Tuning of Fe and V sites in microporous and mesoporous materials for O2 and N2O induced selective oxidation reactions

Tuning of Fe and V sites in microporous and mesoporous materials

for O

and N

O induced selective oxidation reactions

Tuning of Fe and V sites in microporous and mesoporous materials

for O

and N

O induced selective oxidation reactions

Proefschrift

Ter verkrijging van de graad van doctor

aan de Technische Universiteit Delft,

op gezag van de Rector Magnificus prof. dr. ir. J.T. Fokkema,

voorzitter van het College voor Promoties,

in het openbaar te verdedigen op maandag 10 november 2008 om 10.00 uur

door

Wei WEI

scheikundig ingenieur

Geboren te Nan Tong, Jiang Su, P.R. China

Prof. dr. J.A. Moulijn

Copromotor:

Dr. G. Mul

Samenstelling promotiecommissie:

Rector Magnificus voorzitter

Prof. dr. J. A. Moulijn Technische Universiteit Delft, promotor Dr. G. Mul Technische Universiteit Delft, copromotor Prof. dr. F. Kapteijn Technische Universiteit Delft

Prof. dr. ir. B.M. Weckhuysen Universiteit Utrecht

Prof. dr. J.A. Lercher Technische Universiteit München, Germany Prof. dr. J.C. Jansen U-Stellenbosch, RSA

Dr. R.J. Berger Anaproc

Reservelid

Prof. dr. ir. H. van Bekkum Technische Universiteit Delft

NRSC-Catalysis

The research reported in this doctoral dissertation was carried out at the Catalysis Engineering group, DelftChemTech, Faculty of Applied Sciences, Delft University of Technology (Julianalaan 136, 2628 BL, Delft, The Netherlands) with financial support by NRSC-Catalysis (National Research School Combination Catalysis, The Netherlands).

ISBN 978-90-8559-449-9 Copyright © 2008 by Wei WEI

All rights reserved. No part of the material protected by this copyright notice may be reproduced or utilized in any form by any means, electronic or mechanical including photocopying, recording or by any information storage and retrieval system, without written permission from the author.

To my wife Zhiyong

and my sons, Shuya and Jiale

To my parents

1. Introduction... 1

2. Effect of steaming of iron containing AlPO

-5

on the structure and activity in N

O decomposition ... 25

3. Performance of Fe in microporous and mesoporous materials

in N

O induced selective oxidation of propane ... 57

4.

Characterization and performance of vanadia

dispersed in TUD-1 in oxidation reactions ... 81

5. Kinetic study of propane oxidative dehydrogenation

over V-TUD-1 catalyst ... 105

6. Structure and performance of vanadia in (Ti-, Zr-)TUD-1

in propane ODH ... 139

7. Concluding remarks ... 155

Summary ... 163

Samenvatting ... 165

Acknowledgements ... 169

List of publications ... 171

Curriculum vitae ... 175

1

Introduction

Propene is one of the most important feedstocks for the chemical industry. It is mostly produced by naphtha steam cracking, or a refinery process, such as gasoline-making fluidized catalytic cracking (FCC). Propane dehydrogenation is also applied for propene production, but it involves an endothermic reaction at high temperature and the regeneration of the catalysts, as drawbacks. The increasing propene demand requires more dedicated on-purpose technology to produce propene. Propane oxidative dehydrogenation (ODH) provides a new promising route. Recently, Vanadium and Iron based catalysts received a lot of attention for their good catalytic performance, respectively, in reactions involving O2 and N2O as the oxidant. It was shown that the

activity and selectivity of these catalysts largely depend on the nature of active sites, the structure of the support material and reaction conditions. The various technologies for ODH are evaluated in this chapter.

1. Background

As a by-product of ethene production, propene is one of the most important petrochemicals. It was first used in the production of isopropyl alcohol through indirect or direct hydration processes on an industrial scale. After that, many new important areas of application were initiated for propene by the chemical industry. Now, it is the building block for a large number of HPV (High Production Volume) organic chemicals and polymers.

Figure 1 shows the propene consumption in the main applications in Western Europe in 2004. It is obvious that derivatives of propene, such as polypropene, acronitrile, propene oxide, and cumene, have clearly overtaken the classic product isopropyl alcohol in importance [1;2]. Propene demand was approximately 39 million tons in 1995 and it is forecasted to increase to almost 90 million tones by 2015 [3].

The versatility of propene stems from the unique chemical structure. Propene contains both a carbon-carbon double bond and an allylic methyl group (a methyl group adjacent to a double bond) giving chemists, catalyst designers, and engineers two "handles" for carrying out chemical transformations. Thus, the production of propene is an important area in chemical industry.

Figure 1. Western European propene consumption by derivative in 2004 [2].

2. Propene production

Although propene is an important feedstock for the chemical industry, its production mostly relies on naphtha (mixing of hydrocarbons) or ethane steam cracking to ethene, propene, etc. In 2004, steam cracking accounted for 71.6% of the propene

isopropanol 3% Oxo-Alcohols 6% Propene oxide 11% Cumene 7% Acrylonitrile 6% Polypropene 58% Others 9%

Introduction production in Western Europe [2]. As a consequence, the availability of propene is determined largely by the demand for the main product, ethene, although factors such as feedstock and operating conditions have a significant influence on propene yield. In recent years, the demand for propene has increased much more rapidly than for ethene and this situation is expected to continue. In the U.S., the ratio of propene to ethene demand grew from 0.48 in 1980 to 0.61 in 2003. Regions such as the Middle East, Western Europe, and East Asia showed similar patterns over the same period [4]. Furthermore, the cracking furnace is very capital intensive and it is expected that many have to be replaced in the coming decades because of their age.

Propene can also be produced in refinery processes as by-product. For example, gasoline-making fluidized catalytic cracking (FCC) contributes to nearly 30% of the production of propene [4]. But the growth of refinery sources, e.g. FCC, still do not keep pace with this growing propene demand. Other refinery conversion processes, e.g., visbreaking, and coking also contribute to propene production, but the yields are quite low.

To satisfy the growing demand of propene in the modern petrochemical industry, the on-purpose process of propene production, such as propane dehydrogenation, has been developed commercially and accounts for 2-3% of the global production [4].

2 3 2 3 2 3 CH CH CH CH CH H CH − − → = − + (1)

Besides catalytic dehydrogenation of propane, new concepts and strategies, such as oxidative dehydrogenation of propane to produce propene, have emerged in recent years. O H CH CH CH O CH CH CH3 2 3 2 2 2 3 2 2 2 − − + → = − + (2)

Summarizing, all the processes for the production of propene followed by industry and research institutes can be classified into four groups, which will be sequentially dealt with in the following paragraphs.

• Production as by-product of ethene production • Production as by-product of refinery processes • Propane dehydrogenation

1.1 Production as by-product of ethene production

Figure 2 shows a simplified flow scheme of a steam cracker. A mixture of hydrocarbons (naphtha) and steam is passed through tubes placed inside a furnace. The furnace consists of a convection section, in which the hydrocarbon feed and steam are preheated, and a radiation section in which the reactions take place. The hydrocarbons undergo pyrolysis and subsequently the products are rapidly quenched to prevent further reaction to preserve the product distribution. The residence time is less than 1 second [5].

1120K 870K Naphtha (360K)

Dilution steam Convection section 700K Stack High pressure steam Radiation section

Figure 2. Simplified flow scheme of a steam cracker [5].

Cracked gas goes through further processing, such as compression, acid gas removal and gas drying, before the C3 fraction is separated from the depropanizer. The C3 fraction is fed to a selective hydrogenation unit to remove propadiene (H2C=C=CH2) and propyne (CH3-C!CH). This hydrogenation can be performed in the gas or liquid phase over palladium catalysts. The amount of hydrogen added is calculated so that, on the one hand, complete conversion of C3H4 to C3H6 is achieved, and on the other hand, the smallest possible amount of propene is hydrogenated to propane [6].

A typical product distribution shows propene yield is around 1.5% when ethane is used as feedstock [5]. In Europe, naphtha is the primary feedstock for the production of ethene. Due to the high concentration of >C2 hydrocarbons, much higher propene yields are achieved. A propene yield of around 16% is obtained when naphtha is used as feedstock [5]. Obviously, the type of feedstock has a significant influence on the product distribution and propene yields.

Introduction 1.2 Production as by-product of refinery processes

In refineries, propene is also produced from cracking processes. It originates from either catalytic cracking or thermal cracking. These processes are very different from the steam cracking process previously mentioned due to the completely different feed stocks and different production objectives.

The most important refinery process for propene production is the fluid-catalytic cracking (FCC) process, in which catalyst particles are entrained and continuously circulated between the cracker and the regenation unit. This process converts heavy gas oil preferentially into gasoline (>C4) and light hydrocarbons (C1-C4). Conventional FCC units produce about 4 wt-% propene [7]. Adding ZSM-5 additives to the catalysts is a

proven technology to enhance the production of lower olefins at the expense of gasoline. Thermal cracking, which in refineries contributes to propene production, is

employed in coking and visbreaker units. In coking units (delayed coking and fluid coking), residues from the atmospheric and vacuum distillation of the crude oil undergo relatively severe cracking and are thereby converted into gas oil, coke, gasoline, and smaller amounts of cracked gas (6–12 wt-% of C4 and lighter). The cracked gas from the coking unit normally contains 10–15 mol-% C3, mostly propane. In visbreaker units, vacuum residues are subjected to mild cracking, with the objective of reducing the viscosity of the residue oil. Smaller amounts of gas oil, gasoline, and cracked gas (2 – 3 wt-% of C4 and lighter) are formed here. In the case of thermal cracking, visbreaking, and coking processes, the propene yield is lower and the quality often unacceptable.

To separate the C3 fraction, downstream processing of refinery gases generally occurs as follows: first, the light components are separated at ca. 15 bar in a deethanizer, whose upper part is operated as an absorber with gasoline and oil as scrubbing agents. The C5 and heavier hydrocarbons are removed from the bottom product in a connected debutanizer. The C3– C4 fraction obtained as overhead product is normally desulfurized on molecular sieves and dried, before being split into individual fractions in a final depropanizer.

1.3 Propane dehydrogenation processes

Propane dehydrogenation involves catalytic endothermic equilibrium conversion of propane to propene. This technology has its roots in the dehydrogenation of isobutane into isobutene, which is used as a feedstock for methyl tert-butyl ether [4]. From a

thermodynamic viewpoint, the reaction temperature should be high for sufficient conversion of propane, even reaching 700 o_C.

C₃H₈→C₃H₆+H₂ ∆H₀ =125.1kJ/mol (3)

However, at such a high temperature, secondary reactions leading to cracking or coke formation become appreciable. Coke formation wastes feedstock and deactivates the dehydrogenation catalyst. So, it is necessary to regenerate catalysts by combustion of the deposited coke. Moreover, the thermodynamic equilibrium limits the conversion per pass so that a substantial recycle stream is required.

A low partial pressure of alkanes favors the forward reaction, because two molecules are formed for every molecule converted. A process under vacuum would be desirable. In practice, it is more convenient to apply dilution with steam at atmospheric pressure.

Globally seven commercial plants are in operation, of which six use “UOP’s Oleflex continuous moving-bed process” and the other one uses “ABB Lummus’ Catofin cyclic multiple-reactor system”. Other processes on the market include Phillip’s STAR process (Uhde bought this process from Phillips Petroleum in 1999), as well as the FBD-4 process from Snamprogetti Yarsintez [4].Two principal classes of catalysts are applied in these processes: platinum catalysts and chromium oxide catalysts usually supported on alumina.

These processes differ from one another mainly in terms of the reactor technology, the modes of operation, the types of dehydrogenation catalyst, and the methods of catalyst regeneration, which are further explained in the following paragraphs.

1.3.1 Oleflex process

The Oleflex process is an adiabatic process in which the heat of reaction is supplied by reheating the process stream between the different reaction stages. It uses radial flow moving bed reactors in series with inter-stage heaters [5]. The process is operated at a slight positive pressure, and a proprietary platinum catalyst is used. Fresh propane feed is mixed with recycled hydrogen and unconverted propane and admitted to a train of three radial-flow moving-bed catalytic reaction vessels. The process is continuous, and overall selectivities for propene of 89 – 91 % are claimed [6].

Introduction To propene recovery/ purification Fuel Propane a b b b c

Figure 3. UOP Oleflex process: a) Reheat furnace; b) Moving-bed reactors; c) Continuous

catalyst separation system [6].

The Oleflex process was developed from the Pacol process (UOP), which is used to dehydrogenate C10 – C14 paraffins to olefin feedstocks for the production of intermediates for synthetic detergents. So they are operated in the same way. The dehydrogenation catalyst circulates through the reactor section before passing to a separate regeneration vessel where coke is removed from the surface of the catalyst by combustion in air (UOP Continuous Catalyst Regeneration technology). The regenerated catalyst is returned to the first of the dehydrogenation reactors. Propene is recovered by conventional deethanizer – depropanizer splitting [6].

Advantages claimed by the Oleflex process include continuous operation; a uniform, time-invariant catalyst activity profile; and isolation of the oxidative catalyst regeneration phase from the dehydrogenation reactor.

1.3.2 Catofin Process

The Catofin process was widely used in the dehydrogenation of butane to iso-butene, an important primary product in the production of methyl tertiary buty ether (MTBE). Due to the banning of MTBE from fuels for environmental reasons, the application of propane dehydrogenation for the Catofin process was developed. The process utilizes adiabatic, fixed-bed multiple reactors (see Fig. 4), and operates under a slight vacuum at 550 – 750°C . The dehydrogenation catalyst consists of activated alumina pellets impregnated with 18 – 20 wt % chromium [6]. A key feature of the Catofin process is the principle of storing the needed reaction heat in the catalyst bed [5]. The process is cyclic and includes a reaction period, discharge of the reactor, and regeneration of the catalyst in situ. Multiple reactors, usually three to eight are used in

parallel to achieve continuous plant throughput, with the reactor containing the coked catalyst being taken off stream for the regeneration step. The mixed fresh and recycled propane stream is preheated to 600 – 700 °C and fed to the reactor at ca. 30 kPa (one-third atmospheric pressure) [6]. Reactors are alternately on stream for reaction and off stream for regeneration. The combustion of the deposited coke heats the catalyst bed, and this energy is released to the endothermic dehydrogenation reaction when that reactor is in the on-stream period [5]. Due to the formation of coke and reduction of Cr6+ _{to Cr}3+_{, the} catalyst activity changes. During regeneration, coke is burned and the chromium is partially re-oxidized [8]. The overall selectivity of propane to propene is reported to be about 87 %. Propene is recovered by a conventional propane –propene splitter. Concluding, the Catofin process extends catalyst life and saves energy by utilization of the heat generated in the exothermic regeneration step to assist dehydrogenation [6].

Propane feed To Propene recovery/ purification Exhaust Fuel Air Steam b a c d e

Figure 4. Catofin process a) Charge heater; b) Air heater; c) Reactor on purge; d) Reactor on

stream; e) Reactor on regeneration [6].

1.3.3 Phillips STAR process

The Phillips Steam Active Reforming (STAR) process for the dehydrogenation of paraffins differs significantly from other dehydrogenation processes. Here, steam is used as a diluent to maintain an overall positive pressure in the process reactor and simultaneously reduce the partial pressure of the hydrocarbons and hydrogen present. Thus, the equilibrium is shifted toward increased conversion [6]. The preheated feedstream containing steam is admitted to a chain of multiple fixed-bed reactors. Each reactor consists of multiple catalyst-packed tubes in a furnace firebox, which is similar to the steam reformer used for the production of synthesis gas [5]. Reactor operation is cyclic (i.e., one reactor is taken off-line sequentially for catalyst regeneration), whereas

Introduction dehydrogenation is maintained continuously. Catalyst deactivation occurs due to coke deposition, and after about 7 h on-line, an off-line catalyst regeneration by combustion for ca. 1 h is required [6]. Advantages claimed for the STAR process include 80 % yield of propene on propane; isothermal operation that ensures sufficient heat input to promote dehydrogenation; and the steam diluent that reduces the hydrocarbon pressures while maintaining a more practical process pressure. Due to the presence of steam, the catalyst should not deactivate by water adsorption. The applied catalysts contain Pt as active phase on a support of zinc aluminate or magnesium aluminate with calcium aluminate as binder [8].

1.3.4 FBD-4 Process

The FBD-4 (Fluidized-bed reactor and regenerator) is licensed by Snamprogetti Yarsintez. It is similar to older FCC units with continuous catalyst circulation between a fluidized-bed reactor and regenerator [5]. In this process the pure alkane feed bubbles through the staged fluidized bed reactor at 1.1 - 1.5 bar. Since coke is formed on the alumina supported chromium catalyst regeneration of the catalyst is required. Therefore part of the catalyst is continuously moved to a regenerator and back. Typical operating temperatures are 550 – 600 °C. This process is commercially used for the dehydrogenation of isobutane and isopentane [8].

1.3.5 Linde BASF

The Linde process (Linde, BASF) is performed in three fixed bed reactors. Two operate under dehydrogenation conditions while the third is regenerated by combusting the coke on the catalyst with a steam / air mixture. The catalyst in this process is again CrOx /Al2O3. The process is operated at a temperature of 590 °C and is kept at nearly

isothermal conditions, to minimize thermal cracking and coke formation. A pilot plant study for the dehydrogenation of propane has been performed [8].

1.4 Propane oxidative dehydrogenation routes

Although the catalytic dehydrogenation processes have been applied in industry for decades, they still suffer from several limitations:

• Thermodynamic restrictions on conversion and selectivity. Strong endothermic

main reaction and the necessity to supply the heat at high temperature

• Rapid formation of coke and thus continuous necessity for catalyst regeneration

as well as the loss of hydrocarbons in the side reaction

• Difficulty in separating the alkene from the alkane and by-products • Irreversible catalyst deactivation owing to the severe reaction conditions

In all, due to the severity of the reaction conditions and complexity of the process, propane dehydrogenation is very costly. The possibility of developing new lower-cost processes has generated interest in the transformation of light alkanes to valuable alkenes by means of oxidation.

Oxidative dehydrogenation (ODH) is a promising technology to convert light alkanes by exothermic processes.

mol kJ H O H H C O H C_{3 8}+0.5 ₂→ _{3 6}+ ₂ ∆ ₀ =−116.7 / (4)

The deactivation of the catalysts and necessity of catalyst regeneration can be avoided due to the presence of oxygen. The basic concept of oxidative dehydrogenation is that the reaction takes place in the presence of oxygen, which gives rise to the exothermic oxidation to overcome the thermodynamic limitations of a reversible endothermic dehydrogenation. Various oxidants have been proposed in the literature, while oxygen (O2) and nitrous oxide (N2O) are most interesting for the application in the heterogeneous catalysis field.

C3H8+N2O→C3H6+H2O+N2 ∆H0=−198.8kJ/mol ₍₅₎

Vanadium- and iron- based catalysts are mainly investigated to catalyze this reaction, using O2 or N2O, respectively. Currently there is not yet a commercial application on an industrial scale for propane ODH yet.

1.4.1 Propane oxidative dehydrogenation with O2

Oxygen is a conventional and most used oxidant in oxidation reactions. In the presence of oxygen, propane dehydrogenation becomes an exothermal reaction. Compared to the direct propane dehydrogenation, in which huge amounts of heat input are required, ODH could occur at relatively mild conditions. Moreover, oxygen reacts with deposited carbon species over the surface of the catalyst; therefore there is no need to regenerate catalysts.

Introduction On the other hand, the presence of oxygen also induces the formation of carbon oxides, and affects the selectivity to propene negatively. Furthermore, the heat released from the reaction should be removed efficiently. Heat is released not only by ODH, but also by the deep oxidation of intermediate products. So it is important to predict and avoid runaways due to the possible occurrence of hot spots. The flammability of the reaction mixture should also be carefully considered to avoid an explosion.

Vanadium-based catalysts exhibit interesting catalytic properties for partial oxidation of alkanes in the presence of O2. MgO-supported vanadium was reported as a selective catalyst in the ODH of propane and n-butane [9;10], while γ-A12O3 supported

vanadium catalysts were found to present a good selectivity to olefin products for ethane ODH, but a poor selectivity in the ODH of n-butane [11]. The major challenge is to further improve the selectivity, both in regard to reducing carbon oxide formation, and minimizing condensable oxygenated hydrocarbons. Mamedov and Cortés Corberán suggest that a certain extent of reduction is necessary for vanadium to show selectivity in ODH of alkanes. In addition, a good selectivity can be achieved by controlling reaction conditions, such as reaction temperature and alkane to oxygen ratio [12].

Compared to basic MgO, a weaker interaction between V and an acidic A12O3

support is expected leading to less dispersed vanadium species on the surface, which in turn favours the formation of V2O5 crystallites [13;14]. Several investigations have been

performed on the structure of vanadia on alumina-supported catalysts [15-19]. These studies show, in general, that at low loading vanadium form highly dispersed amorphous phases, whose structure changes from isolated tetrahedral vanadium, to polyvanadate species at medium loading. Crystalline V2O5 appears at high loading in addition to the

amorphous vanadia phases.

The binding strength of lattice oxygen has been postulated as another parameter that governs activity and selectivity of these catalysts [20]. VO4 tetrahedra formed at low

vanadium contents are proposed to induce high selectivity to alkenes because three oxygen atoms in these species are bridged between V and the ions of the support oxide [19;21].

Other authors considered that -V2O7- units favour alkane ODH and suggested that

the bridging oxygen between two vanadium ions (V-O-V) plays an important role in propane ODH. The removal of' this oxygen results in a local structural change from the -V2O7- unit to two edge-sharing square-based -VO3- units where vanadium is present as

V4+ _{cations. These stabilized V}4+ _{ions were postulated as responsible for the}

accumulating showing that the direct involvement of the V-O-support bond in the selective oxidation is more likely.

Recently, SiO2-supported vanadia catalysts have been reported to be active and selective for the oxidative dehydrogenation (ODH) of ethane [22], propane [23] and butane [24]. An important dependence between vanadium content and catalytic performance has been observed, in a way that low V-loading catalysts present the highest selectivity to olefins. Moreover, mesoporous SiO2-supported vanadia received much attention, because it shows remarkable activity and selectivity in alkane ODH.

1.4.2 Propane selective oxidation with N2O

2.4.2.1 N2O abatement

N2O is well known to be a greenhouse gas and to contribute to the catalytic destruction of ozone in the stratosphere. The atmospheric concentrations of N2O have been relatively constant for many centuries (~270 ppbv). But during the second part of the 20th century, the N2O atmospheric concentration increased rapidly to around 310 ppbv, which means a 9% increase from pre-industrial levels (285 ppbv) at an annual growth rate of 0.2-0.3% [25].

N2O has 310 and 21 times the global warming potential (GWP) of CO2 and CH4, respectively.[26;27] The Kyoto protocol (December, 1997) set legal binding targets for reducing emissions of 6 greenhouse gases (CO2, CH4, N2O, HFC, PFC, SF6) to be realized in the period 2008-2012 [28]. The European Community committed itself to reducing its emissions of greenhouse gases by 8% during the period 2008-2012 in comparison to their levels in 1990. N2O emissions account for 10% of the total greenhouse gas emissions in the EU, so its reduction plays an important role in reaching the global 8% target [25].

N2O is produced by both natural (65%) and anthropogenic (35%) sources. Biological processes in soils and oceans are the primary natural sources of N2O. Agriculture is a major anthropogenic source, but control of these emissions is difficult due to its very diffuse nature [25]. N2O emissions that can be reduced in the short term are associated with chemical production and the energy industry. Nowadays the major sources of N2O in the chemical industry are the production of nitric acid and adipic acid.

Different (catalytic and thermal) abatement technologies have been successfully developed for adipic acid plants, due the high N2O concentration in the tail-gas

Introduction 40 vol.%) [25;29]. Due to the exothermicity of the decomposition reaction, a large increase in the temperature occurs within the catalyst bed. For instance, the decomposition of 35 vol.% N2O in air leads to an adiabatic temperature rise of the gas of 650 oC. In this temperature window, a large number of catalysts exhibit considerable activity. So, in this case the activity of the catalyst is not a critical factor for the effectiveness of the technology [25].

Only in the particular case of adipic acid production, technologies are commercially available, but their application (extrapolation or adaptation) to other N2O sources, e.g. (single) nitric acid plants and stationary combustion processes, is not straight forward, due to the presence of diluted N2O streams (in the 0.05-0.5 vol% range) at a relative low temperature (typically < 800 K), and the presence of catalyst inhibitors (O2, H2O, SO2, and NOx) [25;29]. Still, novel systems have been developed for abatement of

N2O from nitric acid plants, which will not be further discussed here.

1.4.2.1 Selective oxidation by N2O

N2O is an interesting precursor of active oxygen species. According to the pioneering work of Panov et al., so-called α-type oxygen species can be generated by

N2O decomposition over ZSM-5-type Fe containing catalysts. The feature of this α-type oxygen species is its high reactivity and selectivity for selective oxidation reactions. At temperatures below 300°C, N2O decomposes on these catalysts evolving molecular nitrogen, while the oxygen remains on the catalyst as the α-type oxygen species. Above

300°C, oxygen species start to desorb in the form of molecular oxygen. N2O proved to be a good selective oxidant for the synthesis of phenol from benzene or methanol from methane. Selectivities to phenol up to 99% with a productivity of around 3 mmol phenol per gram of catalyst per hour have been reported using high-silica ZSM-5 zeolites (SiO2/Al2O3 about 100 with Fe below 0.1%) [30-32]. This has been demonstrated for the first time in the pilot adipic acid plant, erected by Solutia and the Boreskov Institute of Catalysis. Adipic acid is an intermediate manufactured by the oxidation of cyclohexanone and/or cyclohexanol by nitric acid. A large outlet stream of N2O with concentrations up to about 35% is available from this process. Since N2O has effect on the global warming and ozone depletion, a catalytic abatement unit is required to decompose N2O in the current commercial process. In the pilot adipic acid plant mentioned above, the waste N2O stream is used to synthesize phenol from benzene, and then phenol is hydrogenated to cyclohexanone that is the raw material for adipic acid. In this way, the waste N2O is

utilized as a valuable oxidant to oxidize benzene to phenol and environmental impact of N2O is reduced. Because "-type oxygen is important in this novel technology, it is called the AlphOxTM process. This process is schematically given in Figure 5.

OH COOH COOH N 2O + OH O N₂O H2 HNO₃ +

Figure 5. The AlphOxTM process for N2O reuse as selective oxidant in the benzene-to-phenol

process.

N2O can also induce the catalytic oxidative dehydrogenation of light alkanes as a potential route to the corresponding alkenes.

Using N2O in this selective oxidation reaction faces the same challenges as using oxygen as oxidant, i.e. to control the formation of carbon oxides and improve the selectivity to propene. It has recently been reported that steam-activated Fe-zeolites show a superior performance in the N2O-mediated propane oxidative dehydrogenation. Initial propene yields of 24% at 500-525 oC were achieved, comparable with the highest reported value over Vanadium-based catalysts with O2 [33]. The remarkable performance of Fe-zeolites has been related to the specificity of N2O as a monooxygen donor and the capability of iron in the zeolite to coordinate reactive atomic oxygen species able to efficiently dehydrogenate propane [33;34]. It leads to a new route to functionalize propane, with utilization of N2O, an environmentally harmful gas. But the major drawback of Fe-zeolites in propane ODH with N2O is the deactivation by coke. It causes a rapid decrease of the propene yield in minutes, although the original catalytic activity was completely recovered after regeneration in pure oxygen [34].

Another important application of N2O as oxidant is in the oxidation of light alkanes (C1-C2) to oxygenates. Bulk FePO4 shows a remarkable selectivity to alcohols (C1-C2) in oxidation of methane and ethane when N2O or mixtures of H2 and O2 are used as oxidants [35]. The total selectivity to CH3OH and HCHO is 20% when N2O is used as oxidant, but the activity is rather low, less than 5% conversion under industrially relevant conditions [35]. The results suggest that the tetrahedrally coordinated iron sites isolated from each other by phosphate groups are the active sites for the selective oxidation of

Introduction methane to methanol by N2O or an H2 and O2 gas mixture. It is proposed that H2 and O2 are activated on the active iron site through the redox reaction between Fe(III) and Fe(II), producing an adsorbed peroxide species responsible for the selective oxidation of methane to methanol [35]. The same iron site is effective for the conversion of methane to methanol by N2O. A monatomic oxygen species Fe-O is likely to be responsible for the selectivity in the oxidation of CH4 to CH3OH when N2O is used as oxidant.

3. Technologies evaluation and research

Table 1 summarizes the different technologies to produce propene with their advantages and drawbacks at the current stage.

Over the years, steam cracking plants have developed in an evolutionary way and plant capacities have increased due to technical improvements, for example, the application of a novel structured packing to increase the tower capacity in a steam cracking plant [5]. In consequence, the production of propene also increased sharply as a co-product of ethene. But the most obvious disadvantage of the steam cracking process is its relatively low selectivity and limited flexibility to adjust product distribution. The co-production of propene from stream crackers is expected to decline as plants are optimized to produce more ethene [36].

Propene production from FCC is desired in order to balance the increasing gap between C3=/C2= demand and the C3=/C2= yield ratio from steam naphtha cracking [3]. Resid Fluid Catalytic Cracking (RFCC) is a new technique that aims to convert heavy oil fractions especially to lower olefins under more severe conditions than conventional FCC. Another potential alternative is to develop novel state-of-the-art catalysts for improved lower olefins production. A new family of catalysts, APEX, developed by Davison Catalysts, will allow refiners to take propene production to a new level. Using proprietary shape-selective zeolite and matrix technologies, APEX catalysts not only produce exceptional yields of propene in the range of 15-20 weight percent, but also demonstrate low coke make and bottoms cracking activity in the presence of contaminant metals. Propene yields of more than 20 weight percent are of particular interest to refiners who are considering revamping their FCC units to operate in a "petrochemicals mode” [37]. Additives and reactor configurations for re-cracking of naphtha olefins to propene and decreasing the undesired dry gas (H2, C1-C2) will also gain importance in the future [7].

Table 1. Summary of catalytic processes for the propene production. Production

/Process

Reactor Catalysts Distinguishing

Aspects

Disadvantages

1. PRODUCTION AS BY-PRODUCT OF OTHER PROCESS ( industral scale)

By-product of ethene production Tubular reactor in furnace None By-product of refinery process

Fluidized bed Zeolites

• Bulk production • Proven tech-nology and process for main products • No extra investment for propene production • Significantly influenced by the type of feedstock and operation condition • Low selectivity • Production determined

by the demand of main products

2. PROPANE DIRECT DEHYDROGENATION (industrial scale)

Oleflex(UOP) Adiabatic moving beds in series with intermediate heating Pt/Al2O3 • Inter-stage heating Catofin (ABB Lummus crest) Parallel adiabatic fixed beds with swing reactor

Cr/Al2O3 • Heat from coke

burning is stored and supply for the reaction STAR (Phillips Petroleum) Tubular reactors in furnace Pt-Sn/Zn-Al2O3 • Relatively low operation temperature FBD-4 (Snamprogetti)

Fluidized bed Cr/Al2O3 • Continuous

catalysts regeneration Linde (BASF) Fixed bed reactors Cr/Al2O3 • Keeping at

isothermal condition

• Difficulty to control side reactions such as thermal cracking and coke depo-sition

• Strong endothermic reaction, heat has to be supplied at high temp-erature

• Large recycle streams required

• Regeneration needed due to rapid catalysts deactivation

3. PROPANE OXIDATIVE DEHYDROGENATION (laboratory-scale)

Propane ODH with O2 as

oxidative agent

Packed bed reactor V/Al2O3

V/SiO2 V/MgO • Direct supply heat of reaction • Relatively mild conditions • Catalyst regeneration not needed • Difficulty of controlling the formation of carbon oxides

• Possibility of runaway • Flammability of the

Introduction

Propane ODH with N2O as

oxidative agent

Packed bed reactor Fe-ZSM-5 FePO4 • Direct supply heat of reaction • Relatively mild conditions • Environmentally harmful N2O utilized. • Difficulty of controlling the formation of carbon oxides

• Possibility of runaway • Catalyst deactivation,

regeneration needed

Commercial dehydrogenation processes are better than steam cracking in terms of higher selectivity to propene. But the conversions are necessarily lower to prevent excessive side reactions. Still, single–pass yields of 45 wt% can be achieved, which is considerably higher than the yields obtained in steam cracking, for example, 15 wt-% propene is typically produced in naphtha cracking [5]. However, as a result of the large recycle streams required, gas compression and purification account for nearly 85% of the total capital costs [5]. Understandably, here the largest savings could be made. Since the low selectivity at high conversion is mainly due to coke formation, monitoring and controlling catalysts deactivation behavior received much attention and research.

The so-called “operando” technology is promising to realize a real-time spectroscopic process control for the dehydrogenation of propane under realistic reaction conditions over Cr/Al2O3 catalyst, by collecting in-situ UV–Vis and Raman spectroscopic data simultaneously [38;39]. It is interesting to find that a small amount of coke improves its activity, most likely by an improved adsorption of propane on the catalyst. But when the amount of coke exceeds a certain value, the activity of the catalyst starts to drop. This turning point is visualized in UV–Vis spectroscopy by a sharp increase in the overall absorbance of the sample and this spectroscopic change can be used as a signal for starting a regeneration cycle. So far this is the most direct method to control a chemical process when coke formation is an important parameter. As a result, the control system can intervene faster to maintain the desired conversion levels and selectivities in the reactor [38]. Furthermore, this new spectroscopic approach can be used to controlling the phase composition of the VPO catalyst in the selective oxidation of butane to maleic anhydride, to prolong the catalyst life, as illustrated in [39]. Apparently, this novel “catalyst diagnostics” strategy has potential industrial impact, which will be proven in the pilot plant soon [39].

In the field of oxidative dehydrogenation, there is no commercialized process yet. Research was done to explore new catalysts to activate the alkane in the presence of N2O or O2. As mentioned in section 2.4.1, steam-activated Fe-ZSM-5 showed anomalous behaviour for decomposing N2O, while bulk FePO4 showed good selectivity in oxidation

of light alkanes to alcohols (C1-C2); it is worth to investigate the combination of the high activity of Fe sites in a zeolite matrix with the functionality of the phosphate groups in N2O decomposition. Thus, so-called aluminophosphate molecular sieves (AlPOs) with isomorphously substituted iron received much attention. This material has been first synthesized by Union Carbide workers [40] who named the materials FAPO-5, which is based on AlPO4-5, a one-dimensional system with a channel diameter of 7.3 Å. The framework structure of AlPO4-5 is based on a strict alternation of AlO4- and PO4+ units. It is a good host for transitional metals, such as Fe [41], Cr [42] and Co [43], by substituting aluminum in the framework. The catalytic behavior of FAPO-5 in N2O decomposition and propane oxidative dehydrogenation with N2O will be presented and discussed in this thesis.

Mesoporous material as supports attracted more and more interest in the last decade. According to IUPAC (International Union of Pure and Applied Chemistry) a mesoporous material is defined as a material with pore diameter between 2 and 50 nm. In 1992, ordered mesoporous silica was first discovered by scientists from Mobil [35;44]. Many silicate and aluminosilicate mesoporous materials were synthesized, such as MCM-41, MCM-48, HMS-1 and SBA-15. Compared to conventional microporous materials, mesoporous materials possess a larger pore size than zeolites. This improves the accessibility to the active sites of the relatively large molecules that are present in crude oil and in production of fine chemicals. At the same time, it facilitates the escape of products, preventing consecutive reactions to unwanted by-products.

In the laboratory of Applied Organic Chemistry and Catalysis, at the Delft University of Technology, Dr. Zhiping Shan succeeded in 1999 to prepare a new mesoporous material. The target was the synthesis of a three-dimensional silica framework with high surface area in a one-pot procedure. Most importantly, the synthesis procedure has to be unique and cost-effective compared to other templated mesoporous materials, such as MCM-41 or MCM-48 [45].

70nm

Figure 6. Computer image of a sponge-like mesoporous material, comprising two endless phases

Introduction In 2001, the first publication about this material appeared [46]. This new material was called TUD-1 (the acronym of Technische Universiteit Delft). In that paper, the authors reported the synthesis of TUD-1 through a new templating method using an inexpensive non-surfactant pathway in a one-pot procedure. The synthesis was mainly related to mixing a silica source (tetraethyl orthosilicate) with diluted triethanolamine to obtain a homogeneous mixture at the molecular level, followed by aging, drying and finally calcination. Hydrothermally stable sponge-like mesopore networks with high surface area (> 600 m2/g) were thus obtained. The porosity parameters (surface area, pore volume and pore size) were found to be tunable [45]. A computer simulation image of a mesoporous material is shown in Figure 6, while a series of more detailed pictures of the 3-D structure of TUD-1 as obtained by HR-TEM are shown in Figure 7.

Figure 7. 3D HRTEM images of a TUD1 particle [47]. The same particle is shown in angles

-180o_{. The size of the particle is 40 nm in height and the mesopores are about 2 nm in diameter.}

Following the success of TUD-1 synthesis, iron, as well as vanadium, titanium and zirconium were incorporated in the TUD-1 framework via hydrothermal treatment. They showed to be promising catalysts with good activity and selectivity in propane oxidative dehydrogenation, as will be discussed in this thesis.

4. Thesis outline

In view of the current state of the art in O2 and N2O mediated selective oxidation of propane, two transition metals, iron and vanadium, are dealt with in this thesis. Iron is incorporated into a microporous AlPO4-5 material and mesoporous TUD-1 material, while vanadium is incorporated into TUD-1, Zr-TUD-1 and Ti-TUD-1. The catalysts were tested in propane oxidation with N2O or O2. A schematic outline of this thesis is shown in Figure 8.

Fe

In chapter 1, a general introduction is given; the current processes for propene production and emerging technology are briefly discussed.

Chapter 2 will show the need of steam treatment for activating FAPO-5, an iron

containing microporous AlPO4-5 material, before its application in N2O decomposition. The iron content is directly related to its activity. The nature of the active site and the change of physicochemical properties due to the steam treatment will be discussed in this chapter.

Figure 8. Schematic content of this thesis.

In chapter 3, Iron was incorporated into microporous AlPO4-5 and into mesoporous TUD-1. These two kinds of catalysts possess different physical and chemical properties, consequently resulting in different behaviour in the N2O mediated oxidative dehydrogenation.

In chapter 4, V-TUD-1 is introduced as a new mesoporous catalyst for propane ODH with O2. The amount of vanadium in TUD-1 is tuned and the catalysts were characterized by means of XRD, HR-TEM, Raman, UV-Vis, and 29_{Si NMR. Propane} oxidative dehydrogenation (ODH) experiments were conducted in a six-flow set-up at different temperatures and space velocity, and will be discussed.

The kinetics of the reaction over the V-TUD-1 sample with 5 wt-% is extensively evaluated in chapter 5.

In chapter 6, the modification of the TUD-1 support by incorporation of Zr or Ti into the V-TUD-1 was investigated. The vanadia structure and catalytic performance in propane oxidative dehydrogenation are largely modified by the presence of Zr or Ti.

Concluding remarks and recommendations are addressed in chapter 7. V AlPO4 TUD-1 N2O decomposition (Zr,Ti)-TUD-1 C3H8 oxidative dehydrogenation with O2 C3H8 oxidative dehydrogenation with N2O

Introduction Reference List

[1] G.Centi, F.Cavani, and F.Trifiro , Selective Oxidation by Heterogeneous Catalysis , Kluwer Academic/Plenum Publishers, New York, 2001.

[2] http://www.petrochemistry.net (2005).

[3] X.Zhao and T.G.Roberie, Ind.Eng.Chem.Res. 38 (1999) 3847-3853. [4] A.H.Tullo, Chemical & Engineering news 81 (2003) 15.

[5] J.A.Moulijn., M.Makkee, and A.Van Diepen , Chemical Process Technology, Wiley, 2001.

[6] P.Eisele, P. and R.Killpack , Ullmann's Encyclopedia of Industrial Chemistry, Wiley-VCH, 2005.

[7] X.Dupain, Chapter 1, PhD Thesis (2006) , Delft University of Technology. [8] S.J.Tinnemans, Chaper 2, PhD Thesis (2006) , Utrecht University.

[9] H.H.Kung and M.A.Chaar. , US Patent(4772319). 1988.

[10] M.C.Kung and H.H.Kung, Journal of Catalysis 134 (1992) 668-677.

[11] P.Concepcion, A.Dejoz, J.M.Lopez Nieto, and Vazquez.M.I., Proceedings of the 14th Ibero-Aamerican Symposium On Catalysis (1994) 769.

[12] E.A.Mamedov and V.Cortes Corberan, Applied Catalysis A: General 127 (1995) 1-40.

[13] T.Blasco, J.M.L.Nieto, A.Dejoz, and M.Vazquez, I, Journal of Catalysis 157 (1995) 271-282.

[14] A.Galli, J.M.Lopez Nieto, and M.I.Vazquez, Catalysis Letters 34 (1995) 51. [15] J.Haber, A.Kozlowska, and R.Kozlowski, Journal of Catalysis 102 (1986) 52-63. [16] G.Bergeret, P.Gallezot, K.V.R.Chary, B.R.Rao, and V.S.Subrahmanyam, Applied

Catalysis 40 (1988) 191-196.

[17] H.Eckert and I.E.Wachs, Journal of Physical Chemistry 93 (1989) 6796. [18] P.J.Andersen and H.H.Kung, Journal of Physical Chemistry 96 (1992) 3114. [19] P.M.Michalakos, M.C.Kung, I.Jahan, and H.Kung, Journal of Catalysis 140

(1993) 226-242.

[20] S.Yoshida, T.Tanaka, Y.Nishima, H.Mizutani, and T.Funabilki, Proceedings of the 9th International Congress on Catalysis 3 (1988) 1473.

[22] M.Puglisi, F.Arena, F.Frusteri, V.Sokolovskii, and A.Parmaliana, Catalysis Letters 41 (1996) 41-43.

[23] J.Le Bars, J.C.Vedrine, A.Auroux, S.Trautmann, and M.Baerns, Applied Catalysis A: General 88 (1992) 179-195.

[24] L.Owens and H.H.Kung, Journal of Catalysis 144 (1993) 202-213.

[25] J.Perez-Ramirez, F.Kapteijn, K.Schoffel, and J.A.Moulijn, Applied Catalysis B: Environmental 44 (2003) 117-151.

[26] O.V.Buyevskaya, D.Wolf, and M.Baerns, Catalysis Today 62 (2000) 91-99. [27] P.Desrosiers, A.Guram, A.Hagemeyer, B.Jandeleit, D.M.Poojary, H.Turner, and

H.Weinberg, Catalysis Today 67 (2001) 397-402.

[28] K.Krantz, S.Ozturk, and S.Senkan, Catalysis Today 62 (2000) 281-289. [29] F.Kapteijn, J.Rodriguez-Mirasol, and J.A.Moulijn, Applied Catalysis B:

Environmental 9 (1996) 25-64.

[30] M.Gubelmann, J.Popa, and P.Tirel, EU patent 341,165 (1989). [31] M.Gubelmann, J.Popa, and P.Tirel, EU patent 406,050 (1991).

[32] A.S.Kharitonov, G.I.Panov, K.G.Ione, V.N.Romannikov, G.A.Sheveleva, L.A.Vostrikova, P.Tirel, and V.I.Sobolev, US Patent 5110,995 (1992). [33] J.Perez-Ramirez, A.Gallardo-Llamas, C.Daniel, and C.Mirodatos, Chemical

Engineering Science 59 (2004) 5535-5543.

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[35] Y.Wang and K.Otsuka, Journal of Molecular Catalysis A: Chemical 111 (1996) 341-356.

[36] P.O.Connor, A.Hakuli, and P.Imhof, Studies in Surface Science and Catalysis 149 (2004) 305.

[37] http://www.chemsystems.com (2005).

[38] T.A.Nijhuis, S.J.Tinnemans, T.Visser, and B.M.Weckhuysen, Chemical Engineering Science 59 (2004) 5487-5492.

[39] S.M.Bennici, B.M.Vogelaar, T.A.Nijhuis, and B.M.Weckhuysen, Angewandte Chemie 46 (2007) 5412-5416.

[40] C.M.Cardile, N.J.Tapp, and N.B.Milestone, Zeolites 10 (1990) 90-94. [41] C.Zenonos, G.Sankar, F.Cora, D.W.Lewis, Q.A.Pankhurst, C.R.Catlow, and

Introduction [42] B.M.Weckhuysen and R.A.Schoonheydt, Zeolites 14 (1994) 360-366.

[43] Y.Yokomori and Y.Kawachi, Zeolites 15 (1995) 637-639.

[44] C.T.Kresge, M.E.Leonowicz, W.J.Roth, J.C.Vartuli, and J.S.Beck, Nature (1992) 71.

[45] M.S.Hamdy, Chapter 1, PhD Thesis (2005) , Delft University of Technology. [46] J.C.Jansen, Z.Shan, Th.Maschmeyer, L.Marchese, W.Zhou, and N.van de Puil,

Chemical Communications (2001) 713-714.

[47] The 3-D HR-TEM images of TUD-1 particle were obtained under the supervision of Prof.K.P.de Jong, Utrecht University, The Netherlands, and were presented at the 13th.ICC,Paris, by Prof.F.Dautzenberg, (2006).

2

Effect of steaming of iron containing AlPO

4

-5

on the structure and activity

in N

₂

O decomposition

∗

∗

∗

∗

Iron-containing microporous AlPO4-5 (FAPO-5) with different Fe loading (1, 2, 4, 9 wt-%) were synthesized,

characterized and evaluated in N2O decomposition before and after steaming at 873 K for 5 h. Similar to the

observations previously reported for Fe-ZSM-5 catalysts, steam treatment leads to a significant improvement in the activity of FAPO-5 catalysts in N2O decomposition. Characterization by UV-Vis, FT-IR and Mössbauer

spectroscopy reveals that migration of Fe from framework to ex-framework positions occurs upon steaming, as well as partial reduction of Fe3+ to Fe2+, explaining the beneficial effect of this procedure on catalyst activity. The relatively low activity of steamed FAPO-5 compared to steamed isomorphously substituted Fe-ZSM-5 is tentatively explained by a negative effect of PO4

3

- groups on the desorption of surface oxygen. The implications of our results for a study recently reported by Sankar and co-workers [1], demonstrating a high performance of FAPO-5 in N2O mediated oxidation of benzene to phenol, are discussed.

∗

W. Wei, J.A. Moulijn, G. Mul, Effect of steaming of iron containing AlPO4-5 on the structure and activity in N2O decomposition, Microporous and Mesoporous Materials, 112 (2008) 193-201

1. Introduction

Nitrous oxide (N2O) is a greenhouse gas and contributes to the destruction of

ozone in the stratosphere. The emission of N2O from industry is mainly from the

manufacture of adipic acid and the production of nitric acid [2]. Thermal and catalytic abatement technologies are commercially available in the case of adipic acid plants, in which the N2O concentration is about 25-40 vol-% [2;3], and are appearing in the case of

nitric acid plants, where N2O concentrations are significantly lower [4]. Rather than just

decomposition of N2O to N2 and O2, N2O can also be used as an oxidant in the selective

oxidation of alkanes to alkenes, alkanes to alcohols, and aromatic hydrocarbons to the corresponding phenols. The most extensively studied system for these reactions is the Fe-ZSM-5 system, which was reported by the group of Panov to yield high selectivities in the oxidation of benzene to phenol [5-7], and methane to methanol [8]. Also the N2O

mediated oxidative dehydrogenation of propane to propene and the formation of oxygenates over Fe-ZSM-5 was reported [9]. Steaming is typically applied to convert isomorphously substituted Fe-ZSM-5 into the active state [9;10]. Besides Fe in ZSM-5 matrices, materials based on combinations of PO43- and Fe3+ ions are also interesting

systems for N2O mediated oxidation reactions. Bulk FePO4 is selective in N2O mediated

oxidation of methane and ethane, involving a selectivity promoting function of the phosphate groups [11], as well as in selective oxidation of benzene to phenol and methane to methanol [12]. Recently Fe substituted AlPO4-5, calcined at relatively high

temperatures (550 oC) was applied in the selective oxidation of benzene, and a high selectivity to phenol and comparable activity to Fe-ZSM-5 was reported [1]. In the latter study the activity was explained by the presence of a coordinatively unsaturated, tetrahedrally Fe3+ site [1].

In the present contribution, we provide further detail on the pioneering work of Sankar and coworkers, by investigation of the structural changes upon calcination and steaming of FAPO-5 with variable Fe-loading. Furthermore, the performance of the FAPO-5 catalysts in N2O decomposition is compared with the performance of

isomorphously substituted Fe-ZSM-5, before and after steam treatment, and discussed on the basis of the nature of the active sites in the respective catalytic formulations.

2. Experiments

2.1. FAPO-5 synthesis

Iron substituted AlPO4-5 catalysts were synthesized according to the recipe of

Catana et al. [13]. An iron source [Fe3+ chloride hexahydrate; Janssen Chimica] was added to a solution of phosphoric acid (85 wt% solution in water; Janssen Chimica) and water under stirring until complete dissolution. Pseudo-boehmite (Catapal; 70% Al2O3,

30% H2O; Vista) was added under vigorous stirring, followed by the dropwise addition of

triethylamine (TEA 99%; Janssen Chimica). All of the reagents were mixed in an ice bath. The obtained gel with a molar composition 1.0 TEA: xFe: (1-x) Al: 1.0 PO4: 80 H2O and

x = 0.000, 0.002, 0.004, 0.008 and 0.032, corresponding to 0, 1, 2, 4, and 9% w/w Fe in the FAPO-5 catalysts, was stirred for 1h and heated in Teflon-lined autoclaves at 150 oC for 20 h. The resulting solids were washed and then dried at 60 oC. All the as-synthesized catalysts were calcined at 500 oC for 10 h applying a ramp rate of 5 oC/min in air to remove the template. Steam treatment was performed with a water partial pressure of 300 mbar at 600 oC for 5 h for all samples, again applying a ramp rate of 5 oC /min to heat the sample from room temperature to 600 oC.

2.2. Characterization of FAPO-5

2.2.1. Characterization techniques

X-ray diffraction patterns for the FAPO-5 catalysts were measured in a Compact X-ray Diffractometer System (Philips PW 1840) equipped with a graphite monochromator, using Cu Kα radiation (λ=0.1541 nm). Data were collected in the 2θ

range of 0 to 60° at a scan rate of 0.02° s-1.

Ar adsorption isotherms at -184 oC were obtained using a Micromeritcs ASAP 2010 apparatus. The pore size distribution was calculated from the adsorption branch of the isotherm using the Satio-Foley (SF) model. Nitrogen ad/desorption isotherms were recorded at -196 oC on a Quantachrome Autosorb-6B using the volumetric technique. Prior to the analysis, the samples were evacuated and preheated at 300 oC overnight. The BET method was used to calculate the surface area (SBET), and the total pore volume (Vtot)

branch of the isotherm using the Barret-Joyner-Halenda (BJH) model, while the t-plot method was used to calculate the micropore volume (Vmicro).

The chemical composition of P in the catalysts was determined by ICP-OES (Perkin Elmer Optima 4300DV). The elemental analyses of the other metal elements, i.e. Fe, and Al were carried out using instrumental neutron activation analysis (INAA), conducted on a THER nuclear reactor with a thermal power of 2 MW and maximum neutron reflux of 2.10 m-2s-1. The method proceeds in three steps: irradiation of the elements with neutrons in the nuclear reactor, followed by a period of decay, and finally a measurement of the radioactivity resulting from irradiation. The energy of the radiation and the half-life period of the radioactivity enable a highly accurate quantitative analysis.

DRS-UV-Vis spectra were collected at ambient temperature on a Varian Cary 1 photo-spectrometer with BaSO4 as a reference.

27

Al and 15P Magic angle Spinning-Nuclear Magnetic Resonance (MAS-NMR) spectra were recorded at 9.4 T on a Bruker Avance 400 spectrometer operating at 104.2 and 161.9 MHz, respectively. The samples were spun in a 4 mm-diameter zirconia rotor with the spin speed set to 11 kHz. The lengths of the r.f. pulses were 1µs for Al and 2.2 µs for P. The acquisition time was 0.2 s for Al and 0.64 s for P. A time interval of 1 s for Al and 10 s for P between successive accumulations was selected in order to avoid saturation effects. The number of accumulations was ±4000 for Al and ±2000 for P. The chemical shift of 27Al and 15P are reported with respect to external standards at 0 ppm referenced to Al(NO3)3, and H3PO4, respectively.

Fourier Transform Infrared (FT-IR) measurements of NO adsorption on FAPO-5 samples were carried out using a Nicolet Magna 550 Fourier transform spectrometer, equipped with a high-temperature cell, in which the as-synthesized, calcined and steamed samples were placed in a holder. The Diffuse Reflectance Infrared Fourier Transformed (DRIFT) spectra were acquired by coaddition of 256 scans with a normal resolution of 4 cm-1. The samples were exposed to a flow of 30 ml/min, 5 vol.% NO in He at 100 oC for 100 min followed by flushing with He for 1 min, prior to recording of the spectra.

Transmission electron microscopy (TEM) was performed using a Philips CM30T electron microscope with a LaB6 filament as the source of electrons, and a Philips CM30UT electron microscope with a FEG as source of electrons, both operated at 300 kV. Samples were mounted on Quantifoil® microgrid carbon polymer supported on a copper grid from an ethanol suspension. TEM-EDX was performed for elemental analysis.

57

Fe Mössbauer spectra were measured on a constant acceleration spectrometer in a triangular mode with a 57Co:Rh source. Spectra for the FAPO-5 samples were obtained at room temperature. The overall spectra were deconvoluted with calculated Mössbauer spectra that consisted of Lorentzian-shape lines. In the case of quadrupole doublets the line widths and the absorption areas of the constituent lines were constrained equal. Positional parameters were not constrained in the fitting procedure. Isomer shift values are reported relative to sodium nitroprusside.

2.2.2. Catalytic testing

The N2O decomposition experiments were carried out in a six-flow reactor set-up

[14], using 50 mg catalyst (180-315µm) and a space velocity (GHSV) of 36000 h-1 at 3 bar. The feed condition used was 4.5 mbar N2O with He as balance gas. Before reaction,

the catalysts were heated in He with a ramp rate of 10oC/min, and then pre-treated in He at 400°C (10 oC/min) for 1 h. N2O, N2 and O2 were analyzed by online GC, using the

procedures described by Perez-Ramirez et al. [14].

In the six-flow reactor, 3 main sections can be distinguished: a) The gas mixing section, in which the reactant gas flow rates are controlled by mass flow controllers. b) The reactor section, that contains six small easily accessible fixed-bed quartz reactors which can be operated in parallel. The catalyst bed is located in the middle of the tube (inner diameter of 4 mm) and held with quartz wool. The total pressure in the reactor was near atmospheric and partial pressures of the reactants were varied by changing the individual flow rates. An extra thermocouple is located in the middle of the catalyst bed to calibrate the real reaction temperature. c) The analysis section contains an online GC (Chrompack CP9001) equipped with a thermal conductivity detector (TCD), using a Poraplot Q column and a Molsieve 5A column. A schematic flow sheet of the six-flow reactor system is shown in Scheme 1. The set-up could be used for high-throughput screening, quantitative testing, stability tests, and kinetic investigations.

MFC MFC MFC MFC MFC MFC MFC P P SM MFC R1 R2 R3 R4 R5 R6 P GC

ANALYSIS

FEED AND MIXING

VENT MFC NO_x SV reference catalyst (SiC) NO/He O2 He MFC MFC N2O/He He H2O LMFC P P SM P GC

ANALYSIS

FEED AND MIXING

VENT NO_x SV reference catalyst (SiC) NO/He O2 He N2O/He He H2O

Scheme 1. Schematic flow sheet of the six-flow reactor system for catalyst screening.

3. Results

3.1. Characterizations of FAPO-5

3.1.1. Chemical composition

The elemental composition of the various FAPO-5 samples, as synthesized, is shown in Table 1. As expected, the ratio of (Fe+Al) to P is close to unity for all samples. Still, with the Fe content increasing from 1 to 9 wt-%, the deviation of the (Fe+Al) to P ratio from unity increases up to 8%, indicating that samples of high Fe loading ( > 4 wt-%) might not consist entirely of well structured FAPO-5, in agreement with the literature [15]. Colour changes from white, to brown, to green, respectively, are consistent with the formation of Fe-rich phases.

Table 1. Chemical elemental composition of as-synthesized AlPO4-5 (denoted as APOas) and

FAPO-5 (denoted as FAPOas) catalysts.

Normalized atom

Samples code Fe(wt-%)

Fe(%) Al(%) P(%) (Fe+Al)/P Colour

APOas 0 0 50.4 49.6 1.02 White 1FAPOas 1.1 1.2 48.9 49.9 1.00 White 2FAPOas 2.1 2.3 48.4 49.3 1.03 White 4FAPOas 3.5 3.8 47.2 49.0 1.04 Brown 9FAPOas 8.7 9.7 42.2 48.2 1.08 Green 3.1.2. Microporous structure

X-ray diffractograms (Figure 1) of the as-synthesized APOas and FAPOas are consistent with the AlPO4-5 structure [16] up to 4 wt-%. In 9FAPOas, the X-Ray pattern

of an additional iron-rich phase, i.e. FeAl2(PO4)2(OH)2 [17] was obtained, obviously due

to the excessive amount of iron present in the system.

0 10 20 30 40 50 60

Diffraction angle 2θθθθ / degrees

4FAPOas 9FAPOas 2FAPOas 1FAPOas APOas

Figure 1. X-ray diffraction patterns of APOas, 1FAPOas, 2FAPOas , 4FAPOas and 9FAPOas. (!)

AlPO4·H2O (") FeAl2(PO4)2(OH)2

By careful inspection of XRD patterns, a slight shift and broadening of the Bragg lines is apparent when the iron content increases from 0 to 4 wt-% (See Figure 2). Such

phenomena are normally due to unit cell size expansion, and decreasing crystallinity: the result of integration of iron into the framework.

In addition, a small shoulder around 2θ =22.3 in the starting-material APOas becomes less visible as the iron loading increases. One possible explanation could also be the incorporation of iron into the framework, which slightly modifies the micro structure.

6.5 8.5 10.5 12.5 14.5

Diffraction angle 2θθθθ / degrees

APOas 1FAPOas 2FAPOas 4FAPOas (100) (110) (200) 19.5 20.5 21.5 22.5

Diffraction angle 2θθθθ / degrees

APOas 1FAPOas 2FAPOas 4FAPOas (210) (002) (300) 25 26 27 28 29 30 31

Diffraction angle 2θθθθ / degrees

APOas 1FAPOas 2FAPOas 4FAPOas (220) (310) (400) 33 34 35 36 37 38 APOas 1FAPOas 2FAPOas 4FAPOas (222) (312) (213)

Diffraction angle 2θθθθ / degrees

Figure 2. X-ray diffraction patterns of APOas, 1FAPOas, 2FAPOas and 4FAPOas.

Upon calcination, there was a slight change in the relative peak intensities of the XRD patterns, but new phases could not be identified in the calcined materials (denoted as FAPOcal). The main change in peak position is observed at the contribution of the 002 plane (at about 21.1 2θ in the calcined samples) indicating a reduction of the d value in the c direction, which is possibly due to the removal of the template from the hexagonal channels(See Figure 3). The shoulder shown at 2θ =22.3 in starting-material APOas disappeared upon calcination, so it could be postulated that this shoulder might also be related to the presence of template.

After steam treatment with a water partial pressure of 300 mbar at 600oC for 5 h, the long-range order of FAPOcal is basically maintained. However, the positions of all the peaks in XRD moved to smaller angle, and the peaks are broadened. This suggests that the unit is expanding along all directions and the structure becomes less crystallized. This might indicate that steam treatment has led to an enlargement of the cell size, modifying the micro structure. For the 9 wt-% sample, the treatment at high temperature did not result in new phases, with the diffraction lines showing the pattern of the FeAl2(PO4)2(OH)2 phase.

0 10 20 30 40 50 60

Diffraction angle 2θθθθ / degrees

APOas APOcal APOste

19 20 21 22 23 24

Diffraction angle 2θθθθ / degrees

APOas APOcal APOste (210) (002) (300) 0 10 20 30 40 50 60

Diffraction angle 2θθθθ / degrees

1FAPOas 1FAPOcal 1FAPOste

19 20 21 22 23 24

Diffraction angle 2θθθθ / degrees

1FAPOas 1FAPOcal 1FAPOste

0 10 20 30 40 50 60

Diffraction angle 2 θθθθ / degrees

2FAPOas 2FAPOcal 2FAPOste

19 20 21 22 23 24

Diffraction angle 2 θθθθ / degrees

2FAPOas 2FAPOcal 2FAPOste