,30] Furthermore, Cohen extended his study to the hydrophobization of amino-containing MOFs through amide coupling, in order to increase their moisture resistance. [31] In a parallel work, Yaghi et al. also reported the functionalization of porous materials [32] and especially the reactivity of ZIF-90 against an amine or a reducing agent. [33] The effect of the functionalization of a MOF by hydrophobic agents on its catalytic activity has, however, not been reported to date.

In contrast to previous studies, our methodology is based on the modification of the environment of the catalytic centers and not on the insertion of new sites onto the MOF structure. We therefore studied the ability of modified zeolitic imidazolate framework (ZIF) materials to accelerate the rate of a reaction involving water formation. We chose the Knoevenagel condensation, which is a crossed aldol condensation of a carbonyl compound with an active methylene compound leading to C=C bond formation. This reaction is widely applied in the synthesis of fine chemicals and is classically catalyzed by bases in solution. [34] This reaction can also be catalyzed by solid bases, such as metal oxides. [34] The water that is produced in the course of the reaction usually competes with substrates for adsorption, however, thereby acting as a poison. Porous solids such as modified SBA-1, [35] zeolites, [36,37] MIL-101 (TOF = 328 h À1 ), [38,39] IRMOF-3 (TOF = 180 h À1 ) or MIL-53, [40] and others [41][42][43] have already been employed as active heterogeneous catalysts for the Knoevenagel reaction.

We report herein the fine tuning of hydrophobic properties of a MOF by post-synthetic modification to optimize its catalytic properties. To our knowledge, this is the first study showing that the engineering of the hydrophobic/hydrophilic environment can enhance the catalytic activity of a MOF by an order of magnitude.

Our work focused on a porous substituted imidazolate material (SIM-1, formulated C 10 H 10 N 4 O 2 Zn) discovered by our group and belonging to the class of ZIFs. [44][45][46] SIM-1 is a robust material, isostructural to ZIF-8 and consisting of ZnN 4 tetrahedra linked by carboxylimidazolates. The aldehyde moiety present on the structure walls allows organic modification in the solid state, such as imine synthesis by condensation with primary amines to give the corresponding imino-functionalized SIM-2. This SIM-1 functionalization by imine condensation proceeds under mild conditions. In a typical experiment, a sample of desorbed SIM-1 (50 mg) was suspended in anhydrous methanol (5 mL) and the desired amine (1 mmol) was added under stirring. The suspension was allowed to react at room temperature for 24 h. After reaction, the solid was centrifuged and washed three times with ethanol and then dried under vacuum, providing the corresponding SIM-2 as a crystalline offwhite powder. In this manner, SIM-1 was treated with the primary amines dodecylamine to give SIM-2(C 12 ) (Scheme 1). The C 12 aliphatic chains present at the surface of the material create a hydrophobic shell surrounding the framework.

Powder XRD analysis of the SIM-2(C 12 ) sample showed a slight loss of crystallinity despite retention of the initial structure. Notably, the porosity of the SIM materials was maintained throughout the modification. As expected, however, the BET surface area decreased from 470 m 2 g À1 for SIM-1 to 112 m 2 g À1 for SIM-2(C 12 ). Liquid 1 H NMR analysis after digestion in acidic DCl-D 2 O-[D 6 ]DMSO solution showed a 22 % modification according to the signal integration. The presence in the NMR spectrum of a new typical peak at 8.09 ppm, corresponding to the imino proton, confirmed the efficiency of the organic transformation. In addition, NMR spectroscopy of the fully imino-functionalized imidazole ligand alone showed no hydrolysis of the imine bond under the same conditions (see the Supporting Information). The low functionalization rate with dodecylamine can be attributed to the high steric hindrance of the alkyl chain, the C 12 imino moiety being nonetheless present on one fifth of the ligands.

We performed water adsorption isotherms to assess the hydrophobic feature of SIM-2(C 12 ). Comparing the data obtained with those of SIM-1, there was a dramatic effect due to the substitution with the C 12 chains (Figure 1). Whereas SIM-1 adsorbed water at low pressure, SIM-2(C 12 ) only adsorbed water close to the condensation point, providing evidence of the water repellence of the porous material.

The catalytic behavior of the SIM materials was then evaluated in the Knoevenagel reaction. We chose substrate combinations of two different representatives of both carbonyl and active methylene compounds that are allowed to react in a sol-vent-free system. In a typical run (Scheme 2), either benzaldehyde (1A; 20 mmol) or acetophenone (1B; 20 mmol) reacted with either ethyl cyanoacetate (2a; 20 mmol) or malonitrile (2b; 20 mmol), with catalysis by SIM material (0.1 mol %, 6-7 mg) to give the corresponding Knoevenagel adduct 3 at 323 K.

According to the catalytic results (Table 1), SIM-1, ZIF-8, and SIM-2(C 12 ) were all active for Knoevenagel condensation with a catalyst loadings as low as 0.1 mol %, leading to high yields of the compound denoted 3Aa. The initial turnover frequencies calculated after 30 min (TOF init ) showed that SIM-2(C 12 ) was the most active, with a TOF init. of 1400 h À1 for 3Aa and 1920 h À1 in the case of 3Ab. In these conditions, we observed no formation of the Tishchenko adduct corresponding to the self-condensation of benzaldehyde. Moreover, even at room temperature, SIM-2(C 12 ) still converted 1A and 2a to give 3Aa in 80 % yield after four hours, whereas SIM-1 was almost inactive under the same conditions. Finally, the activity of SIM-2(C 12 ) was more than four times higher than that found for the reference molecular catalyst, dodecylamine.

SIM-2(C 12 ) was also an efficient catalyst for the reaction involving ketones, such as acetophenone 1B, albeit with a lower activity, as already reported for zeolites. [36,37] Figure 2 shows the time-dependent formation of ethyl cyanocinnamate, 3Aa, The kinetic plot clearly shows that the activity of SIM-2(C 12 ) is much higher than that of its parent SIM-1 or that of dodecylamine and is not the result of the simple addition of the activities of its two components. Furthermore, in these conditions, full conversion was obtained after 12 h for all of the catalysts. Stability is a critical issue for the catalytic application of MOFs. Nevertheless, ZIF-type materials are among the most stable. [47] To prove the stability of our new catalyst, a leaching test was performed. After 15 min, the catalyst was removed by filtration and the evolution of the product 3Aa was followed with the same reaction conditions maintained (Figure 2,a). No further reaction took place after the removal of the catalyst, which demonstrates the absence of leaching of active sites. After washing with ethanol and drying, SIM-2(C 12 ) was introduced in a second catalytic run under the same conditions without significant activity loss, which illustrates that it is possible to reuse the catalyst. In addition, 1 H NMR spectroscopy and PXRD analysis of SIM-2(C 12 ) were performed after catalysis and showed that the dodecylimino groups were retained by the material and that its crystallinity was preserved (see the Supporting Information). However, for real practical applications, further recycling tests would be performed.

Although the mechanism of the catalytic reaction remains unknown, we speculate that the active centers are located at the external surfaces of ZIF-8, SIM-1, and SIM-2(C 12 ) crystallites. Indeed, recent studies performed on ZIF-8 showed that crystalline defects (i.e., monocoordinated imidazoles or imidazolates) or zinc hydroxide surface species, which are Brønsted bases, are involved in acid-base catalysis. [48,49] This hypothesis, however, cannot fully explain the differences observed between SIM-1 and SIM-2(C 12 ) in terms of catalytic activity. The enhanced kinetics can also not be attributed only to the introduction of imino basic sites. We suggest instead that the hydrophobic chains at the surface of the solid, in the vicinity of the catalytic centers, hinder or limit water adsorption, making the centers free to adsorb the substrates.

To demonstrate the synthetic utility of the present hydrophobic SIM-2(C 12 ) catalyst, a rapid synthesis of ethyl 2-cyano-3methyl-5-phenylpent-2-enoate (4) was examined (Scheme 3). Compound 4 is a precursor of citronitrile, a citrus-like odorous compound that is widely used in the cosmetic and fragrance industries. SIM-2(C 12 ) efficiently catalyzes the reaction between benzylacetone and ethyl cyanoacetate to give the Knoevenagel adduct 4 with 95 % yield after three hours at 50 8C.

In summary, we have reported herein the upgrading of a metal-organic framework catalyst by post-synthetic modification. The new catalyst generated showed a tenfold activity increase compared to its parent compounds or other MOFs. We attribute the significant acceleration of the reaction to the modification of the environment of the catalytic sites through the creation of a hydrophobic shell surrounding the structure, as found in enzymes. In principle, the same hydrophobization procedure can be applied to the isostructural ZIF-90. [33] We believe that the development of post-synthetic functionalization routes will make it possible to achieve tailor-made materials that will bridge the gap between MOFs and enzymes.

Experimental Section General

All reactions were carried out in anhydrous solvents and catalytic runs were performed under argon atmosphere. The initial SIM-1 was synthesized according to our previously-published method. [44] All other reagents were commercially available and were used without further purification. NMR spectra were recorded on a Bruker 250 MHz spectrometer. The XRD measurements on the materials were carried out by powder X-ray diffraction (PXRD) using a Bruker D5005 diffractometer equipped with a secondary graphite monochromator and a scintillation counter. Water sorption was carried out on a Belsorp Max apparatus. Catalytic products were identified compared to commercial standards and activities were determined by GC analysis using an HP-5 capillary column.

Typical post-synthetic functionalization procedure

In a typical synthesis, SIM-1 powder (100 mg, ca. 0.7 mmol aldehyde) was suspended in anhydrous methanol (10 mL) and dodecylamine (390 mg, 2.1 mmol) was added. The resultant suspension was stirred at room temperature for 24 h. The solid obtained was washed with dry ethanol (3 10 mL) and centrifuged three times to afford the SIM-2 material as an off-white powder. SIM-2(C 12 ) was obtained with 22 % functionalization yield according to NMR.

Typical catalytic run

SIM-2(C 12 ) (7 mg) was introduced into a 10 mL glass vial and degassed. Benzaldehyde (20 mmol; 2 mL) and ethyl cyanoacetate (20 mmol; 2.1 mL) were then added under argon and the vial was sealed. The resultant suspension was allowed to react at 323 K for 30 min. After quenching by cooling to 0 8C, the suspension was diluted and the product dissolved with Et 2 O and centrifuged to remove the solid catalyst. The supernatant was analyzed by gas chromatography and the yield determined by integration of signals using dodecane as an internal standard.

Scheme 1 .

 1 Scheme 1. Post-synthetic modification of SIM-1 with dodecylamine.

Figure 1 .

 1 Figure 1. Water adsorption isotherms of SIM-1 (squares) and SIM-2(C 12 ) (circles) at 303 K. Closed and open symbols correspond to adsorption and desorption, respectively.

Scheme 2 .

 2 Scheme 2. Knoevenagel condensation catalyzed by SIM materials.

  [a] Reaction conditions (unless otherwise stated): 1 (20 mmol), 2, (20 mmol), catalyst (0.02 mmol), stirred in a 10 mL glass vial under argon at 323 K; [b] determined by GC using dodecane as internal standard; [c] turnover frequencies determined after 30 min and expressed in mol 3 mol catalyst h À1 ; [d] reaction performed at 298 K using 1 mol % catalyst. catalyzed by 0.1 mol % of SIM-1, SIM-2(C 12 ), or dodecylamine at 323 K.

Figure 2 .

 2 Figure 2. Time-dependent formation of 3Aa with 0.1 mol % of SIM-1 (&), SIM-2(C 12 ) (no filtration: *; filtration after 15 min.: *) or dodecylamine (~) as catalysts at 323 K.

Table 1 .

 1 Catalyst screening for Knoevenagel condensation.[a]

	Catalyst	R 1	R 2	t [h]	Yield [%] [b]	TOF init [h À1 ] [c]
	none	H	COOEt	4	0	n.d.
	ZIF-8	H	COOEt	0.5	40	800
	SIM-1	H	COOEt	0.5	6	120
	SIM-1	H	COOEt	4 [d]	3	n.d.
	SIM-2(C 12 )	H	COOEt	0.5	70	1400
	SIM-2(C 12 )	H	COOEt	4 [d]	80	n.d.
	dodecylamine	H	COOEt	0.5	15	300
	dodecylamine	H	COOEt	4 [d]	65	n.d.
	SIM-2(C 12 )	H	CN	0.5	96	1920
	SIM-2(C 12 )C H 3	COOEt	4	45	n.d.
	SIM-2(C 12 )C H 3	CN	4	51	n.d.

  SIM-2(C 12 ):1 H NMR (250 MHz,[ D 6 ]DMSO, 294 K): d = 0.77 (m, 0.93 H, CH 3 ), 1.16 (m, 5.5 H, CH 2 ), 1.61-1.65 (m, 0.57 H, CH 2 ), 2.51 (s, 3H, CH 3 ), 3.57-3.62 (m, 0.62 H, CH 2 ), 8.09 (s, 0.28 H, CHN), 8.90 (s, 0.27 H, CH), 9.20 (s, 1 H, CH), 9.85 ppm (s, 1 H, CHO). NMR spectra, N 2 adsorption isotherms and PXRD patterns of SIM-1 and SIM-2(C 12 ) are available in the Supporting Information.

[a] Dr. J. Canivet, Dr. S.Aguado, C. Daniel, Dr. D. Farrusseng UniversitØ Lyon 1, IRCELYON, Institut de Recherches sur la Catalyse et l'Environnement de Lyon, UMR CNRS 5256 Avenue Albert Einstein 2, 69626 Villeurbanne (France) Fax: (+ 33) 4-72-44-54-36 E-mail: david.farrusseng@ircelyon.univ-lyon1.fr

Acknowledgements

The authors are very grateful to the IRCELYON scientific services.

Keywords: aldol reaction • heterogeneous catalysis • hydrophobic effect • metal-organic frameworks • postsynthetic modification