Classifications

• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
  • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
  • B01J8/24—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique
  • B01J8/26—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations
  • B01J8/28—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles according to "fluidised-bed" technique with two or more fluidised beds, e.g. reactor and regeneration installations the one above the other
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
  • B01J8/005—Separating solid material from the gas/liquid stream
  • B01J8/0055—Separating solid material from the gas/liquid stream using cyclones
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
  • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
  • B01J8/1818—Feeding of the fluidising gas
• • B—PERFORMING OPERATIONS; TRANSPORTING
  • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
  • B01J—CHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
  • B01J8/00—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
  • B01J8/18—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
  • B01J8/1845—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with particles moving upwards while fluidised
  • B01J8/1854—Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with particles moving upwards while fluidised followed by a downward movement inside the reactor to form a loop
• • C—CHEMISTRY; METALLURGY
  • C07—ORGANIC CHEMISTRY
  • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
  • C07C5/00—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms
  • C07C5/32—Preparation of hydrocarbons from hydrocarbons containing the same number of carbon atoms by dehydrogenation with formation of free hydrogen
  • C07C5/327—Formation of non-aromatic carbon-to-carbon double bonds only
  • C07C5/333—Catalytic processes
  • C07C5/3332—Catalytic processes with metal oxides or metal sulfides
• • C—CHEMISTRY; METALLURGY
  • C07—ORGANIC CHEMISTRY
  • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
  • C07C2523/00—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00
  • C07C2523/70—Catalysts comprising metals or metal oxides or hydroxides, not provided for in group C07C2521/00 of the iron group metals or copper
  • C07C2523/74—Iron group metals
  • C07C2523/745—Iron
• • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
  • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
  • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
  • Y02P20/00—Technologies relating to chemical industry
  • Y02P20/50—Improvements relating to the production of bulk chemicals
  • Y02P20/582—Recycling of unreacted starting or intermediate materials
• • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
  • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
  • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
  • Y02P20/00—Technologies relating to chemical industry
  • Y02P20/50—Improvements relating to the production of bulk chemicals
  • Y02P20/584—Recycling of catalysts

Definitions

• this invention pertains to a process of dehydrogenating an alkyl aromatic compound, such as ethylbenzene, to form a vinyl aromatic compound, such as styrene. Additionally, the invention pertains to a process of regenerating a catalyst which is 0 used in the dehydrogenation of the alkyl aromatic compound. In another aspect, this invention pertains to a fluidized bed reactor in which an organic process, such as the aforementioned dehydrogenation process, is conducted.
• alkyl aromatic compounds such as ethylbenzene, isopropylbenzene, diethylbenzene, and p-ethyltoluene
• alkyl aromatic compounds such as ethylbenzene, isopropylbenzene, diethylbenzene, and p-ethyltoluene
• styrene substituted derivatives of styrene, such as ⁇ -methylstyrene, divinylbenzene, and p-methylstyrene.
• Styrene and its substituted derivatives were useful as monomers in the formation of polystyrenes, styrene-butadiene rubbers (SBR), acrylonitrile-butadiene-styrene (ABS), styrene-acrylonitrile (SAN), and unsaturated polyester resins.
• SBR styrene-butadiene rubbers
• ABS acrylonitrile-butadiene-styrene
• SAN styrene-acrylonitrile
• Fluidized bed reactors are important in a wide variety of catalyzed organic o processes, including dehydrogenation processes.
• the chief manufacturing route to vinyl aromatic compounds, such as styrene, is the direct catalytic dehydrogenation of an alkyl aromatic compound, such as ethylbenzene.
• Patents disclosing such a process include, for example, US 4,404,123, US 5,171,914, US 5,510,552, and US 5,679,878.
• the catalyst typically comprises iron oxide 5 and, additionally, may comprise chromium oxide and potassium compounds, such as potassium hydroxide or potassium carbonate, as promoters. Since the process is highly endothermic, energy for the process is obtained by introducing superheated steam into the process reactor, which in the aforementioned disclosures is typically a fixed bed design. Process temperatures are generally between about 550°C and 700°C. Side reactions can be o controlled by maintaining a low partial pressure of ethylbenzene.
• a fixed bed reactor characterized by a stationary bed of catalyst particles, is difficult to heat uniformly to a high temperature.
• the downstream (exit) end of the fixed catalyst bed tends to be colder than the upstream (entrance) end of the catalyst bed.
• the feedstream is typically preheated to a temperature higher 5 than desirable.
• the catalyst bed near the entrance to the reactor may be overheated and typically deteriorates faster than the catalyst farther downstream.
• references such as US 3,651,146 and US 4,471,146, disclose oxidative dehydrogenation processes wherein ethylbenzene is contacted with oxygen in a fluidized bed reactor in the presence of an oxidative dehydrogenation catalyst, for example, an o alkaline earth-nickel phosphate or an alkali metal-chromium oxide composite, to produce styrene.
• a conventional fluidized bed reactor comprises a single reaction zone wherein catalyst particles are disengaged and circulating.
• fluidized bed reactors provide a more isothermal temperature distribution.
• An isothermal catalyst bed is typically less damaging to the catalyst and allows for better product yields.
• Fluidized 5 beds also allow for easy replacement of the catalyst when it is fully deactivated and incapable of regeneration.
• deactivated catalyst can be discharged from the reactor and active catalyst can be added to the reactor without shutting down the chemical process.
• the aforementioned oxidative dehydrogenation process may also be conducted while continuously transporting a portion of the catalyst to a regenerator o for regeneration under oxygen, and then recirculating the regenerated catalyst back to the oxidative dehydrogenation reactor.
• oxidative dehydrogenation processes with co-feeds of alkyl aromatic compound and oxygen may produce low yields of vinyl aromatic product, because oxidation side reactions are more difficult to control.
• safety issues involved with handling and processing mixtures of organic compounds and oxygen are significant.
• the continuous cycling of catalyst between the fluidized bed process reactor and the regenerator requires complex 5 equipment and often requires a highly attrition resistant catalyst particle.
• US 4,152,393 discloses a reactor consisting of a single shell that contains a reaction zone and a regeneration zone arranged in such a manner, specifically, as a collection of concentric walls and paths, that a particulate solid may be o transferred by flow of gases from the regeneration zone to the reaction zone by a first route and then back to the regeneration zone by a second route.
• the gases passing through the regeneration zone are not transferred to the reaction zone, and the gases passing through the reaction zone are not transferred to the regeneration zone. It is taught that this reactor is useful for the ammoxidation of propylene.
• the disclosed reactor may exhibit high slug flow, characterized by gas bubbles flowing along the inner walls of the reactor.
• this invention is a process of dehydrogenating an alkyl aromatic compound over a dehydrogenation catalyst in a single shell, fluidized bed reactor to form a vinyl aromatic compound, and regenerating in situ the dehydrogenation catalyst.
• the process of this invention comprises, in a first step, (a) fluidizing a dehydrogenation catalyst in a single shell, fluidized bed reactor containing a reaction zone and a regeneration zone under fluidization conditions such that the catalyst is circulated within and between the two zones; (b) contacting an alkyl aromatic compound, and optionally, steam with the dehydrogenation catalyst residing in the reaction zone under reaction conditions sufficient to o prepare the corresponding vinyl aromatic compound; and (c) contacting steam with the dehydrogenation catalyst residing in the regeneration zone, the contacting being conducted under regeneration conditions sufficient to regenerate, at least in part, the catalyst.
• the dehydrogenation process of this invention which finds utility in the preparation of vinyl aromatic compounds of industrial significance, such as styrene, p- 5 methylstyrene, ⁇ -methylstyrene, and divinylbenzene, possesses significant advantages over prior art processes.
• the process of this invention does not employ oxygen. Accordingly, safety problems associated with handling and processing mixtures of organic compounds and oxygen, which are found in some prior art processes, are eliminated in the process of this invention.
• the steam to oil ratio is o advantageously lower than that used in prior art fixed bed processes. Accordingly, the process of this invention uses a smaller water recycle and is more energy efficient and economical than prior art processes.
• the process of this invention being conducted in a fluidized bed reactor, is essentially isothermal. Problems associated with non-uniform bed temperatures, such as, overheating and catalyst damage at the upstream end of the catalyst bed and lowered productivity at the downstream end of the catalyst bed, are essentially eliminated. Also, thermal byproduct formation is reduced. As a further advantage, a smaller catalyst bed may be used in the process of this invention, as compared with the size of the catalyst bed used in fixed bed prior art processes, while still achieving comparable running cycles. As another advantage, the process of this invention provides for the continuous in situ regeneration of the dehydrogenation catalyst. There is no necessity in the process of this invention to shut the reactor down for catalyst regeneration or to transport the catalyst out of the fluidized bed reactor to a regenerator.
• the process of this invention possesses a simplicity of design and operation. Moreover, the catalyst used in the process of this invention may not require the high attrition resistance needed for transport reactors. Finally, when the catalyst is incapable of further regeneration, the process of this invention provides for the replacement of deactivated catalyst during continuous operation of the dehydrogenation process. Since the solid catalyst is treated as a fluid, the deactivated catalyst is simply conveyed out of the reactor, and fresh catalyst is conveyed into the reactor, during operation. Thus, regeneration and replacement of the catalyst can both be accomplished without shutting down the dehydrogenation process, thereby resulting in higher productivity. Most advantageously, the process of this invention produces vinyl aromatic compound, preferably styrene, in high yield.
• this invention is a fluidized bed reactor which allows for chemical processing and catalyst regeneration simultaneously.
• the fluidized bed reactor of this invention comprises a single, vertical shell within which the space is divided into a freeboard zone, a reaction zone, and a regeneration zone.
• the reactor also comprises an inlet means for introducing a regeneration feedstream into the regeneration zone, and an inlet means for introducing a reactant feedstream into the reaction zone.
• the reactor further comprises a means for separating the reaction and regeneration zones while allowing for the continuous circulation and large scale backmixing of the catalyst between the two zones.
• one of the inlet means for the reaction or regeneration feed acts as the means for separating the reaction and regeneration zones.
• the reactor of this invention also comprises an outlet means, preferably in the freeboard zone, for removing an effluent stream containing products and any unconverted reactants and regeneration feeds.
• the reactor further contains a means for returning catalyst particles entrained in the effluent stream back to the reactor.
• an inlet means and an outlet means may be present for conveying catalyst into and out of the reactor.
• cross-mixing of the reactant and regeneration feedstreams may occur; although, preferably, the placement of the reaction and regeneration inlets will substantially separate both processes. Nevertheless, with backmixing of gases and solids being allowed in this design, the regeneration and reactant feedstreams should be chemically compatible, as illustrated herein with the instant dehydrogenation-regeneration process.
• the fluidized bed reactor of this invention can be used in a variety of catalyzed organic processes, including, for example, dehydrogenations, oxidations, and halogenations.
• a particularly important dehydrogenation process for which the fluidized bed reactor of this invention can be used, comprises the dehydrogenation of an alkyl aromatic compound, such as ethylbenzene, to a vinyl aromatic compound, such as styrene.
• the fluidized bed reactor of this invention provides a chemical process reaction zone and a catalyst regeneration zone within a single fluidized bed reactor shell. Accordingly, the in situ regeneration of the catalyst can be achieved simultaneously with the desired chemical process.
• the deactivated catalyst need not be transported out of the fluidized bed reactor of this invention to a separate vessel for regeneration; thus, the catalyst is subjected to far less stress and damage than that experienced in transport reactors.
• deactivated catalyst can be replaced on-line without shutting down the chemical process, by simply conveying deactivated catalyst out of the reactor and conveying fresh catalyst into the reactor.
• FIG. 1 shows cross-sectional side and top views of a first preferred embodiment of the fluidized bed reactor of this invention, details of which are set forth hereinafter.
• FIG. 2 shows cross-sectional side and top views of a second preferred embodiment of the fluidized bed reactor of this invention, details of which are set forth hereinafter.
• FIG. 3 is a plot of ethylbenzene conversion and styrene selectivity as a function of run time for an ethylbenzene dehydrogenation and catalyst regeneration process conducted in a pulsed mode reactor.
• this invention is a process of dehydrogenating an alkyl aromatic compound over a dehydrogenation catalyst in a single shell, fluidized bed reactor to form a vinyl aromatic compound, and regenerating in situ the dehydrogenation catalyst.
• the process of this invention comprises (a) fluidizing a dehydrogenation catalyst in a single shell, fluidized bed reactor containing a reaction zone and a regeneration zone under 0 fluidization conditions such that the catalyst is circulated within and between the two zones.
• the process comprises (b) contacting an alkyl aromatic compound, and optionally, steam with the dehydrogenation catalyst residing in the reaction zone under reaction conditions sufficient to prepare the corresponding vinyl aromatic compound.
• a third step which is conducted 5 preferably simultaneously with the first and second steps, the process comprises (c) contacting steam with the dehydrogenation catalyst residing in the regeneration zone, the contacting being conducted under regeneration conditions sufficient to regenerate, at least in part, the catalyst.
• a o portion of the catalyst will be circulating in the reaction zone, while essentially the remainder of the catalyst will be circulating in the regeneration zone, with some co-mixing at the boundary of the two zones.
• catalyst residing in the reaction zone will lose activity and become partially or fully deactivated. Deactivation may be mostly caused by a build-up of coke on the surface of the catalyst. (The invention should 5 not be bound or limited, however, by such a deactivation theory.)
• catalyst in the reaction zone including deactivated catalyst, will circulate into the regeneration zone. Deactivated catalyst residing in the regeneration zone will be reactivated by contact with steam.
• Reactivation results in partial or essentially full recovery of catalyst activity, as compared with the activity of the fresh (unused or "as-synthesized") o catalyst. Thereafter under the fluidization conditions, the regenerated catalyst in the regeneration zone will be recycled to the reaction zone, and the reaction/-regeneration cycle will be repeated again and again.
• the aforementioned description is given as a means of explaining the reaction-regeneration cycle and of defining the words "to regenerate, at least in part, the catalyst.”
• the deactivated catalyst can be replaced simply by conveying it out of the reactor and conveying fresh catalyst into the reactor, preferably simultaneously.
• catalyst replacement can be conducted "on-line" without shutting down the catalytic process, more particularly to this invention, the dehydrogenation process.
• the catalyst can be conveyed into and out of the reactor via air jets or a pneumatic transport loop.
• catalyst can be removed from the reactor via a gravity-driven outlet means at the bottom of the reactor, and catalyst can be added to the reactor from a standpipe inlet means at the top of the reactor.
• the alkyl aromatic compound is ethylbenzene or a substituted derivative of ethylbenzene
• the vinyl aromatic compound produced is styrene or a substituted derivative of styrene
• Any alkyl aromatic compound can be employed in the dehydrogenation process of this invention, provided that the product achieved is a vinyl aromatic compound.
• the aromatic moiety of the vinyl aromatic compound can comprise, for example, a monocyclic aromatic ring, such as phenyl; a fused aromatic ring, such as naphthyl; or a ring assembly, such as biphenylyl.
• the aromatic moiety is a monocyclic aromatic ring, more preferably, phenyl.
• the alkyl moiety of the alkyl aromatic compound can comprise, for example, any saturated straight chain, branched, or cyclic hydrocarbon radical, provided that it can be dehydrogenated in the process of this invention to a vinyl moiety.
• Non-limiting examples of suitable alkyl moieties include ethyl, n-propyl, iso-propyl, n- butyl, iso-butyl, t-butyl, and higher homologues thereof.
• the alkyl moiety is a C2-C1 o alkyl, more preferably, a C2-C5 alkyl, most preferably, ethyl.
• the alkyl aromatic compound may be substituted optionally with two or more alkyl moieties, or substituted with other types of substituents which are essentially inert with respect to the dehydrogenation process of this invention.
• the alkyl aromatic compound is a C -C20 alkyl aromatic compound, more preferably, a Cg-Ci 5 alkyl aromatic compound, and most preferably, ethylbenzene or a substituted derivative thereof.
• the regeneration feedstream typically comprises steam.
• steam may also be incorporated into the dehydrogenation feedstream.
• Any weight ratio of steam to alkyl aromatic compound (steam to oil ratio) is suitable for the process of this invention, provided that the process produces a vinyl aromatic compound.
• the steam to oil ratio is based on the total weight of the steam introduced into the reactor from all sources, including steam from both the dehydrogenation and the regeneration feedstreams.
• the steam to oil weight ratio is greater than about 0.2/1, preferably, greater than about 0.5/1.
• the steam to oil weight ratio is less than about 5.0/1, preferably, less than about 3.0/1, even more preferably, less than about 1.2/1 , and most preferably, less than about 1.0/1.
• the process of this invention operates at lower steam to oil ratios, as compared with prior art processes.
• a low steam to oil ratio advantageously reduces the energy requirement and cost of converting water to steam and reduces the quantity of water recycled to the reactor.
• a sweeping gas may be used in the process of this invention.
• the sweeping gas which may be introduced directly into the freeboard zone of the reactor, primarily functions to remove the product stream from the freeboard zone, where undesirable thermal reactions can occur.
• Any gas which is substantially inert with respect to the dehydrogenation and regeneration processes may be suitably employed as the sweeping gas, including, for example, nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof.
• the concentration of sweeping gas in the freeboard zone can be any concentration, provided that the overall process produces the desired vinyl aromatic compound.
• the concentration of sweeping gas varies depending, for example, upon the specific alkyl aromatic compound and specific process conditions employed, particularly, the temperature and gas velocity.
• the concentration of sweeping gas in the freeboard zone is greater than about 10 volume percent, and preferably, greater than about 20 volume percent. Typically, the concentration of sweeping gas in the freeboard zone is less than about 90 volume percent, and preferably, less than about 70 volume percent.
• the dehydrogenation and/or regeneration feedstream(s) may also contain a diluent.
• the diluent primarily dilutes the reactants and products for improved selectivity or safety considerations. Any gas which is substantially inert with respect to the dehydrogenation and regeneration steps may be suitably employed as the diluent, including, for example, nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof.
• the concentration of diluent in either the dehydrogenation or regeneration feedstream can be any concentration, provided that the overall process produces the desired vinyl aromatic compound.
• concentration of diluent varies depending, for example, upon the specific diluent chosen, the specific alkyl aromatic compound, the specific dehydrogenation or regeneration process conditions, and the specific catalyst and its deactivation properties.
• the concentration of diluent in the dehydrogenation or the regeneration feedstream is greater than about 10 volume percent, preferably, greater than about 20 volume percent.
• the concentration of diluent in either stream is less than about 90 volume percent, preferably, less than about 70 volume percent.
• Oxygen is not required for the process of this invention. Preferably, oxygen is not employed in the process of this invention.
• dehydrogenation catalyst which is capable of catalyzing the dehydrogenation of an alkyl aromatic compound to a vinyl aromatic compound can be used in the process of this invention.
• dehydrogenation catalysts which can be beneficially employed include the catalysts described in the following US patents: US 4,404,123, US 4,503,163, US 4,684,619, US 5,171,914, US 5,376,613, US 5,510,552, and US 5,679,878, which pertain to a variety of iron oxide catalysts, containing, for example, one or more compounds of the alkali metals, preferably, sodium, potassium, and cesium; alkaline earth metals, preferably, calcium; and/or cerium, chromium, zinc, copper, and/or gallium compounds, as well as, the catalyst described in US 3,651,160, which pertains to chromium oxide and alkali metal oxides.
• the catalyst is a dehydrogenation catalyst comprising iron oxide. More preferably, the catalyst comprises (a) at least one iron oxide, (b) at least one carbonate, bicarbonate, oxide or hydroxide of potassium and/or cesium, (c) an oxide, carbonate, nitrate or hydroxide of cerium, (d) optionally, a hydroxide, carbonate, bicarbonate, acetate, oxalate, nitrate, or sulfate of sodium, (e) optionally, a carbonate, sulfate, or hydroxide of calcium, and (f) optionally, one or more binding agents, such as, a hydraulic cement.
• the catalyst comprises (a) at least one iron oxide, (b) at least one carbonate, bicarbonate, oxide or hydroxide of potassium and/or cesium, (c) an oxide, carbonate, nitrate or hydroxide of cerium, (d) optionally, a hydroxide, carbonate, bicarbonate, acetate, oxa
• the more preferred catalyst may additionally comprise one or more oxides selected from zinc, chromium, and copper.
• the more preferred catalyst comprises from 25 to 60 weight percent iron, 5 from 13 to 48 weight percent potassium, and from 1 to 20 weight percent cerium, the weight percentages being calculated as the oxides.
• the dehydrogenation catalyst which is used in the fluidized bed reactor of this invention can possess any particle size or shape, so long as the catalyst is capable of o catalyzing the dehydrogenation of an alkyl aromatic compound to a vinyl aromatic compound.
• the average catalyst particle size is greater than about 20 microns( ⁇ m) in diameter (or cross- sectional dimension), and preferably, greater than about 50 ⁇ m in diameter.
• the average particle size is less than about 1 ,000 ⁇ m, and preferably, less than about 200 ⁇ m.
• the catalyst particle is smooth with rounded 5 edges, is substantially non-cohesive, and possesses an attrition resistance sufficient for use in a fluidized bed reactor.
• One skilled in the art will know whether a particular catalyst has sufficient attrition resistance for use in a fluidized bed reactor.
• the dehydrogenation feedstream may be preheated prior to its introduction into the reaction zone.
• the preheat can be conveniently supplied by o condensing high pressure saturated steam, or alternatively, by combusting a fuel source or process off-gas. Any preheat temperature can be used, provided it lies below the temperature at which thermal cracking of the alkyl aromatic compound becomes measurable. Typical preheat temperatures are greater than about 150°C, preferably, greater than about 250°C, and more preferably, greater than about 350°C. Typical preheat 5 temperatures are less than about 600°C, and preferably, less than about 590°C.
• the regeneration feedstream can be preheated prior to its introduction into the regeneration zone.
• a typical preheat temperature for the regeneration stream is greater than about 200°C, preferably, greater than about 300°C, and more preferably, greater than about 400°C.
• the preheat temperature of the regeneration feedstream is typically less than about 650°C, and o preferably, less than about 630°C.
• a preferred embodiment of the novel reactor which is employed in the process of this invention is the fluidized bed reactor shown in Fig. 1, comprising a single vertical shell whose inner space is functionally divided into a regeneration zone (1), a reaction zone (2), and a freeboard zone (3).
• the regeneration zone located in this preferred embodiment at the bottom of the reactor, comprises the region wherein catalyst is regenerated.
• the reaction zone, situated in this preferred embodiment in the middle zone of 5 the reactor, is the region wherein the catalyzed organic chemical process occurs, such as the dehydrogenation process described herein.
• the freeboard zone located at the top of the reactor, comprises the space above the middle zone up to the top inner wall of the reactor.
• the freeboard zone occupied by gaseous reactants and products, also provides space for expansion of the fluidized bed. Gas phase thermal reactions may occur in the freeboard 0 zone; but process conditions are preferably maintained which minimize these gas phase reactions relative to the catalyzed process occurring in the reaction zone.
• the regeneration zone of Fig. 1 contains an inlet means for introducing the regeneration feedstream, herein steam and, optionally, a diluent into the regeneration zone.
• the inlet means can comprise, for example, a inlet port (4), which exits into a plenum (10) 5 above which a distributor plate or sparger array (9) is situated.
• the freeboard zone contains an inlet means (5) comprising, for example, an inlet port and delivery tube, for introducing the reactant feedstream, herein the dehydrogenation feedstream, into the reaction zone.
• the inlet means for the regeneration zone is shown at the bottom of the figure, and the inlet means for the reaction zone is shown at the top of the figure.
• the o regeneration zone inlet means can be situated anywhere, provided that it exits in the regeneration zone.
• the reaction zone inlet means can be situated anywhere, so long as the reactants are fed into the reaction zone.
• the inlet means (5) terminates in a sparger array or distributor (6) in the reaction zone, located more preferably, at a level above the bottom steam distributor 5 (9).
• the distributor or sparger array (6) is preferably designed to deliver the reaction feedstream in any direction into the reaction zone.
• the distance from the bottom steam distributor (9) to the reactant feedstream distributor (6) can be varied to provide for variable volumes in the regeneration and reaction zones, as desired. The larger the zone, the longer will be the residence time of gases and solids in that zone.
• the o distributor means (6) provides a functional demarcation between the regeneration zone and the reaction zone, such that the regeneration process substantially occurs in the regeneration zone while the catalyzed organic process substantially occurs in the reaction zone, while still allowing for backmixing of solids and gases.
• the gas distributors and sparger arrays can be manufactured, for example, from gas permeable sintered metal, or more preferably, the gas distributors and spargers can be fitted with jets for dispersing the gas.
• the freeboard zone also contains an outlet means (7), such as an outlet port, for the effluent stream, comprising 5 unconverted alkyl aromatic compound, steam, the optional sweeping gas and/or diluent, and products, including the vinyl aromatic compound.
• the outlet means (7) may be connected to a cyclone (shown in Fig. 1 , below outlet 7) for collecting catalyst particles which are entrained with the effluent stream.
• the collected catalyst particles can be recycled to the fluidized bed reactor via inlet means (8), situated at any point along the reactor, but o preferably, as shown in Fig. 1 into the regeneration zone.
• the outlet means (7) is further connected to a separations unit (not shown in Fig. 1), including for example, a condensation means and a distillation train for separating unconverted alkyl aromatic compound and reaction products. Unconverted reactants may be recycled back to the reaction zone via inlet (5).
• the reactor may further comprise a means for measuring 5 the temperature of the catalyst bed and, optionally, a means for heating the reactor (not shown in figures).
• the reaction zone and regeneration zones may also contain baffles (not shown in figures) which function to reduce the formation and size of bubbles, thereby facilitating contact between the gaseous feedstreams and the catalyst.
• the fluidized bed reactor o additionally contains one or more means for enhancing solids circulation and heat transfer.
• the means for enhancing solids circulation comprises one or more draft tubes, optionally, containing internal baffles.
• the means for enhancing solids circulation comprises one or more draft tubes made of heating or cooling elements.
• An embodiment of the invention containing a plurality of draft tubes is shown in 5 Fig 2. (Parts 1-10 of Fig. 2 are identical to parts 1-10 of Fig. 1.)
• the draft tubes (11) may each comprise, for example, concentric cylinders open at both ends, or a bundle or array of heating tubes, or any other design which promotes drafting of the catalyst.
• the draft tube is vertically suspended through the reaction and regeneration zones, preferably, to near the top of the reaction zone.
• the dehydrogenation feedstream is fed through the inlet port (5) into the sparger array (6) and up the inner cylinder of the draft tube (11) into the reaction zone.
• catalyst particles will become entrained in the inner cylinder of the draft tube and transported up to the top of the draft tube. At the top, the catalyst particles will flow over the lip of the inner cylinder and down through the annular region between the two cylinders back into the regeneration zone.
• the reactor may optionally comprise an inlet means and an outlet means (not shown in the figures) for respectively transporting catalyst into and 5 out of the reactor.
• reaction and regeneration zones can be reversed, such that the reaction zone is located at the bottom of the reactor while the regeneration zone is located in the middle of the reactor. (Fig. 1 wherein the reaction zone is located at (1), the regeneration zone is located at (2), and the o inlets thereto are adjusted accordingly.)
• the reactor of this invention can be employed in catalytic processes wherein the reactant feedstream is chemically compatible with the regeneration feedstream.
• the unique reactor of this invention allows for the continuous flow of catalyst particles between reaction and regeneration zones within a single shell fluidized bed reactor. Accordingly, the 5 catalyzed organic process of interest and the catalyst regeneration can be effected simultaneously without transporting catalyst out of the reactor into a separate regenerator.
• the reactor of this invention does not contain complex concentric walls or winding paths through which the catalyst particles traverse. Accordingly, for large scale units, the reactor of this invention does not produce significant slug flow and attrition problems.
• the temperature of the reaction zone wherein the dehydrogenation process occurs can be any operable temperature, provided that a vinyl aromatic compound is produced in the process.
• the operable dehydrogenation temperature will vary with the specific catalyst and specific alkyl aromatic compound employed.
• the dehydrogenation temperature is typically greater than about 5 550°C, and preferably, greater than about 570°C.
• the dehydrogenation temperature is less than about 650°C and, preferably, less than about 610°C. Below about 550°C, the conversion of alkyl aromatic compound may be too low; whereas above about 650°C, thermal cracking of the alkyl aromatic compound and vinyl aromatic product may occur.
• temperature is measured on the catalyst bed in fluidized form. 0 In the regeneration zone, the catalyst is contacted with steam and reactivated.
• the temperature of the regeneration zone can also be varied, so long as the catalyst is at least partially regenerated.
• the regeneration temperature lies below the thermal cracking temperature for the alkyl aromatic reactant and vinyl aromatic product.
• the regeneration temperature is typically greater than about 550°C, and preferably, greater than about 570°C.
• the regeneration temperature is less than about 650°C and, preferably, less than about 610°C. Since the catalyst is recirculating continuously between the reaction and regeneration zones and since the temperatures of the two zones are maintained at closely similar values, the fluidized bed is substantially isothermal throughout both zones.
• the process can be conducted at any operable total pressure, ranging from subatmospheric to superatmospheric, provided that the vinyl aromatic product is produced. If the total reactor pressure is too high, the equilibrium position of the dehydrogenation process may be shifted backwards towards alkyl aromatic compound. On the other hand, an adequate steam pressure is needed to retard coking of the catalyst.
• the process is conducted under a vacuum to maximize the yield of vinyl aromatic product. At the steam to oil mass ratios described hereinbefore, vacuum pressures are sufficient to regenerate the catalyst, at least in part.
• the total pressure in the reactor is greater than about 1 psia (6.9 kPa). More preferably, the total pressure is greater than about 3 psia (20.7 kPa).
• the total pressure is less than about 73 psia (503.3 kPa). More preferably, the total pressure is less than about 44 psia (303.4 kPa). Most preferably, the total pressure is subatmospheric, ranging between about 3 psia (20.7 kPa) and about 13 psia (90.6 kPa).
• the pressure throughout the freeboard, reaction, and regeneration zones may vary depending upon process factors, such as, the weight and buoyancy of the catalyst and frictional effects. Typically, the pressure is somewhat greater at the bottom of the reactor than at the top.
• the space velocity of the dehydrogenation feedstream will depend upon the specific alkyl aromatic compound and catalyst used, the specific vinyl aromatic product formed, the reaction zone dimensions (for example, diameter and height), and the form and weight of the catalyst particles. It is desirable to remove the reactant and products quickly from the freeboard zone, so as to reduce thermal cracking and other undesirable side reactions. Additionally, the gas flow should be sufficient to induce fluidization of the catalyst bed. Generally, the space velocity of the dehydrogenation feedstream varies from the minimum velocity needed to achieve fluidization of the catalyst particles to a velocity just below the minimum velocity needed to achieve pneumatic transport of the catalyst particles. Fluidization occurs when the catalyst particles are disengaged, when the particles move in a fluid-like fashion, and when the bed pressure drop is essentially constant along the bed.
• Pneumatic transport occurs when a substantial quantity of catalyst particles are entrained in the gas flow and transported out of the reactor.
• the space velocity of the dehydrogenation feedstream varies from the minimum bubbling velocity to the minimum turbulent flow velocity. Bubbling occurs when gas bubbles can be seen in the fluidized bed, but little back-mixing of gas and solids occurs.
• Turbulent flow occurs when there is both substantial bubbling and substantial back-mixing of gas and solids. More preferably, the flow is sufficiently high to induce back-mixing.
• the gas hourly space velocity (GHSV), calculated as the total flow of dehydrogenation feedstream, including alkyl aromatic compound and, optionally, steam, sweeping gas, and/or diluent flows, is greater than about 60 ml total feed per ml catalyst per hour (h "1 ), measured at operating conditions.
• the GHSV of the dehydrogenation stream is greater than about 120 h "1 , and more preferably, greater than 720 h "1 .
• the GHSV of the dehydrogenation stream is less than about 12,000 h "1 , preferably, less than about 3,600 h "1 , and more preferably, less than 1,800 h "1 , measured as the total flow at operating process conditions.
• the gas residence time in the reaction zone is greater than about 0.3 seconds (sec), measured at operating conditions.
• the "reaction zone voidage fraction” is the fraction of the reaction zone which is empty.
• the “superficial gas velocity” is the gas velocity through the empty reactor.
• the gas residence time in the reaction zone is greater than about 1 sec, more preferably, greater than about 2 sec, measured at operating conditions.
• the gas residence time in the reaction zone is less than about 60 sec, preferably, less than about 30 sec, and more preferably, less than about 5 sec, measured at operating conditions.
• the gas hourly space velocity of the regeneration feedstream through the regeneration zone can be broadly varied, provided that the catalyst is regenerated, at least in part, and provided that the catalyst particles in the regeneration zone are effectively fluidized.
• the space velocity of the regeneration feedstream can vary from the minimum velocity needed to achieve fluidization of the catalyst particles to a velocity just below the minimum velocity needed to achieve pneumatic transport of the catalyst particles.
• the space velocity of the regeneration feedstream varies from the minimum bubbling velocity to the minimum turbulent flow velocity.
• the gas hourly space velocity (GHSV) calculated as the total of the regeneration feedstream, is greater than about 60 ml total feed per ml catalyst per hour (h "1 ), measured at operating conditions.
• 5 the GHSV of the regeneration stream is greater than about 120 h "1 , and more preferably, greater than about 360 h “1 .
• the gas hourly space velocity of the regeneration stream is less than about 12,000 h "1 , preferably, less than about 3,600 h "1 , and more preferably, less than about 720 h "1 , measured as the total flow at operating conditions.
• the gas residence time calculated as the height of o the regeneration zone times the regeneration zone voidage fraction divided by the superficial gas velocity of the total of the regeneration and reaction feedstreams, is greater than about 0.3 sec, measured at operating conditions.
• the "regeneration zone voidage fraction" is the fraction of the regeneration zone which is empty.
• the gas residence time in the regeneration zone is greater than about 1 sec, and more preferably, greater than about 5 sec. 5
• the gas residence time in the regeneration zone is less than about 60 sec, preferably, less than about 30 sec, and more preferably, less than about 10 sec, measured at operating conditions.
• Ethylbenzene is converted primarily to styrene.
• Ethyltoluene is converted to p-methylstyrene (p-vinyltoluene).
• t-Butylethylbenzene is converted to t- butylstyrene.
• Isopropylbenzene (cumene) is converted to ⁇ -methylstyrene, and diethylbenzene is converted to divinylbenzene.
• Hydrogen is also formed during dehydrogenation.
• conversion of the alkyl aromatic compound in the process of this invention can vary depending upon the specific feed composition, catalyst composition, process conditions, and fluidized bed conditions.
• conversion is defined as the mole percentage of alkyl aromatic compound which is converted to all products.
• the conversion of alkyl aromatic compound is o typically greater than about 30 mole percent, preferably, greater than about 50 mole percent, and more preferably, greater than about 70 mole percent.
• selectivity to products will vary depending upon the specific feed composition, catalyst composition, process conditions, and fluidized bed conditions.
• selectivity is defined as the mole percentage of converted alkyl aromatic compound which forms a specific product, preferably, the vinyl aromatic compound.
• the selectivity to vinyl aromatic compound is typically greater than about 60 mole percent, preferably, greater than about 75 mole percent, and more preferably, greater than about 90 mole percent.
• a fluidized bed reactor [4.25 inches (10.63 cm) inner diameter; 20 inches (50 cm) height)] was constructed as shown in Figure 1.
• the reactor comprised a single, vertical shell functionally divided into three zones: a catalyst regeneration zone (1) at the bottom of the reactor; a freeboard zone (3) at the top of the reactor; and a reaction zone (2) at the middle section between the regeneration and freeboard zones.
• a first inlet port (4) at the bottom of the reactor exited into a plenum area (10) in which a gas distributor (9) was located. This first inlet port was used to distribute a regeneration feedstream into the regeneration zone.
• the second inlet port was connected to an inlet tube which terminated in a sparger array (6) in the reaction zone at a height of 3 inches (7.5 cm) above the bottom distributor plate (9).
• the sparger array constructed as six rows of sintered metal tubing [Inconel, 1/4 inch OD (6.3 mm OD)], was designed to provide a uniform pressure drop throughout the sparger.
• the exit ports in the sparger were positioned horizontally.
• An outlet port (7) was located in the freeboard zone for removing the product stream. Solids entrained in the product stream were collected in a cyclone (located below outlet port 7) and then recirculated to the reactor by means of a third inlet port (8) located in the regeneration zone. Effluent gases were collected downstream of the cyclone.
• the reactor was also equipped with a resistive (electrical) means for heating the reactor and two internal thermocouples (K type) for measuring the fluidized bed temperature in the reaction and regeneration zones.
• the reactor was employed to dehydrogenate ethyl-benzene in the presence of a dehydrogenation catalyst to styrene, while simultaneously and continuously regenerating the dehydrogenation catalyst.
• a dehydrogenation catalyst (2370 g) with a mean particle diameter of 300 ⁇ m and comprising 28.7 percent iron oxide (Fe2 ⁇ 3), 14.3 percent cerium oxide (Ce2 ⁇ 3), 7.6 percent copper oxide (CuO), 31.6 percent potassium carbonate (K2CO3), 0.6 percent chromium oxide (Cr2 ⁇ 3), 9.5 percent zinc oxide (ZnO), and 7.6 percent cements, by weight, was loaded into the reactor.
• the reaction feedstream comprised a mixture of ethylbenzene and steam.
• the regeneration feedstream comprised steam.
• the gas residence time in the regeneration zone was 1.46 seconds; the gas residence time in the reaction zone was 0.67 seconds.
• the reactor temperature and pressure were maintained at 600°C and 15.5 psia (106.9 kPa), respectively.
• Products obtained via exit port (7) were analyzed as noted above. Ethylbenzene conversion was 74.0 mole percent. Selectivity to styrene was 86.0 mole percent. Other products included benzene and toluene. The material balance accounted for 95 weight percent of the organic feed material.
• Liquid ethylbenzene at a feed rate of 2.52 cm-Vmin and liquid water at a feed rate of 2.17 0 cm3/min at room temperature were heated to 500°C and added through the reaction zone to the sparger array. These feed rates correspond to a total steam to oil mass ratio of 2/1 with a superficial velocity of 156 m/min in the regeneration zone and 339.5 m/min in the reaction zone.
• the gas residence time in the regeneration zone was 2.91 seconds; the gas residence time in the reaction zone was 0.78 seconds.
• the reactor temperature and pressure were 5 maintained at 600°C and 15.5 psia (106.9 kPa), respectively. Ethylbenzene conversion was
• Example 1 nitrogen was added as a sweeping gas to the ethylbenzene stream, but no steam was added to the ethylbenzene stream. In contrast in Example 2, no sweeping gas o added to the ethylbenzene stream, and the steam stream was split between the dehydrogenation feed and the regeneration feed.
• Example 2 was compared with Example 1 , it was seen that the conversion of ethylbenzene was higher in Example 2 due to longer residence times in the bed, and the selectivity to styrene was lower due to increased free radical cracking in the freeboard region. 5
• Example 3 The process conditions of Example 3 were closely similar to Example 2, with the following exception. In Example 2 one-half of the total steam was delivered to the regeneration zone, and one-half of the total steam was delivered to the reaction zone. In contrast, in Example 3 all of the steam was delivered to the regeneration zone. When Example 3 was compared with Example 2, it was seen that the conversion of ethylbenzene and the selectivity to styrene were comparable. Little difference was found which depended upon the location of introducing steam.
• Example 4 Example 2 was repeated under closely similar process conditions, except for pressure which was held constant at 5 psia (34.5 kPa).
• the catalyst used in Example 4 had a chemical composition identical to the catalyst of the previous examples; however, the quantity of catalyst used was 1355 g, and the catalyst had a mean particle diameter of 220 ⁇ m.
• Process conditions were as follows: water feed rate to the regenerator zone at 2.9 cm3/min; liquid ethylbenzene and water feed rates to the reaction zone at 2.52 cm-Vmin and 1.45 cm ⁇ /min, respectively; steam to oil mass ratio at 2/1; superficial flow velocity in the regeneration zone at 123 cm/min; superficial flow velocity in the reaction zone at 200 cm/min; and temperature of 600°C.
• the ethylbenzene conversion was 49 mole percent.
• the styrene selectivity was 88 mole percent.
• a material balance accounted for 93 weight percent of the organic feed material.
• Example 4 was repeated under closely similar process conditions, except that the reactor temperature was held constant at 590°C rather than 600°C.
• Process conditions were as follows: water feed rate to the regenerator zone at 2.9 cnvVmin; liquid ethylbenzene and water feed rates to the reaction zone at 2.52 cm ⁇ /min and 1.45 cm-Vmin, respectively; steam to oil mass ratio at 2/1 ; superficial flow velocity in the regeneration zone at 122 cm/min; superficial flow velocity in the reaction zone at 199 cm/min; and pressure of 5 psia (34.5 kPa).
• the ethylbenzene conversion was 50 mole percent.
• the styrene selectivity was 94 mole percent.
• a material balance accounted for 99 weight percent of the organic feed material.
• a comparison of Examples 4 and 5 showed that operating under vacuum with a temperature lower than 600°C provides a further increase in styrene selectivity.
• Example 4 was repeated under closely similar process conditions, except that the reactor temperature was held constant at 580°C rather than 600°C.
• Process conditions were as follows: water feed rate to the regenerator zone at 2.83 cm-Vmin; liquid ethylbenzene and water feed rates to the reaction zone at 2.52 cm-Vmin and 1.45 cm-Vmin, respectively; steam to oil mass ratio at 2/1 ; superficial flow velocity in the regeneration zone at 121 cm/min; superficial flow velocity in the reaction zone at 197 cm min; and pressure of 5 psia (34.5 kPa).
• the ethylbenzene conversion was 44 mole percent.
• the styrene selectivity was 95 mole percent.
• a material balance accounted for 100 weight percent of the organic feed material.
• a comparison of Examples 4, 5, and 6 showed that operating under vacuum with a temperature lower than 600°C provides a further increase in styrene selectivity.
• Example 4 was repeated under closely similar process conditions, except that the steam to oil ratio was 1/1 instead of 2/1.
• Other process conditions were as follows: water feed rate to the regeneration zone at 1.45 cm-Vmin; liquid ethylbenzene and water feed rates to the reaction zone at 2.52 cm- min and 0.73 cm-Vmin, respectively; superficial flow velocity in the regeneration zone at 61.5 cm min; superficial flow velocity in the reaction zone at 107.6 cm/min, pressure of 5 psia (34.5 kPa); and temperature of 600°C.
• the ethylbenzene conversion was 49 mole percent.
• the styrene selectivity was 89 mole percent.
• a material balance accounted for 98 weight percent of the organic feed material.
• a comparison of Examples 4 and 5 with Example 7 showed that lowering the steam to oil ratio from 2/1 to 1/1 did not affect the ethylbenzene conversion and styrene selectivity.
• Example 4 was repeated under closely similar process conditions, except for catalyst particle size and steam to oil ratio.
• the catalyst (1570 g) had a mean particle diameter of 82 ⁇ m, and the steam to oil ratio was 0.5/1
• Other process conditions were as follows: water feed rate to the regeneration zone at 0.8 cm-Vmin; liquid 0 ethylbenzene and water feed rates to the reaction zone at 5.48 cm ⁇ /min and 0.54 cm-Vmin, respectively; superficial flow velocity in the regeneration zone at 33.52 cm/min; superficial flow velocity in the reaction zone at 74.8 cm/min, pressure of 5 psia (34.5 kPa); and temperature of 600°C.
• the ethylbenzene conversion was 54 mole percent.
• the styrene selectivity was 95 mole percent.
• a material balance accounted for 100 weight percent of the 5 organic feed material.
• a pulsed mode reactor was used to study the dehydrogenation of 5 ethylbenzene to styrene as a function of time.
• the reactor was repeatedly cycled through a dehydrogenation step and thereafter a catalyst regeneration step.
• a dehydrogenation catalyst having particles sized between 1.18 mm and 1.70 o mm and comprising 33.2 percent iron oxide (Fe2 ⁇ 3), 17.5 percent cerium oxide (Ce2 ⁇ 3),
• the flow rate of the liquid ethylbenzene measured at ambient temperature and pressure (taken as 23 °C and 1 atm), was 1.16 ml/min. Simultaneously, water preheated to 550°C was fed over the catalyst during the same 2 min period. The flow rate of the water was adjusted to maintain a steam to oil weight ratio of 0.30/1. Total pressure was maintained at 5.0 psia. Thereafter, the ethylbenzene feedstream was stopped, and a regeneration pulse was conducted by feeding the water stream alone, preheated to 550°C, alone over the catalyst for 2 min under the same process conditions. The flow rate of the liquid water during the regeneration pulse was 1 ml/min, measured at 24°C and 1 atm.
• the dehydrogenation pulse was repeated, by reintroducing the ethylbenzene feedstream for 2 min as noted hereinbefore with the continuing water steam. After 2 min, the ethylbenzene feed was again stopped, while the steam stream was continued for a regeneration cycle of 2 min. The dehydrogenation-regeneration cycles were repeated for a total run time of 200 h.
• the product stream was continuously passed through a condenser, separated, and analyzed by conventional methods.
• Example 9 The process of Example 9 was repeated in the same continuous flow, fixed bed reactor under similar process conditions, with the exception that the dehydrogenation was run in continuous mode rather than pulsed mode. Accordingly, there was only one dehydrogenation cycle and no catalyst regeneration cycle. Under these conditions, the catalyst was found to steadily deactivate with a concomitant decrease in ethylbenzene conversion. As the catalyst deactivated, the temperature of the process was increased to maintain a constant ethylbenzene conversion. At a steam to oil ratio of 0.3/1 , the temperature had to be increased at a rate of 0.45°C per min for conversion to be maintained.

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