Method for the catalytic steam reforming of a hydrocarbon

ES 2 344 447 T3

DESCRIPTION

Method and apparatus for obtaining an improved production rate of thermal chemical reactions.

Field of the invention

The present invention relates to a method and apparatus for thermal chemical reactions. The method and apparatus can provide an improved reaction rate for thermal chemical reactions.

Background of the invention

Thermal chemical reactions are those chemical reactions that produce heat (exothermic) or consume it (endothermic). Examples of thermal chemical reactions include hydrocarbon conversion reactions such as steam reforming, water gas exchange reactions, and combustion. These well known reactions are normally carried out in the presence of a catalyst at temperatures up to about 1300 ° C. Because the intrinsic kinetics of a thermal chemical reaction can be much faster than the rate of heat transfer between the reaction vessel and the heat sink, source, or thermal environment, the actual rate of product production (i.e., the observed level) is slower than the intrinsic level. Intrinsic kinetics means the level at which they could theoretically form on the surface of the catalyst.

Limited production rates can result from a longer residence time which is typically seconds to minutes in conventional thermal chemical reaction vessels. As conventionally defined, the residence period is equal to the volume of the reaction zone divided by the inlet volumetric flow rate of the reactants at the temperature and pressure of the reaction system. The reaction zone is the total volume of the catalyst and surrounding area through which the reactants and products flow.

An example of these limited production rates can be seen in the water gas shift reaction, which is conventionally carried out in fixed bed reactors. In the water gas exchange reaction, carbon monoxide and water are converted to carbon dioxide and hydrogen. Conventionally, this reaction suffers from residence times of multiple seconds (a kinetic hindrance) when carried out in fixed-bed reactors. Theoretical kinetics suggests that residence times in the order of milliseconds could theoretically be obtained. There are two aspects of kinetic retardation of conventional reactors. The first is a diffusion limitation as reactants diffuse into and out of a porous catalyst support granulate and the second is a thermal transfer limitation that is a combination of thermal transfer parameters (thermal conductivity and length) of supports. catalyst size and overall reactor geometry (shape, size, and distance to the external heat exchanger). Because the water gas shift reaction is critical for a multi-reactor fuel treatment system that supports distributed energy production through the use of a power cell, there is a need for a smaller and faster reactor. change of water gas.

Another example of a thermal chemical reaction is in the conventional methane steam reforming reactor that produces synthesis gas at an average residence time of several seconds and with an efficiency factor of 0.01 to 0.05 as reported by Adris, A., Pruden, B., Lim, C., J. Grace, 1996, “On the reported attempts to radically improve the performance of the steam metane reforming reactor”, Canadian Journal of Chemical Engineering, 74, 177-186. In a typical industrial operation, the methane to steam ratio is run at 3: 1 to prevent coke formation. Efforts to improve heat transfer between the reaction vessel for this endothermic reaction and the heat source have made only modest improvements in the rate of product production.

Thermal reactions have long been, and continue to be, performed in enormous volumes on production scales requiring very large capital investments, typically greater than $ 100 million. Not surprisingly, there have been extensive efforts, over a long period of time, aimed at improving the speed and efficiency of these reactions. Despite these attempts, a need remains for a method and apparatus that increases the rate of heat transfer between the reaction vessel and the heat sink or source and thus approximates the intrinsic theoretical kinetic rate of reaction and production.

WO 00/06295 discloses methods and apparatus for obtaining an improved production rate per reaction chamber volume of a reaction chamber with an inlet and an outlet for a thermal chemical reaction, wherein a proportion of the improved rate of production per Reaction chamber volume at a conventional rate of production by volume of the conventional reaction chamber for the thermal chemical reaction is at least 2. For example, for steam reforming, the residence time is on the order of seconds whereas with the invention of WO 00/06295, the residence time is less by a factor of 2, on the order of milliseconds to tens or hundreds of milliseconds. In one aspect, the disclosure includes:

(a) a porous insert within the volume of the reaction chamber, wherein a reactive flow passes substantially completely through the porous insert wherein the volume of the reaction chamber with the porous insert has an average porosity of less than 1 and a mass transport distance of the reactants to a catalyst site of no more than 3 mm;

ES 2 344 447 T3 (b) the volume of the reaction chamber with a length parallel to a reactive flux mass, the length less than or equal to 6 inches, and with a height (a thermal distance from the heatsink to the heat source) less than or equal to 2 inches, thus transferring reaction heat at an improved rate of heat transfer through the porous insert; and (c) a thermal transfer chamber in thermal contact with the volume of the reaction chamber, serving as a heat sink or source, the thermal transfer chamber at said improved rate of thermal transfer through a wall between the chamber. heat transfer and reaction chamber, thus obtaining the improved rate of production by volume of the reaction chamber for the thermal chemical reaction wherein a ratio of the improved rate of production by volume of the reaction chamber to a conventional rate of production by volume of the conventional reaction chamber for the thermal chemical reaction it is at least 2.

These characteristics have been found to cooperate with the kinetics of the reaction in terms of heat transfer at a rate sufficient to avoid substantial kinetic hindrance. These characteristics are effective for both catalytic and non-catalytic thermal chemical reactions. For catalytic chemical reactions, the addition of a fine stratum of catalyst (<150 microns, pm, more preferably less than 50 pm) on the porous insert substantially reduces the diffusion paths of reactants to catalyst sites compared to severe limitations. of reactive diffusion within ceramic granules (> 1 mm) as in conventional systems. Thus for catalytic thermal chemical reactions, both kinetic impediments are substantially reduced allowing the realization of the theoretical or quasi-theoretical reaction kinetics. More specifically, a water gas exchange reactor made according to WO 00/06295 is 1/10 to 1/100 the size of conventional processing hardware for the same production output.

Summary of the invention

The present invention further provides methods and apparatus (containers) as set forth in the claims below. These provide an index of heat transfer from a reaction chamber through a wall to a heat transfer chamber (exothermic reaction) or provide heat from a heat transfer chamber through a wall to a reaction chamber (endothermic reaction) at a rate of at least 0.6 W / cc of total reactor volume. An important aspect of this invention is the thermal distance defined in a transverse plane through the container including a thermal transfer chamber, reaction chamber, and a wall between the chambers. The transverse plane is perpendicular to a mass flow direction of the reactive current, and the thermal distance is a distance between a cooler position and a warmer position in the transverse plane. The thermal distance is of a length where the level of heat transfer from (or) the reaction chamber to (or from) a heat transfer chamber (heat exchanger) substantially equals the local heat transfer rate.

The invention includes a process for the catalytic conversion of at least one reagent in a thermal chemical reaction, in which at least one reagent is passed into at least one reaction chamber; heat is transferred to or from the reaction chamber to at least one heat exchanger; and you get at least one product. The reaction chamber contains a catalyst that catalyzes the reaction of the reagent or reagents. The process of the invention has the following characteristics: at steady state, at least 0.6 W / (cc of total reactor volume) of heat is transferred, where the total volume of the reactor is defined as the sum of the volume of the chamber (s ) reaction and heat exchanger chamber (s) including chamber wall volume; the contact time of the reactant with the catalyst is less than about 0.3 seconds; and the pressure drop across the reaction chamber is less than about 15 psig (103 KPa).

An example of a thermal chemical reaction that can be performed using the methods and reactors of the present invention is steam reforming of a hydrocarbon. In this process a feed stream comprising hydrocarbon gas and steam enters a reaction chamber containing a catalyst that catalyzes the reaction of hydrocarbon gas and steam to produce a gaseous mixture comprising at least carbon monoxide and hydrogen gas. This process can produce more than 0.01 SLPM of hydrogen gas per cubic centimeter of total reactor volume.

The present invention also provides a reactor according to claim 12 for the catalytic conversion of at least one reagent in a thermal chemical reaction, comprising: at least one reaction chamber containing a porous catalyst insert; and at least one heat exchanger that is in thermal contact with the reaction chamber. The reaction chamber has a length of less than or equal to 6 inches (150 mm) and a height of less than or equal to 2 inches (51 mm). A porous catalyst insert can comprise a porous metal foam having open cells ranging from about 20 ppi to about 3000 ppi (7.8 ppcm to 1180 ppcm).

The reaction chamber has a height of less than or equal to 2 inches (50.8 mm); and wherein at least one heat exchanger and at least one reaction chamber are configured such that, during steady state operation, at least 0.6 W of heat per cc of the total reactor volume can be transferred between the heat exchanger and the reaction chamber.

ES 2 344 447 T3

The invention also includes a process according to claim 1 for the catalytic conversion of at least one reagent in a thermal chemical reaction in which at least one reagent is passed into at least one reaction chamber containing a catalyst that catalyzes the reaction of al minus one reagent; transferring heat to or from said at least one reaction chamber of or in said at least one heat exchanger; and that obtains at least one product from the reaction chamber; where the thermal transfer phase, in steady state, transfers at least 0.6 W of heat per cc of total reactor volume, so that, in steady state, the catalyst is kept within a temperature range that reduces the formation of al minus one undesirable product of the chemical reaction. The formation of undesirable chemical (s) can be reduced by using a contact time of less than about 0.3 seconds, thus suppressing slow reactions that can form an undesirable chemical reaction product. Unwanted chemicals can result from side reactions or slow side reactions. In the water gas shift reaction, desirable products include carbon dioxide and water, and an undesirable product is methane. In steam reforming a hydrocarbon, desirable products include hydrogen and carbon monoxide and / or carbon dioxide, and an undesirable product is coke.

The subject matter of the present invention is particularly noted and clearly claimed in the concluding portion of this specification. However, both the organization and the method of operation, together with the advantages and additional objects thereof, can be better understood by reference to the following description taken in relation to the attached drawings where the reference characters refer to similar elements .

Brief description of the drawings

Fig. 1a is a cross section of a stacked reaction chamber with heat exchanger chamber.

Fig. 1b is an isometric of a nested reaction chamber with a heat exchanger chamber.

Fig. 2a is a graph of percent selectivity versus residence time for a long contact time water gas change with a porous catalyst powder insert.

Fig. 2b is a graph of percent selectivity versus residence time for a short contact time water gas change with a porous catalyst powder insert.

Fig. 3 is a graph of percent selectivity versus temperature various contact times for water gas change with a porous coated metal foam insert.

Fig. 4 is a graph of methane conversion versus temperature for various contact times for a 2.5: 1 vapor to methane ratio.

Fig. 5a is a graph of conversion and selectivity versus time for steam reforming of n-butane with a porous catalyst insert.

Fig. 5b is a graph of conversion and selectivity versus time for steam reforming of n-butane with a regenerated porous catalyst insert.

Fig. 6 illustrates the design features of a microchannel reactor having multiple reaction chambers and heat exchangers.

Fig. 7 illustrates front (top) and side (bottom) views of a reactor design having porous catalyst inserts within a reaction chamber.

Fig. 8 illustrates a cross-sectional view of a reactor design having cylindrical reaction chambers and heat exchangers.

Fig. 9 is a graph showing the% conversion, selectivity of H₂ and% of H₂ in the effluent of an isooctane steam reforming process.

Fig. 10 is a bar graph showing% conversion, H2 selectivity and% H2 in the effluent from an isooctane steam reforming process at varying ratios of steam to carbon.

Specific description

Referring to Figures 1a and 1b, a vessel 100 for thermal chemical reactions having two chambers 102 and 104 with a wall 106 between them. Either of the two chambers 102, 104 can be the reaction chamber. The bulk flow of reagents within reaction chamber 102 is substantially perpendicular to a plane of cross section 108. Container 100 may have stacked chambers as in Fig. 1a or nested chambers as in Fig. 1b. The reaction in the reaction chamber can be endothermic or exothermic.

In a thermal chemical reaction, the steady state production rate (reaction kinetics) is limited by the rate of heat transfer to either the (endothermic) or the (exothermic) site of the reaction. In the

In the case of exothermic reactions, low heat removal rates can promote unwanted side reactions, or cause thermal hot spots or thermal overflow in the reactor. Commercial exothermic reactors are often operated with low conversion per pass to prevent hot spots and thermal excursions. Improved heat removal would allow safe operation at a higher production rate per unit volume of reactor hardware. In order to obtain an improved heat transfer index and thus an improved production rate, the reaction chamber preferably has a porous insert (not shown) within the volume of the reaction chamber wherein the porous insert within the volume of the reaction chamber has an average porosity of less than 1, a transport distance of the reagent (s) to the catalyst site not greater than 3 mm, and a height (a thermal transport distance from the heat source to the heat sink, not greater than 2 inches (51 mm), thus transferring heat of reaction at an improved heat transfer rate through the porous insert.

The porous insert can be a powder, a porous monolith (including but not limited to metallic or ceramic foam, felt, honeycomb, tube bank, stacked microchannel assembly, and combinations thereof), fibers, pad (eg. steel wool), or combinations thereof. In view of the cost of replacing a spent catalyst with catalytic reactors, it is preferred that the porous insert be removable from the reaction chamber. The porous insert can be arranged to provide single or multiple flow passages for reagents through the volume of the reaction chamber.

Preferably, the volume of the reaction chamber has a length parallel to a bulk reactive flow, the length of less than or equal to 6 inches (150 mm), and has a height, a thermal distance from the heat sink to the source of heat, less than or equal to 2 inches (51 mm). The limited length and height provide short distances that allow for faster heat transfer. Furthermore, the short length reduces the overall pressure drop across the reaction chamber.

The heat transfer chamber (heat exchanger) is in thermal contact with the volume of the reaction chamber, the heat transfer chamber transferring heat at the improved heat transfer rate through the wall 106 between the heat transfer chamber and the chamber of reaction, thus obtaining the rate of improved production by volume of the reaction chamber for the thermal chemical reaction.

For catalytic thermal chemical reactions, a preferred catalyst has a porous support, an interfacial layer deposited with solution thereon, and a catalyst material on the interfacial layer. A more preferred catalyst has a porous support, a buffer layer, an interfacial layer, and a catalyst material. Any layer can be continuous or discontinuous as in the form of spots or dots, or in the form of a layer with gaps or holes.

The porous support can be a porous ceramic or metallic foam. Other suitable porous supports for use in the present invention include carbides, nitrides, and composites. Before depositing the layers, the porous support has a porosity of at least 5% as measured by mercury porosimetry and a mean pore size (sum of pore diameters / number of pores) from 1: m to 1000: m as measured by optical and scanning electron microscopy. Preferably, the porous support has a porosity of from about 30% to about 99%, more preferably from 60% to 98%. The preferred forms of porous supports are foams, felts, pads, and combinations thereof. Foam is a structure with continuous walls that define pores throughout the entire structure. Felt is a structure of fibers with interstitial spaces between them. The pad is a tangled strand structure, like steel wool. Less preferably, the porous supports may also include other porous media such as granules and honeycombs, provided they have said porosity and pore size characteristics. The open cells of a metallic foam preferably range from about 20 pores per inch (ppi) to about 3000 ppi (7.8 to 1180 pores per cm, ppcm) and more preferably about 20 to about 1000 ppi (7.8 to 390 ppcm), even more. preferably about 40 to about 120 ppi (15.7 to 47.2 ppcm). PPI is defined as the largest number of pores per inch (in isotropic materials the direction of the measurement is irrelevant; in anisotropic materials, however, the measurement is made in the direction that maximizes the number of pores). Ppi is measured by scanning electron microscopy. A porous support has been found to provide several advantages, including low pressure drop, improved thermal conductivity over conventional granular ceramic supports, and ease of loading / unloading in chemical reactors.

The buffer layer, if present, has a different composition and / or density from the support and the interfacial layers, and preferably has a coefficient of thermal expansion that is intermediate the coefficients of thermal expansion of the porous support and the interfacial layer. Preferably, the buffer layer is a metal oxide or metal carbide. The applicants found that the vapor-deposited layers are superior because they show better adhesion and resist flaking even after several thermal cycles. Most preferably, the buffer layer is Al₂OR₃, Uncle₂, Yes₂, and ZrO₂ or combinations thereof. More specifically, the Al₂Or is it a-Al₂OR₃, and-Al₂OR₃, and combinations thereof. a-Al₂OR₃ it is most preferred for its excellent resistance to oxygen diffusion. Therefore, it is expected that the resistance against high temperature oxidation can be improved with coated alumina on the porous support. The buffer layer can also be formed from two or more substrates of different compositions. When the porous support is metal, for example a stainless steel foam, a preferred embodiment has a buffer layer formed from two substrates of different compositions (not shown). The first substratum (in contact with the porous support) is preferably TiO₂ because it shows good adhesion to the porous metal support. The second substratum is preferably Al₂OR₃ that is placed on the TiO₂. In a way

ES 2 344 447 T3 of preferred embodiment, the substratum a-Al₂OR₃ It is a dense layer that provides excellent protection for the underlying metal surface. A less dense, high surface area interfacial layer such as alumina can then be deposited as a support for a catalytically active layer.

Typically the porous support has a thermal coefficient of expansion different from that of the interfacial layer. Therefore, for high temperature catalysis (T> 150 ° C) a buffer layer is necessary for the transition between the two coefficients of thermal expansion. The thermal coefficient of expansion of the pad layer can be tailored by controlling the composition to obtain a coefficient of expansion that is compatible with the coefficients of expansion of the porous support and interfacial layers. Another advantage of the buffer layer is that it provides resistance against side reactions such as coking or cracking caused by a bare metal foam surface. For chemical reactions that do not require large surface area supports such as catalytic combustion, the buffer layer stabilizes the catalyst metal due to the strong metal to metal oxide interaction. In chemical reactions that require large surface area supports, the buffer layer provides a stronger bond to the high surface area interfacial layer. Preferably, the tampon layer is free of pin openings and holes - this provides added protection of the underlying support. More preferably, the buffer layer is non-porous. The buffer layer has a thickness that is less than one-half the average pore size of the porous support. Preferably, the buffer layer is between about 0.05 and about 10 µm thick, more preferably less than 5 µm thick. The buffer layer must show thermal and chemical stability at elevated temperatures.

In some embodiments of the present invention, adequate adhesion and chemical stability can be obtained without a buffer layer, so that the buffer layer can be omitted, thus saving costs, supplying extra volume, and further improving the heat transfer of the catalyst. .

The interfacial layer can be composed of nitrides, carbides, sulfides, halides, metal oxides, carbon, and combinations thereof. The interfacial layer provides a high surface area and / or provides a desirable catalyst support interaction for supported catalysts. The interfacial layer can be composed of any material that is conventionally used as a catalyst support. Preferably the interfacial layer is a metal oxide. Examples of metal oxides include, but are not limited to y-Al₂OR₃, Yes₂, ZrO₂, TiO2 tungsten oxide, magnesium oxide, vanadium oxide, chromium oxide, manganese oxide, iron oxide, nickel oxide, cobalt oxide, copper oxide, zinc oxide, molybdenum oxide, tin oxide, oxide calcium, aluminum oxide, lanthanium series oxide (s), zeolite (s) and combinations thereof. The interfacial layer can serve as a catalytically active layer without any additional catalytically active material deposited thereon. Typically, however, the interfacial layer is used in combination with the catalytically active layer. The interfacial layer can also be formed from two or more substrates of different composition. The interfacial layer has a thickness that is less than one-half the size of the average pore of the porous support. Preferably, the thickness of the interfacial layer ranges from about 0.5 to about 100 µm, more preferably from about 1 to about 50 µm. The interfacial layer can be either crystalline or amorphous and preferably has a BET surface area of ​​at least 1 µm.²/ g.

The catalytically active material (when present) can be deposited on the interfacial layer. Alternatively, a catalytically active material can be deposited simultaneously with the interfacial layer. The catalytically active layer (if present) is typically intimately dispersed in the interfacial layer. The fact that the catalytically active layer is "arranged in" or "deposited in" the interfacial layer includes the conventional understanding that the catalytically active microscopic particles are dispersed: on the surface of the support layer (ie, interfacial layer), in cracks in the support layer, and in open pores in the support layer. The catalytically active layer can include: catalyst metals, including but not limited to, noble metal, transition metal, and combinations thereof; Metal oxides, including but not limited to, alkali element oxides, alkaline earth elements, boron, gallium, germanium, arsenic, selenium, tellurium, thallium, lead, bismuth, polonium, magnesium, titanium, vanadium, chromium, manganese, iron , nickel, cobalt, copper, zinc, zirconium, molybdenum, tin, calcium, aluminum, silicon, element (s) of the lanthanium series, and combinations thereof; compounds; zeolite (s); nitride; carbides; sulfides; halides; phosphates; and combinations of any of the above.

To mitigate the limitation of mass transfer of the catalyst structure, the catalyst impregnation preferably forms a porous interfacial layer having a depth of less than 50 µm, preferably less than 20 µm. Therefore, the length of the diffusion path is at least a factor of 5 shorter than for standard catalyst particles. The finer impregnated catalyst structure also improves heat transfer, due to a shorter heat transfer path.

The catalyst structure can be any geometric configuration. Preferably, the catalyst is a porous structure such as a foam, felt, pad, and combinations thereof. The catalyst (including support and catalyst material) is preferably sized to fit within a reaction chamber. The catalyst can be a single piece of contiguous porous material, or many pieces in physical contact. The catalyst is preferred to have contiguous material and contiguous porosity so that molecules can diffuse through the catalyst. The catalyst can be arranged in a reaction chamber such that gases flow substantially through the catalyst (single or multiple pieces) rather than around it. In a preferred embodiment, the cross-sectional area of ​​the catalyst occupies at least 80%, more preferably at least 95% of the cross-sectional area of ​​the reaction chamber. The catalytically active metal can be distributed in

ES 2 344 447 T3 surfaces through the catalyst so that reactants that pass through the catalyst can react anywhere along the passage through the catalyst; this is a significant advantage over granule type catalysts which have a large volume of unused space or inefficiently catalytically used space within the granules. Porous catalyst is also superior over powders because compacted powders can cause severe pressure drop. The catalyst preferably has a surface area, as measured by BET, greater than about 0.5 µm²/ g, more preferably greater than about 2.0 µm₂/ g.

The catalysts for use in the present invention can be characterized by the properties they exhibit. Factors that can be controlled to affect these properties include: selection of porous support, buffer, interfacial, and catalytically active layers; the graduation of the coefficients of thermal expansion, crystallinity, interactions of the metal supports, size of the catalyst, thermal conductivity of the support, porosity, thermal conductance of the reaction chamber, deposition techniques and other factors as they are apparent in view of the descriptions in this one. Certain preferred catalysts of the present invention exhibit one or more of the following properties: adhesion - after 3 thermal cycling in air, the catalyst exhibits less than 2% (by area) flaking as seen by SEM analysis. swept); resistance to oxidation, conversion of reagent (s), contact / residence times, product selectivity, pressure drop and production rates.

A preferred method of making the catalyst has the steps of selecting a porous support, depositing a buffer layer on the porous support, and depositing an interfacial layer on it. Optionally a catalyst layer can be deposited on the interfacial layer or both the interfacial layer and the catalyst layer can be deposited simultaneously on the buffer layer.

Because metal has WEB surfaces that are non-porous and smooth, the deposition of the buffer layer can be prevented. One way to mitigate this problem is to roughen the metal surface using chemical delustration. The adhesion of high surface area gamma-alumina supported metal catalysts to metal foam is significantly improved when the metal foam is roughened by chemical delustration using mineral acid solutions, eg 0.1 to 1 M HCl. The roughened fabric surface also exhibits improved resistance to catalyst stratum peeling under thermal cycling. In a preferred embodiment, where a metallic foam is used as the porous support, the metallic foam is etched prior to vapor deposition of the buffer layer. The delustration is preferably with an acid, for example HCl.

Deposition of the buffer layer is preferably by vapor deposition including but not limited to chemical vapor deposition, physical vapor deposition, or combinations thereof. Surprisingly, it has been found that vapor deposition, which is typically carried out at high temperatures, results in polycrystalline or amorphous phases that provide good adhesion of the buffer layer to the surface of the porous support. The method is particularly advantageous for adhering a metal oxide buffer layer to a porous metal support. Alternatively, the buffer layer can be obtained by coating the solution. For example, the solution coating has the phases of functionalizing the metal surface by exposing the metal surface to water vapor to form surface hydroxyl, followed by a surface reaction and hydrolysis of alkoxides to obtain a coating of metal oxide. This coating of the solution may be preferred as a lower cost method of depositing the buffer layer.

The interfacial layer is preferably formed by vapor or solution deposition using precursors as known to these techniques. Suitable precursors include organometallic compounds, halides, carbonyls, acetonates, acetates, metals, colloidal dispersions of metal oxides, nitrates, sludge, etc. For example, a porous alumina interfacial layer can be dip coated with PQ alumina colloidal dispersion (Nyacol Products, Ashland, MA) followed by drying in a vacuum oven overnight and calcining at 500 ° C for 2 hours.

The catalytically active material can be deposited by any suitable method. For example, the catalyst precursors can be deposited onto colloidal metal oxide particles and coated sludge on a porous buffer coated support, and then dried and reduced.

Certain embodiments of the present invention can be identified in terms of residence or contact time. These terms have well-defined meanings in the art. The contact time is the total volume of the catalyst chambers divided by the total flow rate (defined as F-total) of incoming reagents assuming they are an ideal gas corrected to standard conditions (i.e. the volume of the chamber catalyst / F-total to STP where STP is 273K and 1 atm). The volume of the catalyst chambers includes the volume in immediate and surrounding proximity to the catalyst zone. As an example, if one were to pack a quarter of the channels with dust, then the volume of the catalyst chamber would only include that region where the gas can flow and where the catalyst can contact, i.e. only a quarter of the total volume of the channel. would be included in this calculation. The volume of dead space, i.e. header, footer, etc. is ignored in this calculation. The residence time (which means the average residence time) is the total volume of the catalyst chambers divided by the total flow rate of the inlet reagents, corrected to the actual temperature and the pressure of the reactants in the reactor. (that is, the volume of the catalyst chamber / F-total corrected to real conditions). The F-total at STP is the total reagent volumetric flow rate (includes all reagents, and diluents if present). Inlet gases are typically measured with mass flow controllers, set at standard conditions, ie

ES 2 344 447 T3 tell the user to preset the desired flow rate STP. The F-total corrected to real conditions = F-total-STP X (temperature in K) / 273 x 1 atm / (real P in atm): this value is used to calculate the dwell time or the “real time” within of a reactor. Contact time is preferred by most professionals because it is a convenient method of keeping the variable time fixed while stepping through 10 degree C increments in reaction temperature etc.

A catalytic process comprises the passage of at least one reagent into a reaction chamber comprising the inventive catalyst, conversion of said at least one reagent to at least one product, and the passage of the product out of the reaction chamber. In a preferred embodiment, the catalytic process is carried out in an apparatus having microchannels. The microchannels have at least a dimension of about 1mm or less. Examples of suitable microchannel apparatus and various process related factors are described in US Patent Nos. 5,611,214, 5,811,062, 5,534,328, and US Patent Application Serial Nos. 08 / 883,643, 081938,228, 09 / 375,610, 09 / 123,781, jointly applied for US patent application serial no. 09 / 492,950, 09 / 375,614 (requested on August 17, 1999) and 09 / 265,227 (requested on March 8, 1999). In another preferred embodiment, the catalyst is a monolith - a single contiguous, yet porous piece of catalyst or several contiguous pieces that are stacked together (not a bed of packed powder or granules or a coating on the wall of a microchannel ) that can be easily inserted and removed from a reaction chamber. The catalyst piece or stack of pieces preferably has a width of 0.1mm to about 2cm, with a preferred thickness of less than 1cm, more preferably, about 1 to about 3mm. The inventive catalyst can provide numerous advantages to catalytic processes such as: chemical stability, stability to repeated thermal cyclization, thermostability, efficient loading and unloading of catalysts, high rates of heat transfer and mass transfer, and maintenance of the desired catalytic activity.

In the construction of heat exchanger embodiments, thin sheets or tubes can be used to obtain high heat loads and short contact times. The thickness of the tissue between the reaction channel and the heat exchange channel can vary, but is preferably between about 0.01 inches (0.25 mm) and about 0.25 inches (6.4 mm). The preferred thickness for the heat exchange channel (meaning the thickness of the smallest dimension of the heat transfer channel) preferably ranges from 100 microns to 10 millimeters. In some preferred embodiments, this smallest dimension may be the width of the channel, in other embodiments, the height of the channel. The preferred thickness is 250 microns to 3 millimeters. The flow of the heat transfer fluid can be either countercurrent, crosscurrent, or current parallel to the direction of flow of the reactants. Preferred heat transfer fluids include: a combustion stream (for endothermic reactions), oil (lower temperature reactions), and steam.

The metal surfaces within the microchannel apparatus can be coated with one or the other or both of the interfacial layer and buffer. This can be done using any of the processes described here, preferably by vapor deposition. Preferred coating materials include titania and 5-10% SiO₂/To the₂OR₃. The inner surfaces of the reaction chamber, heat exchanger and other surfaces of microchannel apparatus can be coated. In some embodiments, the walls of a reaction chamber can be coated with an optional buffer layer, an interfacial layer, and a catalytically active material - typically the catalytically active material and the interfacial layer combine to form a supported catalyst. Coatings can also be applied to metal walls in tubes and pipes that form connections to or within microchannel apparatus.

The inventive method is preferably carried out in a reaction chamber in which the catalyst has a thickness of about 2 cm or less and is touching or in close proximity (within about 1 mm) of a wall of the reaction chamber, where the wall of the reaction chamber is in thermal contact with a heat exchanger. Thermal transfer from (or to) the reaction chamber is preferably enhanced by the addition of microchannels in at least one chamber wall through which heat is transferred, preferably, on the opposite reaction chamber wall side to the catalyst structure. The catalyst preferably has relatively large and contiguous pores, such as in a foam, to avoid large pressure drops. Preferably the pore size of the large pores in the catalyst is between about 10 µm and about 300 µm.

The catalytic processes of the present invention include: acetylation, addition reactions, alkylation, dealkylation, hydrodealkylation, reductive alkylation, amination, aromatization, arylation, autothermal reformation, carbonylation, decarbonylation, reductive carbonylation, carboxylation, reductive carboxylation, reductive coupling, condensation, cracking, hydrodisintegration, cyclization, cyclooligomerization, dehalogenation , dimerization, epoxidation, esterification, exchange, Fischer-Tropsch, halogenation, hydrohalogenation, homologation, hydration, dehydration, hydrogenation, dehydrogenation, hydrocarboxylation, hydroformylation, hydrogenolysis, hydrometallization, hydrosilation, hydrolysis, hydrotreatment, hydrodeulferisation / hydrodenitrogenation (HDS / HDN), isomerisation, methane, methanol synthesis, methylation, demethylation, metathesis, metathesis , partial oxidation, polymerization, reduction, steam and carbon dioxide reforming, sulfonation, telomerization, transesterification, trimerization, water gas change (WGS), and reverse water gas change (RWGS).

The reaction process (s) of the present invention can be performed in parallel, with tens, hundreds, thousands, or millions of small reaction chambers, each chamber having an internal diameter of less than approximately 2 inches (50.8 mm). , preferably less than about 2 cm, more preferably ranging from about 1 mm to about 5 mm. The reaction process (s) can also be run

ES 2 344 447 T3 in series. For example, products from one reaction chamber can be fed to another reaction chamber that has the same or different catalyst. A series of reactions could also be carried out by placing a series of different catalysts within the same reaction chamber. The reactive products can be collected and stored, for example in tanks, or consumed immediately in subsequent reactions.

The reactors and methods of the present invention can be characterized by various properties that they exhibit. Heat flow is a particularly important feature in the present invention. For the sum of the volume of the reaction chamber (s) and heat exchanger including the volume of chamber walls, the present invention preferably exhibits a heat flux during steady state operation of at least about 0.6 W / cc. , more preferably about above 1 W / cc still more preferably between about 5 and about 250 W / cc, and in another preferred range, between about 10 and 100 W / cc. These heat fluxes can be obtained at short contact times and low pressure drops through the reaction chamber. The contact time is less than about 0.3 seconds, preferably less than about 0.1 seconds, more preferably less than about 0.05 seconds, and even more preferably less than about 0.01 seconds. The pressure drop across the reaction chamber is preferably less than about 15 psig (pounds per square inch gauge) 103 kPa, more preferably less than about 10 psig 69 kPa, still more preferably less than about 5 psig (35 kPa), and even more preferably less than about 1 psig (6.9 kPa). In the apparatus and methods of the present invention, these high fluxes can be obtained in a wide variety of catalyzed reactions and is not limited to the highly exothermic (combustion) deep oxidation reactions.

It has surprisingly been found that the aforementioned short contact times and high thermal fluxes can be obtained in thermal catalyzed reactions at steady state. Factors that can contribute to shorter contact times and higher heat flux include: catalysts with fast intrinsic kinetics, porous catalysts, thermally conductive supports, the use of microchannel apparatus, short distances for thermal transport in the reaction chamber and / or heat exchanger; short mass transfer distances in the reaction chamber; and selection of heat transfer fluid (s). The transfer of a sufficiently high heat flux is an important aspect of the present invention. The selection of various process factors may depend on the particulars of a given thermal chemical reaction; precise control of all factors is not required in all cases. Guided by the descriptions herein, those skilled in this area of ​​technology can, without undue experimentation, control these factors to obtain the desired level of flow.

The reactors and methods of the present invention can also be identified by their high production rates per reactive unit volume. For example, in a steady state process for steam reforming a gaseous or liquid hydrocarbon, the inventive process preferably produces more than 0.01 standard liters per minute (SLPM) of hydrogen gas per cubic centimeter (cc) of reactor hardware ( the sum of the volume of the reaction chamber (s) and the heat exchanger chamber (s) including the volume of the chamber walls), more preferably more than 0.1 SLPM of hydrogen gas per cc of reactor hardware. Typically, the production rate in the case of hydrogen can range up to about 0.5 SLPM or more of hydrogen gas per cubic centimeter of reactor hardware. In the most general case, the reactors and methods of the present invention can consume more than about 0.01 SLPM of reactive gas per cubic centimeter of reactor hardware. By building multiple reaction chambers and heat exchangers operating in parallel, the reactors and methods of the present invention can produce thousands or millions or more liters of product per day.

The invention also provides a method of suppressing unwanted chemical reactions, especially coke formation. The reactors and methods of the present invention can be characterized by their ability to suppress chemicals that are slower to form, kinetically, or that are more likely to form within a reactor that has a lesser degree of temperature control. For example, in the steady state process for a water gas exchange reactor, mechanization is a common side reaction that does not occur as rapidly as the reaction of interest, for the conversion of carbon monoxide to carbon dioxide. Likewise, coke is a slower forming by-product of steam reforming reactions, with the prevalence for the reaction also being a function of the catalyst surface temperature. In these cases, the reactors and methods of the present invention can withstand the primary reactions of interest (e.g., water gas exchange and steam reforming) while increasing the ability of the system to prevent or suppress the formation of certain undesirable by-products (eg, methane and coke). More generally, the reactors and methods of the present invention can be used to suppress other undesirable by-products where their formation can be prevented by short contact times and / or improved temperature control.

Example 1

An experiment was carried out to demonstrate a thermal chemical reactor using the water gas exchange reaction.

A first porous insert was made with a catalyst material of a stabilized prereduced catalyst of 5% by weight Ru / ZrO₂ (1/8-inch extruded) obtained from Degussa Corporation. The catalyst material was ground and sieved at 65-100 mesh.

ES 2 344 447 T3

A second porous insert was made from 80 pores per inch (ppi) Ni metallic foam machined to fit a 7 mm ID quartz tube, ranging from 0.5 to 2.5 cm in length. The metal foam was washed in an ultrasonic bath with acetone, chloroform, and water successively for 10 minute intervals. It was also deslustered in a 1M HCl solution at 60 ° C for 30 min. The tarnished metallic foam was saturated with a zirconium n-propoxide / 1-propanol solution (Aldrich), followed by ambient hydrolysis with steam for 72 h, and then calcined at 450 ° C for 4 h to form the stratum interfacial. ZrO Coated Metal Foam₂It was saturated with a dilute aqueous RuCb solution (RuCh hydrate, Aldrich). The saturation process was repeated several times until the desired Ru loading was achieved. The coated metal foam supported Ru catalyst was finally dried at 100 ° C under vacuum overnight, followed by calcination at 350 ° C for 1 hr. Before testing, the catalyst was activated with a mixture of 10% H₂/ He at 350 ° C for at least 1 h.

A catalytic plug flow reactor (PFR) system was used to test both porous inserts. The PFR was configured in a mono-zone furnace as the thermal transfer chamber. The reactor system included a steam generator positioned directly before the reactor inlet, a PFR housed within the furnace, and a condenser located at the reactor outlet. The porous insert was packed in a 7 mm ID quartz tube, which was tapered in the center. The feed water was fed into the steam generator using a Cole Parmer syringe pump. The system was fed with carbon monoxide and nitrogen (a diluent) using Matheson mass flow controllers. The mixed feed source flowed through the steam generator before entering the PFR in a downward flow manner. The product gases were routed through the condenser and sent to an online gas chromatograph, where the product vapor was analyzed.

Two thermocouples were placed within the catalytic PFR system. A thermocouple was located above the porous insert. The second thermocouple was placed adjacent to the porous insert outside the quartz tube to measure the temperature of the furnace. A pressure gauge at the inlet of the reactor was used to measure the differential pressure across the porous insert.

Product gases were analyzed immediately at the reactor outlet with a Microsensor Technology Inc., (MTI) M200 Gas Chromatograph. Using a 10-m molecular filter column (argon carrier gas, 100 ° C, 34. 1 psig (235 Pa)) and an 8-m PoraplotU column (helium carrier gas, 65 ° C, 26. 9 psig (185 Pa)) in parallel, the GC analyzes for hydrogen, nitrogen, oxygen, methane, carbon monoxide, air, carbon dioxide, ethane, and ethylene in 75 sec. The M200 used a vacuum pump to extract a small sample from the product flow stream with a 40-sec purge and a 100-millisecond injection time. Water was removed from the gas stream before entering the M200.

The carbon monoxide conversion was calculated based on the moles of material in the gas inlet and outlet stream, as shown in equation 1. The selectivity to carbon dioxide (and hydrogen) or methane was calculated in the equations 2 and 3, respectively.

Xco ⁼ 100 * (nco-in Dco-out) / RcO-in S co2 = 100 * nco2 / (n co2 ^{+ n}CH4)

Sch4 ~ 100 Sco2 (2) (D (3)

Using the first porous insert, the fine catalyst powders (65-100 mesh), the intrinsic reaction kinetics were roughly measured. Contact time ranged from 10 milliseconds to 1 sec. Figures 2a and 2b show the performance for long and short contact times. At 300 ° C and a 3: 1 ratio of steam to carbon, 25 milliseconds on the Ru-based catalyst was enough to convert more than 98% of the carbon monoxide to carbon dioxide and hydrogen. At 50 milliseconds, a CO conversion of 99.8% was measured with 100% selectivity to the desired products (CO₂ and H₂). The equilibrium conversion of CO at 300 ° C and a ratio of steam to carbon of 3: 1 was 99.93%.

Tests with longer contact times (> 100 milliseconds) showed the formation of methane, which has an equilibrium selectivity of 22.82%. The equilibrium selectivity line for carbon dioxide and hydrogen is shown in Fig. 2a. As contact time increased, methane formation also increased. A software package, FACT^TM, was used for all equilibrium calculations.

The results with the second porous insert (coated metal foam) are shown in Fig. 3. At 300 ° C, the CO conversion was less than 10%. However, at 500 ° C and a 3: 1 ratio of steam to carbon, the measured carbon monoxide conversion reached 94% with a contact time of 50 milliseconds. The equilibrium conversion was 94.53% under these conditions. With a contact time as short as 10 milliseconds, the carbon monoxide conversion exceeded 90% and 100% selectivity to carbon dioxide and hydrogen was observed. The selectivity of equilibrium C02 was 93.52% at 500 ° C.

ES 2 344 447 T3

At the 10, 50, and 100 millisecond contact times, the measured selectivity remained near 100%, with methane below the GC limit of detectability. These findings demonstrated that the desired non-equilibrium chemistry was exploited in the coated metallic foam. Unwanted series and parallel slow reaction trajectories, such as methane formation, were effectively stopped.

The second porous insert of the coated metal foam had a higher activation temperature than the first porous insert of catalyst powder for two reasons. First, the catalyst sealant coating had a slightly different composition and structure than the catalyst powders. Independent testing of the catalyst with powders made from the same sealer coating verified the required activation temperature. The other distinction between the two porous inserts was a reduced weight of active catalysts (approximately 10%) in the coated metal foam.

Example 2

An experiment was carried out to demonstrate hydrocarbon steam reforming.

Using the first porous insert (powder) as in Example 1, methane steam reforming was achieved with 100% conversion at 850 ° C in 25 milliseconds on a 5% Rh / gamma-Al catalyst.₂ OR₃(Fig. 4). Using the second porous insert (coated metal foam) as in example 1, with 5% Rh / Al catalyst₂OR₃/ interfacial layers over 80 ppi stainless steel metal foam, reduced operating temperature 100 ° C to achieve the same performance at 750 ° C.

No coke formation was observed during any of the millisecond residence time experiments with lower ratios of vapor to methane (2.5: 1).

The results for other hydrocarbons are shown in Table E2-1 where "time" is the residence time. Data on butane, gasoline and kerosene were obtained using a powder catalyst while data on isooctane were obtained using a foam catalyst.

TABLE E2-1

Preliminary hydrocarbon reforming data based on screening tests with 5% Rh / Ah catalyst

<td>Hydrocarbon</td><td>Temperature (c)</td><td>Time (ms)</td><td>% Conversion</td><td>% H2 Sel</td>

<td>Butane</td><td> 600</td><td> 25</td><td> 100</td><td> 96</td>

<td>Gasoline</td><td> 800</td><td> 50</td><td> 95</td><td> -95</td>

<td>lso-octane^TO</td><td> 700</td><td> 25</td><td> 100</td><td> 89.5</td>

<td>Kerosene</td><td> 600</td><td> 50</td><td> 95</td><td> -98</td>

^TOCatalyst material included a titania buffer layer

Additional data for an n-butane steam reforming experiment is shown in Fig. 5a. A porous insert consisted of 80 ppi of stainless steel with an interfacial layer of alumina and a catalyst material of rhodium (15.6 wt% Rh in 17.1 wt% alumina, equilibrium stainless steel foam, and no buffer layer) . Conditions were 650 ° C at a dwell time of 95 ms with a vapor to carbon ratio of 3.58: 1. The pressure drop increased from negligible to above 7 psig (48 kPa), attributed to cracking and detachment of the interfacial layer and catalyst. The catalyst was regenerated in air to remove the deposited carbon. Fig. 5b shows poorer performance. The pressure drop increased to above 7 psig (48 kPa) after only 5 operating hours in two days.

Example 3

A preferred, contemplated embodiment of the invention is shown in Fig. 6. The distance from the heat source to the heat sink is about 1 centimeter or less.

ES 2 344 447 T3

This distance is a function of the heat load, the selection of heat transfer fluid (s), and the effective thermal conductivity of the porous catalyst insert. The porous catalyst insert can have a porosity greater than 95%, which creates an effective thermal conductivity nearly two orders of magnitude lower than the pure or alloyed metal that forms the porous support.

Thin sheets or tubes can be used to obtain high thermal loads and short contact times. The thickness of the tissue between the reaction channel and the heat exchange channel can vary, but is preferably between about 0.01 inches (0.25 mm) and about 0.25 inches (6.4 mm). The preferred thickness for the heat exchange channel preferably ranges from 100 microns to 10 millimeters. The preferred thickness is 250 microns to 3 millimeters. The flow of the heat transfer fluid can be either countercurrent, crosscurrent, or current parallel to the direction of the reagent flow.

The thickness for the catalyst preferably ranges from 100 microns to 10 millimeters. The preferred thickness is 250 microns to 1 millimeter. The catalyst can be composed of a single contiguous monolithic porous catalyst, or it can be created by placing multiple porous monoliths contiguous with each other. Porous monoliths can also be inserted with a gap between the porous monolith catalysts, or with a smaller monolith (as shown in Fig. 7) contiguous to and / or between larger monoliths. Preferably, one or more monoliths of equal dimensions are contiguous with each other; This design option will favor the contact of the reactants with the catalyst surface.

The volumetric heat flux using short contact time reactions within a device that can facilitate a high heat load should easily exceed 0.6 W / cc. Thermal loads in the 1 to 10 W / cc range have been shown to be well within the range of this catalyst and reactor. Higher thermal loads, up to 100 W / cc, can also be achieved if the contact time is less than 25 milliseconds and if the distance between the heat source and the heat sink is on the order of one millimeter (approximately 0.5 to about 5 mm).

Example 4

A compact reactor that transferred approximately 10 W / cc to 16 W / cc was demonstrated for isooctane steam reforming. The steam reforming reaction uses steam to transform hydrocarbons into CO and H₂. For the case of steam reforming of isooctane, the stoichiometry of the reaction is:

<img file="ES2344447T3_D0001.tif" />

and the standard heat of reaction is strongly endothermic (AH_r ° = + 1275 kJ / mol isooctane). Isooctane was chosen to simulate gasoline because it represents the mid-range molecular weight of the gasoline hydrocarbon mixture. The steam reforming reaction is often run at higher steam concentrations than the reaction stoichiometric shown above to avoid coke formation and to improve conversion. Vapor concentration is usually given in the form of the ratio of vapor to carbon (S: C), the ratio of vapor molecules per carbon atom in reactive food. A ratio greater than one indicates a higher vapor concentration than stoichiometric, and under these conditions the water gas shift reaction is also carried out as a side reaction.

A microchannel isooctane steam reformer was built, with a total volume of practically 30 cubic centimeters. This reactor has integrated heat exchange to add the heat needed for this highly endothermic reaction. The heat exchange fluid used to heat the reactor zone was flue gas set at 750 ° C and a total flow rate of approximately 100 SLPM. This design was built to process a stream of hydrogen gas large enough to produce 1.0 kilowatt of equivalent electrical power (kW_and) in a PEM fuel cell. The reactor configuration is shown in Fig. 6. There were 12 reactive channels interspersed among 13 heat exchange channels. The reaction channels had a thickness of approximately 0.03 inch (0.76 mm), a length of 1.1 inch (28 mm), and a channel height of 1 inch (25 mm). The heat exchange channels were identical in length and height. The thickness of the heat exchange channels was approximately 0.02 inches (0.50 mm). The tissue thickness between the reaction and the heat exchange channel was approximately 0.19 inches (4.8 mm).

The catalyst was 15% by weight of Rh₂OR₃ in an aluminum-containing spinel encased in a porous stainless steel foam. The specific details of the catalyst preparation are as follows. The Gamma alumina support (Strem) was calcined at 500 ° C for 5 hours. The incipient moisture method was used to impregnate the magnesium nitrate solution on the alumina support to achieve 5% by weight of MgO. The modified support was dried at 110 ° C under vacuum for 4 hours followed by calcination at 900 ° C for 2 hours to form the aluminum-containing spinel support. The spinel support was impregnated with a rhodium nitrate solution (Engelhard) using the incipient moisture technique to achieve Rh loading.₂OR₃ desired. After drying at 110 ° C under vacuum for 4 hours, the supported Rh powder catalyst was calcined at 500 ° C for 3 hours. Powdered catalyst was ball milled overnight and dip coated

ES 2 344 447 T3 in an 80 ppi (346 ppcm) stainless steel foam (Astromet, Cincinnati, Ohio). Prior to dip coating, the stainless steel foam was coated with the alumina-titania buffer layer using CVD.

The data of the first hour of operation of the steam reforming of the isooctane in microchannels are shown in Fig. 9. The liquid isooctane inflow rate was 2.54 ml / min. The reactive zone was operated at 650 ° C and one atmosphere. The reactants had a 6: 1 vapor to carbon ratio, which created a total contact time of approximately 22 milliseconds within the total sum of the reaction channels. The reactor was able to achieve isooctane conversions ranging from 86.5% to 95%, thus requiring approximately 300 W of thermal energy. The hydrogen selectivity was in the range of 85 to 90%. The results show that the microchannel isooctane steam reforming reactor can supply the heat necessary to sustain this reaction with a high throughput rate per unit volume of the reactor. The reactive secondary pressure drop across this device was about 6.9 kPa (1.0 psi). The expected fuel cell output under these conditions was sufficient to produce a 0.5-kW electrical output._and of a PEM fuel cell. This demonstration continued for another hour and then stopped. The volumetric heat flux of the reactor was around 10 W / cc.

Three other demonstrations were run using this device, passing the device through four thermal cycles. These four thermal cycles put the device through more than 12 hours in total online service. The results of the final three thermal cycles in the form of a bar graph in Fig. 10. All results in Fig. 10 they are at one atmosphere, a temperature range of 630 to 670 ° C and a contact time of 22 milliseconds within the sum of the total reaction channels. The results for the equivalent 0.5-kW power output_and and the 6: 1 ratio of steam to carbon for these tests are consistent with the values ​​shown in Fig. 9. As the ratio of steam to carbon is decreased from 6: 1 to 5.7: 1, 5: 1, 4.06: 1, and finally at 2.98: 1, the isooctane conversion is reduced, but the hydrogen selectivity remains stable. At a constant residence time, the effect of reducing the vapor to carbon ratio is the same as increasing the isooctane flow rate. So, as the percentage of conversion decreases with the reduction of the ratio of steam to carbon, the amount of isooctane being converted increases, resulting in a net increase in the actual level at which hydrogen is generated. This is seen in the highest equivalent electrical power output listed with these values. The last set of bar data at the far right of Figure 10 are the results acquired with a vapor to carbon ratio of 5.7: 1 and about half the originating contact time, resulting in an output of 1.0-kW equivalent power_and. Under these conditions, nearly 500 W of thermal energy was required to convert about 75% of the isooctane vapor input stream at 5.04 ml / min. This device demonstrated a volumetric heat flux greater than 16 W / cc.

Conclution

While a preferred embodiment of the present invention has been shown and described, it will be apparent to those skilled in the art that many changes and modifications can be made without departing from the invention in its broadest aspects as defined by the appended claims.

References cited in description

This list of references cited by the applicant has been compiled exclusively for the information of the reader. It is not part of the European patent document. It has been made with the greatest diligence; However, the EPO does not assume any responsibility for eventual errors or omissions.

Patent bibliography cited in the description • WO 0006295 A [0007] [0007] [0008] • US 5611214 A [0038] • US 5811062 A [0038] • US 5534328 A [0038] • US 883643 A [0038] • US 081938228 B [0038] • US 09375610 B [0038] • US 09123781 B [0038] • US 492950 A [0038]

ES 2 344 447 T3 • US 09375614 B, 1999 [0038] • US 09265227 B, 1999 [0038]

Bibliography of non-patents cited in the description • Adris, A .; Pruden, B .; Lim, C .; J. Grace. On the reported attempts to radically improve the performance of the steam methane reforming reactor. Canadian Journal of Chemical Engineering, 1996, vol. 74, 177-186 [0005]