CN1204095C - Dehydrogenation of alkyl aromatic compound and catalyst generation in fluidized bed reactor - Google Patents

Abstract

A process of preparing a vinyl aromatic compound, such as styrene. The process involves fluidizing a dehydrogenation catalyst in a single shell fluidized bed reactor containing a reaction zone and a regeneration zone; contacting an alkyl aromatic compound, such as ethylbenzene, with the dehydrogenation catalyst in the dehydrogenation zone so as to produce the vinyl aromatic compound, such as styrene; and regenerating the catalyst <i>in situ</i> by contacting steam with the deactivated catalyst in the regeneration zone. A fluidized bed reactor is described, characterized by a freeboard zone, a reaction zone, and a catalyst regeneration zone, all within a single shell.

Description

Dehydrogenation of Alkyl Aromatic Compounds and Catalyst Regeneration in a Fluidized Bed Reactor

This application claims priority to US Provisional Application Serial No. 60/172,274, filed December 17,1999.

In one aspect, the invention relates to a process for the dehydrogenation of an alkylaromatic compound, such as ethylbenzene, to a vinylaromatic compound, such as styrene. In addition, the present invention relates to a catalyst regeneration process for the dehydrogenation of alkylaromatic compounds. In another aspect, the invention relates to a fluidized bed reactor in which an organic process, such as the aforementioned dehydrogenation process, is carried out.

The dehydrogenation of alkyl aromatic compounds, such as ethylbenzene, cumene, diethylbenzene, p-ethyltoluene, can be applied to the preparation of styrene and substituted derivatives of styrene, such as α-methylstyrene, di Vinylbenzene and p-methylstyrene. Styrene and its substituted derivatives are also available as polystyrene, styrene-butadiene rubber (SBR), acrylonitrile-butadiene-styrene (ABS), styrene-acrylonitrile (SAN) and unsaturated Monomer of polyester resin.

Fluidized bed reactors are important for a number of catalytic organic processes including dehydrogenation processes.

For vinyl aromatic compounds, such as styrene, the main manufacturing route is the direct catalytic dehydrogenation of alkyl aromatic compounds, such as ethylbenzene. Patents disclosing this method include, for example, US Patent 4,404,123, US Patent 5171,914, US Patent 5,510,552, and US Patent 5679,878. The catalyst generally comprises iron oxide and, additionally, possibly chromium oxide and a potassium compound, such as potassium hydroxide, or potassium carbonate as cocatalyst. Because the process is highly endothermic, energy is absorbed by introducing a stream of high heat vapor into the process reactor, the aforementioned process reactor representing a fixed bed design. The process temperature is generally between 550°C and 700°C. Side reactions are controlled by keeping the partial pressure of ethylbenzene low.

The dehydrogenation process described above using a fixed bed reactor has a number of disadvantages. First, fixed bed reactors, featuring a fixed bed of catalyst particles, are difficult to uniformly heat to high temperatures. In the endothermic dehydrogenation of ethylbenzene to styrene, the downstream (outlet end) of the fixed catalytic bed is cooler than the upstream catalytic bed (inlet end). The feed stream is generally preheated to a higher temperature than required because the difference in temperature may result in reduced conversion of ethylbenzene at the downstream outlet end of the reactor. As a result, the catalyst bed near the inlet end of the reactor may overheat and generally fail faster than the catalyst at the farther downstream end. In this case, large catalyst beds are required to withstand long run cycles. Another point, in the process of dehydrogenating ethylbenzene to styrene in a fixed bed, during the dehydrogenation process, steam and ethylbenzene enter in parallel to promote the in-situ regeneration of the catalyst. Generally, a high steam to ethylbenzene weight ratio is required, generally higher than 1.2 to 2.0 or possibly higher. The disadvantage of this process is that it requires high energy input and large amounts of recirculated water. (The steam to alkyl aromatic compound weight ratio is hereinafter referred to as "gas to oil ratio") When the overall bed activity decreases beyond the operable point, the catalyst must be replaced. As a further disadvantage, fixed bed reactors are typically shut down for several weeks due to catalyst change.

Reference is also made to oxidative dehydrogenation processes as disclosed in U.S. Patent No. 3,651,160 and U.S. Patent No. 4,471,146, wherein in the process ethylbenzene is Styrene is formed on contact with oxygen in a fluidized bed reactor. A conventional fluidized bed reactor comprises a single reaction zone where catalyst particles are separated and circulated. Fluidized bed reactors provide a more isothermal temperature profile compared to fixed beds. An isothermal catalyst bed generally results in less catalyst damage and higher product yields. When the catalyst fails completely and cannot be regenerated, the fluidized bed reactor can easily replace the catalyst. Because the catalyst can be viewed as a fluid, spent catalyst can be drained from the reactor and active catalyst can be added to the reactor without stopping the chemical process. The above-mentioned oxidative dehydrogenation process can also be carried out when a part of the catalyst is continuously sent to the regeneration device for regeneration in an oxygen environment, and then the regenerated catalyst is recycled into the oxidative dehydrogenation reactor. Because the oxidation side reaction is difficult to control, the oxidative dehydrogenation process can produce a small amount of vinyl aromatic compound product at the downstream outlet port as the alkyl aromatic compound and oxygen co-enter. Additionally, the safety issues involved in handling and processing mixtures of organic compounds and oxygen are important. As a further disadvantage, the continuous circulation of the catalyst between the fluidized bed reactor and the regeneration unit requires complex equipment and often highly friction-resistant catalyst particles.

Some patents, such as U.S. Patent 4,229,604 and U.S. Patent 5,510,553, disclose ethylbenzene in the absence of oxygen but in the presence of oxidative dehydrogenation catalysts, such as magnesium oxide-modified silicon, or metal oxide-supported reducible vanadium oxides. Dehydrogenation to styrene requires the use of a transport reactor. These processes reduce the risk of using mixtures of alkylaromatic compounds and oxygen; however, the useful life of such catalysts is short-lived. Therefore, the catalyst must be continuously circulated between the process reactor and the regeneration unit, where the catalyst is regenerated in the presence of oxygen. As mentioned above, the continuous circulation of catalyst between the fluidized bed reactor and the regeneration unit requires complex equipment and often requires highly friction-resistant catalyst particles.

The aforementioned descriptions indicate a need to improve the dehydrogenation process of catalysts. It will have many advantages, such as, discovery of the process of dehydrogenation of alkylaromatics, such as ethylbenzene, to vinylaromatics, such as styrene, providing efficient in-situ catalyst regeneration at economical gas-to-oil ratios. It would be very beneficial if, during this process, an active catalyst replaced the spent catalyst without having to shut down the reactor or use complex transport devices. It is also beneficial if the isothermal bed temperature can be maintained. Oxygen will complicate safety issues and handling procedures, and it will also be beneficial if the process does not require oxygen. Finally, it would be more desirable if the process had all the aforementioned characteristics and could achieve high yields of vinyl aromatic compounds.

On the other hand, U.S. Patent 4,152,393 discloses a reactor consisting of a single shell containing a reaction zone and a regeneration zone, which are arranged in a certain way, in particular, as having concentric walls and rails The collection of solid particles can pass from the regeneration zone to the reaction zone through the air flow through the first path, and then return to the regeneration zone through the second path. Gases passing through the regeneration zone are no longer transferred to the reaction zone, and gases passing through the reaction zone are no longer transferred to the regeneration zone. Such reactors are known to be very useful in the ammoxidation of propylene. The disclosed reactors will exhibit high stagnation flow, characterized by gas bubbles flowing along the inner walls of the reactor. The disadvantage is that stagnant flow reduces the contact of the gas phase reactants with the solid catalyst particles, thus reducing the yield of the process. As a further disadvantage, this reactor has a narrow bending space and multiple gas injections in this space, which can lead to a high degree of friction of the catalyst particles. US Patent 6,048,459 discloses a method of fluidizing a granular bed material which involves collecting and moving up a portion of the fluid above the fluidized bed and circulating the ascending fluid down the fluidized bed through a downcomer positioned in the bed. The granular material zone can be extended below the fluidized bed for use, especially for treating long-term accumulated anaerobic sludge.

In one aspect, the present invention is a method for catalyzing the dehydrogenation of alkyl aromatic compounds to generate vinyl aromatic compounds through a dehydrogenation catalyst in a single-shell fluidized bed reactor and regenerating the dehydrogenation catalyst in situ. The method of the present invention comprises, the first step: (a) comprises the fluidized dehydrogenation catalyst in the single-shell fluidized bed reactor of a reaction zone and a regeneration zone, under fluidized state, catalyst is in zone and two Interval cycle. (b) contacting the alkyl aromatic compound, optionally the vapor stream, with a dehydrogenation catalyst residing in the reaction zone under reaction conditions sufficient to produce the corresponding vinyl aromatic compound. (c) contacting the dehydrogenation catalyst residing in the regeneration zone with a stream of steam under regeneration conditions sufficient to regenerate, at least partially regenerate, the dehydrogenation catalyst.

The dehydrogenation process of the present invention has more advantages than the previous process, and it is applied to the preparation of industrially important vinyl aromatic compounds, such as styrene, p-methylstyrene, α-methylstyrene, diethylene Benzene. First, the method of the present invention does not require oxygen. The corresponding safety problems associated with the handling and processing of mixtures of organic compounds and oxygen found in some previous processes are eliminated in the process of the present invention. Second, in the method of the present invention, the gas to oil ratio is preferably lower than the gas to oil ratio used in the previous fixed bed process. Therefore, the process of the present invention uses less water for circulation, is more energy efficient and economical than prior art processes. An additional advantage is that the process of the invention, carried out in a fluidized bed reactor, is substantially isothermal. Problems associated with non-uniform bed temperature, such as overheating and catalyst damage at the inlet end upstream of the catalytic bed, and low productivity at the outlet end downstream of the catalytic bed, are substantially eliminated. Furthermore, the formation of thermal by-products is reduced. As a further advantage, the process of the present invention can use fewer catalytic beds than fixed bed catalytic bed sizes used in prior art processes, yet still achieve comparable operating cycles. Another advantage is that the process of the present invention provides for continuous in situ regeneration of the dehydrogenation catalyst. The process of the present invention eliminates the need to shut down the reactor for regeneration of the catalyst or for removal of the catalyst from the fluidized bed to a regeneration unit. Therefore, the method of the present invention possesses simple design and operation. Furthermore, the catalysts used in the process of the invention are not required to be highly friction-resistant for transfer to the reactor. Finally, when the catalyst cannot be further regenerated, the method of the present invention provides for replacement of the spent catalyst but the dehydrogenation process can continue. Because the solid catalyst is treated as a fluid, the spent catalyst is simply transported out of the reactor and fresh catalyst is transported in without stopping the operation. Therefore, catalyst regeneration and replacement can be done simultaneously without stopping the dehydrogenation process, resulting in higher yields. Most advantageously, the process of the present invention produces high yields of vinyl aromatic compounds, preferably high yields of styrene.

In another aspect, the invention relates to fluidized bed reactors that allow simultaneous chemical processing and catalyst regeneration. The fluidized bed reactor of the present invention includes a single-layer vertically placed shell, and the inner space is divided into a dilute phase zone, a reaction zone and a regeneration zone. The reactor also includes an inlet device to direct a feed stream into the regeneration zone, and an inlet device to direct a reactant feed stream into the reaction zone. The reactor further includes means to separate the reaction zone from the regeneration zone while allowing continuous circulation and extensive backmixing of the catalyst between the two zones. In a preferred embodiment, one of the inlets for the reaction or regeneration feed serves as a means of separating the reaction zone and the regeneration zone. The reactor of the present invention also includes an outlet device, preferably in the dilute phase zone, for withdrawing an effluent comprising product and some unconverted reactants and regeneration feedstock. Optionally, the reactor further comprises a device for re-injecting the catalyst particles entrained in the effluent into the reactor. Optionally, inlet and outlet devices can feed catalyst into or out of the reactor.

In the reactor of the present invention, the reactant and regeneration feed streams may be cross-mixed; although, preferably, the reaction and regeneration inlets physically separate the two processes. However, back-mixing of gases and solids is allowed in this design, and the regeneration and reactant feed streams can be chemically compatible, as exemplified here, by a transient dehydrogenation-regeneration process.

The fluidized bed reactor of the present invention can be used in a variety of catalytic organic processes including, for example, dehydrogenation, oxidation, halogenation. The fluidized bed reactor of the present invention can be used in a particularly important dehydrogenation process involving the dehydrogenation of alkyl aromatic compounds, such as ethylbenzene, to vinyl aromatic compounds, such as styrene. Advantageously, the fluidized bed reactor of the present invention provides a chemical processing reaction zone and a catalyst regeneration zone within a single layer fluidized bed reactor shell. Thus, the catalyst can be simultaneously regenerated in situ during the desired chemical process. Spent catalyst need not be conveyed out of the fluidized bed reactor of the present invention to be regenerated in a separate regeneration vessel facility. Therefore, the catalyst here is subject to much less pressure and losses than in a transport reactor. As another advantage, spent catalyst can be replaced on-line without shutting down the chemical reaction, by simply leading the spent catalyst out of the reactor and introducing fresh catalyst into the reactor. These advantages provide for the advances required for fluidized bed processes.

Figure 1 shows a side and top view of a first preferred embodiment of a fluidized bed reactor according to the invention, the details of which follow.

Figure 2 shows a side and top view of a second preferred embodiment of a fluidized bed reactor according to the invention, the details of which follow.

Figure 3 is a schematic representation of ethylbenzene conversion and styrene selectivity versus reaction time during ethylbenzene dehydrogenation and catalyst regeneration occurring in a pulsed reactor.

In one aspect, the present invention is a method for catalytic dehydrogenation of alkyl aromatic compounds under a dehydrogenation catalyst in a single-shell fluidized bed reactor to form vinyl aromatic compounds, and in-situ regeneration of the catalyst. The process of the present invention comprises (a) a fluidized dehydrogenation catalyst in a single-shell fluidized bed reactor comprising a reaction zone and a regeneration zone, under fluidized conditions, the catalyst can be circulated within the zone and between the two zones, The second step is carried out simultaneously with the first step, the method comprising (b) under reaction conditions sufficient to prepare the corresponding vinyl aromatic compound, the vinyl aromatic compound, optionally, the vapor stream, and the desorbed gas residing in the reaction zone hydrogen catalyst contact. A third step, preferably performed simultaneously with the first and second steps, comprises (c) contacting steam with a dehydrogenation catalyst residing in a regeneration zone under conditions sufficient to regenerate, at least partially regenerate, the catalyst ongoing.

In the fluidized bed process of the present invention, at any one time a portion of the catalyst is circulated in the reaction zone, substantially the remainder of the catalyst is circulated in the regeneration zone, and some is co-mixed at the edge of the two zones. Over time, the catalyst residing in the reaction zone will become deactivated and become partially or totally inactive. The failure is mainly caused by the combined coke on the surface of the catalyst. (However, the invention should not be bound or limited by such failure theory.) In the fluidized state, the catalyst in the reaction zone, including the spent catalyst, will circulate into the regeneration zone. Spent catalyst residing in the regeneration zone will be reactivated by exposure to steam. The result of the reactivation reaction is a partial or substantially complete restoration of catalyst activity compared to fresh (unused or "as-synthesized") catalyst. Thereafter, in the fluidized state, the catalyst regenerated in the regeneration zone will be circulated to the reaction zone, and the reaction-regeneration cycle will be repeated again and again. The preceding description explains the reaction-regeneration cycle and defines the phrase "regenerate, at least partially regenerate, the catalyst".

It should be understood that after repeated reaction and regeneration, the catalyst will eventually not be regenerated to practical levels of activity, even with the regeneration process described herein. When this occurs, spent catalyst can simply be exported out of the reactor and replacement fresh catalyst introduced into the reactor, preferably simultaneously. In the reactor of the present invention, the replacement of the catalyst can be carried out "on-line" without shutting down the catalytic process, in particular the dehydrogenation process in the present invention. Catalyst can be fed into or out of the reactor via gas flow or pneumatic transfer loops. Alternatively, the catalyst may be removed from the reactor by a gravity driven discharge at the bottom of the reactor and added to the reactor from a standpipe inlet at the top of the reactor.

In a preferred aspect of the invention, the alkylaromatic compound is ethylbenzene or a substituted derivative of ethylbenzene and the resulting vinylaromatic compound is styrene or a substituted derivative of styrene.

Provided that the desired product is a vinyl aromatic compound, any alkyl aromatic compound can be subjected to the dehydrogenation process of the present invention. The aromatic portion of the vinyl aromatic compound may include, for example, a monocyclic aromatic ring, such as phenyl; a fused aromatic ring, such as naphthalene; or a bicyclic ring, such as biphenyl. Preferred aromatic moieties are monocyclic aromatic rings, more preferably phenyl. Provided that the process of the present invention can be used to dehydrogenate the alkylaromatic compound to form an alkenyl moiety, the alkyl moiety of the alkylaromatic compound may comprise, for example, any saturated straight chain, branched chain, or cyclic hydrocarbon group. Non-limiting examples of suitable alkyl moieties include ethyl, n-propyl, isopropyl, n-butyl, isobutyl, t-butyl and higher analogs thereof. Preferably, the alkyl moiety is C ₂ -C ₁₀ alkyl, more preferably C ₂ -C ₅ alkyl, most preferably ethyl. The alkylaromatic compound can be optionally substituted with two or more alkyl moieties, or substituted with other types of substituents which are substantially inactive to the dehydrogenation substitution process of the present invention. Alkylaromatic compounds suitable for use in the process of the present invention include, but are not limited to, ethylbenzene, diethylbenzene, ethyltoluene, ethylxylene, cumene, tert-butylethylbenzene, ethylnaphthalene, ethylbenzene Benzene, and higher corresponding alkylates, wherein the preferred alkyl aromatic compounds are C ₈ -C ₂₀ alkyl aromatic compounds, more preferably C ₈ -C ₁₅ alkyl aromatic compounds, most preferably B Substituted derivatives of benzene or ethylbenzene.

In the process of the present invention, regeneration of the feedstream generally involves steam. Optionally, steam can be combined with the dehydrogenation feed stream. Alternatively, any steam to alkyl aromatic weight ratio (gas to oil ratio) is suitable in the process of the invention if the process produces vinyl aromatic compounds. The gas to oil ratio is known to be based on the total weight of all sources of vapor streams entering the reactor, including dehydrogenation vapor streams and regeneration feedstock streams. Generally, the gas-to-oil weight ratio exceeds about 0.2/1, preferably exceeds about 0.5/1, generally, the gas-to-oil weight ratio is less than about 5.0/1, preferably less than about 3.0/1, more preferably Less than about 1.2/1, most preferably less than about 1.0/1. Generally, compared with the prior art, the method of the present invention can be operated under the ratio of lower gas volume and oil volume, which advantageously reduces the energy consumption and expense of converting water into steam, and reduces circulation to the reactor. amount of water.

Alternatively, the method of the invention uses a purge gas. The purge gas is introduced directly into the dilute phase zone of the reactor and its primary function is to remove vapor products from the dilute phase zone where unwanted thermal reactions would occur. Any gas that is substantially inert to the dehydrogenation and regeneration process may be suitable as a purge gas, including, for example, nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof. The purge gas concentration in the dilute phase zone can be any concentration provided the overall process produces the desired vinyl aromatic compound. Generally, the concentration of the purge gas is varied based on, for example, the particular alkylaromatic compound and the particular process conditions employed, especially temperature and gas flow rate. Generally, the concentration of purge gas in the dilute phase zone is greater than about 10 volume percent, preferably greater than about 20 volume percent. Generally, the concentration of purge gas in the dilute phase zone is less than about 90 volume percent, preferably less than about 70 volume percent.

Optionally, the dehydrogenation and/or regeneration feedstream may include a diluent. Diluents mainly dilute reactants and products to improve selectivity or for safety considerations. Any gas that is substantially inert to the dehydrogenation and regeneration steps is suitable as a diluent, including, for example, nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof. The diluent concentration in the dehydrogenation or regeneration feed stream can be any concentration as long as the overall process produces the desired vinyl aromatic compound. Generally, diluent concentration changes are based on, for example, specific selected diluents, specific alkylaromatic compounds, specific dehydrogenation or regeneration process conditions, and specific catalyst and failure characteristics. Generally, the diluent concentration in the dehydrogenation or regeneration feed stream is greater than about 10 volume percent, preferably greater than about 20 volume percent. Generally, the concentration of diluent in any gas stream is less than about 90 volume percent, preferably less than about 70 volume percent. When steam is used as the diluent, as described above, the gas to oil weight ratio determines the concentration of the steam in the dehydrogenation feed stream.

The process of the present invention does not require oxygen. Preferably, the process of the invention does not use oxygen.

Any dehydrogenation catalyst capable of catalyzing the dehydrogenation of alkyl aromatic compounds to vinyl aromatic compounds can be used in the process of the present invention. There are countless examples of dehydrogenation catalysts that can be adapted for use, including catalysts described in the following U.S. Patents: U.S. Patent 4,404,123, U.S. Patent 4,503,163, U.S. Patent 4,684,619, U.S. Patent 5,171,914, U.S. Patent 5,376,613, U.S. Patent 5,510,552, U.S. Patent 5679, 878, these patents pertain to a number of iron oxide catalysts including, for example, one or more compounds of alkali metals, preferably sodium, potassium, and cesium; alkaline earth metals, preferably calcium; and/or cerium, chromium, zinc, Copper, and/or gallium compounds, and the catalysts described in US Patent No. 3,651,160, are chromium oxides and alkali metal oxides. Preferred catalysts are dehydrogenation catalysts comprising iron oxide. More preferred catalysts comprise (a) at least one iron oxide (b) at least one carbonate, bicarbonate, oxide or hydroxide of potassium and/or cesium, (c) an oxide of cesium, Carbonate, nitrate or hydroxide, (d) optionally, a sodium hydroxide, carbonate, bicarbonate, acetate, oxalate, nitrate or sulfate, (e ) optionally, a calcium carbonate, sulfate, hydroxide, and (f) optionally, one or more binders, eg, hydraulic binders. As a further option, more preferred catalysts may additionally include one or more oxides selected from zinc, chromium, copper. In general, more preferred catalysts include 25-60 weight percent iron, 13-48 weight percent potassium, and 1-20 weight percent cesium, calculated as oxides. These ratios and other suitable catalyst component ratios are described in the aforementioned US patents.

The dehydrogenation catalyst used in the fluidized bed reactor of the present invention can be particles of any size or shape, as long as the catalyst can catalyze the dehydrogenation of alkyl aromatic compounds to form vinyl aromatic compounds. Generally, the average catalyst particle size is greater than about 20 [mu]m in diameter (or cross-sectional dimension). A diameter greater than about 50 μm is preferred. Generally, the average catalyst particle size is less than about 1000 μm, preferably less than about 200 μm. Preferred catalyst particles are non-angular and rounded, completely non-sticky, and resistant to friction sufficiently for use in fluidized bed reactors. Those skilled in the art know whether a catalyst has sufficient frictional resistance to be used in a fluidized bed reactor.

If desired, the dehydrogenation feed stream can be preheated before entering the reaction zone. Preheating may conveniently be performed using condensed high pressure saturated steam, or alternatively, by burning fuel or process off-gas. Preheating to any temperature is possible as long as it is below the temperature at which thermal cracking of the alkylaromatic compound can be measured. Typically the preheat temperature is greater than about 150°C, preferably greater than about 250°C, more preferably greater than about 350°C. Typically the preheat temperature is less than about 600°C, preferably less than about 590°C. Likewise, the regeneration feed stream can also be preheated before entering the regeneration zone. Typically the preheat temperature of the regeneration stream is greater than about 200°C, preferably greater than about 300°C, more preferably greater than about 400°C. The preheat temperature of the regeneration stream is generally less than about 650°C, preferably less than about 630°C.

A preferred embodiment of a novel reactor used in the process of the present invention is the fluidized bed reactor shown in Figure 1, comprising a single-layer vertical shell, the inner space of which is functionally divided into a regeneration zone (1), a reaction zone (2 ), dilute phase region (3). The regeneration zone of this embodiment is at the bottom of the reactor and includes the area where the catalyst is regenerated. The reaction zone in this preferred embodiment is located in the central region of the reactor, where catalytic organic chemical processes take place, such as the dehydrogenation process described herein. The dilute phase zone, located at the top of the reactor, includes the space from the middle to the top of the inner wall of the reactor. The dilute phase zone, occupied by gaseous reactants and products, also provides space for the expansion of the fluidized bed. The gas phase thermal reaction takes place in the dilute phase region; however, the conditions of the present catalytic reaction are preferably maintained to minimize the gas phase reaction relative to the catalytic process taking place in the reaction zone.

Figure 1 The regeneration zone, containing the regeneration feed stream, here means the inlet means for introducing steam and optionally a diluent into the regeneration zone. The inlet means may comprise, for example, an inlet (4), the outlet of which is a plenum zone (10) above which is a distribution plate or spray array (9). The dilute phase zone comprises an inlet means (5) comprising, for example, an inlet and a transfer pipe which introduces a reaction feed stream, here a dehydrogenation feed stream, into the reaction zone. In Figures 1 and 2, the inlet means for the regeneration zone are shown at the bottom of the figure and the inlet means for the reaction zone are shown at the top of the figure. In fact, the inlet device to the regeneration zone can be at any location, provided it exits in the regeneration zone. Similarly, the inlet means to the reaction zone can be located anywhere as long as the reactants can enter the reaction zone.

Preferably, the inlet means (5) terminates in the distribution plate or injection array (6) of the reaction zone, positioned preferably above the vapor distributor (9) at the bottom. The distribution plate or spray array (6) is preferably designed to deliver reactant feed flow in any direction to the reaction zone. The distance from the bottom vapor distributor (9) to the reactant feed stream distributor (6) can be varied as required to provide regeneration and reaction zones of different volumes. The larger the zone, the longer the residence time for gases and solids to reside in that zone. As described herein, the distributor means (6) provides a functional demarcation between the regeneration zone and the reaction zone, such that when the catalytic organic process is substantially carried out in the reaction zone, the regeneration process is substantially occurring in the regeneration zone, while Back mixing of solids and gases is still allowed. The gas distributor and jet array may be made, for example, of gas permeable sintered metal, or more preferably, the gas distributor and jet array are adapted to disperse the gas with the jet. The dilute phase zone also includes outlet means (7), e.g., an outlet, for venting the gas stream, including unconverted alkylaromatic compounds, steam, optional purge gas and/or diluent, and products, including vinyl Aromatic compounds. The outlet device (7) may be connected with a cyclone dust collector (shown in FIG. 1, at the lower end of the outlet 7), and the cyclone dust collector is used to collect catalyst particles entrained by the exhaust airflow. The collected catalyst particles can be circulated in the fluidized bed reactor through the inlet device (8). The inlet means can be located at any point along the reactor, but preferably, see Figure 1, it is located in the regeneration zone. The outlet device is further connected to a separation unit (not shown in FIG. 1 ), including, for example, a concentration device and a still for separating unconverted alkyl aromatic compounds and reaction products. Unconverted reactants can be recycled to the reaction zone through inlet (5). In addition to the above, the reactor further includes means for measuring the temperature of the catalyst bed and, optionally, a means for heating the reactor (not shown in the figure). The reaction and regeneration zones may also include diverters (not shown) whose function is to reduce bubble formation and bubble size and facilitate contact between the gaseous feed stream and the catalyst.

In another preferred embodiment of the present invention, the fluidized bed reactor further includes one or more means for increasing solids circulation and heat transfer. In a preferred embodiment, the means for increasing solids circulation includes one or more draw tubes, optionally including internal diverters. Alternatively, the means for increasing solids circulation includes one or more traction tubes that heat or cool the constituent elements. One embodiment of the present invention comprises a plurality of traction tubes, shown in Figure 2 (section 1-10 of Figure 2, same as section 1-10 of Figure 1). Each drag tube (11) comprises, for example, a concentric cylinder open at both ends, or a bundle or array of heated tubes, or any other design that facilitates catalyst drag. Typically, the draw tube is suspended vertically through the reaction zone and the regeneration zone, preferably near the top of the reaction zone. The dehydrogenation feedstock stream enters the jet array tube (6) through the inlet hole (5) and goes up to the inner cylinder of the draw tube (11) to enter the reaction zone. As a result of the fluidization conditions, catalyst particles will be entrained into the inner cylinder of the draw tube and transported to the top end of the draw tube. At the top, the catalyst particles will flow to the edge of the inner cylinder and descend through the annulus between the two cylinders back to the regeneration zone.

In addition to the above, the reactor may optionally include an inlet means and an outlet means (not shown) which convey catalyst into and out of the reactor, respectively.

In another embodiment of the fluidized bed, the reaction zone and regeneration zone can be reversed so that the reaction zone can be positioned at the bottom of the reactor and the regeneration zone in the middle of the reactor. (Fig. 1, wherein the reaction zone is positioned at (1), the regeneration zone is positioned at (2), and the inlet is adjusted accordingly).

The reactor of the present invention may be used in catalytic processes where the reactant feed stream is chemically compatible with the regeneration feed stream. The unique reactor of the present invention allows for a continuous flow of catalyst particles between the reaction zone and the regeneration zone, which are contained within a single shell of the fluidized bed reactor. Thus, the desired catalytic organic process and catalytic regeneration can be achieved simultaneously without the need to export the catalyst out of the reactor and into a separate regenerator. The reactor of the present invention does not contain complex concentric walls and convoluted labyrinths for the transport of catalyst particles. Thus, for large-scale reactions, the reactor of the present invention does not present significant problems of stagnant flow and friction.

The temperature in the reaction zone, in the dehydrogenation process, can be any operable temperature if vinyl aromatic compounds can be produced in the process. Operable dehydrogenation temperatures vary with the particular catalyst and the particular alkylaromatic compound employed. For the preferred catalysts comprising iron oxide, the dehydrogenation temperature is generally greater than about 550°C, preferably greater than about 570°C. Generally, the dehydrogenation temperature is less than about 650°C, preferably less than about 610°C. Below about 550°C, the conversion of the alkylaromatics would be too low, while above 650°C, thermal cracking of the alkylaromatics and vinylaromatic products would occur. In the present invention, the temperature is measured on the catalyst bed in fluidized form.

In the regeneration zone, the catalyst is contacted with steam and reactivated. The temperature of the regeneration zone can also be varied as long as the catalyst can be at least partially regenerated. Generally, the regeneration temperature is lower than the thermal cracking temperature of the alkylaromatic compounds and vinylaromatic products. For the preferred iron oxide containing catalysts, the regeneration temperature is generally greater than about 550°C, preferably greater than about 570°C. Generally, the regeneration temperature is less than about 650°C, preferably less than about 610°C. All fluidized beds are essentially two-zone isothermal due to the continuous recirculation of the catalyst in the reaction and regeneration zones and the temperature of the two zones being maintained at similar values.

The process can be carried out at any operable total pressure, between subatmospheric and superatmospheric, so long as the vinyl aromatic product can be produced. If the total reactor pressure is too high, the equilibrium point of the dehydrogenation process can return to the alkylaromatics. On the other hand, sufficient vapor pressure is required to delay coking of the catalyst. Preferably, the process is performed under vacuum to maximize the yield of vinyl aromatic product. At the aforementioned gas-to-oil ratios, the vacuum pressure is sufficient to regenerate the catalyst, at least partially. Preferably, the total pressure in the reactor is greater than 1 psig (6.9 kPa). More preferably, the total air pressure is greater than about 3 psig (20.7 kPa). Preferably, the total air pressure is less than about 73 psig (503.3 kPa). More preferably, the total air pressure is less than about 44 psig (303.4 kPa). Most preferably the total gas pressure is sub-atmospheric, between about 3 psig (20.7 kPa) and 13 psig (90.6 kPa). The pressure through the dilute phase, reaction, and regeneration zones varies according to process factors such as catalyst weight and buoyancy and frictional influences. Generally, the pressure at the bottom of the reactor is slightly higher than at the top.

The space velocity of the dehydrogenation feed stream depends on the particular alkylaromatic compound and catalyst used, the particular vinylaromatic compound formed, the size (eg, diameter and height) of the reaction zone, and the shape and weight of the catalyst particles. Rapid removal of reactants and products from the dilute phase region is required to reduce thermal cracking and other unwanted side reactions. Additionally, the gas flow is sufficient to cause fluidization of the catalyst bed. Generally, the space velocity of the dehydrogenation feed stream is varied from the minimum velocity required to achieve fluidization of the catalyst particles to a somewhat lower velocity required to achieve the minimum velocity required to achieve pneumatic transport of the catalyst particles. Fluidization occurs when the catalyst particles separate, the particles move in a fluid-like manner, and the bed pressure decreases at a uniform rate along the bed to be substantially constant. Pneumatic transport occurs when sufficient catalyst particles are entrained into the gas stream and out of the reactor. Preferably, the space velocity of the dehydrogenation feed stream varies from a minimum foaming rate to a minimum turbulence rate. Bubbling occurs when gas bubbles can be seen in the fluidized bed, but there is little backmixing of gas and solids, and turbulence occurs when there is enough foaming and enough backmixing of gas and solids. More preferably, the flow rate is high enough to cause back mixing.

As a general guideline, under operating conditions, the gaseous hourly space velocity (GHSV) is greater than about 60 ml total loading/ml catalyst/hour, as measured by the total flow rate of the dehydrogenation feed stream, which includes alkylaromatic compounds And, optionally, gas flow, purge gas, and/or dilution gas flow. Preferably, the GHSV of the dehydrogenation flow rate is greater than about 120 h ^-1 , more preferably greater than 720 h ^-1 . Generally, the GHSV of the dehydrogenation gas stream is less than about 12000h ^-1 , preferably less than about 3600h ^-1 , more preferably less than about 1800h ^-1 , which measures the total flow at operating process conditions.

As a general guideline, the gas residence time in the reaction zone is greater than about 0.3 seconds (sec) under operating conditions. The gas residence time can be calculated by multiplying the reaction zone height by the reaction zone void fraction divided by the surface gas of the regeneration and reaction feed streams. rate to measure. The "reaction zone void portion" is the portion of the reaction zone that is vacant. The surface gas velocity is the velocity of the gas as it passes through an empty reactor. Preferably, the gas residence time in the reaction zone is greater than about 1 second, more preferably greater than about 2 seconds, measured under operating conditions. Generally, the gas residence time in the reaction zone is less than about 60 seconds, preferably less than about 30 seconds, more preferably less than about 5 seconds, measured under operating conditions.

The hourly gas space velocity of the regeneration feed stream through the regeneration zone can vary widely if the catalyst is regenerated, at least in part, if the catalyst particles in the regeneration zone can be effectively fluidized. The space velocity of the regeneration feed stream can be lower from the minimum velocity at which fluidization of the catalyst particles can be achieved to the minimum velocity at which catalyst gas pressure can be conveyed. Preferably, the space velocity of the regeneration feed stream varies from a minimum foaming rate to a minimum turbulence rate. Typically, the gaseous hourly space velocity (GHSV) is greater than 60 ml total loading/ml catalyst/hour under operating conditions, as measured by the total flow rate of the regeneration feed stream. Preferably, the GHSV of the regeneration gas stream is greater than about 120 h ^-1 , more preferably greater than about 360 h ^-1 . Generally, the GHSV of the regeneration gas stream is less than about 12000h ^-1 , preferably less than about 3600h ^-1 , more preferably less than about 720h ^-1 , measured as a total flow at operating process conditions.

In the regeneration zone, the gas residence time in the regeneration zone measured under operating conditions is greater than about 0.3 seconds (sec). The gas residence time can be calculated by multiplying the height of the regeneration zone by the regeneration zone void fraction divided by the total regeneration and reaction feed flow The velocity of the surface gas is calculated. The blanket void portion is the empty blanket section. The preferred gas residence time in the regeneration zone is greater than about 1 second, more preferably greater than about 5 seconds. Generally, the gas residence time in the regeneration zone is less than about 60 seconds, preferably less than about 30 seconds, more preferably less than about 10 seconds, measured under operating conditions.

Vinyl aromatic compounds are produced when alkyl aromatic compounds and optionally vapors are contacted with a dehydrogenation catalyst in the manner previously described. Ethylbenzene is mainly converted to styrene. Ethyltoluene is converted to p-methylstyrene (p-vinyltoluene). tert-butylethylbenzene is converted to tert-butylstyrene. Cumene (cumene) is converted to α-methylstyrene and diethylbenzene is converted to divinylbenzene. Hydrogen gas is formed simultaneously during dehydrogenation. Another product is a minor product including benzene and toluene.

The conversion of alkylaromatic compounds in the method of the present invention varies according to the specific feed composition, catalyst composition, process conditions and fluidized bed conditions. For the purposes of this invention, "conversion" is defined as the mole percent of alkylaromatic compound converted to all products. In the present process, the conversion of the alkylaromatic compound is generally greater than about 30 mole percent, preferably greater than about 50 mole percent, and more preferably greater than about 70 mole percent.

Likewise, product selectivity will vary depending on the particular feed composition, catalyst composition, process conditions and fluid bed conditions. For the purposes of the present invention, "selectivity" is defined as the mole percent of converted alkyl aromatic compounds that form a specific product, preferably vinyl aromatic compounds. In the inventive process it is meant that the selectivity to vinyl aromatic compounds, preferably styrene or substituted derivatives of styrene is generally higher than about 60 mole percent, preferably more than about 75 mole percent, more preferably Preferably, the molar percentage exceeds about 90%.

The invention is further illustrated by the following examples, which are purely illustrative of the invention. Other embodiments of the invention will be apparent to those skilled in the art from description or practice of the invention disclosed herein. Measures of selectivity were corrected for deviations from 100% of organic material equilibrium.

Example 1

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 consisted of a single vertical shell divided into three zones by function: Catalyst regeneration zone (1) at the bottom of the reactor; dilute phase zone (3) at the top of the reactor; and reaction zone (2) in the middle part of the regeneration and dilute phase zone. The first inlet (4 ) is located at the bottom of the reactor where it leads to the plenum area (10), where a gas distributor is installed. The first inlet is used to distribute the regeneration raw material stream to the regeneration zone. The second inlet (5 ), located in the dilute phase zone, used to deliver the dehydrogenation feed stream to the reaction zone (2). The second inlet is connected to an inlet pipe, which terminates in a row of jet array tubes (6), and the jet array tubes are arranged 3 inches (7.5 cm) above the distribution plate (9) at the bottom of the reaction zone. The spray array tube is composed of six rows of sintered metal tubing (Inconel, 1/4 inch OD (6.3 mm OD)), designed to Provide a constant pressure drop throughout the jet array. Place the jet outlet (7) of the jet array horizontally. The outlet (7) is located in the dilute phase zone for removal of the product stream. Entrained solids in the product stream are cleaned with a cyclone The dust collector (located below the outlet 7) collects and circulates into the reactor through the third inlet (8) located in the regeneration zone. The outflow gas is collected by the downstream cyclone collector. The reactor is equipped with an impedance (resistance) Apparatus were added to heat the reactor and two internal thermocouples (Type K) were used to measure the fluidized bed temperature in the reaction and regeneration zones.

In the presence of a dehydrogenation catalyst, the reactor is used for a dehydrogenation reaction, converting ethylbenzene to styrene, while simultaneously and continuously regenerating the dehydrogenation catalyst. A dehydrogenation catalyst (2370 grams) having an average particle diameter of 300 microns and comprising 28.7% by weight of iron oxide (Fe ₂ O ₃ ), 14.3% of cerium oxide (Ce ₂ O ₃ ), 7.6% of copper oxide ( CuO), 31.6% potassium carbonate (K ₂ CO ₃ ), 0.6% chromium oxide (Cr ₂ O ₃ ), 9.5% zinc oxide (ZnO), and 7.6% binder, were loaded into the reactor. The reaction feed stream includes a mixture of ethylbenzene and vapour. The regeneration feed stream includes steam, gaseous products with five columns installed in parallel (2.7% Carbowax 1540 on Porasil C; 3% Carbowax 1540 on Porasil C; 27% Bis(EE)A on Chromosorb R PAW; Porapak Q; and two 13X molecular sieve chromatography column) Carle gas chromatograph analysis. The liquid product was analyzed with a HP 5890 gas chromatograph equipped with a J&W DB-5 column. Nitrogen was used as internal standard for gas analysis; heptane was used as internal standard for liquid analysis. Sampling was taken over a period of six hours, including four or more samples taken every 30 minutes during the last hours of operation. The ethylbenzene conversion and styrene selectivity results shown here are the average of four or more samples.

In the reactor described above, water with an injection rate of 4.3 cubic centimeters per minute at room temperature is heated to 600°C, fed into the plenum zone (10) through the inlet (4), and passed through the distribution plate (9 ) into the regeneration zone (1) at the bottom of the reactor. At room temperature, liquid ethylbenzene is mixed with nitrogen at a flow rate of 2.5 cubic centimeters per minute and an inflow rate of 1088 cubic centimeters per minute, heated to 500 ° C, and fed into the dehydrogenation reaction zone through the inlet (5) and the spray arrangement tube (6) . The outflow rate corresponds to the following indicators: total gas to oil ratio 2/1, superficial velocity 1.86 m/min in the regeneration zone and 237 m/min in the reaction zone. The gas residence time in the regeneration zone was 1.46 seconds, the gas residence time in the reaction zone was 0.67 seconds, and the temperature and pressure of the reactor were maintained at 600°C and 15.5 psig (106.9 kPa), respectively. The product obtained via outlet (7) was analyzed with the method mentioned above. Ethylbenzene was converted to 74.0 mole percent. The selectivity to styrene was 86.0 mole percent. Additional products include benzene and toluene. The material balance is up to 95% by weight of the organic material fed.

Example 2

Using the reactor and catalyst of Example 1, water with a feed flow rate of 2.17 cc/min was heated to 600°C and fed to the distribution plate in the regeneration zone. At room temperature, liquid ethylbenzene was heated to 500°C at an inflow rate of 2.52 cc/min and liquid water at an inflow rate of 2.17 cc/min and fed through the reaction zone to the sparger array. The outflow rate corresponds to the following indicators: total steam to oil ratio of 2/1, 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, and the temperature and pressure of the reactor were maintained at 600°C and 15.5 psig (106.9 kPa), respectively. Ethylbenzene was converted to 85 mole percent. The selectivity to styrene was 69 mole percent. The material balance was 96% by weight of the organic material fed.

In Example 1, nitrogen was added to the ethylbenzene stream as a purge gas, but no vapor was added to the ethylbenzene stream. In Example 2, by contrast, no purge gas was added to the ethylbenzene stream, and the vapor stream was divided into feeds for dehydrogenation and feeds for regeneration. When Example 2 is compared with Example 1, it can be seen that the conversion rate of ethylbenzene is higher in Example 2, because there is a longer residence time in the bed, and the selectivity of styrene is lower, because is the increased free radical cleavage in the dilute phase region.

Example 3

Using the reactor and catalyst of Example 1, water with a feed flow rate of 4.3 cc/min was heated to 600°C and fed to the distribution plate in the regeneration zone. At room temperature, liquid ethylbenzene was heated to 500°C at an inflow rate of 2.49 cc/min and fed through the reaction zone to the spray train. The outflow rate corresponds to the following indicators: a total gas to oil ratio of 2/1, a superficial velocity of 309 m/min in the regeneration zone and 417 m/min in the reaction zone. The gas residence time in the regeneration zone was 1.47 seconds, the gas residence time in the reaction zone was 0.63 seconds, and the temperature and pressure of the reactor were maintained at 600°C and 15.5 psig (106.9 kPa), respectively. Ethylbenzene was converted to 85 mole percent. The selectivity to styrene was 72 mole percent. The material balance is up to 98% by weight of the incoming organic material.

Except the difference below, the process condition of embodiment 3 is very similar to embodiment 2, and in embodiment 2, 1/2 total steam is sent to regeneration zone, and other 1/2 total steam is sent to reaction zone. In contrast, in Example 3 all vapors are sent to the regeneration zone. When embodiment 3 is compared with embodiment 2, it can be seen that the conversion rate of ethylbenzene is equivalent to the selectivity of styrene, and there are slight differences according to the position of introducing steam.

Example 4

Example 4 closely replicates the process conditions of Example 2, except that the pressure is kept constant at 5 psi (34.5 kPa). The catalyst used in Example 4 has the same chemical composition as other previous examples; but the amount of catalyst used is 1355 grams, and the average particle diameter of the catalyst is 220 microns. The process conditions are as follows: The rate at which water is fed into the regeneration zone It is 2.9 cubic centimeters per minute, and the rates of liquid ethylbenzene and water fed into the reaction zone are 2.52 cubic centimeters per minute and 1.45 cubic centimeters per minute; the ratio of gas volume and oil volume is 2/1; the superficial velocity in the regeneration zone is 123 cm /min, 200 cm/min in the reaction zone; and a temperature of 600°C. The conversion of ethylbenzene was 49 mole percent. The selectivity to styrene was 88 mole percent. The material balance was 93% by weight of the organic material fed.

Comparing Example 2 with Example 4 shows that significantly high styrene selectivities can be obtained by operating the fluidized bed reactor under vacuum conditions. The local lower pressure at the feed of ethylbenzene somewhat lowers the overall conversion.

Example 5

Example 5 repeats the process conditions of Example 4 very similarly, except that the temperature of the reactor is kept constant at 590°C instead of 600°C. The process conditions are as follows: the rate of water supply into the regeneration zone is 2.9 cubic centimeters per minute; the supply rate of liquid ethylbenzene and water into the reaction zone is 2.52 cubic centimeters per minute and 1.45 cubic centimeters per minute; gas volume and oil volume The ratio was 2/1; the superficial velocity was 122 cm/min in the regeneration zone and 199 cm/min in the reaction zone; and the pressure was 5 psi (34.5 kPa). The conversion of ethylbenzene was 50 mole percent. The selectivity to styrene was 94 mole percent. The material balance accounts for 99% by weight of the incoming organic material. Comparing Example 4 with Example 5 shows a greater increase in selectivity to styrene by operating under vacuum at temperatures below 600°C.

Example 6

Example 6 repeats the process conditions of Example 4 very similarly, except that the temperature of the reactor is kept constant at 580°C instead of 600°C. The process conditions are as follows: the rate of water supply into the regeneration zone is 2.83 cubic centimeters per minute; the supply rate of liquid ethylbenzene and water into the reaction zone is 2.52 cubic centimeters per minute and 1.45 cubic centimeters per minute; gas volume and oil volume The ratio was 2/1; the superficial velocity was 121 cm/min in the regeneration zone and 197 cm/min in the reaction zone; and the pressure was 5 psi (34.5 kPa). The conversion of ethylbenzene was 44 mole percent. The selectivity to styrene was 95 mole percent. The mass balance represents 100% by weight of the organic material fed. Comparing Examples 4, 5 and Example 6 shows a greater increase in selectivity to styrene by operating under vacuum at temperatures below 600°C.

Example 7

Example 7 repeats the process conditions of Example 4 very similarly, except that the gas to oil ratio is 1/1 instead of 2/1. Other technological conditions are as follows: the rate that water is fed into regeneration zone is 1.45 cubic centimeters/minute; The feed rate that liquid ethylbenzene and water flow into reaction zone is respectively 2.52 cubic centimeters/minute and 0.73 cubic centimeters/minute; The surface velocity is 61.5 cm/min and the superficial velocity in the reaction zone is 107.6 cm/min; and the pressure is 5 psi (34.5 kPa) and the temperature is 600°C. The conversion of ethylbenzene was 49 mole percent. The selectivity to styrene was 89 mole percent. The material balance is up to 98% by weight of the incoming organic material.

Comparing Examples 4, 5 and Example 7, it is shown that the reduction of the gas and oil ratio from 2/1 to 1/1 does not affect the conversion of ethylbenzene and the selectivity of styrene.

Example 8

Example 8 repeats the process conditions of Example 4 very similarly except for the particle size of the catalyst and the gas to oil ratio. For Example 8, the average particle diameter of the catalyst (1570 grams) was 82 microns, the ratio of gas volume to oil volume was 0.5/1, and other process conditions were as follows: the rate of water feed into the regeneration zone was 0.8 cubic centimeters per minute; Ethylbenzene and water flow into the feed rate of reaction zone and are respectively 5.48 cubic centimeters per minute and 0.54 cubic centimeters per minute; The superficial flow velocity in regeneration zone is 33.52 centimeters per minute, and the superficial flow velocity in reaction zone is 74.8 centimeters per minute; And the pressure is 5 psi (34.5 kPa) and the temperature is 600°C. The conversion of ethylbenzene was 54 mole percent. The selectivity to styrene was 95 mole percent. The mass balance represents 100% by weight of the organic material fed.

Comparing Examples 4, 7 and Example 8, it shows that the ratio of gas volume and oil volume to 0.5/1 can obtain high styrene selectivity. Additionally, reducing the particle diameter of the catalyst from 220 microns to 82 microns increased the conversion of ethylbenzene. This result is likely attributed to the fact that smaller catalyst particles tend to produce smaller bubble diameters at equilibrium with better mass transport in fluidized bed reactors.

Example 9

A pulse reactor was used to study the effect of time on the dehydrogenation of ethylbenzene to styrene. In the pulse reactor, the dehydrogenation step followed by the catalyst regeneration step was cycled continuously. The test results obtained with the pulsed reactor showed the expected results for the fluidized bed reactor.

A particle size between 1.18mm and 1.70mm and comprising 33.2% by weight iron oxide (Fe ₂ 0 ₃ ), 17.5% cerium oxide (Ce ₂ O ₃ ), 7.8% copper oxide (CuO) , 36% potassium carbonate (K ₂ CO ₃ ), 0.6% chromium oxide (Cr ₂ O ₃ ), and 4.7% binder, the dehydrogenation catalyst was loaded in a continuous flow, fixed-bed reactor [304 stainless steel , Chart 40, 1" (2.5 cm) OD x 36" (90 cm) L]. The catalyst bed occupies 7 inches (17.5 cm) of the reactor length. The space above the bed was filled with ceramic bell saddle filler (1/4 inch, 0.6 cm). A metal partition was installed under the bed. The temperature of the reaction was measured by thermocouples mounted on the catalyst bed. The dehydrogenation pulse was accomplished by feeding ethylbenzene over the catalyst preheated to 550°C for 2 minutes. The outflow rate of liquid ethylbenzene was 1.16 ml/min, measured at ambient temperature and pressure (23°C and 1 atmosphere). At the same time, water preheated to 550 °C covered the catalyst in the same 2 minutes. The water feed rate was adjusted to maintain a gas to oil weight ratio of 0.30/1. The total pressure was maintained at 5.0 psig. Thereafter the feed flow of ethylbenzene was stopped and the regeneration pulse was measured by feeding only a stream of water preheated to 550°C under the same process conditions, covering the catalyst alone for 2 minutes. The inflow rate of liquid water during the regeneration pulse was 1 ml/min, measured at 24°C and 1 atmosphere. After regeneration, the dehydrogenation pulse was repeated by introducing the ethylbenzene feed stream for an additional 2 min with continuous water flow as previously described. After 2 minutes, the ethylbenzene feed was stopped again, and the vapor flow continued to feed the regeneration cycle for 2 minutes, and the dehydrogenation-regeneration cycle was repeated for a total time of 200 hours. The product stream is passed continuously through the condenser, separated, and analyzed by conventional methods.

The results of the pulsed process are shown in Figure 3, which plots ethylbenzene conversion and styrene selectivity versus time at constant pressure (5.0 psig) and temperature (550°C). A surprising finding was that ethylbenzene conversion increased slightly with time. The selectivity to styrene, however, remained constant throughout the process with values greater than 95 mole percent. The pulse reactor results indicated that the dehydrogenation catalyst could be cycled through the dehydrogenation-regeneration steps of the fluidized bed reactor for a long period of time without appreciable deactivation.

Comparative Example 1

The procedure of Example 9 was repeated with the same continuous flow and fixed bed reactor under the same reaction conditions, except that the dehydrogenation was performed in a continuous manner instead of a pulsed manner. Therefore, there is only one dehydrogenation cycle and no catalyst regeneration cycle. Under these conditions, a gradual deactivation of the catalyst was found to be accompanied by a decrease in ethylbenzene conversion. When the catalyst is deactivated, the temperature of the reaction process is increased to keep the conversion of ethylbenzene constant. At a gas to oil ratio of 0.3/1, the temperature had to be increased at a rate of 0.45°C per minute to keep the ethylbenzene conversion constant. When comparing Comparative Example 1 and Example 9, it was found that in the pulse reactor, the service life of the catalyst can be significantly extended at normal temperature and pressure, and at the same time, if the regeneration reaction is not carried out, the catalyst will be deactivated very quickly, and An increase in temperature is required to maintain constant ethylbenzene conversion. The pulse reactor results indicated that the dehydrogenation catalyst could be cycled through the dehydrogenation-regeneration cycle steps in the fluidized bed reactor for a long period of time without significant deactivation.

Claims (47)

1. A method for dehydrogenating a _C2-10 alkyl aromatic compound to generate a vinyl aromatic compound through a dehydrogenation catalyst, and regenerating the dehydrogenation catalyst in situ, the method comprising (a) comprising a reaction in a fluidized state A dehydrogenation catalyst is fluidized in a single-shell fluidized bed reactor in the zone and a regeneration zone, so that the catalyst circulates in the zone and between the two zones, (b) dehydrogenation containing _C2-10 alkyl aromatic compounds contacting a hydrogen feedstream with a dehydrogenation catalyst residing in the reaction zone under reaction conditions sufficient to produce the corresponding vinyl aromatic compound; and (c) under conditions sufficient to regenerate, at least partially regenerate, said catalyst will contain vapor The regeneration feed stream contacts the dehydrogenation catalyst residing in the regeneration zone. 2. The method according to claim 1, wherein the dehydrogenation feed stream in the step (b) may also contain steam. 3. The method according to claim 1, wherein said alkyl aromatic compound is a C ₈ -C ₂₀ alkyl aromatic compound. 4. The method according to claim 3, wherein said alkylaromatic compound is ethylbenzene, cumene, diethylbenzene or ethyltoluene. 5. The method according to claim 1, wherein said fluidized bed reactor further comprises a dilute phase zone, and the purge gas is fed into the dilute phase zone. 6. The method according to claim 1, wherein the weight percentage of total vapor to alkyl aromatic compound is greater than 0.2/1 and less than 5.0/1. 7. The method according to claim 1, wherein the weight percentage of the total vapor to the alkyl aromatic compound is greater than 0.2/1 and less than 1.2/1. 8. The method of claim 1, wherein said diluent gas is added with the dehydrogenation feed stream, or with the regeneration feed stream, or with both feed streams. 9. The method of claim 8, wherein said diluent gas is selected from the group consisting of nitrogen, argon, helium, carbon dioxide, steam, and mixtures thereof. 10. The method of claim 8, wherein said diluent comprises from greater than 10 volume percent to less than 90 volume percent of the dehydrogenated or regenerated feedstock stream, or both streams independently. 11. The method according to claim 1, wherein the dehydrogenation feed stream is preheated to a temperature greater than 150°C and less than 600°C. 12. The method according to claim 1, wherein the regeneration raw material stream is preheated to a temperature greater than 200°C and less than 650°C. 13. The method according to claim 1, wherein the temperature of said reaction zone and/or regeneration zone is greater than 550°C and less than 650°C. 14. The method according to claim 1, wherein the total pressure of the reactor is greater than 6.9Kpa and less than 503.3Kpa. 15. The method according to claim 1, wherein measured under operating conditions, calculated by the total flow rate of the dehydrogenation raw material stream, the method is carried out under the condition that the hourly space velocity of the gas is greater than 60h ^-1 and less than 12000h ^-1 . 16. The method of claim 1, wherein the method is performed with a residence time of the total gas flow in the reaction zone greater than 0.3 seconds and less than 60 seconds, measured under operating conditions. 17. The method according to claim 1, wherein measured under operating conditions, calculated by the total flow rate of the regeneration raw material flow, the method is carried out under the condition that the hourly space velocity of the gas is greater than 60h ^-1 and less than 12000h ^-1 . 18. The method of claim 1, wherein the method is performed with a total blanket gas flow residence time greater than 0.3 seconds and less than 60 seconds, measured under operating conditions. 19. The method of claim 1, wherein said dehydrogenation catalyst comprises iron oxide. 20. The method of claim 19, wherein said dehydrogenation catalyst further comprises at least one or more compounds selected from the group consisting of alkali metals, alkaline earth metals, chromium, gallium, cerium, zinc and copper. 21. The method of claim 19, wherein said dehydrogenation catalyst comprises (a) at least one iron oxide, (b) at least one potassium and/or cesium carbonate, bicarbonate, oxide or Hydroxide, (c) an oxide, carbonate, nitrate or hydroxide of cesium. 22. The method of claim 21, wherein said dehydrogenation catalyst further comprises (d) a sodium hydroxide, carbonate, bicarbonate, acetate, oxalate, nitrate or Sulfates. 23. The method of claim 22, wherein said dehydrogenation catalyst further comprises (e) a calcium carbonate, sulfate, or hydroxide. 24. The method of claim 23, wherein said dehydrogenation catalyst further comprises (f) one or more oxides of zinc, chromium and copper. 25. The method of claim 24, wherein said dehydrogenation catalyst further comprises (g) a binder. 26. The method of claim 1, wherein the conversion of said alkylaromatic compound is greater than 30 mole percent. 27. The method of claim 1, wherein said vinyl aromatic compound selectivity is greater than 60 mole percent. 28. The method of claim 1, wherein said vinyl aromatic compound is styrene, divinylbenzene, alpha-methylstyrene or vinyl toluene. 29. The method of claim 1, wherein the average particle size of said dehydrogenation catalyst is greater than 20 microns and less than 1000 microns. 30. The method according to claim 1, wherein said fluidized bed reactor comprises a single-layer vertical shell containing a dilute phase zone, a reaction zone and a regeneration zone; an inlet device for leading the regeneration feed stream into the regeneration zone; and An inlet device for introducing the reaction feed stream into the reaction zone, and one of the inlet devices into the reaction zone or the regeneration zone can separate the two zones while allowing catalyst particles to circulate between the two zones; the fluidized bed reactor further includes a Outlet device for discharging the liquid stream. 31. The method of claim 30, wherein said fluidized bed reactor further comprises an inlet means for returning catalyst entrained in the effluent stream to the reactor. 32. The method of claim 30, wherein said dehydrogenation catalyst further comprises an inlet and an outlet means for feeding the catalyst into and out of the reactor. 33. A method according to claim 30, wherein said means for separating the reaction zone from the regeneration zone comprises a spray array or distributor. 34. The method of claim 1 comprising (a) fluidizing a dehydrogenation catalyst in a fluidized state in a single-shell fluidized bed reactor comprising a reaction zone and a regeneration zone, such that Catalyst can be circulated within and between the two zones. (b) contacting ethylbenzene, cumene, diethylbenzene or ethyltoluene with a dehydrogenation catalyst residing in the reaction zone, the catalyst comprising iron oxide, and said contacting is at a weight ratio of steam to ethylbenzene Greater than 0.2/1 and less than 3.0/1, the temperature is greater than 570°C and less than 610°C, and the total reactor pressure is greater than 20.7 kPa and less than 303.4 kPa, (c) the dehydrogenation catalyst that will reside in the regeneration zone The catalyst is at least partially regenerated by contacting with a steam-containing regeneration feed stream at a temperature greater than 570°C and less than 610°C to regenerate. 35. The method according to claim 34, wherein step (b) may also contain steam. 36. The method according to claim 34, wherein step (b) may also contain diluent gas. 37. The method of claim 34, wherein said regeneration feed stream further comprises a diluent. 38 A fluidized bed reactor for organic processes that catalyze in situ catalyst regeneration comprising a single vertical shell containing a dilute phase zone, a reaction zone and a regeneration zone; An inlet device, an inlet device for introducing the reactant raw material flow into the reaction zone, one of the inlet devices can separate the reaction zone and the regeneration zone, while allowing catalyst particles to circulate between the two zones; fluidized bed reactor further comprising an outlet means for discharging an effluent stream; and further comprising at least one of said inlet means having means for increasing solids circulation, said means comprising a draw tube. 39. The method of claim 38, wherein said fluidized bed reactor further comprises an inlet means for returning catalyst entrained in the effluent stream to the reactor. 40. The method of claim 38, wherein said pull tube comprises an internal diverter. 41. A fluidized bed reactor according to claim 38, wherein the means for directing the reactant feed stream comprises a sparger array or distributor. 42. A fluidized bed reactor as claimed in claim 38, wherein the means for directing the flow of regeneration feed comprises a sparger array or distributor. 43. The fluidized bed reactor according to claim 38, wherein said means for separating the reaction zone from the regeneration zone is selected from a spray array or a distributor. 44. A fluidized bed reactor as claimed in claim 38, wherein said means for increasing solids circulation comprises draw tubes made of heating or cooling elements. 45. A fluidized bed reactor according to claim 44, further comprising inlet means and outlet means for introducing and removing catalyst into and out of the reactor. 46. The fluidized bed reactor of claim 44, further comprising at least one means for measuring the temperature of the fluidized bed. 47. The fluidized bed reactor of claim 46, further comprising a heating means.

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