US3607960A - Thermal hydrodealkylation process - Google Patents

Abstract

In a process for the thermal hydrodealkylation of an alkyl aromatic, such as toluene, to produce benzene, a mixture of diphenyls and high-boiling polymers is separated as bottoms from the effluent of a primary hydrodealkylation zone. This mixture is mixed with recycle hydrogen and flash vaporized, the vapor comprising polyphenyls and hydrogen being passed to an auxiliary reactor wherein the polyphenyls are converted to benzene. The flash vaporization selectively rejects substantially all of the high-boiling polymer compounds which would otherwise tend to form coke in the auxiliary reactor.

Description

United States Patent 3,152,980 10/1964 Coonradtetal....

Delbert L. Button Allison Park, Pa.;

Lawrence J. Kirby, Tokyo, Japan June 18, 1970 Sept. 21, 1971 Gull Research & Development Company Pittsburgh, Pa.

Continuation-impart of application Ser. No. 723,998, Apr. 25, 1968, now abandoned.

Inventors Appl. No. Filed Patented Assignee THERMAL HYDRODEALKYLATION PROCESS 8 Claims, 1 Drawing Fig.

[1.8. CI ..260/672 NC,

203/88, 208/48 R, 208/78, 208/99, 208/102, 260/668 F, 260/672 R, 260/674 R Int. Cl B0ld 3/06, C07c 3/58, ClOq 9/16 Field of Search 260/672 NC References Cited UNITED STATES PATENTS MAKE'UP HYDROGEN POLYMER PURGE 3,213,150 10/1965 Cabbage 260/667 3,288,875 11/1966 Payne et al 260/672 3,291,849 12/1966 King et al 260/672 3,296,120 l/1967 Doelp et al 208/143 3,296,323 1/1967 Myers et a1... 260/672 3,390,200 6/1968 Sze 260/672 OTHER REFERENCES Fowle & Pitts Thermal Hydrodealkylation Chem. Eng.

Progress 58 (4) 37- 40 (April 1962) Primary Examiner-Delbert E. Gantz Assistant Examiner-G. E. Schmitkons Attorneys-Meyer Neishloss, Deane E. Keith and Thomas G..

Ryder ABSTRACT: In a process for the thermal hydrodealkylation of an alkyl aromatic, such as toluene, to produce benzene, a mixture ofdiphenyls and high-boiling polymers is separated as bottoms from the effluent of a primary hydrodealkylation zone. This mixture is mixed with recycle hydrogen and flash vaporized, the vapor comprising polyphenyls and hydrogen being passed to an auxiliary reactor wherein the polyphenyls are converted to benzene. The flash vaporization selectively rejects substantially all of the high-boiling polymer compounds which would otherwise tend to form coke in the auxiliary reactor.

VAPOR STREAM HYDROGEN PLANT FUEL GAS 73 PURGE BENZENE PRODUCT THERMAL HYDRODEALKYLATION PROCESS RELATED APPLICATIONS This application is a continuation-in-part of Ser. No. 723,998, filed on Apr. 25, 1968 and now abandoned and is related to Ser. No. 596,125, filed Nov. 22, 1966 and now abandoned and to Ser. No. 816,145, filed Mar. 14, 1969.

This invention relates to a process for the thermal hydrodealkylation of alkyl aromatic compounds. More-particularly, this invention relates to such a process wherein higher benzene selectivity and increased benzene yields are obtained.

Toluene can be dealkylated to benzene by subjecting it in the presence of hydrogen to an elevated temperature and elevated pressure for a controlled length of time. As a result of such reaction conditions, the methyl group is cleaved from the toluene and replaced by hydrogen. The mechanism probably involves the generation of methyl andphenyl radicals and the combination of these radicals with-hydrogento formmethane and benzene. The overall hydrodealkylation reaction is highly exothermic.

The yield of benzene in the thermal hydrodealkylation of an alkyl aromatic may be increasedsomewhat by conducting the thermal hydrodealkylation of the alkyl aromatic in the presence of diphenyl. However, the presence of diphenyl during the hydrodealkylation of an alkyl aromatic such as toluene depresses the dealkylation reaction rate of'toluene to benzene. Moreover, in order to obtain a high selectivity to benzene, e.g. about 98 percent, when thermal hydrodealkylating a mixture of toluene and diphenyl (selectivity to benzene being defined as the ratio of the amount of benzene which is actually obtained to that which theoretically could be obtained if all of the alkyl aromatic which has reacted were converted to benzene), the process must be operated at not more than about 75 percent conversion per pass. This requires that a relatively high amount of the unconverted reactant material be recycled to the reactor for conversion to benzene. The necessity for recycling a large amount'of reactant material at a given temperature requires athermal hydrodealkylation unit and recovery apparatus of increasedsize for a given benzene product capacity, and also requires the recycle of extra hydrogen commensurate with the total amount of fresh and recycle aromatics used.

It is an object of this invention to provide an improved thermal hydrodealkylation process.

it is another object of this invention to provide a thermal hydrodealkylation process for alkyl aromatics which results in increased yields of benzene without a decrease in the reaction rate of the alkyl aromatic to benzene.

It is a further object of this invention to provide a thermal hydrodealkylation process for alkyl aromatics in which the main dealkylation reactors can be operated-at conversions per pass of 80 percent or higher and still attain a high benzene selectivity for the overall process.

It is still another object to provide a means for'separating polyphenyls from high boiling compounds which have been separated from the effluent from a primary hydrodealkylation rea ti n zone earlier in the operatingsequence and'then introducing the polyphenyls to an auxiliary reaction zone for conversion to benzene.

These and other objects are attained by the practice of this invention which, briefly, comprises subjecting a gaseous mixture comprising at least one alkyl aromatic and hydrogen in a primary reaction zone to reaction temperatures in the range of I from about1000 to 1800 F. An effluent comprising unreacted hydrogen and alkyl aromatic, benzene product, polyphenyls (primarily diphenyl) and higher boiling polymers is removed from the primary reaction zone. This effluent is then passed to a separation zone in which a mixture comprising unreacted alkyl aromatic, polyphenyls and higher boiling polymers are removed. The mixture is mixed with hydrogencontaining gas and then flash, vaporized to separate as a vapor constituents having a boiling point below about 750 to800 F. at atmospheric pressure (e.g. hydrogen, alkyl aromatics and polyphenyls) from the liquid higher boiling polymers boiling atatmospheric pressure at about 750 to 800 F. and above. The vaporous constituents are then subjected to temperatures in the range of from about 900 to 1500 F. in an auxiliary reaction zone to convert at least a portion of the polyphenyls to benzene. The flash vaporization of the feed to the auxiliary reactor pennits the selective rejection of high-boiling polymer compounds which cause coke formation in the preheater for the auxiliary reaction zone. The benzene product from the effluents from the primary reaction zone and from the auxiliary reaction zone is recovered. Ina preferred embodiment, the effluent from the auxiliary reaction zone is combined with the effluent from the primary reaction zone before the effluent from the primary reaction zone is passed to the separation zone and the benzene product is recovered from the separation zone. Thus, the product separation equipment normally used for the separation of the products and unconverted reactants from the primary reaction zone is also used to separate the products and unconverted reactants obtained from the auxiliary reaction zone.

This use of the auxiliary reaction zone provides an integrated process'for converting to benzene a portion of the diphenyl content contained inthe polymer or polyphenyl bottoms formed during the thermal hydrodealkylation reaction, thus increasing the total yield of benzene and overall process selectivity to benzene; This increase in the total yield of benzene is achieved without depressing the reaction rate of alkyl aromatic to benzene in the main reaction zone. Thus, the process of this invention-will provide favorably high benzene selectivity when the primary reaction zone is operated at a minimum conversion level of about 50 percent. Moreover, in the process of this invention, abenzene selectivity of about 98 percent may be obtained when operating the primary reaction zone at a conversion in excess of percent, normally between and percent conversion per pass.

The preferred embodiment of this invention will be further illustrated with reference to the accompanying drawing. In the interest of simplification of the drawing presented herein, numerous valves, pumps, and other related pieces of process equipment have been omittedfrom the figure. However, it is to be understood that the addition of these omitted items may be accomplished without changing the nature and scope of this invention.

Referring to the. drawing, the alkyl aromatic feed stock is fed into the system by way of line 11. The alkyl aromatic can be, for example, toluene, m-xylene, o-xylene, p-xylene, mixed xylenes, ethylbenzene, propylbenzene, butylbenzene and other C and C alkylbenzenes and mixtures of any of these. The feed stock can also contain up to 10 percent by weight of heavy paraffins-containingfrom six to 12 carbon atoms, as described in US. Pat. No. 3,363,019.

Line 11 is provided with pump 12, for compressing the feed to an elevated pressure. The compressed feed is then passed by line 13 to a heat exchanger 14 in which it is indirectly heated with hot product effluent obtained as hereinafter described.

Makeup hydrogen-containing gas at an elevated pressure is introducedthrough line 15. Hydrogen can be added to the feed at any point before the feed stream reaches process heater 22 but preferably is added before process temperatures of 600 F. are obtained. A portion of this hydrogen-containing.

gas can be passed through line 16 and combined with the feed to the auxiliary reactor-94 hereinafter described. The remaining portion of the hydrogen-containing makeup gas is combined with the allcyl aromatic feed in line 21.

Hydrogen-containing recycle gas is introduced into line 15 from line 17 through line 18. The makeup hydrogen gas stream, as well as the hydrogen-containing recycle gas stream, need not be pure hydrogen. These streams can contain between about 40 and percent hydrogen by volume. Preferably, the makeup hydrogen gas stream passed through line 16 contains 90 percent or more hydrogen; and the combined makeup hydrogen and hydrogen-containing recycle gas which is combined with the alkyl aromatic feed in line 21 contains from 45 to 95 percent by volume of hydrogen.

The reactant feed stream comprising the alkyl aromatic and hydrogen can contain a hydrogen to aromatic hydrocarbon mol ratio within the range of from about 1.5 to 20.0 and, preferably, from about 3 to 10. the reactant feed stream is heated in the heat exchanger 14. The reactant stream is then passed through line 19 to a second indirect heat exchanger 20 wherein the reactant stream is further indirectly heated with reaction effluent. The preheated reactant stream is passed from the exchanger 20 by line 21 wherein it is combined with makeup hydrogen from line 15 as hereinabove described, and then is passed to a heater or furnace 22 wherein final heating of the reactant feed stream up to the reaction temperature is accomplished. The reactor feed stream heated to reaction temperature in the heater or furnace 22 is then passed by line 23 to the first reactor 24. An effluent is recovered through the top of the reactor 24 by line 25. This effluent optionally can be quenched to a lower temperature by direct mixing with a cool hydrogen-containing recycle stream obtained from a highpressure flash drum more fully described hereinafter and introduced by line 26. The effluent in line 25 is thereafter introduced into the bottom of a second reactor 27.

The thermal hydrodealkylation reaction which occurs in reactors 24 and 27 is conducted at temperatures in the range of from about 1000 to 1800 F. and a pressure of from about 100 to 1000 p.s.i.g. with a contact time or residence time of the reactants in the reactor of from about to 600 seconds. In a preferred embodiment of this invention, the reaction is conducted at temperatures in the range of from about 1050 to 1400 F. and a pressure of from about 400 to 600 p.s.i.g. for from about 10 to 100 seconds.

An effluent is recovered from the top of the reactor 27 by line 28. The effluent in line 28 then optionally can be quenched to a temperature below the reaction temperature by direct mixing with a portion of the cool recycle stream obtained from the high pressure flash drum hereinabove mentioned and introduced by a line 29, as described in U.S. Pat. No. 3,188,359. Alternatively, the effluent in line 28 can be quenched by direct mixing with a portion of the liquid phase obtained from the high pressure flash drum. Cooling also can be accomplished by insertion of a steam generated in line 28 or in heat exchangers 20, 30 and 31, as more fully described elsewhere, or by a combination of this procedure together with direct quenching as described above.

As previously mentioned, the quenching of the effluent from reactors 24 and 27 is optional. Therefore, lines 26 and 29 can be omitted if a quench is not employed at these points. Moreover, quenching at these points also can be accomplished by means of liquids or gases other than the hydrogencontaining recycle stream illustrated.

A portion of the effluent in line 28 is passed by line 29 to a reboiler 30 associated with the bottom of a product stripper tower more fully discussed hereinafter, to provide the heat duty of the stripper tower. In reboiler 30, the effluent gives up a portion of its heat by indirect heat exchange with a liquid stream withdrawn from the lower portion of the stripper tower.

The remaining effluent in line 28 is passed to a reboiler 31 associated with a flash drum, more fully described hereinafter, to provide the heat duty for the flash drum. The effluent is removed from reboiler 31 by line 32 and is combined with the effluent removed from reboiler 30 by line 33.

The effluent in line 32 is combined with the effluent in line 96 from the auxiliary reactor (more fully described hereinafter) and the combined effluents are passed by line 34 to indirect heat exchanger 20, wherein they give up a portion of their heat to preheat the feed in line 19. From indirect heat exchanger 20, the stream is passed to a hydrogenation chamber 35 wherein it is subjected to mild hydrogenation conditions as described in U.S. Pat. Nos. 3,310,593 and 3,310,594. Any additional hydrogen required to satisfy the hydrogenation requirements can be supplied in the form of makeup gas or recycle gas.

In chamber 35, any materials containing aliphatic unsaturation in the liquid product stream are hydrogenated to saturated products, thereby facilitating subsequent fractionation. The thus-treated liquid product stream is then passed by line 36 to heat exchanger 14 wherein the effluent gives up additional heat to the reactant feed stream in line 13, thereby being cooled. Accordingly, the hot effluent recovered from reactor 27 supplies the heat duty of the flash tower and the product stripper in addition to supplying the major portion of the heat to bring the reactant feed stream up to reaction temperature.

The hot effluent can then be passed from the exchanger 14 by line 37 to a suitable cooler 38. Cooler 38 can be any suitable arrangement of coolers comprising a water cooler, air cooler, or a combination thereof which will sufficiently cool the effluent for passage by line 39 to a high-pressure flash drum 40 maintained at a pressure of about 400 p.s.i.g. and a temperature of about F.

In high-pressure flash drum 40, a vaporous stream comprising hydrogen, methane and small amounts of entrained benzene product is separated from a major benzene liquid product stream. The vaporous stream is removed from drum 40 by line 41 and separated into two streams with the major portion thereof being passed by line 42 to a recycle compressor 43 and the minor portion of the stream being passed for further treatment by line 44 as, discussed hereinafter.

The recycle gas stream is compressed in compressor 43 to an elevated pressure suitable for recycle to the reactors thereby raising the temperature of this stream. The thus-compressed recycle stream is passed by line 17 optionally to line 29 and/or line 26 for use as quench material in the reactor effluent streams as discussed above; and to line 45 for use as quench material in the effluent stream for the auxiliary reactor as hereinafter described. Another portion of this recycle stream can be passed by line 18 to line 15 wherein it can be combined with hydrogen-rich makeup gas. The combined gas stream is thereafter combined with the hydrocarbon feed to be dealkylated prior to the heat exchange steps hereinabove discussed. Optionally, another portion of this recycle stream is passed by line 17 to line 97 and subsequently into lines 16, 85, 88 and 89 and introduced into the auxiliary reactor 94 as an additional or alternative source of hydrogen for this reactor.

Optionally, a vaporous stream of minor portion in line 44 recovered from the high pressure flash drum can be further treated to obtain maximum recovery of entrained benzene product material. To accomplish this end, the vaporous stream in line 44 is passed to an indirect heat exchanger 46 wherein it is cooled to a temperature of about 65 F. by indirect heat exchange with refrigeration flash vapors obtained as hereinafter described. The vaporous stream cooled in the indirect heat exchanger 46 is then passed by line 47 through a refrigeration cooler 48 to further cool the vaporous stream to a temperature of about 40 F. The thus-cooled vaporous stream is then passed by line 49 to a separator 50 maintained at a temperature of about 40 F. and a pressure of about 380 p.s.i.g.

In separator drum 50, a vapor stream, referred to herein as refrigeration flash vapors, is separated and recovered from a liquid benzene stream. The refrigeration flash vapors of reduced temperature are passed by line 51 to the heat exchanger 46 to precoo] the vaporous stream in line 44 as described above. The refrigeration flash vapors are recovered from heat exchanger 46 by line52 and passed to a hydrogen plant, not shown, for manufacturing fresh hydrogen.

The liquid benzene stream separated in drum 50 is withdrawn and passed by line 53 to line 54 wherein it is combined with the liquid stream recovered from the high-pressure separator drum 40. The thus-combined stream is then passed to the upper portion of a product stripper tower 55.

Stripper tower 55 is maintained at a temperature in the range of from about 1 F. to about 450 F. and a pressure in the range of from about 290 p.s.i.g. to about 400 p.s.i.g. with heat being supplied to the lower portion of the stripper tower by passing a liquid stream withdrawn from the lower portion thereof by line 56 to heat exchanger and thereafter returning the heated withdrawn stream to the tower by line 57 to supply the heat duty of the stripper tower.

ln stripper tower 55 a vaporous stream containing a small amount of benzene is recovered from the liquid product introduced thereto by line 54 and removed from the upper portion of the tower by line 58. The vaporous stream in line 58 can be passed through refrigeration drum 48 to cool this stream to about F. from whence it is withdrawn and passed by line 59 to separator drum 60 maintained at a temperature of about 40 F. and a pressure of about 290 p.s.i.g. ln separator drum 60, a vapor stream is separated from a liquid stream comprising benzene, the vaporous stream is removed therefrom by line 61 and the liquid stream is removed.

therefrom by line 62.

A stripped liquid product stream comprising benzene is recovered from the bottom of the stripper tower by line 63 and is passed through line 63 to line 64 and then into fractionator 65. The liquid stream in line 62 recovered from separator drum is also connected to line 64 in order that this recovered liquid material may be passed. to fractionator 65.

Fractionator tower 65 is designed to withdraw a benzene product stream from the upper portion thereof by line 66 which is provided with cooler 67 for cooling the benzene product stream to a temperature of about 100 F. To assure recovery of a high purity benzene product stream from the fractionator, the benzene stream iswithdrawn from the fractionator at about the fifth tray and any lower boiling materials are withdrawn from the top of the towerby line 68, cooled in cooler 69 to a temperature of about 180 F. and then passed to separator 70. All or a portion of this material is employed as a cool reflux stream and is withdrawn from separator 70 and returned to the top portion of the fractionator above the point of withdrawal of benzene product material by line 71. Line 72, which is connected to line 70, is provided for withdrawing any excess reflux material from the fractionation system. Any lighter than benzene material is withdrawn from separator 70 by line 73.

A liquid stream comprising unconverted alkyl aromatics, polyphenyl-type aromatics, and higher-boiling, condensed aromatics which are chiefly polymer bottoms is removed from the bottom of fractionator 65 by line 74. The polyphenyls include diphenyl, methyl diphenyl, C diphenyls and triphenyls. In order to supply the heat duty of the fractionator, which is maintained at a temperature in the range of from about 210 F. to about 425 F a portion of the stream 74 is passed by line 75 through heat exchanger 76 and then returned to the fractionator 65 by line 77. Heat for the heat exchanger 76 is supplied by introducing steam through line 78. Alternatively, heat can be supplied to heat exchanger 77 by indirect heat exchange with the effluent from hydrogenation reactor 27.

The remainder of the liquid stream from line 74 is passed by line 79 to flash drum 80 which is operated at a pressure of about 15 p.s.i.g. and a temperature of about 360 F. The heat duty for the flash drum 80 is supplied by passing a liquid stream withdrawn from the lower portion thereof by line 81 to heat exchanger 31 and thereafter returning the heated withdrawn stream by line 82 to the flash drum 80.

In the flash drum 80, unconverted alkyl aromatics are removed overhead by line 83, cooled to about 100 F. and recycled to the feed in line 11. A portion of the entire part of the cooled liquid stream from line 83 can be employed as additional feed to the auxiliary reactor 94 by passing it through line 84 into line 85. A stream comprising polyphenyls and high-boiling materials is removed from the flash drum 80 by line 86 and is passed to line 85.

if the conversion level in the system is sufficiently high, i.e., above about 94 percent, recovery of unconverted alkyl aromatics for recycle to the thermal hydrodealkylation reactors may not be justified. In that instance, the flash drum 80, reboiler 31 and lines 81, 82, 83 and 86 can be eliminated and the liquid stream withdrawn from the bottom of fractionator 65 by line 79 can be passed directly to line 85.

Makeup hydrogen is introduced into line through line 16 as hereinbefore described. The gaseous stream can be passed to heat exchanger 87 whereinit is indirectly heated with hot effluent from the auxiliary reactor as hereinafter described. Thereafter, the stream is passed by line 88 to furnace 22 wherein it is further heated. The stream is then passed from furnace 22 by line 89 to flash vaporization unit 90,

The flash vaporization unit 90 is operated at a temperature of from about 450 to 750 F. and a pressure of about 300 to 800 p.s.i.g. in a preferred embodiment of the invention, this unit is operated at a temperature within the range of 500 to 650 F. and a pressureof 400 to 600 p.s.i.g. The hydrogen to aromatic hydrocarbon mol ratio of the feed to the flash vaporization unit may be within the range of 1.5 to 20 and, preferably, from about 4 to 10. i

High-boiling hydrocarbons, preferably those boiling about 750 F. at atmospheric pressure are withdrawn as a'liquid from the flash vaporization unit 90 by line 91 and are purged from the system. An all vapor feed is withdrawn from the top of flash vaporization unit 90 and is passed by line 92 to furnace 22 wherein it is further heated. The stream is then passed from furnace 22 by line 93 to auxiliary reactor 94 wherein at least a portion of the polyphenyls is converted to benzene. The overall reaction which takes place in the auxiliaryreactor as exemplified by C diphenyl is represented bythe following equation:

The reaction in the auxiliary reactor 94 can be conducted thermally at temperatures'in the range of from about 900 to 1500 F. and preferably, between l000 to 1350 F., at pressure of from about 300 to 800 p.s.i.g. and, preferably, between 400 and 600 p.s.i.g. with a contact time or residence time of from about 10 to 200 seconds and, preferably, from 20 to l20 seconds. The hydrogen to aromatic hydrocarbon mol ratio in the auxiliary reactor may be from 1.5 to 20 and, preferably, from about 4 to l0. While it is not necessary to use a catalyst in the auxiliary reactor, nonhydrogenation-dehydrogenation catalysts, such as natural clays, aluminas, silica-aluminas and sulfided metal (e.g., Ni) supported catalysts can be used. When a catalyst is employed in this auxiliary reactor, relatively less severe reaction conditions than those indicated above for thermal conversion can be used. The auxiliary reactor 94 is preferably sized to provide a sufficient reactant holding time to obtain a maximum temperature of about 1300 F. at about 400 p.s.i.g. The auxiliary reactor is advantageously operated at a lower temperature and so as to permit a longer contact time than is the case with the main reactors, since the selectivity to benzene for a feed stock containing chiefly diphenyls is improved at lower temperatures than those favoring the selectivity to benzene of a feed stock containing chiefly toluene or similar alkylbenzenes.

An effluent. is recovered from auxiliary reactor 94 by line 95. The effluent in line 95 is quenched with recycle gas from line 45. The quenched effluent is then passed to indirect heat exchanger 87, wherein it gives up a portion of its heat to pre heat the stream in line 85. The effluent is passed from heat exchanger 87 through line 96 and is combined with the product effluent stream from primary reactors 24 and 27.

It is preferred that the hydrogen-containing gas introduced through line 16 which is combined with the polyphenyls in line 85 be of relatively high purity, e.g., 80 percent or more by volume of hydrogen. Therefore, makeup. hydrogen gas is preferred instead of recyclehydrogen-containing gas which can contain only from 50 to- 80 percent by volume of hydrogen. However, a mixture of recycle gas and makeup hydrogen introduced through lines 15, 17, 97 and 16 can be used, and if necessary, recycle gas alone introduced through lines 17, 97 and 16 can be employed. The use of high purity hydrogen at this stage of the process tends to increase the selectivity which is obtained in the auxiliary reactor 94. Moreover, the use of high-purity hydrogen has the further advantage that it reduces the total volume of reactant gas in the auxiliary reactor which, in turn, reduces the required size of such auxiliary reactor. Furthermore, when high-purity gas is used in the auxiliary reactor 94, a relatively greater amount of hydrogen gas is recovered from this reactor in the product recovery system and it subsequently enriches the recycle hydrogen stream which is recycled to primary reactors 24 and 27 and thereby reduces the normal requirement of makeup hydrogen supplied to these reactors.

As a safety factor in the refrigeration section of the process herein described, provision is made for introducing a portion of the alkyl aromatic feed, when necessary, to the vaporous streams in lines 44 and 58 by way of lines 98 and 99 to avoid freezing of any benzene material, cooled in refrigeration exchanger 48.

The use of the auxiliary reactor as described herein results in 3 to 4 percent higher overall yields of benzene being obtained than in a conventional thermal hydrodealkylation process wherein only fresh and unconverted alkyl benzenes are charged to a main reaction zone. Moreover, there is obtained a high benzene selectivity even at high conversions per pass through the main reactors. Therefore, the main reaction zones are normally operated at conversions of from 85 to 95 percent per pass and the additional reaction zone is normally operated at conversions of from 60 to 90 percent per pass. These advantages are obtainable without the additional cost or complexity of adding any extraneous material, without increased hydrogen consumption that is unproductive of additional benzene product, and without formation of products requiring different separation facilities. Furthermore, the auxiliary reactor does not influence the performance of the primary reactors and, if it must be shut down for maintenance, this can be done without affecting the operation of the primary reactors.

The practice of this invention, i.e., the flash vaporization of the charge to the auxiliary reactor, selectively rejects substantially all of the high-boiling coke-forming polymer compounds and provides an all vapor feed to the furnace which heats the feed to the auxiliary reactor. This, coke formation in the furnace is minimized resulting in fewer shutdowns of the auxiliary reactor.

When the flash vaporization unit 90 is omitted and the feed containing high-boiling materials is fed directly to the furnace for the auxiliary reactor, coking of the preheater coils may result. This coking is believed to be caused by the buildup of the higher boiling point polymer compounds in the charge to the auxiliary reactor. The buildup of these compounds raises the dew point of the charge to the auxiliary reactor above the temperature at which liquid hydrocarbons form coke, i.e., about 650 to about 700 F. While this elevation of dew-point may be partially avoided by rejecting an aliquot portion of the polymer bottoms continuously bled from the fractionator 65, even a 15 percent mol rejection of polymer yields a charge to the auxiliary reactor of about 694 F. dew point.

For example, when mol percent of polymer bottoms is rejected from the fractionator 65 and the flash vaporization unit 90 is omitted, the charge to the auxiliary reactor has a dew point of about 715 F. However, when no polymer is rejected from the fractionator 65 and the flash vaporization unit 90 is operated to reject 10 mol percent of the polymer, the charge to the auxiliary reactor has a dew point of only about 595 F. At a 5 mol percent polymer rejection rate from the flash vaporization unit 90, the dew point is only raised to about 610 F. Thus, lees polymer need be rejected to minimize the risk of coking in the preheater for the auxiliary reactor. The dew points discussed above were measured at 460 p.s.i.g.

Furthermore, the use of the flash vaporization unit results in a better selectivity-conversion relationship at a given polymer rejection rate. Thus, when the flash vaporization unit is omitted and 10 percent of the polymer bottoms are rejected from the bottom of the fractionator 65, and the system is operated at 85 percent toluene conversion, there is obtained an overall selectivity to benzene of about 97.9 percent. By contrast, when the flash vaporization technique of this invention is employed and the system is operated at 5 percent polymer rejection rate and at 85 percent toluene conversion, the overall selectivity is about 98.3 percent; and at percent toluene conversion, the selectivity is about 98 percent. It will be seen, therefore, that the use of the flash vaporization technique of this invention and the selective rejection of heavy ends not only insures a noncoking, all vapor feed to the auxiliary reactor but also reduces the loss of benzene precursors by permitting the unit to operate at a low rejection rate.

The present invention, therefore, rejects preferentially only the heavy ends of the polymer, insures an all vapor feed to the preheater coils of the auxiliary reactor, avoids formation of coke in the preheater for the auxiliary reactor to allow for longer operating periods without shutdown, reduces the amount of polymer that must be rejected and provides higher selectively.

The following example illustrates the practice of the invention.

EXAMPLE 1 This example illustrates a continuous process in which all amounts are expressed in mols per hour of material. A feed comprising 2161.9 mols of hydrogen, 1083.8 mols of methane, 72.0 mols of C H 8.1 mols of C 36.9 mols of benzene and 501.6 mols of toluene (of which 441.5 mols of toluene are fresh feed and the remainder are recycle) is preheated in a furnace to a temperature of about 1215" F. at about 460 p.s.i.g. and is charged to a first primary reactor where the temperature increases to about 1285 F. The first primary reactor effluent is fed to the second primary reactor where the temperature rises to'about 1340 F. The nominal residence time of the feed in the primary and secondary reactors is about 56 seconds. The'effluent from the secondary reactor comprises 1663.0mols of hydrogen, 1600.0 mols of methane, 59.4 mols of C H 439.8 mols of benzene 75.2 mols of toluene, and 9.4 mols of polymer, giving a net selectivity 94.5 percent for this primary reactor system. After quenching with recycle gas, the effluent is passed through a series of heat exchangers wherein the temperature is lowered to about 500 F. before entering the hydrogen treater for product purification. The treated effluent is further cooled to about 100 F. and then flashed in a high-pressure separator. The liquid is charged to a stripper and the gas returns in part to a recycle gas compressor, and in part is rejected as fuel. The liquid is stabilized in the stripper and charged to a fractionator operated at a temperature of 308 F. and a pressure of 15 p.s.i.g. wherein it is separated into two fractions. One fraction comprises substantially pure benzene in an amount of 402.9 mols. The second fraction is withdrawn from the bottom of the fractionator and comprises 1.9 mols of benzene, 75.2 mols of toluene and 9.4 mols of polymer. This fractionator bottom stream is mixed with hydrogen and charged to an auxiliary reaction system, which operates at conditions resulting in a 75 percent conversion of diphenyl compounds, a 15 percent conversion of fluorene-type compounds, and a negligible conversion of condensed aromatic compounds such as antracene.

The unconverted polymer is recycled by mixing it with the stream containing polymer recovered as fractionator bottoms from the primary reaction system and charging the mixture to the auxiliary reaction system.

Upon the attainment of a steady state composition, the gross charge comprises 1.9 mols of benzene, 75.6 mols of toluene, and 27.1 mols of polymer. The gross charge is passed to a flash drum which is operated at a temperature of 360 F.

and a pressure of 15 p.s.i.g. A vaporous stream is removed from the flash separator comprising 1.5 mols of benzene and 57.5 mols of toluene. This stream is cooled and recycled to the feed to the primary reactor. A liquid stream is removed from the bottom of the flash separator comprising 0.4 mols of benzene, 18.1 mols of toluene and 27.1 mols of polymer, the polymer containing naphthalene, C diphenyl, C fluorene, C diphenyl, C fluorene, C antracene-phenanthrene, and C, aromatics including pyrene, chrysene, and other condensed-ring aromatics. This bottoms stream. is mixed with recycle hydrogen comprising 270.7 mols of hydrogen and 30.1 mols of methane and the mixture is charged to a flash drum heated to a temperature of 610 F. and maintained at a pressure of 460 p.s.i.g. A liquid polymer is purged from the bottom of this separator comprising 0.1 mols of toluene and 1.36 mols of high-boiling polymer. A vapor feed is removed from the top of the flash drum comprising 270.7 mols of hydrogen, 30.1 mols of methane, 0.4 mol of benzene, 18.0 mols of toluene and 25.7 mols of condensed aromatics and polyphenyls. This vapor feed is then preheated to about 1230 F. and is charged to an auxiliary reactor wherein the nominal residence time or holding time is about 1 10 seconds to an average reactor temperature of 1250 F. and is charged to an auxiliary reactor wherein the nominal residence time or holding time is about 110 seconds to an average reactor temperature of 1250 F. and a pressure of about 460 p.s.i.g. The effluent from this reactor comprises 238.4 mols of hydrogen, 47.4 mols of methane, 31.8 mols of benzene, 2.7 mols of toluene and 17.7 mols of aromatic polymer. This effluent is quenched, cooled to 500 F., passed through a hydrogen treater for product purification, further cooled to 100 F., and then charged to a high-pressureseparator. The liquid is charged to a stripper and the gas returns in part to a recycle compressor and in part is rejected fuel. The liquid is stabilized in the stripper and charged to a fractionatorwherein it is separated into two fractions. One fraction comprises substantially pure benzene in the amount of 31.1 mols. The second fraction, which is withdrawn from the bottom of the fractionator. comprises a trace of benzene, 0.4 mol of toluene and i7.7 mols of aromatic polymer. This bottoms product is recycled by combining it with the bottoms product from the primary reaction system to produce a mixture comprising 1.9 mols of benzene, 75.6 mols of toluene, and 27.1 mols of polymer, which is charged to the auxiliary reaction system. The selectivity to benzene without the auxiliary reactor is about 94.5 mol percent. The overall selectivity to benzene for the combined reaction (Le. for the reactions which take place in the primary reactors and the reactions which take place in the auxiliary reactors) based on the fresh feed to the entire unit is about 98.3 mol percent.

Obviously, many modifications and variations of the invention as hereinabove set forth can be made without departing from the spirit and scope thereof.

What we claim is:

1. In a process for the thermal hydrodealkylation of an alkyl aromatic to produce benzene which comprises:

a. subjecting a gaseous mixture comprising at least one alkyl aromatic and hydrogen in a primary reaction zone to temperatures in the range of from about 1000 to 1800 F. to effect a minimum conversion level of at least about 50 percent;

b. recovering an effluent comprising unreacted hydrogen zone; the improvement which comprises: l. mixing the bottoms mixture from the separation zone with hydrogen-containing gas and continuously flash vaporizing said mixture to separate as a vapor constituents having a normal boiling point below about 750 F. and rejecting liquid higher boiling polymers; and

I1. subjecting said vapor in an auxiliary reaction zone to temperatures in the range of from about 900 to 1500 F. whereby at least a portion of said diphenyls is converted to benzene.

2. The process of claim 1 wherein the effluent from said auxiliary reaction zone is combined with the efi'luent from said primary reaction zone before said effluent from said primary reaction zone is passed to said separation zone, and the benzene product is thereafter recovered from the combined effluents.

3. The process of claim 2 wherein the reaction in said primary reaction zone is conducted at a pressure of from about to 1000 p.s.i.g. for from about 10 to 600 seconds and the hydrogen to aromatic hydrocarbon mol ratio is within the range of from about 1.5 to 20.0; said flash vaporizationis conducted at a pressure of from about 300 to 800 p.s.i.g. and temperatures in the range of from about 450 to 750 F.; and wherein the reaction in said auxiliary reaction zone is conducted at a pressure of from about 300 to 800 p.s.i.g. for from about 10 to 200 seconds and the hydrogen to hydrocarbon mol ratio is within the range of from about 1.5 to 20.0.

4. The process of claim 2 wherein the reaction in said primary reaction zone is conducted at temperatures in the range of from about 1050 to 1400 F. and a pressure of from about 400 to 600 p.s.i.g. for from about 10 to 100 seconds and the hydrogen to aromatic hydrocarbon mol ratio is within the range of from about 3 to 10; said flash vaporization is conducted at a pressure of from about 400 to 600 l p.s.i.g. and a temperature of from about 500 to 650 F.; and wherein the reaction in said auxiliary reaction zone is conducted at temperatures in the range of from about l000 to 1300 F. and a pressure of from about 400 to 600 p.s.i.g. for from about 20 to seconds and the hydrogen to aromatic hydrocarbon mol ratio is within the range of from about 4 to 10.

5. The process of claim 2 wherein at least a portion of said unreacted alkyl aromatic is separated from the mixture obtained from the separation zone and recycled to said primary reaction zone.

6. The process of claim 2 wherein said alkyl aromatic is toluene.

7. The process of claim 1 wherein a catalyst is employed in the auxiliary reaction zone.

8. The process of claim I wherein a noncatalytic auxiliary reaction zone is employed.

Claims (7)

1. 2. The process of claim 1 wherein the effluent from said auxiliary reaction zone is combined with the effluent from said primary reaction zone before said effluent from said primary reaction zone is passed to said separation zone, and the benzene product is thereafter recovered from the combined effluents.
2. 3. The process of claim 2 wherein the reaction in said primary reaction zone is conducted at a pressure of from about 100 to 1000 p.s.i.g. for from about 10 to 600 seconds and the hydrogen to aromatic hydrocarbon mol ratio is within the range of from about 1.5 to 20.0; said flash vaporization is conducted at a pressure of from about 300 to 800 p.s.i.g. and temperatures in the range of from about 450* to 750* F.; and wherein the reaction in said auxiliary reaction zone is conducted at a pressure of from about 300 to 800 p.s.i.g. for from about 10 to 200 seconds and the hydrogen to hydrocarbon mol ratio is within the range of from about 1.5 to 20.0.
3. 4. The process of claim 2 wherein the reaction in said primary reaction zone is conducted at temperatures in the range of from about 1050* to 1400* F. and a pressure of from about 400 to 600 p.s.i.g. for from about 10 to 100 seconds and the hydrogen to aromatic hydrocarbon mol ratio is within the range of from about 3 to 10; said flash vaporization is conducted at a pressure of from about 400 to 600 l p.s.i.g. and a temperature of from about 500* to 650* F.; and wherein the reaction in said auxiliary reaction zone is conducted at temperatures in the range of from about 1000* to 1300* F. and a pressure of from about 400 to 600 p.s.i.g. for from about 20 to 120 seconds and the hydrogen to aromatic hydrocarbon mol ratio is within the range of from about 4 to 10.
4. 5. The process of claim 2 wherein at least a portion of said unreacted alkyl aromatic is separated from the mixture obtained from the separation zone and recycled to said primary reaction zone.
5. 6. The process of claim 2 wherein said alkyl aromatic is toluene.
6. 7. The process of claim 1 wherein a catalyst is employed in the auxiliary reaction zone.
7. 8. The process of claim 1 wherein a noncatalytic auxiliary reaction zone is employed.

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