KR890001849B1 - Process dehydrogenation of alkylaromatic hydrocarbon - Google Patents

Abstract

Alkylaromatic hydrocarbon dehydrogenation process contacting a reactant stream contg. the alkylaromatic hydrocarbon with a dehydrogenation catalyst at sub-atmospheric pressure to give a reactor effluent stream, from which prods. are recovered by; (A) cooling, without significant condensation, by indirect heat exchange against the reactant stream; (B) compresing to a pressure below 1 atm. by a 1st mechanical compressor; (C) partially condensing; and (D) sepg. the resultant effluent stream in sepn. zone, with pressure maintained 1 atm. absolute by 2nd mechanical compressor, into a vapour comprising H2 and a liq. from which the unsatd. hydrocarbon is recovered.

Description

Dehydrogenation Process of Alkyl Aromatic Hydrocarbons

Figure shows the dehydrogenation process flow diagram of ethylbenzene to produce styrene.

The present invention relates to a catalytic dehydrogenation process of hydrocarbons and especially of alkylaromatic hydrocarbons. The present invention therefore relates to the synthesis of aromatic compounds, such as steylene, by dehydrogenation taking place in many successive steps.

The present invention relates directly to an improved method of maintaining subatmospheric pressure in a dehydrogenation reactor.

Dehydrogenation of hydrocarbons is well known in the prior art, whereby the cyclic and aromatic hydrocarbons are converted into the corresponding saturated products. For example, dehydrogenation is carried out commercially to produce styrene from ethylbenzene to meet the demand of such polymer precursors. The produced styrene itself is polymerized or copolymerized with butadiene, isoprene, acrylonitrile and the like. Processes for dehydrogenation of alkylaromatic hydrocarbons are sometimes combined with alkylation processes to produce alkylaromatic hydrocarbons.

US Patent No. 3,515,766 to W.N. Root and US Patent No. 3,409,689 to D.J.Ward suitably show conventional methods of catalytic aromatic dehydrogenation of alkylaromatics including ethylbenzene. This reference describes the mixing of superheated steam with incoming hydrocarbons and the addition of additional superheated steam between successive layers of the dehydrogenation catalyst with the reactants. It also shows the overall process flow which may include the main process.

The use of subatmospheric pressure in the reaction zone of the dehydrogenation process of alkylaromatic hydrocarbons is specified in US Pat. Nos. 3,868,428 and 4,288,234. For example, the latter patent states that it operates at a pressure of 2-25 psia for dehydrogenation. It also shows the use of compressors on the discharge side of the dehydrogenation zone. However, this compressor is described as being located downstream of the cooling zone where the liquid aromatic hydrocarbon product stream is removed by condensation.

Therefore, this patent is similar to US Pat. No. 3,702,346, in which a vacuum pump controls steam from a phase separation vessel that receives the elution flow of a partially condensed dehydrogenation zone.

The production of stalene by dehydrogenation at low pressures of ethylbenzene is described in detail in Volume 41 p 519-528 of The Transactions of The American Institute or Chemical Engineers (1945).

U. S. Patent No. 3,978, 150 appropriately shows a paraffin dehydrogenation process in which the reactor is operated at subatmospheric pressure. This pressure is maintained by the vacuum source, which is mechanically located in the line carrying the reactor elution flow. The heat exchanger is located between the reactor and the vacuum source. This patent points out that sub-barometric separation of hydrocarbons can be carried out.

The present invention relates to a process for dehydrogenation of alkylaromatic hydrocarbons which can more economically achieve the desired subatmospheric pressure in the reaction zone. This improvement is achieved by installing a compressor in the elution line of the reaction zone in that water is present in the initial condensate upstream of the heavier C ₆ ⁺ hydrocarbons and in the elution flow in the dehydrogenation reaction zone. The embodiment of the present invention is therefore characterized by the process of dehydrogenation of alkylaromatic hydrocarbons consisting of a step of contacting a reactant stream comprising alkylaromatic hydrocarbons with a catalyst maintained under dehydrogenation conditions, which is a subatmospheric pressure and thus alkyl Production of gaseous carbohydrate elution streams containing aromatic hydrocarbons and unsaturated hydrocarbons, hydrogen and steam; Indirect heat exchange to the reactant stream to cool the carbohydrate zone elution stream without affecting condensation; Compressing the molten metal flow in the dehydrogenation zone by a first compressor to a high pressure equal to or less than one absolute air pressure; Partially condense the elution flow in the dehydrogenation zone; Separating the generated mixed phase dehydrogenation zone elution flow in the gas-liquid separation zone into a gas phase process flow including a hydrogenated liquid phase process flow while maintaining a pressure lower than one absolute pressure through the use of a second compressor; Recovering unsaturated hydrocarbon products from the liquid phase process stream. As shown in the figure, the inflow stream comprising ethylbenzene is indirectly heated by a heat exchanger in the influent eluent heat exchanger 2 to mix with the superheated stream vapor from line 3 through line 1. do.

This mixture of steam and ethylbenzene enters the bottom of reactor 5 via line 4. The reactants are sent up through the reactor to form three or more different layers of dehydrogenation catalyst maintained at subatmospheric dehydrogenation conditions. The reaction is heated in the reactor by a method not shown.

Contacting ethylbenzene with the dehydrogenation catalyst contained in the reactor produces a reaction zone elution stream carried through line 6 containing unreacted ethylbenzene, styrene, steam and hydrogen. This elution flow is cooled in the inlet elution heat exchanger (2) and through the steam generator (7) which extracts only detectable heat from the elution flow and does not condense water, ethylbenzene or styrene. The elution flow in the vapor phase is compressed in the first compressor 8 to a higher pressure of less than one absolute atmospheric pressure. Elution flow continues through line 6 and partially condenses in condenser 9 to produce a mixed phase stream. This stream is divided into liquid and gaseous hydrocarbons of water, including dissolved hydrocarbons in gas-liquid separator 10. The gas phase comprising hydrogen and low temperature water vapor is removed from the separator using a second compressor 12 via line 1. The hydrocarbons flow out through line 13 and pass through a fractionation distillation unit for recovery of the generated styrene. The condensed water is discharged from the separator via line 14.

The flow charts used for the purposes of the present invention do not limit the scope of the invention and other process flows herein or other variations not described herein may be implemented.

Processes for the dehydrogenation of aromatic hydrocarbons are widely used commercially. Large amounts of styrene, for example, are produced by dehydrogenation of ethylbenzene. The resulting styrene is itself polymerized or copolymerized with butadiene, isoprene, acrylonitrile and the like. Other hydrocarbons dehydrogenated in the same manner include diethylbenzene, ethyltoluene, propylbenzene and isopropylbenzene. However, since most of the commercial dehydrogenation processes described are used for dehydrogenation of ethylbenzene, the object of the present invention is represented by dehydrogenation of ethylbenzene. It does not depart from the scope of the present invention and has other ring structures including the alkylaromatic hydrocarbons or bicyclic compounds described above.

Dehydrogenation favors waterways as the number of molecules in the influent flow increases. Therefore, dehydrogenation is affected by pressure reduction. Therefore, most commercial dehydrogenation processes operate at relatively low pressures for hydrocarbon conversion processes and it is known that the use of subatmospheric pressure is preferred in conventionally known reference books. Despite this, conventional commercial processes usually use pressures close to atmospheric pressure rather than subatmospheric pressure in the dehydrogenation zone. The basic reason for this is that it is very difficult and expensive to maintain the dehydrogenation zone at or below 0.75 absolute atmospheres of the subatmospheric pressure.

One of the major causes of economically desirable low pressures is the reaction flow through the reactor and the pressure drop in the indirect heat exchanger of the dehydrogenation zone. More specifically, it is common to maintain the dehydrogenation reaction zone at low pressure by the outflow of gas flow from the product settler (gas-liquid separator), such as by a steam jet ejector.

Condensation significantly reduces the amount of gas remaining, so in this respect the vacuum is activated and it is essential that there is a reactor elution flow to the product separator. On a typical commercial process scale, diluents are present, such as evidence of high velocity flow of incoming hydrocarbons and of the entire apparatus, including delivery lines and heat exchangers. The pressure drop due to this high velocity flow is such that the product settler is at a gentle subatmospheric pressure and the reaction zone itself is at or above atmospheric pressure. This effect can be reduced but not eliminated by usefully designed instruments. Therefore, it is necessary in this way to maintain a commercially impractical vacuum in the product separator for supporting the reactor at subatmospheric pressure. It is therefore an object of the present invention to dehydrogenate an alkylaromatic hydrocarbon in which the reaction zone is maintained at subatmospheric pressure.

Another object of the present invention is a dehydrogenation process of ethylbenzene or ethyltoluene, where it is more economical to maintain the reaction zone at a pressure lower than atmospheric pressure.

In this process, a compression device is used at an upstream point of a product settler or gas-liquid separator to maintain a low pressure in the reaction zone.

More specifically these compression devices are located upstream of the direct or indirect heat exchanger is used to condense the C ₆ ⁺ hydrocarbons present in the reaction zone heuyi most elution flow. Such a compression device is suitably located downstream of the indirect heat exchanger used to recover the heat detectable from the reaction zone elution flow, but upstream of the device close to the reaction zone used to condense the product hydrocarbons. This location is important because the indirect heat exchanger used to influence condensation is an important cause of the total pressure drop across the reactant flow and path. By installing the compressor in this position, the pressure drop in the condenser can be avoided.

The compressor used upstream of the condenser is not the only device used to maintain the reduced pressure in the reaction zone. This compressor is therefore operated in combination with a second compressor used to remove gas flow from a gas-liquid separator conventionally used in the same way as the method. This downstream compressor is applied to maintain the low pressure of the discharge portion of the upstream compressor, thereby reducing the compression ratio across the upstream compressor and reducing the cost of the upstream compressor being operated. Therefore, the pressure of the discharge portion of the first or upstream compressor is less than one absolute air pressure. The gas-liquid separation zone is also kept below 1 absolute atmospheric pressure.

The dehydrogenation zone elution flow is cooled by an indirect heat exchanger before entering the first compressor. This heat is preferably used to heat the incoming hydrocarbons entering the dehydrogenation zone and produce steam in the same manner as shown in the figures. Initial heat exchange and cooling of the dehydrogenation zone elution stream are formed without significant condensation of water and C ₆ ⁺ hydrocarbons in the dehydrogenation zone elution stream. As used herein, "tapped condensation" refers to condensation of at least 5 mole percent of a particular compound or class of compounds. Therefore, the elution flow in the dehydrogenation zone is entirely gaseous when entering the first or upstream compressor. Initial indirect heat exchange recovers the high temperature detectable heat present in the elution flow to cool the elution flow thereby increasing the density of the gas phase flow and lowering the temperature of the first compressor. The elution flow is preferably cooled at least 400 ° F. and preferably 600 ° F. before passing through the first compressor. Compressors of a commercially suitable form are used as the first or second compression device with the use of the preferred centrifugal compressor.

A suitable embodiment of the present invention is to contact a reactant stream comprising alkylaromatic hydrocarbons with a multilayer of dehydrogenation catalyst maintained at subatmospheric pressure and dehydrogenation conditions in which steam is present, whereby alkyl aromatic hydrocarbons, unsaturated hydrocarbon products, hydrogen and steam To form a dehydrogenation zone effluent stream, including, to cool the dehydrogenation zone effluent stream by indirect heat exchange to the reactant stream without significant condensation of water and alkylaromatic hydrocarbons, and to reduce the dehydrogenation zone elution flow in the first compressor to 1 atmosphere Pressure in the gas-liquid separation zone, which partially condenses the dehydrogenation zone elution stream by indirect heat exchange, thereby creating a process flow of the mixed phase and maintained at a pressure below 1 absolute atmospheric pressure in the use of the second compressor. Mixed process flows include hydrogen and liquid process flows. Is separated by gas phase process flow; A dehydrogenation process of an alkylaromatic hydrocarbon consisting of recovering unsaturated product hydrocarbons from the liquid phase process stream and releasing the gaseous process flow through the use of a second compressor.

As used herein, "dehydrogenation zone" refers to an overall reactor system comprising a dehydrogenation catalyst. These catalysts may be divided into 10 or more separation beds or the dehydrogenation zone may consist of two or three catalyst beds with intermediate additions and apparatus for oxygen feed streams and vapor mixtures. Suitable systems for this are shown in US Pat. Nos. 3,498,755, 3,515,763, and 3,751,232. The catalyst layer is contained in the separation reaction vessel and has a cylindrical and cyclic form.

It is preferable to use the catalyst layer of the arrangement state piled up the whole single container. Other dehydrogenation catalysts are used in other layers as specified in US Pat. No. 3,223,743. Dehydrogenation catalysts generally consist of one or more metals selected from Group VI and Group V of the periodic table. One form of the alkylaromatic dehydrogenation catalyst consists of 85% by weight of ferric oxide, 2% chromia, 12% potassium hydroxide and 1% sodium hydroxide. Commercially used second dehydrogenation catalysts consist of 87-90% ferric oxide, 2-3% chromium oxide and 8-10% potassium oxide.

The third type of catalyst consists of 90% by weight iron oxide, 4% chromia and 6% potassium carbonate. Methods of preparing suitable catalysts are well known in the art. This technique is described in U.S. Patent No. 3,387,053, which contains 35% by weight of iron oxide as active catalyst and 1-8% by weight of zinc or copper oxide and 0.5-50% by weight of chalcery accelerator and stabilizer. A preparation of a catalyst composition comprising wt% chromium oxide and a binding agent is described.

Dehydrogenation conditions generally include a temperature of 538 ° -1000 ° C (1000 ° -1832 ° F), preferably 565 ° -675 ° C (1050 ° -1250 ° F). When ethylbenzene is dehydrogenated, the space velocity, the rate of the vapor mixture and the internal temperature are adjusted according to the elution of the catalyst bed having a temperature of about 593 ° C.

The temperature required for the effective operation of the characteristic dehydrogenation process depends on the activity of the incoming hydrocarbons and the catalyst used. The pressure maintained in the dehydrogenation zone is 100-750 mmHg, preferably 250-700 mmHg. The operating pressure in the dehydrogenation zone is measured at the inner, middle and outer zones to represent the average pressure. The combined inflow flow is released to the dehydrogenation zone at an hourly liquid space velocity of about 0.1-2.0 hours ^-1 based on 60 ° F liquid hydrocarbon emissions.

The alkylaromatic hydrocarbon to be dehydrogenated is mixed with the superheated vapor to hinder the temperature drop effect of the exothermic dehydrogenation reaction. The presence of steam interferes with the accumulation of carbon deposition and thus favors the stability of the dehydrogenation catalyst. Preferably, the vapor is mixed with the other components of the inflow stream at a vapor rate of 0.8-1.7 pounds per pound of incoming hydrocarbons. Other steam is added after one or more consecutive layers, if necessary. However, the dehydrogenation zone elution flow should include less than about 3 pounds of steam per hydrocarbon product and preferably less than 2 pounds of steam per mole of hydrocarbon production.

The elution flow removed from the dehydrogenation zone is rapidly heat exchanged for the dual effect of lowering its temperature and preventing heat recovery to prevent styrene polymerization. The elution flow is heat exchanged for the steam flow and the reactant streams of this or other processes are used as heat sources for fractionation. Commercially, the elution flows through various heat exchangers whereby many streams are heated. This heat exchange fulfills the limitations described above. The heat exchange performed downstream of the first compressor cools the dehydrogenation elution flow to allow sufficient condensation of at least 95 mole percent influent and produced hydrocarbons and at least 95 mole percent water vapor. It is not desirable to use a quench zone through this condensation. In essence, all other styrene or other hydrocarbon products, most of the other compounds present in the water and elution streams, are converted to liquid by this condensation. This creates a mixed phase flow through the phase separation vessel. This process can be easily separated by adding hydrocarbons from water and hydrogen present in the elution stream. Styrene present in the dehydrogenation zone elution stream becomes part of the hydrocarbon stream exiting the separation vessel and is passed to a suitable separation unit. Styrene is recovered from the hydrocarbon stream using one of several fractionation systems in accordance with known techniques.

This fractionation yields an additional stream comprising relatively pure ethylbenzene and benzene and toluene, which are recycled. These two aromatic hydrocarbons are by-products of the dehydrogenation reaction. These are part of, but not the entire process, of US Pat. No. 3,409,689 and UK Pat. No. 1,238,602. Styrene is recovered as a third flow and released from the process. If desired, fractionation is used to recover styrene. For example, US Pat. No. 3,784,620 specifies the separation of styrene and ethylbenzene by the use of polyamide permeable membranes such as nylon-6 and nylon 6.10. U.S. Patent No. 3,513,213 specifies a separation method of a liquid-liquid extraction method using silver fluoroborate anhydride as a solvent. Similar methods of separation using Cuprus flowoborate and Cuprus flowophosphate are described in US Pat. Nos. 3,517,079 and 3,517,080 and 3,517,081.

Methods for recovering styrene through the use of fractionation are described in several references, including US Pat. No. 3,525,776. In this reference the hydrocarbonaceous phase removed from the phase separation zone is sent to the first column referred to as the benzenetoluene column. The column is operated under subatmospheric pressure, allowing operation at lower temperatures and reducing the rate of styrene polymerization. Various blockers such as sulfur atoms, 2, 4-dinitrophenol or a mixture of N-nitrosodiphenylamine and dinitroso-ortho-cresol are also injected into the column for the same purpose. Sulfur atoms are also injected into this column by recycling at least a portion of the high molecular weight material separated from the bottom stream of the styrene purified column. Further details are set forth in US Pat. Nos. 3,476,656, 3,408,263 and 3,398,063. The separation of benzene and toluene from the benzene-toluene column content produces an overhead stream free of styrene and ethylbenzene. This stream contains at least 95 mole percent benzene and toluene. The bottom of the benzene-toluene column is recycled into the second fractionation column where ethylbenzene is removed as a top product. The bottom flow of this column is purified to give styrene.

The exotherm of the dehydrogenation reaction passes through the bed of the dehydrogenation catalyst resulting in rapid cooling of the reactants. This reduces the conversion that can occur. For this reason, most commercial processes involve internal stage heating between the separation layers of the dehydrogenation catalyst. This heating is done by indirect heat exchange or the addition of very hot gases which are superheated steam. Continuous heating by indirect heat exchange is also used. In a more limited embodiment of the process, at least a portion of the internal stage heating is obtained by catalytically promoting oxidation. This not only reheats the incoming hydrocarbons, but also reduces the hydrogen concentration in the reaction stream and thereby promotes increased conversion. The introduction of heat by hydrogen combustion reduces the heat that must be supplied by other devices. In the use of superheated steam, the use of partial hydrogen combustion reduces the amount of steam required. This in turn reduces the cost of the operating process because less superheated steam must be produced and less water vapor must be condensed to remove from the dehydrogenation zone elution flow.

Oxygen consumed during hydrogen combustion is mixed into the reactant stream at the point of internal stage heating as part of the oxygen feed flow. The oxygen supply flow is air but is preferably a gas having a higher oxygen content than air. The oxygen supply flow has a nitrogen content of less than 10 mole percent and the use of pure oxygen is preferred if economically possible. The desired oxygen concentration in the oxygen supply flow is an economic problem and is determined by comparing the benefits of pure oxygen with the cost of obtaining oxygen. The basic advantage of the presence of nitrogen is to dilute the gas flow, including hydrogen removed from the product separation vessel, to increase the input reduction of the catalyst bed as nitrogen passes through the dehydrogenation zone and to maintain absolute pressure in the dehydrogenation zone. On the one hand, the presence of nitrogen acts as a diluent to affect the level of equilibrium conversion.

The conversion catalyst used in this process to promote the hydrogenation of the internal stage is a commercially suitable catalyst suitable for the required stability and activity, and the hydrocarbon product or influent has a high selectivity of hydrogenation compared to oxidation. In other words, the oxidation catalyst must have a high selectivity of hydrogen oxidation while oxidizing only a small amount of influent or hydrocarbons produced. The oxidation catalyst has a different composition from the dehydrogenation catalyst. Preferred oxidation catalysts include Group VIII metals or metal cations with crystal ion diameters greater than 1.35 A, both being present in small amounts in the refractory solid support. Preferred Group VIII metals are platinum and palladium, but the use of ruthenium, rhodium, osmium and iridium is also possible. Group VIII metals are present in amounts of 0.01-5.0% by weight of the post-treated catalyst. The metal or metal cation having a diameter of at least 1.35 mm 3 is selected from Group IA or Group IIA and is present in an amount of 0.01-2.0% by weight of the post-treated catalyst. The component of this catalyst is barium, but the use of other metals, including rubidium or cesium, is also possible.

Suitable solid supports are aluminas having a surface area of 1-300 m 2 / g, a bulk density of 0.2-1.5 g / cc, and an average pore size of at least 20 mm 3. The metal-containing component is immersed in an aqueous solution, soaked into solid particles of a solid support, dried and quenched in air at 500-600 ° C. The support is spherical, pallet, or injection. The total amount of oxidation catalyst present in the dehydrogenation zone is 30% by weight or less and more preferably 5-15% by weight of the total dehydrogenation catalyst amount.

The conditions used during contacting the reactant stream with the other layers of the oxidation catalyst apply to a wide range by the dehydrogenation conditions mentioned above. The preferred external temperature of the oxidation catalyst layer is the internal temperature of the downstream layer of the dehydrogenation catalyst. The temperature increasing across the layer of the oxidation catalyst is 80 ° C. or less. The hourly liquid space velocity based on liquid hydrocarbon release at 60 ° F. is 2-10 hours ⁻¹ . All of the oxygen entering the oxidation catalyst layer is consumed in the layer of the oxidation catalyst and the elution flow of the oxidation catalyst layer contains less than 0.1 mole percent oxygen. The total moles of oxygen released to the dehydrogenation zone is less than 60% of the total hydrogen moles available in the dehydrogenation zone for combustion and therefore depends on the conversions obtained in the dehydrogenation zone and some of the hydrogen is lost into the solution or open gas stream. do. Useful hydrogen is the sum of the hydrogen recycled to the dehydrogenation zone and hydrogen is produced everywhere in the last layer of the dehydrogenation catalyst. Oxygen released to the dehydrogenation zone is 20-50 mole percent of the useful hydrogen mentioned above. As used herein, “substantially all” refers to the majority of the indicated compound compounds acting in the manner described and this ratio is 90 mole percent more preferably 95 mole percent. As mentioned earlier, this process does not limit the production of styrene and is used to produce paramethyl styrene by dehydrogenation of ethyltoluene or for the production of other unsaturated hydrocarbon products.

Claims (11)

The reactant stream comprising alkyl aromatic hydrocarbons is contacted with a dehydrogenation catalyst maintained at quasi-pressure dehydrogenation conditions, thereby forming a gaseous dehydrogenation zone elution stream comprising alkylaromatic hydrocarbons, unsaturated hydrocarbon products and hydrogen: Cool the dehydrogenation zone elution flow without condensation by indirect heat exchange to the reactant stream: compress the dehydrogenation zone elution stream to less than one absolute pressure by means of a first compressor: partially condense the dehydrogenation zone elution stream and Decomposition of the mixed phase dehydrogenation zone in a gas-liquid separation zone maintained at a pressure of 1 atmosphere or less using a second compressor, is separated into a gas phase process flow comprising hydrogen and a liquid process flow. Recovering the unsaturated hydrocarbon product from the alkyl Aromatic Dehydrogenation Process. The process of claim 1 wherein the dehydrogenation zone elution stream is cooled to at least 204 ° C. before being compressed in the first compressor. 3. The process of claim 2 wherein the alkyl aromatic hydrocarbon is ethylbenzene and the unsaturated hydrocarbon product is styrene. 3. The process of claim 2 wherein the alkylaromatic hydrocarbon is ethyltoluene and the unsaturated hydrocarbon product is methylstyrene. The process of claim 2, wherein the reactant stream comprises steam. 6. The process according to claim 5, wherein the reactant stream is in contact with at least two separation layers of the dehydrogenation catalyst and the reactant stream is heated by oxidation of hydrogen at the midpoint between the two layers of the dehydrogenation catalyst. A reactant stream comprising alkylaromatic hydrocarbons is contacted with a multilayer of dehydrogenation catalyst maintained at subatmospheric pressure and dehydrogenation conditions in the presence of steam, thereby dehydrogenation zone elution stream comprising alkylaromatic hydrocarbons, unsaturated hydrocarbon products, hydrogen and steam To cool the dehydrogenation zone elution flow by indirect heat exchange to the reactant stream without noticeable condensation, and to compress the dehydrogenation zone elution flow to below 1 absolute atmospheric pressure by means of a first compressor: Partially condensed by indirect heat exchange, thereby creating a process flow of the mixed phase: using a second compressor to maintain the process flow of the mixed phase in the gas-liquid separation zone maintained at a pressure of less than one absolute atmospheric pressure. Gas phase process flow, including hydrogen and liquid process streams, Separation and: a dehydrogenation step of the alkyl aromatic hydrocarbons are composed of a step of recovering the unsaturated product hydrocarbon from the liquid phase process stream. 8. The process of claim 7, wherein the reactant stream is heated by oxidation of catalytically promoted hydrogen at the midpoint between the separation layers of the dehydrogenation catalyst. The process of claim 8, wherein the unsaturated hydrocarbon product is paramethyl styrene. The process of claim 8, wherein the unsaturated hydrocarbon product is styrene. The process of claim 8, wherein the dehydrogenation zone is cooled to at least 204 ° C. before being compressed in the first compressor.

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