RU2234491C2 - Method of dehydrogenation of alkyl aromatic compounds, method of dehydrogenation of ethyl benzene or substituted ethyl benzene and reactor for dehydrogenation of alkyl aromatic compounds - Google Patents

Abstract

FIELD: production of aromatic compounds.

SUBSTANCE: reactor contains fluidized layer with dehydrogenation catalyst for obtaining vinyl aromatic compound, such as styrene and regeneration of catalyst through interaction of vapor with de-activated catalyst in regeneration zone. Reactor is characterized by availability of zone above layer, reaction zone and catalyst regeneration zone located inside single housing.

EFFECT: enhanced efficiency of catalyst regeneration.

38 cl, 3 dwg, 10 ex

Description

In one aspect, the invention relates to a method for dehydrogenating an alkyl aromatic compound such as ethylbenzene to form a vinyl aromatic compound such as styrene. Additionally, this invention relates to a method for the regeneration of a catalyst that is used in the dehydrogenation of an alkyl aromatic compound. In another aspect, the invention relates to a fluidized bed reactor in which an organic process is carried out, such as said dehydrogenation process.

The dehydrogenation of alkyl aromatic compounds, such as ethylbenzene, isopropylbenzene, diethylbenzene and p-ethyl toluene, is used in the preparation of styrene and substituted styrene derivatives such as α-methyl styrene, divinylbenzene and p-methyl styrene. Styrene and its substituted derivatives are useful as monomers for the formation of polystyrenes, styrene-butadiene rubbers (SBC), acrylonitrile-butadiene-styrene (ABS), styrene-acrylonitrile (SAN) and resins based on unsaturated polyester.

Fluidized bed reactors are important in many diverse catalyzed organic processes, including dehydrogenation processes.

The main route for the production of 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 Pat. Since the process is highly endothermic, energy for the process is provided by introducing superheated steam into the process reactor, which usually has a fixed bed in these references. Process temperatures are usually between about 550 ° C and 700 ° C. Adverse reactions may be limited by maintaining a low partial pressure of ethylbenzene.

The use of a fixed bed reactor for these dehydrogenation processes has some disadvantages. First, a fixed bed reactor characterized by a fixed bed of catalytic particles is difficult to uniformly heat to a high temperature. In the endothermic dehydrogenation of ethylbenzene to styrene, the downstream end (upstream) of the fixed catalyst bed tends to be colder than the upstream end (upstream) of the catalyst bed. Since the temperature drop can reduce the conversion of ethylbenzene at the downstream end of the reactor, the feed stream is usually preheated to a higher temperature than desired. As a result, the catalytic bed near the inlet of the reactor may turn out to be superheated and usually breaks down faster than the catalyst further downstream. Under such conditions, a large catalytic bed is required to withstand a long cycle. On the other hand, in processes for the dehydrogenation of ethylbenzene to styrene in a fixed bed, steam is supplied simultaneously with ethylbenzene to promote on-site catalyst regeneration during the dehydrogenation process. Typically, a high mass ratio of steam to ethylbenzene is required, usually from more than 1.2 to 2.0 and possibly higher, which adversely imposes high energy costs and large water recirculation on the process. The mass ratio of steam to alkylaromatic compound is hereinafter referred to as the "steam to oil ratio". When the overall activity of the bed falls below the point of practicality, the catalyst should be replaced. As an additional drawback, fixed bed reactors are usually shut down for several weeks to replace the catalyst.

Other sources, such as US Pat. get styrene. A conventional fluidized bed reactor contains a single reaction zone where the catalyst particles are separated and circulated. Compared to fixed beds, fluidized bed reactors provide a more isothermal temperature distribution. The isothermal catalytic layer typically does less damage to the catalyst and allows higher product yields. Fluidized beds also make it easy to replace the catalyst when it is completely deactivated and unable to regenerate. Since the catalyst is treated as a liquid, the deactivated catalyst can be discharged from the reactor, and the active catalyst can be added to the reactor without stopping the chemical process. Said oxidative dehydrogenation process can also be carried out by continuously transferring part of the catalyst to a regenerator for regeneration in an oxygen atmosphere and then recycling the regenerated catalyst back to the oxidative dehydrogenation reactor.

The disadvantage is that the processes of oxidative dehydrogenation with the simultaneous supply of alkyl aromatic compounds and oxygen can give low yields of vinyl aromatic product, because it is more difficult to regulate the side reactions of oxidation. In addition, safety issues arising from the transportation and processing of mixtures of organic compounds and oxygen are significant. As an additional drawback, the continuous circulation of the catalyst between the fluidized bed process reactor and the regenerator requires sophisticated equipment and often requires catalyst particles with high abrasion resistance.

Some patents, such as US Pat. Nos. 4,229,604 and 5,510,553, disclose the use of a transport reactor for dehydrogenating ethylbenzene to styrene in the absence of oxygen and in the presence of an oxidative dehydrogenation catalyst such as magnesia-modified silica or reducible vanadium oxide on a metal oxide support.

These processes eliminate the danger of using mixtures of alkyl aromatic compounds and oxygen, however, the service life of such catalysts is short. Therefore, the catalyst must be continuously circulated between the process reactor and the regenerator, where the catalyst is regenerated in an oxygen atmosphere. As noted above, continuous catalyst recycling between the process reactor and the regenerator requires sophisticated equipment and often requires catalyst particles with high abrasion resistance.

The above description confirms the need for improved catalytic dehydrogenation processes. It would be useful, for example, to open the process of dehydrogenation of an alkyl aromatic compound, such as ethylbenzene, to a vinyl aromatic compound, such as styrene, which provides for effective regeneration of the catalyst in situ with a cost-effective steam to oil ratio. It would be useful if during the process the deactivated catalyst could be replaced by an active catalyst without the need to shut down the reactor or use sophisticated conveying equipment. It would be useful if the isothermal temperature of the layer could be ensured. It would also be useful if the process did not require oxygen, which complicates safety and transportation procedures. And finally, the process would be more desirable if it possessed all of these features and also gave a high yield of vinyl aromatic compounds.

In another aspect, US Pat. No. 4,152,393 discloses a reactor consisting of a single vessel that contains a reaction zone and a regeneration zone arranged in such a way, specifically as a set of concentric walls and passages, that the powdered solid can be carried by a gas stream from the regeneration zone to the reaction zone in the first way and then back to the regeneration zone in a second way. Gases passing through the regeneration zone are not transferred to the reaction zone, and gases passing through the reaction zone are not transferred to the regeneration zone. The patent states that such a reactor is suitable for ammoxidation of propylene.

The specified reactor can detect a high slag stream, characterized by gas bubbles flowing along the inner walls of the reactor. The slag stream adversely reduces the contact between the gas phase reactants and the solid catalyst particles, thereby reducing the process productivity. As an additional disadvantage, this reactor has narrow curved gaps and numerous gas nozzles inside these gaps, which can lead to severe abrasion of the catalyst particles. US Pat. No. 6,048,459 discloses a method for fluidizing a powder layer material, which comprises collecting and raising a portion of a fluid above a fluidized bed and recirculating the raised fluid above the bed through an overflow tube located within the bed. The area of the powdery material can be distributed under the fluidized bed for use, in particular for the treatment of anaerobic sludge for a long time.

In one aspect, the invention relates to a method for dehydrogenating an alkyl aromatic compound over a dehydrogenation catalyst in a single-shell fluidized bed reactor to form a vinyl aromatic compound and regenerate the dehydrogenation catalyst in situ. The method of the invention comprises, in a first step (a) a fluidization of 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 circulates inside and between the two zones; (b) the interaction of the alkyl aromatic compound and possibly steam with a dehydrogenation catalyst in the reaction zone under reaction conditions sufficient to obtain the corresponding vinyl aromatic compound, and (c) the interaction of steam with the dehydrogenation catalyst located in the regeneration zone, the interaction is carried out under regeneration conditions sufficient to regenerate the catalyst, at least in part.

The dehydrogenation process of this invention, which is used to produce vinyl aromatic compounds of industrial importance, such as styrene, p-methylstyrene, α-methylstyrene and divinylbenzene, has significant advantages over previously known processes. First, the method of this invention does not use oxygen. Therefore, the safety problems associated with transported and processed mixtures of organic compounds and oxygen, which are found in some previously known processes, are excluded in the method according to this invention. Secondly, in the method according to this invention, the ratio of steam to oil is favorably lower than the ratio used in fixed-bed processes of the prior art. Therefore, the method according to this invention uses less water recirculation and is more energy efficient and economical than the prior art processes. As an additional advantage, the process of this invention, carried out in a fluidized bed reactor, is substantially isothermal. Problems associated with uneven bed temperatures, such as overheating and damage to the catalyst at the upstream end of the catalyst bed and reduced performance at the downstream end of the catalyst bed, are substantially eliminated. In addition, the thermal formation of a by-product is reduced. As an additional advantage in the method of this invention, a smaller catalyst bed can be used compared to the size of the catalyst bed used in the fixed bed of prior art processes while still achieving comparable process cycles. As another advantage, the process of this invention provides continuous regeneration of the dehydrogenation catalyst in situ. When carrying out the method of the invention, it is not necessary to interrupt the operation of the reactor for regenerating the catalyst or for transferring the catalyst from the fluidized bed reactor to the regenerator. Thus, the method according to this invention is simple to implement and structurally. Moreover, the catalyst used in the process may not be required to have the high abrasion resistance required for transfer reactors. And finally, when the catalyst is not capable of further regeneration, the method according to this invention provides for the replacement of the deactivated catalyst during the continuous implementation of the dehydrogenation process. Since the solid catalyst is treated as a fluid, the deactivated catalyst is simply transported from the reactor and the fresh catalyst is transported to the reactor during operation. Thus, both regeneration and replacement of the catalyst can be carried out without stopping the dehydrogenation process, thereby increasing productivity. Most advantageously, when the method according to this invention receive vinylaromatic compound, preferably styrene, in high yield.

In another aspect, the present invention relates to a fluidized bed reactor that allows a chemical process and catalyst regeneration to be carried out simultaneously. The fluidized bed reactor contains a single vertical casing, the space inside of which is divided into a zone above the bed, a reaction zone and a regeneration zone. The reactor also contains an inlet device for introducing a stream of starting material for regeneration into the regeneration zone and an inlet device for introducing a stream of starting material of the reagent into the reaction zone. The reactor also contains devices for the separation of reaction zones and regeneration, while allowing continuous circulation and reverse mixing of the catalyst on an industrial scale between the two zones. In a preferred embodiment, one of the inlets for the feed streams for the reaction or for regeneration acts as a device for separating the reaction and regeneration zones. The reactor of this invention also contains exhaust devices, preferably in the area above the bed, to remove the effluent containing products and any non-converted reactants and starting materials for regeneration. Optionally, the reactor further comprises devices for returning catalyst particles trapped in the effluent back to the reactor. Optionally, inlet devices and outlet devices for transporting the catalyst to and from the reactor may also be present.

In the reactor of the present invention, cross-mixing of the streams of the starting material of the reactant and the material for regeneration can occur, although, preferably, the arrangement of the inlets of the materials for the reaction and regeneration should essentially separate both processes. However, although this design allows for back-mixing of gases and solids, the feed streams for regeneration and reaction will be chemically compatible, as explained here, with the current dehydrogenation-regeneration process.

The fluidized bed reactor of this invention can be used with a variety of catalyzed organic processes, including, for example, dehydrogenation, oxidation and halogenation processes. A particularly important dehydrogenation process, for which a fluidized bed reactor of the present invention can be used, comprises dehydrogenation of an alkyl aromatic compound such as ethylbenzene to a vinyl aromatic compound such as styrene. Advantageously, the fluidized bed reactor of the present invention provides a chemical reaction zone and a catalyst regeneration zone within a single fluidized bed reactor vessel. Thus, in situ regeneration of the catalyst can be achieved simultaneously with the desired chemical process. It is not necessary to transport the deactivated catalyst from the fluidized bed reactor of the present invention to a separate regeneration vessel, so that the catalyst is subjected to significantly less stress and damage than is the case in transfer reactors. As another advantage, the deactivated catalyst can be replaced on-line without interrupting the chemical process by simply transporting the deactivated catalyst from the reactor and transporting fresh catalyst to the reactor. These advantages provide the necessary improvements in the design of a fluidized bed process.

1 shows side and top cross-sectional views of a first preferred embodiment of a fluidized bed reactor of the present invention, the details of which are explained below.

FIG. 2 shows side and top cross-sectional views of a second preferred embodiment of a fluidized bed reactor of the present invention, the details of which are explained below.

Figure 3 is a graph of ethylbenzene conversion and styrene selectivity as a function of experimental time for the process of ethylbenzene dehydrogenation and catalyst regeneration carried out in a pulsating reactor.

In one aspect, the present invention relates to a method for dehydrogenating an alkyl aromatic compound over a dehydrogenation catalyst in a single-shell fluidized bed reactor to form a vinyl aromatic compound and regenerate it in place of the dehydrogenation catalyst. The method 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 fluidization conditions such that the catalyst is circulated inside and between the two zones. In a second step, which is carried out simultaneously with the first step, the method comprises (b) reacting the alkyl aromatic compound and optionally steam with a dehydrogenation catalyst in the reaction zone under reaction conditions sufficient to obtain the corresponding vinyl aromatic compound. In the third stage, which is carried out simultaneously with the first and second stages, the method comprises (c) reacting steam with a dehydrogenation catalyst located in the regeneration zone, the interaction is carried out under regeneration conditions sufficient to regenerate at least a portion of the catalyst.

In the fluidized bed process of this invention, at any given point in time, part of the catalyst will circulate in the reaction zone, while essentially the rest of the catalyst will circulate in the regeneration zone with some mutual mixing at the boundary of the two zones. During the period of residence in the reaction zone, the catalyst will lose activity and become partially or completely deactivated. Decontamination can mainly be caused by coke buildup on the catalyst surface. The invention, however, should not be bound or limited by such a theory of decontamination. Under fluidization conditions, the catalyst in the reaction zone, including the deactivated catalyst, will circulate in the regeneration zone. The deactivated catalyst located in the regeneration zone will be reactivated in contact with the vapor. Reactivation results in a partial or substantially complete restoration of the activity of the catalyst compared with the activity of fresh unused or “just synthesized” catalyst. From now on, under fluidized 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 above description is given as a means to explain the reaction-regeneration cycle and to define the words “regenerate, at least in part, the catalyst”.

It should also be understood that after repeated reactions and regenerations, the time comes when the catalyst cannot be further regenerated to a practical level of activity even with the regeneration procedure described here. When this happens, the deactivated catalyst can be replaced simply by transporting it from the reactor and transporting fresh catalyst to the reactor, preferably at the same time. In the reactor of this invention, the catalyst can be replaced on-line without interrupting the catalytic process, more specifically for the present invention, the dehydrogenation process. The catalyst can be transported to and from the reactor through air nozzles or a pneumatic transfer circuit. Alternatively, the catalyst can be removed from the reactor via gravity-driven exhaust devices at the bottom of the reactor, and the catalyst can be added to the reactor from an inlet device such as a pressure pipe at the top of the reactor.

In a preferred aspect of the invention, the alkyl aromatic compound is ethylbenzene or a substituted ethylbenzene derivative, and the resulting vinyl aromatic compound is styrene.

Any alkyl aromatic compound may be used in the dehydrogenation process of this invention provided that the product obtained is a vinyl aromatic compound. The aromatic part of the vinyl aromatic molecule may contain, for example, a monocyclic aromatic ring such as phenyl; a fused aromatic ring such as naphthyl; or ring assemblies such as biphenyl. Preferably, the aromatic part of the molecule is a monocyclic aromatic ring, more preferably phenyl. The alkyl portion of the molecule of the alkyl aromatic compound may contain, for example, any saturated linear, branched or cyclic hydrocarbon radical, provided that it can be dehydrogenated in the process of this invention to form the vinyl portion of the molecule. Non-limiting examples of suitable alkyl moieties include ethyl, n-propyl, isopropyl, n-butyl, isobutyl, tert-butyl and higher homologues thereof. Preferably, the alkyl part of the molecule is C ₂ -C ₁₀ alkyl, more preferably C ₂ -C ₅ alkyl, most preferably ethyl. The alkyl aromatic compound may be optionally substituted with two or more alkyl groups or substituted with other types of substituents that are substantially inert with respect to the dehydrogenation process of this invention. Examples of alkyl aromatic compounds that are advantageously used in the method of this invention include, but are not limited to, ethylbenzene, diethylbenzene, ethyltoluene, ethylxylene, isopropylbenzene, tert-butyl ethylbenzene, ethylnaphthalene, ethylbiphenyl and their higher alkyl homologs. Preferably, the alkyl aromatic compound is a C ₈ -C ₂₀ alkyl aromatic compound, more preferably a C ₈ -C ₁₅ alkyl aromatic compound and most preferably ethylbenzene or a substituted derivative thereof.

In the method of this invention, the feed stream for regeneration typically comprises steam. Optionally, steam may also be introduced into the dehydrogenation feed stream. Any mass ratio of steam to alkyl aromatic compound (steam to oil ratio) is suitable for the process of this invention, provided that a vinyl aromatic compound is obtained in the process. It should be noted that the ratio of steam to oil is given based on the total mass of steam introduced into the reactor from all sources, including steam from both feed streams for dehydrogenation and for regeneration. Usually the mass ratio of steam to oil is more than about 0.2 / 1, preferably more than about 0.5 / 1. Typically, 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. Typically, the method of this invention is carried out at lower steam to oil ratios compared to prior art processes. The low ratio of steam to oil favorably reduces the energy demand and the cost of converting water to steam and reduces the amount of water recycled to the reactor.

The purification gas may optionally be used in the method of this invention. The cleaning gas, which can be introduced directly into the zone above the reactor bed, operates mainly so as to remove the product stream from the zone above the bed where undesirable thermal reactions can occur. Any gas that is substantially inert with respect to dehydrogenation and regeneration processes may suitably be used as a cleaning gas, including, for example, nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof. The concentration of the cleaning gas in the region above the layer can be any concentration, provided that the complete process gives the desired vinyl aromatic compound. Typically, the concentration of the cleaning gas varies depending, for example, on the particular alkyl aromatic compound used and the specific process conditions, in particular the temperature and gas velocity. Typically, the concentration of the cleaning gas in the region above the bed is greater than about 10 volume percent and preferably more than about 20 volume percent. Typically, the concentration of the cleaning gas in the region above the bed is less than about 90 volume percent and preferably less than about 70 volume percent.

Optionally, the dehydrogenation and / or regeneration feed stream (s) may also contain a diluent. The diluent primarily dilutes reagents and products for reasons of enhanced selectivity or safety. Any gas that is substantially inert with respect to the dehydrogenation and regeneration steps may suitably be used as a diluent, including, for example, nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof. The concentration of diluent in the feed stream for either dehydrogenation or regeneration may be any concentration, provided that the complete process gives the desired vinyl aromatic compound. Typically, the concentration of diluent varies depending, for example, on the particular diluent selected, the particular alkyl aromatic compound, the specific conditions of the dehydrogenation or regeneration process, and the particular catalyst and its deactivation properties. Typically, the concentration of diluent in the dehydrogenation or regeneration feed stream is greater than about 10 volume percent and preferably more than about 20 volume percent. Typically, the diluent concentration in one of the two streams is less than about 90 volume percent, preferably less than about 70 volume percent. When steam is used as a diluent, then the steam to oil ratio described above determines the concentration of steam in the dehydrogenation feed stream.

Oxygen is not required for the process of this invention. Preferably, oxygen is not used in the method of this invention.

Any dehydrogenation catalyst that 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. Non-limiting examples of dehydrogenation catalysts that can be advantageously used include those described in the following US patents: 4,404,123, 4,503,163, 4,684,619, 5,171,914, 5,376,613, 5,510,552 and 5,679,878, which relate to a variety of iron oxide catalysts containing, for example, one or several alkali metal compounds, preferably sodium, potassium and cesium; alkaline earth metals, preferably calcium, and / or compounds of cerium, chromium, zinc, copper and / or gallium, as well as the catalyst described in US patent 3651160, which relates to chromium oxide and alkali metal oxides. Preferably, the catalyst is a dehydrogenation catalyst containing iron oxide. More preferably, the catalyst contains (a) at least one iron oxide, (b) at least one carbonate, bicarbonate, potassium and / or cesium oxide or hydroxide, (c) cerium oxide, carbonate, nitrate or hydroxide, (d) optionally hydroxide, carbonate, bicarbonate, acetate, oxalate, nitrate or sodium sulfate, (e) optionally carbonate, sulfate or calcium hydroxide, and (f) optionally one or more binding agents, such as hydraulic cement. As an additional feature, a more preferred catalyst may further comprise one or more oxides selected from oxides of zinc, chromium and copper. Typically, a more preferred catalyst contains from 25 to 60 weight percent of iron, from 13 to 48 weight percent of potassium and from 1 to 20 weight percent of cerium, weight percent based on oxides. These proportions and other proportions of suitable catalyst components are described in the aforementioned US patents.

The dehydrogenation catalyst used in the fluidized bed reactor of this invention can have particles of any size or shape as long as the catalyst is capable of catalyzing the dehydrogenation of an alkyl aromatic compound to a vinyl aromatic compound. Typically, the average particle size of the catalyst is more than about 20 micrometers (μm) in diameter (or in cross-sectional size) and preferably more than about 50 μm in diameter. Typically, the average particle size is less than about 1000 microns and preferably less than about 200 microns. Preferably, the catalyst particle is smooth with rounded edges and substantially non-cohesive and has an abrasion resistance sufficient for use in a fluidized bed reactor. One of skill in the art will know if a particular catalyst has sufficient abrasion resistance for use in a fluidized bed reactor.

If desired, the dehydrogenation feed stream may be preheated before being introduced into the reaction zone. Preheating may be suitably provided by condensation of the high pressure saturated steam or, alternatively, by burning a fuel source or process off-gas. Any heating temperature can be used provided that it is lower than the temperature at which thermal cracking of the aromatic compound becomes measurable. Typical heating temperatures are higher than about 150 ° C, preferably higher than about 250 ° C, and more preferably higher than about 350 ° C. Typical heating temperatures are lower than about 600 ° C and preferably lower than about 590 ° C. Likewise, the regeneration feed stream can be preheated before being introduced into the regeneration zone. A typical temperature for preheating the regeneration stream is higher than about 200 ° C, preferably higher than about 300 ° C, and more preferably higher than about 400 ° C. The temperature of the heating stream for regeneration is usually lower than about 650 ° C and preferably lower than about 630 ° C.

The preferred embodiment of the new reactor used in the method of this invention is the fluidized bed reactor shown in FIG. 1, containing a single vertical vessel, the interior of which is divided into a regeneration zone (1), a reaction zone (2) and a zone above the bed (3). The regeneration zone located in this preferred embodiment at the bottom of the reactor contains a region where the catalyst is regenerated. The reaction zone located in this preferred embodiment in the middle zone of the reactor is the region where the catalyzed organic chemical process, such as the dehydrogenation process described herein, takes place. The zone above the layer located in the upper part of the reactor occupies the space above the middle zone up to the upper inner wall of the reactor. The area above the bed occupied by gaseous reagents and products also provides space for expansion of the fluidized bed. Gas phase thermal reactions can occur in the region above the bed, but the process conditions are preferably maintained such as to minimize these gas phase reactions with respect to the catalyzed process occurring in the reaction zone.

The regeneration zone in figure 1 contains an inlet device for introducing a stream of source material for regeneration, here is a pair, and, optionally, a diluent, in the regeneration zone. The inlet device may comprise, for example, an inlet (4) that leads to a discharge part (10), over which a distribution plate or a plurality of sprinklers (9) is located. The area above the layer contains an inlet device (5), containing, for example, an inlet and a supply pipe for introducing a stream of the starting material of the reagent, here, a stream of the starting material for dehydrogenation, into the reaction zone. 1 and 2, the inlet device for the regeneration zone is shown in the lower part of the figure, and the inlet device for the reaction zone is shown in the upper part of the figure. In practice, the inlet of the regeneration zone may be located elsewhere, provided that it leads to the regeneration zone. Similarly, the inlet of the reaction zone may be located elsewhere as the reactants enter the reaction zone.

Preferably, the inlet (5) terminates in a plurality of sprinklers or a distributor (6) in a reaction zone located more preferably above the bottom of the steam distributor (9). A distributor or a plurality of sprinklers (6) are preferably designed to deliver a feed stream of the reaction material in any direction to the reaction zone. The distance from the bottom steam distributor (9) to the distributor (6) of the reagent feed stream can be varied to provide variable volumes in the regeneration and reaction zones when necessary. The larger the zone, the longer the residence time of gases and solids in this zone. As described herein, dispenser (6) provides a functional distinction between the regeneration zone and the reaction zone so that the regeneration process essentially occurs in the regeneration zone, while the organic process essentially occurs in the reaction zone, although solid mixing is still allowed substances and gases. Gas distributors and a plurality of sprayers may be made, for example, of gas-permeable sintered metal, or, more preferably, gas distributors and sprayers may be provided with nozzles for dispersing gas. The area above the layer also contains an outlet (7), such as an outlet, for an exhaust stream containing unconverted alkyl aromatic compound, steam, an optional cleaning gas and / or diluent, and products, including vinyl aromatic compound. An outlet (7) may be attached to the cyclone (shown in FIG. 1 below the outlet 7) to collect catalyst particles that are entrained in the effluent. The collected catalyst particles can be returned to the fluidized bed reactor through an inlet (8) located at any point along the reactor, but preferably, as shown in FIG. 1, in the regeneration zone. The outlet device (7) can be further connected to a separation unit (not shown in FIG. 1), containing, for example, condensation devices and a distillation unit for separating the unconverted alkyl aromatic compound and reaction products. Non-converted reagents can be recycled back to the reaction zone through the inlet (5). In addition to the specified reactor may also contain means for measuring the temperature of the catalytic layer and, optionally, means for heating the reactor (not shown). The reaction zone and the regeneration zone may also contain reflective walls (not shown), the function of which is to reduce the formation of bubbles and their size, thereby facilitating contact between the gaseous streams of the starting material and the catalyst.

In another preferred embodiment of the invention, the fluidized bed reactor further comprises one or more devices for improving solids circulation and heat transfer. In a preferred embodiment, the device for improving the circulation of solids contains one or more suction tubes, optionally containing internal baffles. Alternatively, the device for improving the circulation of solids contains one or more suction tubes made of heating or cooling elements. An embodiment of the invention, comprising a plurality of suction tubes, is shown in FIG. 2 (parts 1-10 of FIG. 2 are identical to parts 1-10 of FIG. 1). The suction pipes (11) may each contain, for example, concentric cylinders open at both ends, or a bundle or set of heating pipes or any other structure that promotes the purging of the catalyst. Typically, the suction pipe is vertically suspended, passing through the reaction and regeneration zones to a place near the top of the reaction zone. The dehydrogenation feed stream is fed through an inlet (5) to a plurality of sprinklers (6) and up to the inner cylinder of the suction pipe (11) in the reaction zone. Under the influence of fluidization conditions, the catalyst particles will fall into the inner cylinder of the suction pipe and transferred to the top of the suction pipe. At the top, the catalyst particles will flow through the outlet of the inner cylinder and down through the annular gap between the two cylinders back to the regeneration zone.

In addition to the above, the reactor may optionally comprise an inlet device and an outlet device (not shown) for transporting the catalyst to and from the reactor, respectively.

In another embodiment of the fluidized bed reactor, the reaction and regeneration zones can be arranged in the reverse order, so that the reaction zone is at the bottom of the reactor, while the regeneration zone is in the middle of the reactor. (Figure 1, where the reaction zone is at (1), the regeneration zone is at (2) and the inlets therein are adapted accordingly.)

The reactor of this invention can be used in catalytic processes where the stream of reactant feed is chemically compatible with the feed of regeneration. The unique reactor of this invention provides continuous flow of catalyst particles between the reaction and regeneration zones within the single-shell fluidized bed reactor. Thus, the catalyzed organic process of interest and catalyst regeneration can be carried out simultaneously without transferring the catalyst from the reactor to a separate regenerator. The reactor of this invention does not contain complex concentric walls or winding paths through which catalyst particles move. Thus, for plants of an industrial scale, the reactor of this invention will not create significant problems with slag flow and abrasion.

The temperature of the reaction zone where the dehydrogenation process takes place can be any operating temperature, provided that a vinyl aromatic compound is formed in the process. The operating dehydrogenation temperature will vary depending on the particular catalyst used and the particular alkyl aromatic compound. For a preferred catalyst containing iron oxide, the dehydrogenation temperature is usually higher than about 550 ° C and preferably higher than about 570 ° C. Typically, the dehydrogenation temperature is lower than about 650 ° C and preferably lower than about 610 ° C. Below about 550 ° C, the conversion of the alkyl aromatic compound may be too small, while above 650 ° C, thermal cracking of the alkyl aromatic compound and vinyl aromatic product can occur. In the present invention, the temperature is measured on a catalyst bed in a fluidized form.

In the regeneration zone, the catalyst is in contact with steam and reactivated. The temperature of the regeneration zone may also vary as long as the catalyst is at least partially regenerated. Typically, the regeneration temperature is lower than the thermal cracking temperature of the alkyl aromatic reagent and the vinyl aromatic product. For a preferred catalyst containing iron oxide, the regeneration temperature is usually higher than about 550 ° C and preferably higher than about 570 ° C. Typically, the regeneration temperature is lower than about 650 ° C and preferably lower than about 610 ° C. Since the catalyst is continuously circulating 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 over both zones.

The process can be carried out at any operating total pressure ranging from below atmospheric to above atmospheric, provided that a vinyl aromatic product is formed. If the total pressure in the reactor is too high, the equilibrium position of the dehydrogenation process may shift back towards the alkyl aromatic compound. On the other hand, adequate vapor pressure is needed to slow down the coking of the catalyst. Preferably, the process is carried out in vacuo to maximize the yield of the vinyl aromatic product. In the steam to oil mass ratios described hereinbefore, pressures below atmospheric are sufficient to regenerate the catalyst at least partially. Preferably, the total pressure in the reactor is higher than about 1 psi. inch (6.9 kPa). More preferably, the total pressure in the reactor is higher than about 3 psi. inch (20.7 kPa). Preferably, the total pressure in the reactor is lower than about 73 psi. inch (503.3 kPa). More preferably, the total pressure is lower than about 44 psi. inch (303.4 kPa). Most preferably, the total pressure is below atmospheric between about 3 psi. inch (20.7 kPa) and about 13 psi inch (90.6 kPa). The pressure over the zone above the bed and the reaction and regeneration zones may vary depending on process factors, such as mass and buoyancy of the catalyst and frictional effects. Typically, the pressure is slightly higher at the bottom of the reactor than at the top.

The volumetric flow rate of the starting material for dehydrogenation will depend on the particular alkyl aromatic compound and catalyst used, the specific vinyl aromatic product formed, the size of the reaction zone (e.g., diameter and height) and the shape and mass of the catalyst particles. It is desirable to quickly remove the reagent and products from the zone above the bed so as to reduce thermal cracking and other undesirable side reactions. In addition, the gas flow should be sufficient to cause fluidization of the catalytic bed. Typically, the volumetric flow rate of the starting material for dehydrogenation varies from the minimum speed necessary to achieve fluidization of the catalyst particles to a speed slightly lower than the minimum speed necessary to achieve pneumatic transfer of the catalyst particles. Fluidization occurs when the catalyst particles are not adhered to each other, when the particles move like a fluid, and when the pressure drop in the bed is substantially constant along the bed. Pneumatic transfer occurs when a significant amount of catalyst particles is captured by the gas stream and carried out of the reactor. Preferably, the volumetric flow rate of the dehydrogenation feed is varied from a minimum bubbling rate to a minimum turbulent flow rate. Bubbling occurs when gas bubbles can be seen in the fluidized bed, but a slight back-mixing of gas and solids occurs. Turbulent flow occurs when both substantial bubbling and significant back-mixing of gas and solids occur. More preferably, the flow rate should be high enough to cause back mixing.

As a general guideline, the gas hourly space velocity (COSG) measured under operating conditions, calculated as the total flow rate of the feed for dehydrogenation, including the costs of the alkyl aromatic compound and, optionally, steam, cleaning gas and / or diluent, more than about 60 ml of the total feed material per ml of catalyst per hour (h ^-1 ). Preferably, the HOSG of the dehydrogenation stream is greater than about 120 h ^-1 and more preferably more than 720 h ^-1 . Typically, the CLR of the dehydrogenation stream is less than about 12,000 h ⁻¹ , preferably less than about 3600 h ^−1, and more preferably less than 1800 h ⁻¹ , measured as a total flow rate under process conditions.

As a main guideline, the gas residence time in the reaction zone, measured under operating conditions, calculated as the height of the reaction zone, multiplied by the void fraction of the reaction zone, divided by the gas flow rate per unit cross-section of the flow stream, the sum of the flows of the starting materials for regeneration and reaction is more than about 0, 3 seconds (s). The "void fraction of the reaction zone" is the part of the reaction zone that is empty. "Gas flow per unit cross-section of the flow" is the gas flow through an empty reactor. Preferably, the gas residence time in the reaction zone, measured under operating conditions, is more than about 1 s, more preferably more than about 2 s. Typically, the gas residence time in the reaction zone, measured under operating conditions, is less than about 60 seconds, preferably less than about 30 seconds, and more preferably less than about 5 seconds.

The hourly space velocity of the gas flow of the feed material for regeneration through the regeneration zone can vary widely, provided that the catalyst particles in the regeneration zone are effectively fluidized. In addition, the volumetric flow rate of the starting material for regeneration can vary from the minimum speed necessary to achieve fluidization of the catalyst particles, to a speed slightly lower than the minimum speed necessary to achieve pneumatic transfer of catalyst particles. Preferably, the volumetric flow rate of the starting material for regeneration varies from a minimum bubbling rate to a minimum turbulent flow rate. Typically, measured under operating conditions, hourly gas volumetric velocity (COSH), calculated as the total feed stream for regeneration, is greater than about 60 ml of all feed material per ml of catalyst per hour (h ^-1 ). Preferably, the HOSG of the stream for regeneration of more than about 120 h ^-1 and more preferably more than about 360 h ^-1 . Typically, measured as the total flow under operating conditions, the hourly space velocity of the gas stream for regeneration is less than about 12000 h ^-1 , preferably less than about 3600 h ^-1 and more preferably less than 720 h ^-1 .

In the regeneration zone, the gas residence time measured under operating conditions, calculated as the height of the regeneration zone, multiplied by the void fraction of the regeneration zone, divided by the gas flow rate per unit section of the sum of the flows of the starting materials for regeneration and reaction, more than about 0.3 s. The "void fraction of the regeneration zone" is a part of the regeneration zone that is empty. Preferably, the residence time of the gas in the regeneration zone is more than about 1 s and more preferably more than about 5 s. Typically, the gas residence time in the regeneration zone, measured under operating conditions, is less than about 60 seconds, preferably less than about 30 seconds, and more preferably less than about 10 seconds.

When the alkyl aromatic compound and, optionally, the vapor is brought into contact with the dehydrogenation catalyst in the manner described above, the vinyl aromatic compound is obtained. Ethylbenzene is converted primarily to styrene. Ethyl toluene is converted to p-methyl styrene (p-vinyl toluene). Tert-butyl ethylbenzene is converted to tert-butyl styrene. Isopropylbenzene (cumene) is converted to α-methylstyrene and diethylbenzene is converted to divinylbenzene. Hydrogen is also formed during dehydrogenation. Other products formed in lower yields include benzene and toluene.

The conversion of the alkyl aromatic compound in the process of this invention may vary depending on the specific composition of the starting material, the composition of the catalyst, process conditions, and fluidized bed conditions. For the purposes of this invention, “conversion” is defined as the molar percentage of the alkyl aromatic compound that has been converted to all products. In this process, the conversion of the alkyl aromatic compound is usually more than about 30 mol. percent, preferably more than about 50 mol. percent and more preferably more than about 70 mol. percent.

Similarly, selectivity to products will vary depending on the specific composition of the starting material, the composition of the catalyst, process conditions, and fluidized bed conditions. For the purposes of this invention, “selectivity” is defined as the molar percentage of the converted alkyl aromatic compound that forms a particular product, preferably a vinyl aromatic compound. In the process of this invention, selectivity to a vinyl aromatic compound, preferably to styrene or a substituted styrene derivative, usually more than about 60 mol. percent, preferably more than about 75 mol. percent and more preferably more than about 90 mol. percent.

The invention will be further explained by considering the following examples, which are intended only to illustrate the application of the invention. Other embodiments of the invention will be apparent to those skilled in the art from consideration of this specification or practice of the invention disclosed herein. Selectivity measurements are adjusted for deviations from 100 percent in the organic material balance.

Example 1

The fluidized bed reactor [4.25 inches (10.63 cm) inner diameter, 20 inches (50 cm) height] is constructed as shown in FIG. The reactor contains a single vertical vessel functionally divided into three zones: a catalyst regeneration zone (1) at the bottom of the reactor, a zone (3) above the bed at the top of the reactor, and a reaction zone (2) in the middle part between the regeneration zone and the zone above the bed. The first inlet (4) at the bottom of the reactor leads to the discharge region (10), where the gas distributor (9) is located. The specified first inlet is used to distribute the flow of the source material for regeneration in the regeneration zone. A second inlet (5) located in the region above the bed is used to introduce a dehydrogenation feed stream into the reaction zone (2). A second inlet is connected to the inlet pipe, which terminates in a plurality of sprinklers (6) in the reaction zone at a height of 3 inches (7.5 cm) above the bottom distribution plate (9). A plurality of sprinklers in the form of six rows of sintered metal pipes [Inconel, 1/4 inch WD (6.3 mm WD)] are designed to provide a uniform pressure drop across the sprinkler. Outlets in the sprayer are placed horizontally. The outlet (7) is located in the area above the layer to remove the product stream. Solids entrained in the product stream are collected in a cyclone (located below outlet 7) and then returned to the reactor through a third inlet (8) located in the regeneration zone. Exhaust gases are collected downstream of the cyclone. The reactor is also equipped with (electrical) resistances for heating the reactor and two internal thermocouples (K type) for measuring the temperature of the fluidized bed in the reaction and regeneration zones.

The reactor is used to dehydrogenate ethylbenzene in the presence of a dehydrogenation catalyst to styrene while simultaneously and continuously regenerating the dehydrogenation catalyst. A dehydrogenation catalyst (2370 g) having an average particle diameter of 300 μm and containing 28.7 percent iron oxide (Fe ₂ O ₃ ), 14.3 percent cerium oxide (Ce ₂ O ₃ ), 7.6 percent copper oxide (CuO) , 31.6 percent of potassium carbonate (K ₂ CO ₃ ), 0.6 percent of chromium oxide (Cr ₂ O ₃ ), 9.5 percent of zinc oxide (ZnO) and 7.6 percent of cement by weight are charged to the reactor. The reaction feed stream contains a mixture of ethylbenzene and steam. The feed stream for regeneration contains steam. Gaseous products are analyzed using a Carle gas chromatograph equipped with a parallel row of five columns (2.7 percent Carbowax® 1540 on Porasil C; 3 percent Carbowax® 1540 on Porasil C; 27 percent Bis (EE) A on Chromosorb® PAW; Porapak® Q and two columns of 13X molecular sieves). Liquid products were analyzed using an HP 5890 gas chromatograph equipped with a DB-5 J&W column. Nitrogen is used as the internal standard for gas analysis, while heptane is used as the internal standard for liquid analysis. Sampling is carried out for six hours, taking four or more samples every 30 minutes during the last few hours of operation. The results presented here for the conversion of ethylbenzene and selectivity for styrene are the average of four or more samples taken.

In the reactor described above, water at a feed rate of 4.3 cm ³ / min at room temperature is heated to 600 ° C and added through the inlet (4) to the discharge region (10) and through the distribution plate (9) into the regeneration zone (1) at the bottom of the reactor. Ethylbenzene liquid at a feed rate of 2.5 cm ³ / min and nitrogen gas at a feed rate of 1088 cm ³ / min at room temperature are mixed together, heated to 500 ° C and added to the dehydrogenation reaction zone through the inlet (5) and a lot of sprayers ( 6). The feed rates correspond to the mass ratio of total steam to oil 2/1 with a surface speed of 1.86 m / min in the regeneration zone and 237 m / min in the reaction zone. The gas residence time in the regeneration zone is 1.46 s, the gas residence time in the reaction zone is 0.67 s. The temperature and pressure in the reactor are maintained at 600 ° C. and 15.5 psi. inch (106.9 kPa), respectively. Products obtained through the outlet (7) are analyzed as described above. The conversion of ethylbenzene is 74.0 mol. percent. The styrene selectivity is 86.0 mol. percent. Other products contain benzene and toluene. Material balance is calculated for 95 weight percent organic feed.

Example 2

Using the reactor and catalyst of Example 1, water at a feed rate of 2.17 cm ³ / min is heated to 600 ° C. and added to the distribution plate in the regeneration zone. Ethylbenzene liquid at a feed rate of 2.52 cm ³ / min and liquid water at a feed rate of 2.17 cm ³ / min at room temperature are heated to 500 ° C. and added through a reaction zone to a plurality of sprinklers. The indicated feed rates correspond to the mass ratio of total steam to oil 2/1 with a surface speed of 156 m / min in the regeneration zone and 339.5 m / min in the reaction zone. The gas dwell time in the regeneration zone is 2.91 s, the gas dwell time in the reaction zone is 0.78 s. The temperature and pressure in the reactor are maintained at 600 ° C. and 15.5 psi. inch (106.9 kPa), respectively. The conversion of ethylbenzene is 85 mol. percent. The styrene selectivity is 69 mol. percent. Material balance is calculated for 96 weight percent organic feed.

In Example 1, nitrogen was added as a cleaning gas to the ethylbenzene stream, but no vapor was added to the ethylbenzene stream. In contrast, in Example 2, no cleaning gas was added to the ethylbenzene stream, and the steam stream was distributed between the dehydrogenation feed and the regeneration feed. When Example 2 is compared with Example 1, it is obvious that the conversion of ethylbenzene is higher in Example 2, due to longer residence times in the layer, and selectivity for styrene is lower due to enhanced free radical cracking in the region above the layer.

Example 3

Using the reactor and catalyst of Example 1, water at a feed rate of 4.3 cm ³ / min is heated to 600 ° C. and added to the distribution plate in the regeneration zone. Ethylbenzene liquid at a feed rate of 2.49 cm ³ / min at room temperature is heated to 500 ° C. and added through a reaction zone to a plurality of sprinklers. The indicated feed rates correspond to the mass ratio of total steam to oil 2/1 with a surface speed of 309 m / min in the regeneration zone and 417 m / min in the reaction zone. The gas residence time in the regeneration zone is 1.47 s, the gas residence time in the reaction zone is 0.63 s. The temperature and pressure in the reactor are maintained at 600 ° C. and 15.5 psi. inch (106.9 kPa), respectively. The conversion of ethylbenzene is 85 mol. percent. The styrene selectivity is 72 mol. percent. Material balance is calculated for 98 weight percent organic feed.

The process conditions of example 3 are closely similar to the conditions of example 2 with the following exception. In Example 2, half of all the steam is delivered to the regeneration zone and half of all the steam is delivered to the reaction zone. In contrast, in Example 3, all the steam is delivered to the regeneration zone. When example 3 is compared with example 2, it is obvious that the conversion of ethylbenzene and selectivity for styrene are comparable. There is a slight discrepancy, which depends on the injection site.

Example 4

Example 2 is repeated under closely similar process conditions except that the pressure is kept constant at about 5 psi. inch (34.5 kPa). The catalyst used in example 4 has a chemical composition identical to the catalyst from the previous examples, however, the amount of catalyst used is 1355 g, and the catalyst has an average particle diameter of 220 μm. The process conditions are as follows: flow rate of water into the regeneration zone 2.9 cm ³ / min; the feed rate of liquid ethylbenzene and water into the reaction zone is 2.52 cm ³ / min and 1.45 cm ³ / min, respectively; mass ratio of steam to oil 2/1; surface flow velocity in the regeneration zone 123 cm / min; surface flow rate in the reaction zone 200 cm / min and a temperature of 600 ° C. The conversion of ethylbenzene is 49 mol. percent. The styrene selectivity is 88 mol. percent. Material balance is calculated for 93 weight percent organic feed.

A comparison of examples 2 and 4 shows that significantly higher styrene selectivities can be achieved by operating a fluidized bed reactor under vacuum. The lower partial pressure of the ethylbenzene feed slightly reduces the total conversion.

Example 5

Example 4 is repeated under closely similar process conditions, except that the temperature of the reactor is maintained constant at 590 ° C rather than 600 ° C. The process conditions are as follows: flow rate of water into the regeneration zone 2.9 cm ³ / min; the feed rate of liquid ethylbenzene and water into the reaction zone is 2.52 cm ³ / min and 1.45 cm ³ / min, respectively; mass ratio of steam to oil 2/1; surface flow velocity in the regeneration zone 122 cm / min; surface flow rate in the reaction zone 199 cm / min and a pressure of 5 psi inch (34.5 kPa). The conversion of ethylbenzene is 50 mol. percent. The styrene selectivity is 94 mol. percent. Material balance is calculated for 99 weight percent organic feed. A comparison of examples 4 and 5 shows that operation under vacuum at a temperature lower than 600 ° C provides a further increase in styrene selectivity.

Example 6

Example 4 is repeated under closely similar process conditions, except that the temperature of the reactor is maintained constant at 580 ° C rather than 600 ° C. The process conditions are as follows: the rate of water supply to the regeneration zone is 2.83 cm ³ / min; the feed rate of liquid ethylbenzene and water into the reaction zone is 2.52 cm ³ / min and 1.45 cm ³ / min, respectively; mass ratio of steam to oil 2/1; surface flow velocity in the regeneration zone 121 cm / min; surface flow rate in the reaction zone of 197 cm / min and a pressure of 5 psi inch (34.5 kPa). The conversion of ethylbenzene is 44 mol. percent. The styrene selectivity is 95 mol. percent. Material balance is calculated for 100 mass percent of organic starting material. A comparison of examples 4, 5 and 6 shows that operation under vacuum at a temperature lower than 600 ° C provides a further increase in styrene selectivity.

Example 7

Example 4 is repeated under closely similar process conditions except that the mass ratio of steam to oil is 1/1 instead of 2/1. Other process conditions are as follows: water flow rate to the regeneration zone 1.45 cm ³ / min; the feed rate of liquid ethylbenzene and water into the reaction zone is 2.52 cm ³ / min and 0.73 cm ³ / min, respectively; surface flow rate in the regeneration zone 61.5 cm / min; surface flow rate in the reaction zone 107.6 cm / min; pressure 5 psi inch (34.5 kPa) and a temperature of 600 ° C. The conversion of ethylbenzene is 49 mol. percent. The styrene selectivity is 89 mol. percent. Material balance is calculated for 98 weight percent organic feed.

A comparison of examples 4 and 5 with example 7 shows that a decrease in the ratio of steam to oil from 2/1 to 1/1 does not affect the conversion of ethylbenzene and selectivity for styrene.

Example 8

Example 4 is repeated under closely similar process conditions, with the exception of the particle size of the catalyst and the ratio of steam to oil. For example 8, the catalyst (1570 g) has an average particle diameter of 82 μm, and the ratio of vapor to oil is 0.5 / 1. Other process conditions are as follows: the rate of water supply to the regeneration zone is 0.8 cm ³ / min; the feed rates of liquid ethylbenzene and water into the reaction zone are 5.48 cm ³ / min and 0.54 cm ³ / min, respectively; surface flow velocity in the regeneration zone 33.52 cm / min; surface flow rate in the reaction zone 74.8 cm / min; pressure 5 psi inch (34.5 kPa) and a temperature of 600 ° C. The conversion of ethylbenzene is 54 mol. percent. The styrene selectivity is 95 mol. percent. Material balance is calculated for 100 mass percent of organic starting material.

A comparison of examples 4 and 7 with example 8 shows that high selectivity for styrene can be achieved with a steam to oil ratio of 0.5 / 1. In addition, a decrease in the average particle diameter of the catalyst from 220 μm to 82 μm results in an increase in the conversion of ethylbenzene. It is most likely that this result is a result of improved mass transfer, since smaller catalyst particles tend to decrease the diameters of the equilibrium bubbles in the fluidized bed reactor.

Example 9

A pulsating reactor is used to study the dehydrogenation of ethylbenzene to styrene as a function of time. In a pulsating mode, the reactor is recycled through the dehydrogenation step and then the catalyst regeneration step. Experiments in pulsating reactors show that such results can be expected in a fluidized bed reactor.

A dehydrogenation catalyst having a particle size between 1.18 mm and 1.70 mm and containing 33.2 percent iron oxide (Fe ₂ O ₃ ), 17.5 percent cerium oxide (Ce ₂ O ₃ ), 7.8 percent copper oxide (CuO), 36.0 percent potassium carbonate (CaCO ₃ ), 0.6 percent chromium oxide (Cr ₂ O ₃ ) and 4.7 percent cement by weight are charged to a continuous flow reactor with a fixed bed [304 stainless steel inventory 40, 1 inch (2.5 cm) ID × 36 inches (90 cm) length]. The catalyst bed occupies 7 inches (17.5 cm) of reactor length. The space above the layer is filled with Berl ceramic saddle nozzles (1/4 inch, 0.6 cm). Below the layer is a metal spacer. The reaction temperature is measured in a thermocouple pocket embedded in a catalyst bed. The dehydrogenation pulse is carried out by passing ethylbenzene, preheated to 550 ° C, over the catalyst for 2 minutes. The flow rate of liquid ethylbenzene, measured at ambient temperature and pressure (taken as 23 ° C and 1 atm), is 1.16 ml / min. At the same time, water preheated to 550 ° C is passed over the catalyst for the same 2-minute period. The water flow rate is regulated to provide a mass ratio of steam to oil of 0.30 / 1. Total pressure support 5.0 psi inch (34.5 kPa). After that, the supply of ethylbenzene is stopped and a regeneration pulse is carried out, passing only a stream of water, previously heated to 550 ° C, over the catalyst for 2 minutes under the same process conditions. The flow rate of liquid water during a regeneration pulse, measured at 24 ° C and 1 atm, is 1 ml / min. After regeneration, the dehydrogenation pulse is repeated by re-introducing the ethylbenzene stream for 2 minutes, as previously indicated, with a continuous supply of water vapor. After 2 minutes, the ethylbenzene supply was again stopped, while the steam supply was continued for the regeneration cycle for 2 minutes. The dehydrogenation-regeneration cycles are repeated throughout the entire experiment 200 hours. The product stream is continuously passed through a condenser, separated and analyzed by conventional methods.

The pulsed process results are shown in FIG. 3, which graphically represents the conversion of ethylbenzene and selectivity for styrene as a function of test time at constant temperature (550 ° C.) and pressure (5 psi, 34.5 kPa). It was unexpectedly found that the conversion of ethylbenzene slightly increases with time. The styrene selectivity remains constant at a value of more than 95 mol. percent throughout the experience. The results from the pulsating reactor show that the dehydrogenation catalyst can be circulated through dehydrogenation-regeneration steps in a fluidized bed reactor for extended periods of time without significant decontamination.

Comparative experiment 1.

The process of Example 9 is repeated in the same continuous flow and fixed bed reactor under similar process conditions, except that the dehydrogenation is carried out in a continuous mode rather than a pulsed mode. Thus, there is only one dehydrogenation cycle and no catalyst regeneration cycle. It was found that under such conditions, the catalyst is gradually deactivated with a concomitant decrease in the conversion of ethylbenzene. When the catalyst is deactivated, the process temperature is raised to maintain a constant conversion of ethylbenzene. With a steam / oil ratio of 0.3 / 1, the temperature should be increased at a rate of 0.45 ° C. per min to maintain the conversion. When comparative experiment 1 is compared with experiment 9, it is found that the catalyst life achieved in the pulsating reactor is significantly extended at constant temperature and pressure, whereas without regeneration, the catalyst is rapidly deactivated and requires an increase in temperature to ensure a constant conversion. The results from the pulsating reactor show that the dehydrogenation catalyst can be circulated through dehydrogenation-regeneration steps in a fluidized bed reactor for extended periods of time without significant decontamination.

Claims (38)

1. A method for dehydrogenating alkyl aromatic compounds over a dehydrogenation catalyst to form vinyl aromatic compounds and regenerating the dehydrogenation catalyst in situ, comprising (a) fluidizing the dehydrogenation catalyst in a single-shell fluidized bed reactor containing a reaction zone and a regeneration zone under fluidization conditions, so that the catalyst is circulated inside and between the two zones; (b) reacting a dehydrogenation feed stream containing an alkyl aromatic compound and possibly steam with a dehydrogenation catalyst in the reaction zone under reaction conditions sufficient to produce the corresponding vinyl aromatic compound; and (c) reacting the feed stream steam-containing regeneration material with a dehydrogenation catalyst located in the regeneration zone under sufficient regeneration conditions to regenerate the catalyst, at least in part. 2. The method according to claim 1, where the alkyl aromatic compound is a C ₈ -C ₂₀ alkyl aromatic compound. 3. The method according to claim 2, where the alkyl aromatic compound is ethylbenzene or substituted ethylbenzene. 4. The method according to claim 2, where the alkyl aromatic compound is selected from isopropylbenzene, diethylbenzene and ethyl toluene. 5. The method according to claim 1, where the fluidized bed reactor further comprises a zone above the bed, and a cleaning gas is added to the zone above the bed. 6. The method according to claim 1, where the mass ratio of total steam to the alkyl aromatic compound is more than about 0.2 / 1 and less than about 5.0 / 1. 7. The method according to claim 1, where the mass ratio of total steam to the alkyl aromatic compound is more than about 0.2 / 1 and less than about 1.2 / 1. 8. The method according to claim 1, where the dilution gas is supplied with a stream of starting material for dehydrogenation, or served with a stream of starting material for regeneration, or served with both streams. 9. The method of claim 8, wherein the dilution gas is selected from nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof. 10. The method of claim 8, wherein the diluent comprises from more than about 10 vol.% To less than about 90 vol.% Of the feed stream for dehydrogenation or for regeneration, or both, independently. 11. The method according to claim 1, where the flow of the starting material for dehydrogenation is preheated to a temperature of more than about 150 ° C and less than about 600 ° C. 12. The method according to claim 1, where the stream of starting material for regeneration is preheated to a temperature of more than about 200 ° C and less than about 650 ° C. 13. The method according to claim 1, where the temperature in the reaction and / or regeneration zones is more than about 550 ° C and less than about 650 ° C. 14. The method according to claim 1, where the total pressure in the reactor is more than about 1 psi. inch (6.9 kPa) and less than about 73 psi inch (503.3 kPa). 15. The method according to claim 1, where the process is carried out at a gas hourly space velocity calculated as the total flow of the starting material for dehydrogenation, more than about 60 h ^-1 and less than about 12000 h ^-1 , measured under operating conditions of the process. 16. The method according to claim 1, where the process is carried out at a residence time of the entire gas stream in the reaction zone of more than about 0.3 s and less than about 60 s, measured under operating conditions of the process. 17. The method according to claim 1, where the process is carried out at a gas hourly space velocity calculated as the total flow rate of the source material for regeneration, more than about 60 h ^-1 and less than about 12000 h ^-1 , measured under operating conditions of the process. 18. The method according to claim 1, where the process is carried out at a residence time of the gas in the regeneration zone of more than about 0.3 s and less than about 60 s, measured under operating conditions of the process. 19. The method according to claim 1, where the dehydrogenation catalyst contains iron oxide. 20. The method according to claim 19, where the dehydrogenation catalyst further comprises at least one or more compounds selected from compounds of alkali metals, alkaline earth metals, chromium, gallium, cerium, zinc and copper. 21. The method according to claim 19, where the dehydrogenation catalyst contains (a) at least one iron oxide, (b) at least one carbonate, bicarbonate, oxide or hydroxide of potassium and / or cesium, (c) oxide, carbonate, nitrate or cerium hydroxide, (d) possibly hydroxide, carbonate, bicarbonate, acetate, oxalate, nitrate or sodium sulfate, (e) possibly carbonate, sulfate or calcium hydroxide, (f) possibly one or more zinc, chromium and copper compounds and (g) possibly cement. 22. The method according to claim 1, where the conversion of the alkyl aromatic compound is more than about 30 mol.%. 23. The method according to claim 1, where the selectivity for vinyl aromatic compound is more than about 60 mol.%. 24. The method according to claim 1, where the vinyl aromatic compound is styrene or a substituted derivative of styrene. 25. The method according to paragraph 24, where the substituted styrene is selected from divinylbenzene, α-methylstyrene and vinyltoluene. 26. The method according to claim 1, where the average particle size of the dehydrogenation catalyst is more than about 20 microns and less than about 1000 microns. 27. The method according to claim 1, where the fluidized bed reactor contains a single vertical body, covering the area above the bed, the reaction zone and the regeneration zone; an inlet for introducing a stream of starting material for regeneration into the regeneration zone and an inlet for introducing a stream of starting material of reagent into the reaction zone, one of these inlet devices in the reaction or regeneration zones can separate the two zones, while allowing the circulation of catalyst particles between the two zones; and further comprises outlet devices for the effluent, and possibly inlet devices for returning to the reactor a catalyst entrained in the effluent, and possibly inlet and outlet devices for transporting the catalyst to and from the reactor. 28. The method according to item 27, where the device for separating reaction zones and regeneration contains many sprinklers or dispenser. 29. A method of dehydrogenating ethylbenzene or substituted ethylbenzene over a dehydrogenation catalyst to form styrene or substituted styrene and regenerating the dehydrogenation catalyst in situ, comprising (a) fluidizing the dehydrogenation catalyst in a single-bed fluidized bed reactor containing a reaction zone and a regeneration zone, under conditions that circulate the catalyst inside and between the two zones; (b) the interaction of ethylbenzene or substituted ethylbenzene and possibly steam, and possibly dilution gas with a dehydrogenation catalyst in the reaction zone, the catalyst containing iron oxide, and the interaction is carried out with a mass ratio of steam to ethylbenzene of more than about 0.2 / 1 and less than about 3.0 / 1, a temperature of more than about 570 ° C. and less than about 610 ° C., and a total reactor pressure of more than about 3 psi. inch (41 kPa) and less than about 44 psi inch (302 kPa), and (c) the interaction of the dehydrogenation catalyst located in the regeneration zone with a regeneration feed stream containing steam and possibly a diluent at a temperature of more than about 570 ° C and less than about 610 ° C so to regenerate the catalyst, at least in part. 30. A fluidized bed reactor for the dehydrogenation of alkyl aromatic hydrocarbons with in-place catalyst regeneration, comprising a single vertical vessel spanning the zone above the bed, the reaction zone, and the regeneration zone; an inlet device for introducing a stream of starting material for regeneration into the regeneration zone and an inlet device for introducing a stream of starting material of the reagent into the reaction zone, one of the inlet devices being able to separate the reaction and regeneration zones, while allowing the circulation of catalyst particles between the two zones; and further comprises outlet devices for the effluent, and possibly inlets for returning to the reactor a catalyst entrained in the effluent. 31. The fluidized bed reactor of claim 30, wherein the device for introducing a stream of reactant feed material comprises a plurality of sprinklers or a distributor. 32. The fluidized bed reactor of claim 30, wherein the means for introducing a feed stream for regeneration comprises a plurality of sprinklers or a distributor. 33. The fluidized bed reactor according to claim 30, wherein the device for separating reaction and regeneration zones is selected from a plurality of sprinklers or dispenser. 34. The fluidized bed reactor of claim 30, further comprising a device for improving the circulation of solids. 35. The fluidized bed reactor according to clause 34, where the device for improving the circulation of solids contains a suction pipe, possibly containing internal reflective walls. 36. The fluidized bed reactor according to clause 34, where the device for improving the circulation of solids contains a suction pipe made of heating or cooling elements. 37. The fluidized bed reactor according to clause 34, further comprising inlet devices and outlet devices for transporting the catalyst into and out of the reactor. 38. The fluidized bed reactor according to clause 34, further comprising at least one device for measuring the temperature of the fluidized bed and, possibly, heating devices.

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