CA1226587A - Alkylaromatic hydrocarbon dehydrogenation process - Google Patents
Alkylaromatic hydrocarbon dehydrogenation processInfo
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- CA1226587A CA1226587A CA000463924A CA463924A CA1226587A CA 1226587 A CA1226587 A CA 1226587A CA 000463924 A CA000463924 A CA 000463924A CA 463924 A CA463924 A CA 463924A CA 1226587 A CA1226587 A CA 1226587A
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Abstract
"ALKYLAROMATIC HYDROCARBON DEHYDROGENATION PROCESS"
ABSTRACT
A hydrocarbon conversion process is disclosed for the de-hydrogenation of alkylaromatic hydrocarbons such as ethylbenzene or ethyltoluene. A subatmospheric pressure is maintained in the reaction zone by the use of two separate compressors. The first compressor pressurizes the vapor phase reactor effluent prior to the final indirect heat exchange step(s) used to partially condense this stream. The second compressor maintains the vapor-liquid separator which receives the partially condensed reactor eff1uent stream at a subatmospheric pressure. This facilitates operation of the reactor at a subatmospheric pressure.
ABSTRACT
A hydrocarbon conversion process is disclosed for the de-hydrogenation of alkylaromatic hydrocarbons such as ethylbenzene or ethyltoluene. A subatmospheric pressure is maintained in the reaction zone by the use of two separate compressors. The first compressor pressurizes the vapor phase reactor effluent prior to the final indirect heat exchange step(s) used to partially condense this stream. The second compressor maintains the vapor-liquid separator which receives the partially condensed reactor eff1uent stream at a subatmospheric pressure. This facilitates operation of the reactor at a subatmospheric pressure.
Description
"ALKYLAROMATIC HYDROCARBON DEHYDROGENATION_PROCESS"
F LO OF THE INVENTION
The invention relates to the processing of hydrocarbons and more specifically to a process for the catalytic dehydrogenation of alkylaromatic hydrocarbons. The invention therefore relates to the synthesis of aromatic compounds, such as styrenes by dehydrogena-lion, with the dehydrogenation occurring in a number of sequential stages. The invention directly relates to an improved method for maintaining a sub atmospheric pressure in the dehydrogerlation reactor.
PRIOR ART
___ The dehydrogenation of hydrocarbons is well described in the prior art, with both cyclic and aromatic hydrocarbons being lo thereby converted to the corresponding less saturated products. For instance dehydrogenation is performed commercially for the production of styrenes from ethylben~ene to fulfill the sizable demand for this polymer precursor. The product styrenes may be polymerized with it self or it may be copolymerized with butadiene, isoprene, acryloni-trite, etc. Processes for the dehydrogenation of alkylaromatic hydrocarbons are often integrated with an alkylation process which produces the alkylaromatic hydrocarbons.
US. Patent No. 3,~15,766 issued to W. N. Root et at and US. Patent No. 3140916B9 issued to D. I. Ward are pertinent for their showing of typical prior art catalytic steam dehydrogenation processes for alkylaromatics including ethylbenzene. These referent cues describe the admixture of superheated steam into the feed hydra-carton and the admixture of additional amounts of superheated steam I;
..~,~, I
with the reactants between sequential beds of dehydrogenation gala-lust. These references also show an overall process flow into which the subject process could be integrated.
The use of a sub atmospheric pressure in the reaction zone of an alkylaromatic hydrocarbon dehydrogenation process it known as shown by the teachings of US. Patent Nos. 3,868,428 and 4,2~8t~34.
For instance, the latter of these patents teaches that the dodder-genation zone is operated at a pressure in the order of prom 2 to 25 Asia. This reference is also pertinent for its showing of the use of a compressor on the discharge side of the dehydrogenati.on zone.
However, this compressor is described as being located downstream of a cooling zone in which a liquid phase crude aromatic hydrocarbon product stream is removed by condensation. The reference therefore seems to be similar in this regard to US. Patent No. 3,702~346, in which a vacuum pump removes vapors from the phase separation vessel which receives the partially condensed dehydrogenation zone effluent stream.
The production of styrenes by the low pressure dehydrogena-lion of ethylbenzene is described in detail in an article at pages 519-528 of volume 41 of The Transactions of The American Institute at (1945).
US. Patent No. 3~978~150 is pertinent for it showing ox a paraffin dehydrogenation process in which the reactor is operated at a sub atmospheric pressure. This pressure is maintained by vacuum source which may by mechanical in nature, located in the line carry-in the reactor effluent stream. A heat exchanger is located between the reactor and thy vacuum source. This patent also indicates that a sub atmospheric pressure separation of the product hydrocarboll could Lo be perfDnmed.
BRIEF SUMMARY OF THE INVENTION
The invention provides a process for the dehydrogenation of alkylaromatic hydrocarbons in which it is more economical to achieve a desired sub atmospheric pressure in the reaction zone.
This improvement is accomplished by placing a compressing means in the effluent line of the reaction zone at a point upstream of the initial condensation of the heavier C6~ hydrocarbons and water pros-en in the dehydrogenation reaction Noah effluent stream. A broad embodiment of the invention may therefore be characterized as a pro-lo cuss for the dehydrogenation of an alkylaromatic hydrocarbon which comprises the steps of contacting a reactant stream comprising an alkylaromatic hydrocarbon with a catalyst maintained at dehydrogena-lion conditions which include a sub atmospheric pressure dud thereby forming a vapor-phase dehydrogenation zone effluent stream comprise in the alkylaromatic hydrocarbons, an unsaturated product hydrocar-bony hydrogen and steam; cooling the dehydrogenation zone effluent : stream without affecting significant condensation by indirect heat - exchange against the reactant stream; compressing the dehydrogena~ion zone effluent stream to a higher pressure which is less than one atom-sphere absolute by means of a first mechanical compressing jeans;
partially condensing the dehydrogenation zone effluent stream; sepal rating the resultant mixed phase dehydrogenation zone effluent stream in a vapor-liquid separation one maintained at a pressure less than one atmosphere absolute through the use of a second mechanical come pressing means, into a vapor-phase process stream comprising hydrogen and a liquid-phase process storyline; and recovering the unsaturated product Herr-~22~587 carbon from the liquid phase process stream.
BRIEF DESCRIPTION OF THE DRAWING
The Drawing is a schematic illustration of the reaction section of a process for the dehydrogenation of ethylbenzene to pro-dupe styrenes Referring now to the Drawing, a feed stream comprise in ethylbenzene is heated by indirect heat exchange in the feed effluent heat exchange means 2 and then continues through line 1 to be admixed with a stream of superheated steam from line 3. This ad-mixture of steam and ethylbenzene passes into the bottom of the react ion 5 through line 4. The reactants pass upward through the reactor preferably making three or more passes through different beds of de-hydrogenation catalyst maintained at dehydrogenation conditions which include a sub atmospheric pressure. The reactants may be heated with-in the reactor by means not shown.
The contacting of the ethylbenzene with the dehydrogenation catalyst contained within the reactor produces a reaction zone efflu-en stream carried by line 6 which comprises unrequited ethylbenzene, styrenes steam and hydrogen. This effluent stream is cooled in the feed effluent heat exchange means 2 and by passage through a steam generator 7, with this heat exchange extracting only sensible heat from the effluent stream and producing no significant condensation of water/ ethylbenzene or styrenes The still vapor-phase effluent stream is then compressed in the first compressor 8 to a higher pros-sure which preferably lo still less than 1 atmosphere absolute. The effluent stream continues through line 6 and is then partially con-dented in the condenser 9 to produce a mixed-phase stream. This stream is separated in the vapor-liquid separator 10 into a first liquid phase of water containing dissolved hydrocarbons, a hydrocar-bun liquid phase and a vapor phase. The vapor phase comprising hydra-gun and low temperature water vapor is removed from the separator through line 11 through the use of a second compressor 12. The hydra-carbons are withdrawn through line 13 and preferably passed to the appropriate fractional distillation facilities for the recovery of the product styrenes The condensed water is decanted from the sepal rotor through line 14.
This illustration of one possible process flow which may be utilized with the subject invention is not intended to thereby limit the scope of the invention, which may be practiced with the other process flows set out herein or in other variations not de-scribed herein.
DETAILED DESCRIPTION
Processes for the dehydrogenation of aromatic hydrocarbons are in widespread commercial use. For instance, large quantities of styrenes are produced by the dehydrogenation of ethylbenzene. The no-sultan styrenes may be polymerized with itself or it may be copolyml-prized with butadiene, isoprene~ acrylonitrile, etc. Other hydrocar-buns which may be dehydrogenated in much the same manner include do-ethylbenzene, ethyltoluene, propylbenzene and isopropylbenaene. How-ever since the great majority of the present commercial dehydrogena-lion processes are employed for the dehydrogenation of ethylbenzene, the following description ox the subject invention will be presented primarily in terms of the dehydrogenatian of ethylben~ene. This is not intended to exclude from the scope of the subject invention those other alkylaromakic hydrocarbons set out above or those having differ-en ring structures including bicyclic compounds.
The dehydrogenation reaction releases hydrogen thereby in-creasing the number of molecules present compared to the number in the feed stream. The dehydrogenation reaction is therefore favorably influenced by a decrease in pressure. Accordingly most commercial dehydrogenation processes operate at a relatively low pressure or hydrocarbon conversion processes, and as pointed out in the previous-lye cited references, it is known that the use of sub atmospheric pros-surges would be desirable. Despite this knowledge, it is believed that the present commercial operations normally employ a pressure in the dehydrogenation zone which is close to atmospheric pressure rather than significantly below it. The basic reasons for this are the ox-pens and difficulty of maintaining the dehydrogenation zone at sub-atmospheric pressures on the order of 0.7~ atmosphere absolute or be-low. One of the major causes of this difficulty in economically ox-twining the desired low pressures is the pressure drop which is in-hornet in the foe of the reactants through the reactor and the sub sequent indirect heat exchange means of the dehydrogenation zone.
More specifically, it is common practice to maintain the dehydrogena-ZOO lion reaction zone at a low pressure by withdrawing a vapor stream from the product settler (vapor-liquid separator), as by a steam jet ejector. The vacuum is "pulled" at this point since the condensation has greatly reduced the amount of vapor which remains, and it is neck essay to have the reactor effluent flow toward the product swooper-ion. In a commercial scale process, there is a very high flow rate of the weed hydrocarbon and any delineate such as steam through the total apparatus including the transfer lines and heat exchangers.
The cumulative pressure drops in these devices due to thus high flow I
rate are such that even though the products settler is at a mildly sub atmospheric pressure, the reaction zone itself will be at a pros-sure equal to or above one atmosphere absolute. This effect can be reduced by judicious equipment design but cannot be eliminated.
Therefore, it becomes necessary to maintain a co~ercially imprecate-eel vacuum in the product separator to hold the reactor at a subitem-spheric pressure in this manner.
It is therefore an objective of the subject invention to provide a process for the dehydrogenation of alkylaromatic hydrocar-buns in which the reaction zone is maintained at a sub atmospheric pressure. It is a further objective of the subject invention to pro-vise a process for the dehydroyenation of ethylbenzene or ethyltolu-one in which it is more economical to maintain the reaction zone at a pressure which is significantly below atmospheric pressure.
In the subject process a mechanical compression means is utilized at a point upstream of the products settler or Yapor-liquid separator to maintain the desire low pressure within the reaction zone. More specifically this mechanical compression means it toga-ted upstream of the indirect or direct heft exchange means utilized to condense the very great majority of the I hydrocarbons present in the reaction zone effluent stream. This compression means is preferably located downstream of any indirect heat exchange means which are utilized to recover sensible heat from the reaction zone efflu nut stream but is located upstream that is closer to the react I lion zone, of the means utilized to condense the product hydrocarbon.This location is significant since the preferred indirect heat ox-change means utilized to effect the condensation is the source of a very significant amount of the total pressure drop through the react lent flow path. The provision of a compression means at this toga-lion thereby avoids the pressure drop through this condensing means.
The compression means utilized upstream of the condensing means is not the sole means utilized to maintain the desired reduced pressure within the reaction zone. This compression means therefore operates in cooperation with a second compression means which prefer-ably is utilized to remove the vapor Stream from the vapor-liquid separator much in the same manner emp10yed in the prior art. This downstream compression means functions to maintain a low pressure on the discharge side of the upstream compression means, thereby red-cuing the compression ratio across the upstream compression means and reducing the utilities cost of operating the upstream compression means. It is therefore preferred that the pressure on the discharge side ox the first or upstream compression means is less than 1 atom-sphere absolute. It is also preferred that the vapor-liquid swooper-lion zone is maintained at a pressure less than I atmosphere absolute.
The dehydrogenation zone effluent stream is tooled by in-direct heat exchange means prior to entering the first compression means. Preferably, this heat is utilized to heat the feed hydr~car-bun entering the dehydrogenation zone and to generate steam in the Fanner illustrated in the Drawing. This initial heat exchange and cooling of the dehydrogenation zone effluent stream is to be per-formed without any significant condensation of either water or C6+
hydrocarbons present in the dehydrogenation zone effluent stream.
I As the term "significant condensation" is used herein, it is intended to indicate the condensation of more than S mole percent of any par titular compound or class of compounds. It is therefore preferred that the dehydrogenation zone effluent stream is a totally vapor-phase stream when it enters the first or upstream compression jeans.
The initial indirect heat exchange recovers the valuable high them-portray sensible heat present in the effluent stream and cools the effluent stream thereby increasing the density of this vapor-phase stream and lowering the operating temperature of the first compress soon means. It is preferred that the effluent stream is cooled at least 400 and more preferably at least 600 Fahrenheit degrees before being passed into the first compressing means. Any commercially suitably type of mechanical compression means may be employed as the first or second compression means, with the use of centrifugal come pressers being preferred.
A preferred embodiment of the invention may therefore be characterized as a process for the dehydrogenation of an alkylaroma-tic hydrocarbon which comprises the steps of contacting a reactant stream comprising an alkylaromatic hydrocarbon with a plurality of beds of dehy~rogenation catalyst maintained at dehydrogenation condo-lions which include a subatmcspheric pressure and the presence of steam and thereby forming a dehydrogenation zone effluent stream come prosing the alkylaromatic hydrocarbon, an unsaturated product hydra-carbon, hydrogen and steam; cooling the dehydrogenation zone efflu-en stream by indirect heat exchange against the reactant stream without significant condensation of water or the alkylaromatic hydra-carbon; compressing the dehydrogenation zone effluent stream in a first mechanical compressing means to a higher pressure less than 1 Z5 atmosphere absolute; partially condensing the dehydrogenation zone effluent stream by indirect heat exchange and whereby producing a mixed-phase process stream; separating the mixed-phase process stream in a vapor-liquid separation zone, maintained at a pressure Tess than l atmosphere absolute through the use of a second mechanical compress sing means, into a vapor phase process stream comprising hydrogen and a liquid-phase process stream; and recovering the unsaturated product hydrocarbon from the liquid-phase process stream, and with-drawing the vapor-phase process stream through the use of a second compressing means.
As used herein the tern "dehydrogenation zone" is intended to refer to the total reactor system which contains the dehydrGgena-lion catalyst. This catalyst may be divided into ten or more sepal rate beds, but the dehydrogenation zone preferably comprises two or three catalyst beds with means for the intermediate addition and ad-mixture of steam and possibly an oxygen supply steam. Suitable soys-terms for this may be patterned after those presented in US. Patent Nos. 3,498,755, 3,515,763 and 3,751,232. The catalyst beds may be contained in separate reaction vessels and may have either a Solon-Dracula or an annular shape. The use of radial flow catalyst beds in a stacked configuration in a single overall vows is preferred.
Different dehydrogenation catalysts may be used in different beds as described in US. Patent No. 3,223,743. Dehydrogenation catalysts generally consist of one or more metallic components selected prom Groups VI and VIII of the Periodic Table. One typical catalyst for the dehydrogenation of alkylaromatics comprises 85~ by weight ferris oxide, 2% Crimea 12% potassium hydroxide and 1% sodium hydroxide.
A second dehydrngenation catalyst, which is used commercially, con-sits of 87-90% eureka oxide, 2-3X chromium oxide and from 8-lQ~
potassium oxide. A third typical catalyst comprises 90~ by weight iron oxide I Crimea and I potassium carbonate. Methods for pro-paring suitable cdtdlysts are well known in the art. This is demon-I it striated by the teachings of US. Patent No. 3,3~7,053, which describes the manufacture of a catalytic composite of at least 35 wt.% iron ox-ire as an active catalytic agent, from about 1-8 wt.% zinc or copper oxide, about 0.5-50 wt.% of an alkali promoter, and from about 1-5 wt.% chronic oxide as a stabilizer and a binding agent.
Dehydrogenation conditions in general include a tempera-lure of about 538 to about 1000C (1000-1832F~ and preferably about 565 to about 675C (1050-1250F). When ethylbenzene is ye-in dehydrogenated, the space velocity, the rate of steam admixture and the inlet temperature are preferably adjusted to result in the effluent of each catalyst bed having a temperature of about 593C.
The temperature required for efficient operation of any specific de-hydrogenation process will depend on the feed hydrocarbon and the activity of the catalyst employed. The pressure maintained within I the dehydrogenation zone may range from about 100 to about 750 mm Hug, with a preferred range of pressures being from 250 to 7Q0 mm Hug. The operating pressure within the dehydrogenation zone is measured at the inlet, midsection, and outlet of the zone to thereby provide an approximately average pressure. The combined feed stream is charged to the dehydrogenation zone at a liquid hourly space velocity, based on liquid hydrocarbon charge at 60F, of about 0.1 to about 2.0 hurl and preferably from 0.2 to 1.0 ho 1.
The alkylaromatic hydrocarbon to be dehydrogenated is prey-drably admixed with superheated steam to counteract the temperature lowering effect of the endotherm-c dehydrogenation reaction. the presence of steam has also been described as benefiting the stabile lay of the dehydrogenation catalyst by preventing the accumulation of carbon deposits. Preferably, the steam is admixed with the other ~2~i58~
components of the feed stream at a rate of about 0.8 to about 1.7 pound of steam per pound of feed hydrocarbon. Other quantities of steam may be added after one or more subsequent beds if desired.
However, the dehydrogenation zone effluent stream should contain less than about 3 pounds of steam per pound of product hydrocarbon and preferably less than 2 pounds of steam per pound of product ho-drocarbon.
The effluent stream removed from the dehydrogenation zone is normally quickly heat exchanged for the dual purposes of lowering its temperature to prevent polymerization of the styrenes and for the recovery of heat. The effluent stream may be heat exchanged against a stream of steam, a reactant stream of this or another process or used as a heat source for fractionation, etc. Commercially the of-fluent stream is often passed through several heat exchangers there-by heating a number of different streams. This heat exchange is per-formed subject to the constraints set out above. The heat exchange performed downstream of the first compression means should cool the dehydrogenation zone effluent stream sufficiently to affect the con-sensation of at least 95 mole percent of the feed and product hydra-carbons and also at least 95-mole percent of the water vapor. The use of a quench zone to accomplish this condensation is not preferred.
Essentially all of the styrenes or other product hydrocarbon, most water and other readily condensable compounds present in the efflu-en stream are thereby converted to liquids. This produces a mixed-phase stream which us passed into a phase separation vessel. This procedure allows the facile crude separation by d contusion of the hydrocarbons from the water and hydrogen present in the effluent stream. The styrenes present in the dehydrogenat;on zone effluent stream becomes part of a hydrocarbon stream which it withdrawn from the separation vessel and transferred to the proper separation facile-ties. Preferably, the styrenes is recovered from the hydrocarbon stream by using one of the several fractionation systems known in the art. This fractionation will preferably yield a relatively pure stream of ethylbenzene, which is recycled, and an additional stream comprising Bunsen and Tulane. These two aromatic hydrocarbons are by-products of the dehydrogenation reaction. They may be recycled in part as taught in US. Patent No. 3,409,689 and British Patent 1,238,602 or entirely rejected from the process. Styrenes is recov-eyed as a third stream, which is withdrawn from the process. If de-sired, methods other than fractionation may be used to recover the styrenes For instance, US. Patent No. 3,784,620 teaches the swooper-lion of styrenes and ethylbenzene through the use of it polyamide per-I mention membrane such as nylon-6 and nylon 6,10. US. Patent No.
3,513,213 teaches a superior method employing liquid-liquid extract lion in which an hydrous silver fluoroborate it used as the solvent.
Similar separator methods utilizing cuprous fluoroborates and cuprous fluorophosphates are described in US. Patent Nos. 3,517,079; 3,517,080 and 3,517,081.
The recovery ox styrenes through the use of fractionation is described in several references including ITS Patent No. 3,~25~776.
In this reference, the hydrocarSonaceous phase removed from the phase separation zone is passed into a first column referred to as a Hanson-Tulane column. This column is operated at a sub atmospheric pressure to allow its operation at lower temperatures and hence reduce the rite of styrenes polymerization. Various inhibitors such as elemental sulfur, 2,4-d~n~tropheniol or a mixture of Nutrias diphenyl amine and a dinitroso-o-cresol are injected into the column for this same purpose. Preferably, sulfur is also introduced into this column by returning at least a portion of the high molecular weight material separated from the bottoms stream of a styrenes purification column.
A more detailed description of this is contained in Us Patent Nos.
3,476,656; 3,408~263 and 3,398,063. There is effected within the benzene-toluene column a separation of Bunsen and Tulane from the effluent to produce an overhead stream which is substantially free of styrenes and ethylbenzene. This stream preferably contains at least 95 mole percent Bunsen and Tulane. The bottoms of the bent zene-toluene column is passed into a second fractionation column from which ethylbenzene is removed as an overhead product and rely-clod. The bottoms stream of this column is then purified to obtain the styrenes The endothermic nature of the dehydrogenation reaction no-suits in a quick cooling of the reactants as they pass through the beds of dehydrogenation catalyst. This reduces the achievable con-version. For this reason the predominant commercial processes in-elude some form of interstate heating between the separate beds of dehydrogenation catalyst. This heating may be by indirect heat ox-change or by the addition of a very hot vapor which is normally superheated steam. Continuous heating by indirect heat exchange may also be employed. In more limited embodiments of the subject pro-cuss, at least a portion of the interstate heating is obtained by the catalytically promoted oxidation of hydrogen. This not only no-heats the feed hydrocarbon but also reduces the hydrogen concentra-lion in the reactant stream and thereby promotes increased confer-soon. Any heat input by hydrogen combustion reduces the heat which I
I
must be supplied by other means. In the preferred use of superheated steam, the use of partial hydrogen combustion therefore reduces the amount of steam which is required. This in turn reduces the utile-ties cost of operating the process since less superheated steam must be produced and less water vapor must be condensed for removal from the dehydrogenation zone effluent stream.
The oxygen consumed during the hydrogen combustion is prey-drably admixed into the reactant stream at the point of interstate heating as part of an oxygen supply stream. The oxygen supply stream may be air but is preferably a gas having a higher oxygen content Han air. It is preferred that the oxygen supply stream has a vitro-gun content less than 10 mole percent, with the use of substantially pure oxygen being preferred if it is economically viable. The pro-furred oxygen concentration in the oxygen supply stream is primarily lo a matter of economics and would be determined by a comparison of the advantage of having pure oxygen to the cost of obtaining the oxygen.
The basic disadvantages of the presence of nitrogen are the dilution of the hydrogen-containing gas stream removed from the product sepal ration vessel and the fact that the nitrogen passes through the dewy-drogenation Noah thereby increasing the pressure drop through the catalyst bed and the absolute pressure being maintained within the dehydrogenation zone. On the other hand, the presence of nitrogen favorably affects the equilibrium conversion level by acting as a : delineate.
I The oxidation catalyst employed in the subject process Jo pronto the ;nterstage hydrogen oxidation may be any commercially suitable catalyst which meets the required standards for stability and activity and which possesses high selectivity for the oxidation _15_ of hydrogen as compared with the oxidation of the feed or product ho-drocarbon. That is, the oxidation catalyst must have a high select tivity for the oxidation of hydrogen with only small amounts of the feed or product hydrocarbon being oxidized. The oxidation catalyst will have a different composition than the dehydrogenation catalyst.
The preferred oxidation catalyst comprises a Group VIII noble metal and a metal or metal cation which possesses a crystal ionic radius greater than 1.35 A, with both of these materials being present in small amounts on a refractory solid support. The preferred GroLIp VIII metals are platinum and palladium, but the use of ruthenium, rhodium, osmium and iridium is also contemplated. The Group VIII
metal is preferably present in an amount equal to 0.01 to 5.0 wt.%
of the finished catalyst. The metal or metal cation having a radius greater than 1.35 A is preferably chosen from Groups IA or IDA and is present in an amount equal to about 0.01 to about 20 wt.% of the finished catalyst. Iris component of the catalyst is preferably barium, but the use of other metals including rubidium or sesame is also contemplated.
The preferred solid support is alumina having a surface area between 1 and 300 mug an apparent bulk density of between about 0.2 and 1.5 g/cc, and an average pore size greater than I A.
The metal-contalning components are preferably impregnated into solid particles of the solid support by immersion in an aqueous solution followed my drying and calcination at a temperature of from about ~500 to 600C in air. The support may be in the form of spheres.
pellets or extradites. The total amount of oxidation catalyst pros-en within the dehydrogenaeion zone is preferably less than 30 White of the total amount of dehydrogenation catalyst and more preferably ~22~ 37 is between 5 and 15 White of this total amount of dehydrogenation catalyst.
The conditions utilized during the contacting of the react lent streams with the different beds ox oxidation catalyst will be set to a large extent by the previously referred' to dehydrogenation conditions. The preferred outlet temperature of any bed of oxide-lion catalyst is the preferred inlet of the immediately downstream bed of dehydrogenation catalyst. The temperature rise across any bed of oxidation catalyst is preferably less than By Centigrade de-greet. The liquid hourly space velocity, based on the liquid hydra-carbon charge at 60~F, is preferably between 2 and 10 hurl It is preferred that substantially all of the oxygen which enters a bed of oxidation catalyst is consumed within that bed of oxidation catalyst and that the effluent stream of any bed of oxidation catalyst contains less than 0.1 mole percent oxygen. The total moles of oxygen charged to the dehydrogenation zone is preferably less than 60~' of the total moles of hydrogen available within the dehydrogenat;on zone or come bastion and is therefore dependent on the conversion achieved in the dehydrogenation zone and the amount of hydrogen lost in solution or in any off-gas streams. This available hydrogen is the sum of any hydrogen recycled to the dehydrogenation Noah and the hydrogen pro-duped in all but the last bed ox dehydrogenation catalyst. Prefer-by the oxygen charged to the dehydrogenation zone is equal to about 2Q to So mole percent of the thus-defined available hydrogen. As used herein, the term "substantially all" is intended Jo indicate a major fraction ox the indicated chemical compound(s) which navy been acted upon in lie manner described, wit this major fraction prefer-by being over 90 mole percent and more preferably over 95 mole per-cent. As previously mentioned, the subject process is not limited to the production of styrenes and may be used to produce paramethyl-styrenes by dehydrogenating ethyltoluene or for the production of other unsaturated product hydrocarbons.
F LO OF THE INVENTION
The invention relates to the processing of hydrocarbons and more specifically to a process for the catalytic dehydrogenation of alkylaromatic hydrocarbons. The invention therefore relates to the synthesis of aromatic compounds, such as styrenes by dehydrogena-lion, with the dehydrogenation occurring in a number of sequential stages. The invention directly relates to an improved method for maintaining a sub atmospheric pressure in the dehydrogerlation reactor.
PRIOR ART
___ The dehydrogenation of hydrocarbons is well described in the prior art, with both cyclic and aromatic hydrocarbons being lo thereby converted to the corresponding less saturated products. For instance dehydrogenation is performed commercially for the production of styrenes from ethylben~ene to fulfill the sizable demand for this polymer precursor. The product styrenes may be polymerized with it self or it may be copolymerized with butadiene, isoprene, acryloni-trite, etc. Processes for the dehydrogenation of alkylaromatic hydrocarbons are often integrated with an alkylation process which produces the alkylaromatic hydrocarbons.
US. Patent No. 3,~15,766 issued to W. N. Root et at and US. Patent No. 3140916B9 issued to D. I. Ward are pertinent for their showing of typical prior art catalytic steam dehydrogenation processes for alkylaromatics including ethylbenzene. These referent cues describe the admixture of superheated steam into the feed hydra-carton and the admixture of additional amounts of superheated steam I;
..~,~, I
with the reactants between sequential beds of dehydrogenation gala-lust. These references also show an overall process flow into which the subject process could be integrated.
The use of a sub atmospheric pressure in the reaction zone of an alkylaromatic hydrocarbon dehydrogenation process it known as shown by the teachings of US. Patent Nos. 3,868,428 and 4,2~8t~34.
For instance, the latter of these patents teaches that the dodder-genation zone is operated at a pressure in the order of prom 2 to 25 Asia. This reference is also pertinent for its showing of the use of a compressor on the discharge side of the dehydrogenati.on zone.
However, this compressor is described as being located downstream of a cooling zone in which a liquid phase crude aromatic hydrocarbon product stream is removed by condensation. The reference therefore seems to be similar in this regard to US. Patent No. 3,702~346, in which a vacuum pump removes vapors from the phase separation vessel which receives the partially condensed dehydrogenation zone effluent stream.
The production of styrenes by the low pressure dehydrogena-lion of ethylbenzene is described in detail in an article at pages 519-528 of volume 41 of The Transactions of The American Institute at (1945).
US. Patent No. 3~978~150 is pertinent for it showing ox a paraffin dehydrogenation process in which the reactor is operated at a sub atmospheric pressure. This pressure is maintained by vacuum source which may by mechanical in nature, located in the line carry-in the reactor effluent stream. A heat exchanger is located between the reactor and thy vacuum source. This patent also indicates that a sub atmospheric pressure separation of the product hydrocarboll could Lo be perfDnmed.
BRIEF SUMMARY OF THE INVENTION
The invention provides a process for the dehydrogenation of alkylaromatic hydrocarbons in which it is more economical to achieve a desired sub atmospheric pressure in the reaction zone.
This improvement is accomplished by placing a compressing means in the effluent line of the reaction zone at a point upstream of the initial condensation of the heavier C6~ hydrocarbons and water pros-en in the dehydrogenation reaction Noah effluent stream. A broad embodiment of the invention may therefore be characterized as a pro-lo cuss for the dehydrogenation of an alkylaromatic hydrocarbon which comprises the steps of contacting a reactant stream comprising an alkylaromatic hydrocarbon with a catalyst maintained at dehydrogena-lion conditions which include a sub atmospheric pressure dud thereby forming a vapor-phase dehydrogenation zone effluent stream comprise in the alkylaromatic hydrocarbons, an unsaturated product hydrocar-bony hydrogen and steam; cooling the dehydrogenation zone effluent : stream without affecting significant condensation by indirect heat - exchange against the reactant stream; compressing the dehydrogena~ion zone effluent stream to a higher pressure which is less than one atom-sphere absolute by means of a first mechanical compressing jeans;
partially condensing the dehydrogenation zone effluent stream; sepal rating the resultant mixed phase dehydrogenation zone effluent stream in a vapor-liquid separation one maintained at a pressure less than one atmosphere absolute through the use of a second mechanical come pressing means, into a vapor-phase process stream comprising hydrogen and a liquid-phase process storyline; and recovering the unsaturated product Herr-~22~587 carbon from the liquid phase process stream.
BRIEF DESCRIPTION OF THE DRAWING
The Drawing is a schematic illustration of the reaction section of a process for the dehydrogenation of ethylbenzene to pro-dupe styrenes Referring now to the Drawing, a feed stream comprise in ethylbenzene is heated by indirect heat exchange in the feed effluent heat exchange means 2 and then continues through line 1 to be admixed with a stream of superheated steam from line 3. This ad-mixture of steam and ethylbenzene passes into the bottom of the react ion 5 through line 4. The reactants pass upward through the reactor preferably making three or more passes through different beds of de-hydrogenation catalyst maintained at dehydrogenation conditions which include a sub atmospheric pressure. The reactants may be heated with-in the reactor by means not shown.
The contacting of the ethylbenzene with the dehydrogenation catalyst contained within the reactor produces a reaction zone efflu-en stream carried by line 6 which comprises unrequited ethylbenzene, styrenes steam and hydrogen. This effluent stream is cooled in the feed effluent heat exchange means 2 and by passage through a steam generator 7, with this heat exchange extracting only sensible heat from the effluent stream and producing no significant condensation of water/ ethylbenzene or styrenes The still vapor-phase effluent stream is then compressed in the first compressor 8 to a higher pros-sure which preferably lo still less than 1 atmosphere absolute. The effluent stream continues through line 6 and is then partially con-dented in the condenser 9 to produce a mixed-phase stream. This stream is separated in the vapor-liquid separator 10 into a first liquid phase of water containing dissolved hydrocarbons, a hydrocar-bun liquid phase and a vapor phase. The vapor phase comprising hydra-gun and low temperature water vapor is removed from the separator through line 11 through the use of a second compressor 12. The hydra-carbons are withdrawn through line 13 and preferably passed to the appropriate fractional distillation facilities for the recovery of the product styrenes The condensed water is decanted from the sepal rotor through line 14.
This illustration of one possible process flow which may be utilized with the subject invention is not intended to thereby limit the scope of the invention, which may be practiced with the other process flows set out herein or in other variations not de-scribed herein.
DETAILED DESCRIPTION
Processes for the dehydrogenation of aromatic hydrocarbons are in widespread commercial use. For instance, large quantities of styrenes are produced by the dehydrogenation of ethylbenzene. The no-sultan styrenes may be polymerized with itself or it may be copolyml-prized with butadiene, isoprene~ acrylonitrile, etc. Other hydrocar-buns which may be dehydrogenated in much the same manner include do-ethylbenzene, ethyltoluene, propylbenzene and isopropylbenaene. How-ever since the great majority of the present commercial dehydrogena-lion processes are employed for the dehydrogenation of ethylbenzene, the following description ox the subject invention will be presented primarily in terms of the dehydrogenatian of ethylben~ene. This is not intended to exclude from the scope of the subject invention those other alkylaromakic hydrocarbons set out above or those having differ-en ring structures including bicyclic compounds.
The dehydrogenation reaction releases hydrogen thereby in-creasing the number of molecules present compared to the number in the feed stream. The dehydrogenation reaction is therefore favorably influenced by a decrease in pressure. Accordingly most commercial dehydrogenation processes operate at a relatively low pressure or hydrocarbon conversion processes, and as pointed out in the previous-lye cited references, it is known that the use of sub atmospheric pros-surges would be desirable. Despite this knowledge, it is believed that the present commercial operations normally employ a pressure in the dehydrogenation zone which is close to atmospheric pressure rather than significantly below it. The basic reasons for this are the ox-pens and difficulty of maintaining the dehydrogenation zone at sub-atmospheric pressures on the order of 0.7~ atmosphere absolute or be-low. One of the major causes of this difficulty in economically ox-twining the desired low pressures is the pressure drop which is in-hornet in the foe of the reactants through the reactor and the sub sequent indirect heat exchange means of the dehydrogenation zone.
More specifically, it is common practice to maintain the dehydrogena-ZOO lion reaction zone at a low pressure by withdrawing a vapor stream from the product settler (vapor-liquid separator), as by a steam jet ejector. The vacuum is "pulled" at this point since the condensation has greatly reduced the amount of vapor which remains, and it is neck essay to have the reactor effluent flow toward the product swooper-ion. In a commercial scale process, there is a very high flow rate of the weed hydrocarbon and any delineate such as steam through the total apparatus including the transfer lines and heat exchangers.
The cumulative pressure drops in these devices due to thus high flow I
rate are such that even though the products settler is at a mildly sub atmospheric pressure, the reaction zone itself will be at a pros-sure equal to or above one atmosphere absolute. This effect can be reduced by judicious equipment design but cannot be eliminated.
Therefore, it becomes necessary to maintain a co~ercially imprecate-eel vacuum in the product separator to hold the reactor at a subitem-spheric pressure in this manner.
It is therefore an objective of the subject invention to provide a process for the dehydrogenation of alkylaromatic hydrocar-buns in which the reaction zone is maintained at a sub atmospheric pressure. It is a further objective of the subject invention to pro-vise a process for the dehydroyenation of ethylbenzene or ethyltolu-one in which it is more economical to maintain the reaction zone at a pressure which is significantly below atmospheric pressure.
In the subject process a mechanical compression means is utilized at a point upstream of the products settler or Yapor-liquid separator to maintain the desire low pressure within the reaction zone. More specifically this mechanical compression means it toga-ted upstream of the indirect or direct heft exchange means utilized to condense the very great majority of the I hydrocarbons present in the reaction zone effluent stream. This compression means is preferably located downstream of any indirect heat exchange means which are utilized to recover sensible heat from the reaction zone efflu nut stream but is located upstream that is closer to the react I lion zone, of the means utilized to condense the product hydrocarbon.This location is significant since the preferred indirect heat ox-change means utilized to effect the condensation is the source of a very significant amount of the total pressure drop through the react lent flow path. The provision of a compression means at this toga-lion thereby avoids the pressure drop through this condensing means.
The compression means utilized upstream of the condensing means is not the sole means utilized to maintain the desired reduced pressure within the reaction zone. This compression means therefore operates in cooperation with a second compression means which prefer-ably is utilized to remove the vapor Stream from the vapor-liquid separator much in the same manner emp10yed in the prior art. This downstream compression means functions to maintain a low pressure on the discharge side of the upstream compression means, thereby red-cuing the compression ratio across the upstream compression means and reducing the utilities cost of operating the upstream compression means. It is therefore preferred that the pressure on the discharge side ox the first or upstream compression means is less than 1 atom-sphere absolute. It is also preferred that the vapor-liquid swooper-lion zone is maintained at a pressure less than I atmosphere absolute.
The dehydrogenation zone effluent stream is tooled by in-direct heat exchange means prior to entering the first compression means. Preferably, this heat is utilized to heat the feed hydr~car-bun entering the dehydrogenation zone and to generate steam in the Fanner illustrated in the Drawing. This initial heat exchange and cooling of the dehydrogenation zone effluent stream is to be per-formed without any significant condensation of either water or C6+
hydrocarbons present in the dehydrogenation zone effluent stream.
I As the term "significant condensation" is used herein, it is intended to indicate the condensation of more than S mole percent of any par titular compound or class of compounds. It is therefore preferred that the dehydrogenation zone effluent stream is a totally vapor-phase stream when it enters the first or upstream compression jeans.
The initial indirect heat exchange recovers the valuable high them-portray sensible heat present in the effluent stream and cools the effluent stream thereby increasing the density of this vapor-phase stream and lowering the operating temperature of the first compress soon means. It is preferred that the effluent stream is cooled at least 400 and more preferably at least 600 Fahrenheit degrees before being passed into the first compressing means. Any commercially suitably type of mechanical compression means may be employed as the first or second compression means, with the use of centrifugal come pressers being preferred.
A preferred embodiment of the invention may therefore be characterized as a process for the dehydrogenation of an alkylaroma-tic hydrocarbon which comprises the steps of contacting a reactant stream comprising an alkylaromatic hydrocarbon with a plurality of beds of dehy~rogenation catalyst maintained at dehydrogenation condo-lions which include a subatmcspheric pressure and the presence of steam and thereby forming a dehydrogenation zone effluent stream come prosing the alkylaromatic hydrocarbon, an unsaturated product hydra-carbon, hydrogen and steam; cooling the dehydrogenation zone efflu-en stream by indirect heat exchange against the reactant stream without significant condensation of water or the alkylaromatic hydra-carbon; compressing the dehydrogenation zone effluent stream in a first mechanical compressing means to a higher pressure less than 1 Z5 atmosphere absolute; partially condensing the dehydrogenation zone effluent stream by indirect heat exchange and whereby producing a mixed-phase process stream; separating the mixed-phase process stream in a vapor-liquid separation zone, maintained at a pressure Tess than l atmosphere absolute through the use of a second mechanical compress sing means, into a vapor phase process stream comprising hydrogen and a liquid-phase process stream; and recovering the unsaturated product hydrocarbon from the liquid-phase process stream, and with-drawing the vapor-phase process stream through the use of a second compressing means.
As used herein the tern "dehydrogenation zone" is intended to refer to the total reactor system which contains the dehydrGgena-lion catalyst. This catalyst may be divided into ten or more sepal rate beds, but the dehydrogenation zone preferably comprises two or three catalyst beds with means for the intermediate addition and ad-mixture of steam and possibly an oxygen supply steam. Suitable soys-terms for this may be patterned after those presented in US. Patent Nos. 3,498,755, 3,515,763 and 3,751,232. The catalyst beds may be contained in separate reaction vessels and may have either a Solon-Dracula or an annular shape. The use of radial flow catalyst beds in a stacked configuration in a single overall vows is preferred.
Different dehydrogenation catalysts may be used in different beds as described in US. Patent No. 3,223,743. Dehydrogenation catalysts generally consist of one or more metallic components selected prom Groups VI and VIII of the Periodic Table. One typical catalyst for the dehydrogenation of alkylaromatics comprises 85~ by weight ferris oxide, 2% Crimea 12% potassium hydroxide and 1% sodium hydroxide.
A second dehydrngenation catalyst, which is used commercially, con-sits of 87-90% eureka oxide, 2-3X chromium oxide and from 8-lQ~
potassium oxide. A third typical catalyst comprises 90~ by weight iron oxide I Crimea and I potassium carbonate. Methods for pro-paring suitable cdtdlysts are well known in the art. This is demon-I it striated by the teachings of US. Patent No. 3,3~7,053, which describes the manufacture of a catalytic composite of at least 35 wt.% iron ox-ire as an active catalytic agent, from about 1-8 wt.% zinc or copper oxide, about 0.5-50 wt.% of an alkali promoter, and from about 1-5 wt.% chronic oxide as a stabilizer and a binding agent.
Dehydrogenation conditions in general include a tempera-lure of about 538 to about 1000C (1000-1832F~ and preferably about 565 to about 675C (1050-1250F). When ethylbenzene is ye-in dehydrogenated, the space velocity, the rate of steam admixture and the inlet temperature are preferably adjusted to result in the effluent of each catalyst bed having a temperature of about 593C.
The temperature required for efficient operation of any specific de-hydrogenation process will depend on the feed hydrocarbon and the activity of the catalyst employed. The pressure maintained within I the dehydrogenation zone may range from about 100 to about 750 mm Hug, with a preferred range of pressures being from 250 to 7Q0 mm Hug. The operating pressure within the dehydrogenation zone is measured at the inlet, midsection, and outlet of the zone to thereby provide an approximately average pressure. The combined feed stream is charged to the dehydrogenation zone at a liquid hourly space velocity, based on liquid hydrocarbon charge at 60F, of about 0.1 to about 2.0 hurl and preferably from 0.2 to 1.0 ho 1.
The alkylaromatic hydrocarbon to be dehydrogenated is prey-drably admixed with superheated steam to counteract the temperature lowering effect of the endotherm-c dehydrogenation reaction. the presence of steam has also been described as benefiting the stabile lay of the dehydrogenation catalyst by preventing the accumulation of carbon deposits. Preferably, the steam is admixed with the other ~2~i58~
components of the feed stream at a rate of about 0.8 to about 1.7 pound of steam per pound of feed hydrocarbon. Other quantities of steam may be added after one or more subsequent beds if desired.
However, the dehydrogenation zone effluent stream should contain less than about 3 pounds of steam per pound of product hydrocarbon and preferably less than 2 pounds of steam per pound of product ho-drocarbon.
The effluent stream removed from the dehydrogenation zone is normally quickly heat exchanged for the dual purposes of lowering its temperature to prevent polymerization of the styrenes and for the recovery of heat. The effluent stream may be heat exchanged against a stream of steam, a reactant stream of this or another process or used as a heat source for fractionation, etc. Commercially the of-fluent stream is often passed through several heat exchangers there-by heating a number of different streams. This heat exchange is per-formed subject to the constraints set out above. The heat exchange performed downstream of the first compression means should cool the dehydrogenation zone effluent stream sufficiently to affect the con-sensation of at least 95 mole percent of the feed and product hydra-carbons and also at least 95-mole percent of the water vapor. The use of a quench zone to accomplish this condensation is not preferred.
Essentially all of the styrenes or other product hydrocarbon, most water and other readily condensable compounds present in the efflu-en stream are thereby converted to liquids. This produces a mixed-phase stream which us passed into a phase separation vessel. This procedure allows the facile crude separation by d contusion of the hydrocarbons from the water and hydrogen present in the effluent stream. The styrenes present in the dehydrogenat;on zone effluent stream becomes part of a hydrocarbon stream which it withdrawn from the separation vessel and transferred to the proper separation facile-ties. Preferably, the styrenes is recovered from the hydrocarbon stream by using one of the several fractionation systems known in the art. This fractionation will preferably yield a relatively pure stream of ethylbenzene, which is recycled, and an additional stream comprising Bunsen and Tulane. These two aromatic hydrocarbons are by-products of the dehydrogenation reaction. They may be recycled in part as taught in US. Patent No. 3,409,689 and British Patent 1,238,602 or entirely rejected from the process. Styrenes is recov-eyed as a third stream, which is withdrawn from the process. If de-sired, methods other than fractionation may be used to recover the styrenes For instance, US. Patent No. 3,784,620 teaches the swooper-lion of styrenes and ethylbenzene through the use of it polyamide per-I mention membrane such as nylon-6 and nylon 6,10. US. Patent No.
3,513,213 teaches a superior method employing liquid-liquid extract lion in which an hydrous silver fluoroborate it used as the solvent.
Similar separator methods utilizing cuprous fluoroborates and cuprous fluorophosphates are described in US. Patent Nos. 3,517,079; 3,517,080 and 3,517,081.
The recovery ox styrenes through the use of fractionation is described in several references including ITS Patent No. 3,~25~776.
In this reference, the hydrocarSonaceous phase removed from the phase separation zone is passed into a first column referred to as a Hanson-Tulane column. This column is operated at a sub atmospheric pressure to allow its operation at lower temperatures and hence reduce the rite of styrenes polymerization. Various inhibitors such as elemental sulfur, 2,4-d~n~tropheniol or a mixture of Nutrias diphenyl amine and a dinitroso-o-cresol are injected into the column for this same purpose. Preferably, sulfur is also introduced into this column by returning at least a portion of the high molecular weight material separated from the bottoms stream of a styrenes purification column.
A more detailed description of this is contained in Us Patent Nos.
3,476,656; 3,408~263 and 3,398,063. There is effected within the benzene-toluene column a separation of Bunsen and Tulane from the effluent to produce an overhead stream which is substantially free of styrenes and ethylbenzene. This stream preferably contains at least 95 mole percent Bunsen and Tulane. The bottoms of the bent zene-toluene column is passed into a second fractionation column from which ethylbenzene is removed as an overhead product and rely-clod. The bottoms stream of this column is then purified to obtain the styrenes The endothermic nature of the dehydrogenation reaction no-suits in a quick cooling of the reactants as they pass through the beds of dehydrogenation catalyst. This reduces the achievable con-version. For this reason the predominant commercial processes in-elude some form of interstate heating between the separate beds of dehydrogenation catalyst. This heating may be by indirect heat ox-change or by the addition of a very hot vapor which is normally superheated steam. Continuous heating by indirect heat exchange may also be employed. In more limited embodiments of the subject pro-cuss, at least a portion of the interstate heating is obtained by the catalytically promoted oxidation of hydrogen. This not only no-heats the feed hydrocarbon but also reduces the hydrogen concentra-lion in the reactant stream and thereby promotes increased confer-soon. Any heat input by hydrogen combustion reduces the heat which I
I
must be supplied by other means. In the preferred use of superheated steam, the use of partial hydrogen combustion therefore reduces the amount of steam which is required. This in turn reduces the utile-ties cost of operating the process since less superheated steam must be produced and less water vapor must be condensed for removal from the dehydrogenation zone effluent stream.
The oxygen consumed during the hydrogen combustion is prey-drably admixed into the reactant stream at the point of interstate heating as part of an oxygen supply stream. The oxygen supply stream may be air but is preferably a gas having a higher oxygen content Han air. It is preferred that the oxygen supply stream has a vitro-gun content less than 10 mole percent, with the use of substantially pure oxygen being preferred if it is economically viable. The pro-furred oxygen concentration in the oxygen supply stream is primarily lo a matter of economics and would be determined by a comparison of the advantage of having pure oxygen to the cost of obtaining the oxygen.
The basic disadvantages of the presence of nitrogen are the dilution of the hydrogen-containing gas stream removed from the product sepal ration vessel and the fact that the nitrogen passes through the dewy-drogenation Noah thereby increasing the pressure drop through the catalyst bed and the absolute pressure being maintained within the dehydrogenation zone. On the other hand, the presence of nitrogen favorably affects the equilibrium conversion level by acting as a : delineate.
I The oxidation catalyst employed in the subject process Jo pronto the ;nterstage hydrogen oxidation may be any commercially suitable catalyst which meets the required standards for stability and activity and which possesses high selectivity for the oxidation _15_ of hydrogen as compared with the oxidation of the feed or product ho-drocarbon. That is, the oxidation catalyst must have a high select tivity for the oxidation of hydrogen with only small amounts of the feed or product hydrocarbon being oxidized. The oxidation catalyst will have a different composition than the dehydrogenation catalyst.
The preferred oxidation catalyst comprises a Group VIII noble metal and a metal or metal cation which possesses a crystal ionic radius greater than 1.35 A, with both of these materials being present in small amounts on a refractory solid support. The preferred GroLIp VIII metals are platinum and palladium, but the use of ruthenium, rhodium, osmium and iridium is also contemplated. The Group VIII
metal is preferably present in an amount equal to 0.01 to 5.0 wt.%
of the finished catalyst. The metal or metal cation having a radius greater than 1.35 A is preferably chosen from Groups IA or IDA and is present in an amount equal to about 0.01 to about 20 wt.% of the finished catalyst. Iris component of the catalyst is preferably barium, but the use of other metals including rubidium or sesame is also contemplated.
The preferred solid support is alumina having a surface area between 1 and 300 mug an apparent bulk density of between about 0.2 and 1.5 g/cc, and an average pore size greater than I A.
The metal-contalning components are preferably impregnated into solid particles of the solid support by immersion in an aqueous solution followed my drying and calcination at a temperature of from about ~500 to 600C in air. The support may be in the form of spheres.
pellets or extradites. The total amount of oxidation catalyst pros-en within the dehydrogenaeion zone is preferably less than 30 White of the total amount of dehydrogenation catalyst and more preferably ~22~ 37 is between 5 and 15 White of this total amount of dehydrogenation catalyst.
The conditions utilized during the contacting of the react lent streams with the different beds ox oxidation catalyst will be set to a large extent by the previously referred' to dehydrogenation conditions. The preferred outlet temperature of any bed of oxide-lion catalyst is the preferred inlet of the immediately downstream bed of dehydrogenation catalyst. The temperature rise across any bed of oxidation catalyst is preferably less than By Centigrade de-greet. The liquid hourly space velocity, based on the liquid hydra-carbon charge at 60~F, is preferably between 2 and 10 hurl It is preferred that substantially all of the oxygen which enters a bed of oxidation catalyst is consumed within that bed of oxidation catalyst and that the effluent stream of any bed of oxidation catalyst contains less than 0.1 mole percent oxygen. The total moles of oxygen charged to the dehydrogenation zone is preferably less than 60~' of the total moles of hydrogen available within the dehydrogenat;on zone or come bastion and is therefore dependent on the conversion achieved in the dehydrogenation zone and the amount of hydrogen lost in solution or in any off-gas streams. This available hydrogen is the sum of any hydrogen recycled to the dehydrogenation Noah and the hydrogen pro-duped in all but the last bed ox dehydrogenation catalyst. Prefer-by the oxygen charged to the dehydrogenation zone is equal to about 2Q to So mole percent of the thus-defined available hydrogen. As used herein, the term "substantially all" is intended Jo indicate a major fraction ox the indicated chemical compound(s) which navy been acted upon in lie manner described, wit this major fraction prefer-by being over 90 mole percent and more preferably over 95 mole per-cent. As previously mentioned, the subject process is not limited to the production of styrenes and may be used to produce paramethyl-styrenes by dehydrogenating ethyltoluene or for the production of other unsaturated product hydrocarbons.
Claims (11)
. .
1. A process for the dehydrogenation of an alky1aromatic hydrocarbon which comprises the steps of:
(a) contacting a reactant stream comprising an alkylaroma-tic hydrocarbon with a dehydrogenation catalyst maintained at dehy-drogenation conditions which include a subatmospheric pressure and thereby forming a vapor-phase dehydrogenation zone effluent stream comprising the alkylaromatic hydrocarbon, an unsaturated product hy-drocarbon and hydrogen;
(b) cooling the dehydrogenation zone effluent stream with-out significant condensation by indirect heat exchange against the reactant stream;
(c) compressing the dehydrogenation zone effluent stream to a higher pressure less than 1 atmosphere absolute by means of a first mechanical compressing means;
(d) partially condensing the dehydrogenation zone effluent stream;
e) separating the resultant mixed-phase dehydrogenation zone effluent stream in a vapor-liquid separation zone, maintained at a pressure less than 1 atmosphere absolute through the use of a second mechanical compressing means, into a vapor-phase process stream comprising hydrogen and a liquid-phase process stream; and, (f) recovering the unsaturated product hydrocarbon from the liquid-phase process stream.
(a) contacting a reactant stream comprising an alkylaroma-tic hydrocarbon with a dehydrogenation catalyst maintained at dehy-drogenation conditions which include a subatmospheric pressure and thereby forming a vapor-phase dehydrogenation zone effluent stream comprising the alkylaromatic hydrocarbon, an unsaturated product hy-drocarbon and hydrogen;
(b) cooling the dehydrogenation zone effluent stream with-out significant condensation by indirect heat exchange against the reactant stream;
(c) compressing the dehydrogenation zone effluent stream to a higher pressure less than 1 atmosphere absolute by means of a first mechanical compressing means;
(d) partially condensing the dehydrogenation zone effluent stream;
e) separating the resultant mixed-phase dehydrogenation zone effluent stream in a vapor-liquid separation zone, maintained at a pressure less than 1 atmosphere absolute through the use of a second mechanical compressing means, into a vapor-phase process stream comprising hydrogen and a liquid-phase process stream; and, (f) recovering the unsaturated product hydrocarbon from the liquid-phase process stream.
2. The process of Claim 1 further characterized in that the dehydrogenation zone effluent stream is cooled by at least 400 Fahrenheit degrees prior to being compressed in the first compress-ing means.
3. The process of Claim 2 further characterized in that the alkylaromatic hydrocarbon is ethylbenzene and the unsaturated product hydrocarbon is styrene.
4. The process of Claim 2 further characterized in that the alkylaromatic hydrocarbon is ethyltoluene and the unsaturated product hydrocarbon is methyl styrene.
5. The process of Claim 2 further characterized in that the reactant stream comprises steam.
6. The process of Claim 5 further characterized in that the reactant stream is contacted with at least two separate beds of dehydrogenation catalyst and in that the reactant stream is heated at an intermediate point between two beds of dehydrogenation cata-lyst by the oxidation of hydrogen.
7. A process for the dehydrogenation of an alkylaromatic hydrocarbon which comprises the steps of:
a) contacting a reactant stream comprising an alkylaroma-tic hydrocarbon with a plurality of beds of dehydrogenation catalyst maintained at dehydrogenation conditions which include a subatmo-spheric pressure and the presence of steam and thereby forming a de-hydrogenation zone effluent stream comprising the alkylaromatic hy-drocarbon, an unsaturated product hydrocarbon, hydrogen and steam;
b) cooling the dehydrogenation zone effluent stream by indirect heat exchange against the reactant stream without signifi-cant condensation;
c) compressing the dehydrogenation zone effluent stream in a first mechanical compressing means to a higher pressure less than 1 atmosphere absolute;
d) partially condensing the dehydrogenation zone effluent stream by indirect heat exchange and thereby producing a mixed-phase process stream;
(e) separating the mixed-phase process stream in a vapor-liquid separation zone, maintained at a pressure less than 7 atmo-sphere absolute through the use of a second mechanical compressing means, into a vapor-phase process stream comprising hydrogen and which is withdrawn through the second compressing means and a liquid-phase process stream; and, (f) recovering the unsaturated product hydrocarbon from the liquid-phase process stream.
a) contacting a reactant stream comprising an alkylaroma-tic hydrocarbon with a plurality of beds of dehydrogenation catalyst maintained at dehydrogenation conditions which include a subatmo-spheric pressure and the presence of steam and thereby forming a de-hydrogenation zone effluent stream comprising the alkylaromatic hy-drocarbon, an unsaturated product hydrocarbon, hydrogen and steam;
b) cooling the dehydrogenation zone effluent stream by indirect heat exchange against the reactant stream without signifi-cant condensation;
c) compressing the dehydrogenation zone effluent stream in a first mechanical compressing means to a higher pressure less than 1 atmosphere absolute;
d) partially condensing the dehydrogenation zone effluent stream by indirect heat exchange and thereby producing a mixed-phase process stream;
(e) separating the mixed-phase process stream in a vapor-liquid separation zone, maintained at a pressure less than 7 atmo-sphere absolute through the use of a second mechanical compressing means, into a vapor-phase process stream comprising hydrogen and which is withdrawn through the second compressing means and a liquid-phase process stream; and, (f) recovering the unsaturated product hydrocarbon from the liquid-phase process stream.
8. The process of Claim 7 further characterized in that the reactant stream is heated at an intermediate point between sepa-rate beds of dehydrogenation catalyst by the catalytically promoted oxidation of hydrogen.
9. The process of Claim 8 further characterized in that the unsaturated product hydrocarbon is para-methylstyrene.
10. The process of Claim 8 further characterized in that the unsaturated product hydrocarbon is styrene.
11. The process of Claim 8 further characterized in that the dehydrogenation zone is cooled at least 400 Fahrenheit degrees prior to being compressed in the First compresslng means.
Priority Applications (2)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA000463924A CA1226587A (en) | 1984-09-25 | 1984-09-25 | Alkylaromatic hydrocarbon dehydrogenation process |
| IN756/DEL/84A IN162092B (en) | 1982-11-22 | 1984-09-26 |
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA000463924A CA1226587A (en) | 1984-09-25 | 1984-09-25 | Alkylaromatic hydrocarbon dehydrogenation process |
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| Publication Number | Publication Date |
|---|---|
| CA1226587A true CA1226587A (en) | 1987-09-08 |
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ID=4128773
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA000463924A Expired CA1226587A (en) | 1982-11-22 | 1984-09-25 | Alkylaromatic hydrocarbon dehydrogenation process |
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| Country | Link |
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| CA (1) | CA1226587A (en) |
-
1984
- 1984-09-25 CA CA000463924A patent/CA1226587A/en not_active Expired
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