US3591651A

US3591651A - Combination thermal hydrodealkylation diphenyl hydrogenation process

Info

Publication number: US3591651A
Application number: US816145*A
Authority: US
Inventors: Norman L Carr; Sheldon J Kramer; Donald L Stahlfeld
Original assignee: Gulf Research and Development Co
Current assignee: Chevron USA Inc; Gulf Research and Development Co
Priority date: 1969-03-14
Filing date: 1969-03-14
Publication date: 1971-07-06
Anticipated expiration: 1988-07-06
Also published as: FR2034929A2; JPS518929B1; FR2034929B2; ES376755A2; GB1292701A

Abstract

IN A PROCESS FOR THE THERMAL HYDRODEALKYLATION OF ALKYL AROMATIC COMPOUNDS, WHEREIN POLYPHENYL REACTANTS ARE SUBJECTED TO CONVERSION TO BENZENE IN AN AUXILIARY REACTION ZONE, HYDROGENATION OF THE EFFLUENTS OF THE PRIMARY AND AUXILIARY REACTION ZONES RESULTS IN PROLONGED COKE-FREE OPERATION.

Description

COMBINATION THERMAL HYDRODEALKYLATION DIPHENYL HYDROGENATION PROCESS Filed March 14. 1969 J1y6,1971 ,CARR E-rAL 3,591,651

United States Patent Oice 3,591,651 Patented July 6, 1971 U-S. Cl. 260-672 9 Claims ABSTRACT OF THE DISCLOSURE In a process for the thermal hydrodealkylation of alkyl aromatic compounds, wherein polyphenyl reactants are subjected to conversion to benzene in an auxiliary reaction zone, hydrogenation of the effluents of the primary and auxiliary reaction zones results in prolonged coke-free operation.

This application is a continuation-in-part of Ser. No. 596,125, filed Nov. 22, 1966, now abandoned.

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 temperatures within the range of from about 1000 to about 1800 F. 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 and phenyl radicals and the combination of these radicals with hydrogen to form methane and benzene. The overall hydrodealkylation reaction is highly exothermic.

The yield of benzene in the thermal hydrodealkylation of an alkyl aromatic may be increased somewhat 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 thermally 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 requires a thermal hydrodealkylation unit and recovery apparatus of increased size for a given benzene product capacity, and also requires the recycle of a greater quantity of 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 may be operated at conversions per pass of percent or higher and still attain a high benzene selectivity for the overall process.

It is a further object of this invention to provide a thermal hydrodealkylation process for alkyl aromatics which results in prolonged coke-free operation.

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 an-d hydrogen in a primary reaction zone to reaction temperatures in the range of from about 1000" to 1800 F., and advantageously in the range of about 1050" to 1400 F. An effluent comprising unreacted hydrogen and alkyl aromatic, benzene product, polyphenyls (primarily diphenyl), and fused ring constituents (such as naphthalene, anthracene and tribenzene) is removed from the primary reaction zone. This eluent is then passed to a separation zone in which polyphenyls and fused ring constituents are removed as bottoms. The removed polyphenyl-rich stream is mixed with hydrogen, and preferably with additional alkyl aromatics, and the mixture is subjected to temperatures in the range of from about 900 to 1500 F., and especially advantageously in the range of about 1000 to 1350 F., in an auxiliary reaction zone to convert at least a portion of the polyphenyl-rich stream to benzene. The benzene product from the effluents from the primary reaction zone and from the auxiliary reaction zone is recovered. In a preferred embodiment, the effluent from the auxiliary reaction zone is combined with the eflluent 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.

By subjecting at least the diphenyls boiling range components of the eluent from the auxiliary reaction zone, and preferably the polymer bottoms from both the primary and auxiliary reaction zones to hydrogenation conditions, as herein described, prior to recycling the unconverted material from the auxiliary reactor effluent to lthe auxiliary reactor, prolonged coke-free operation is maintained. Maximum utilization of fractionating equipment, maximum freedom of coke and best benzene product quality are obtained by subjecting the combined effluents from the primary and auxiliary reaction zones to hydrogenation, as described herein.

This invention provides an integrated process for converting to benzene a portion of the diphenyl content contained in the 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. Moreover, in the process of this invention, an overall benzene selectivity of about 98 percent may be obtained when operating the primary reaction zone at a conversion in excess of 80 percent, normally between 85 and 95 percent conversion per pass. More particularly, an overall benzene selectivity of 98 percent or more can be obtained when the temperatures in the primary reaction zone are in the range of about 1050L7 to 1400 and the residence time of the reactants in this zone is in the range of about 10 to 100 seconds, and preferably at least 30 seconds, and the combination of temperatures and residence times in the primary reaction zone are so selected as to produce a conversion in excess of 80 percent per pass with a benzene selectivity of at least 92 percent, and when the temperatures in the auxiliary reaction zone are in the range of about 1000 to 1350 F., the residence time of the reactants in this zone is in the range of about 20 to 120 seconds, preferably at least 60 seconds, and the combination of temperatures and residence time in the auxiliary reaction zone is so selected as to produce a conversion of at least 75 percent per pass with a benzene selectivity greater than 94 percent.

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 omitted from 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 C9 and C10 alkylbenzenes and mixtures of any of these. The feed stock may also contain up to 10 percent by weight of heavy parafflns containing from 6 to 12 carbon atoms, as described in copending application, Ser. No. 596,158, led Nov. 22, 1966, now Pat. No. 3,340,318.

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.

Make-up hydrogen-containing gas at an elevated pressure is introduced to the process through line 15. A portion of this hydrogen-containing gas is passed through line 16 and combined with the feed to the auxiliary reactor 85 as hereinafter described. The remaining portion of the hydrogen-containing make-up gas is combined with the alkyl aromatic feed in line 13.

Hydrogen-containing recycle gas is introduced into line 15 from line 17 through line 18. The make-up hydrogen gas stream, as well as the hydrogen-containing recycle gas stream, need not be pure hydrogen. These streams may contain between about 40 and 100 percent hydrogen by volume. Preferably, the make-up hydrogen gas stream passed through line 16 contains 90 percent or more hydrogen; and the combined make-up hydrogen and hydrogen-containing recycle gas which is combined with the alkyl aromatic feed in line 13 contains from 45 to 95 percent by volume of hydrogen.

The reactant feed stream comprising the alkyl aromatic and hydrogen may contain a hydrogen-to-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 to a temperature of about 340 F. 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 eflluent to a temperature of about 940 F. The preheated reactant stream is passed from the exchanger 20 by line 21 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 ellluent optionally may be quenched to a lower temperature by direct mixing with a cool hydrogen-containing recycle stream obtained from a high pressure flash drum more fully described hereinafter and 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 1i800 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 10 to 600 seconds. In a preferred em- CFI bodiment 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, and preferably at least 30 seconds. Advantageously, the temperature is sufficiently low and the contact time is sufficiently long to provide a conversion of at least percent, preferably to 95 percent, per pass, and a benzene selectivity of at least 92 percent.

An eflluent is recovered from the top of the reactor 27 by line 28. The eflluent in line 28 then optionally may 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 hereinbefore mentioned and introduced by a line 29, as described in Pat. No. 3,188,359. Alternatively, the etiluent in line 28 can be quenched by direct mixing with a portion of the liquid phase obtained from the high pressure flash drum, as described in copending application, Ser. No. 586,823, filed Oct. 14, 1966, and now abandoned. Cooling also may be accomplished by insertion of a steam generator in line 28 or in heat exchangers 20, 31 and 33, 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 eflluent from reactors 24 and 27 is optional. Therefore, lines 26 and 29 may be omitted if a quench is not employed at these points. Moreover, quenching at these points also may be accomplished `by means of liquids or gases other than the hydrogen-containing recycle stream illustrated.

The effluent in line 28 after quenching is passed to indirect heat exchanger 20, wherein the eflluent gives up a portion of its heat to preheat the feed in line 19, thereby cooling the effluent to a temperature of about 700 F. The partially cooled eflluent is then passed by line 30 to a reboiler 31 associated with the bottom of a fractionator, more fully discussed hereinafter, to provide the heat duty of the fractionator. In reboiler 31, the effluent gives up a portion of its heat and is cooled to a temperature of about 580 F. by indirect heat exchange with a liquid stream withdrawn from the lower portion of the fractionator. The thus cooled effluent is passed from reboiler 31 by line 32 to reboiler 33 associated with the bottom of a product stripper tower more fully discussed hereinafter. In reboiler 33, the effluent is further cooled to a temperature of about 495 F., from which it is withdrawn and passed by line .34 to a hydrogenation chamber 35 wherein the eflluent is subjected to hydrogenation conditions as hereinafter described.

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 to a temperature of about 345 F. Accordingly, the hot eflluent recovered from reactor 27 supplies the heat duty of the fractionator 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 is then passed from the exchanger 14 by line 37 to a suitable cooler 38. Cooler 38 may be any suitable arrangement of coolers comprising a water cooler, air cooler, or a combination thereof which will sufficiently cool the eflluent 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 of about 525 p.s.i.g. suitable for recycle to the reactors, thereby raising the temperature of this stream to about 130 F. The thus-compressed recycle stream is passed by line 17 optionally to line 29 and/ or line 26 for use as yquench material in the reactor effluent streams as discussed above; and to line 45 for use as quench material in the effluent stream from the auxiliary reactor as hereinafter described. Another portion of this recycle stream is passed by line 18 to line 15 wherein it is combined with hydrogen-rich make-up gas. The combined gas stream is thereafter combined with the hydrocarbon feed to be dealkylated prior to the heat exchange steps hereinbefore discussed. Optionally, another portion of this recycle stream is passed by line 17 to line 90 and subsequently into lines 16, 81, 83 and 84 4and introduced into the auxiliary reactor 85 as an additional or alternative source of hydrogen for this reactor.

The vaporous stream of minor portion in line 44 recovered from the high pressure flash drum is 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 0f 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 precool the vaporous stream in line 44 as described above. The refrigeration flash vapors are recovered from heat exchanger 46 by line 52 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 110 F. to about 450 F. and a pressure in the range of from about 290 p.s.i.g. to about 320 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 33 and thereafter returning the heated withdrawn stream to the tower by line 57 to supply the heat duty of the stripper tower.

In stripper tower 55 a vaponous stream comprising about 2 mol percent 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 is passed through refrigeration drum 48 to cool this stream to about 40 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. In 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 55 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 60 is also connected to line 64 in order that this recovered liquid material may be passed to fractionator 65.

Fractionator 65 is maintained at a temperature in the range of from about 210 F. to about 425 F. The heat duty of the fractionator 65 is supplied by withdrawing a liquid stream from the lower portion of the tower by line 66, passing the thus-withdrawn material through reboiler 31 and returning the heated liquid stream to the tower by line 67. Provision is also made for recycling a portion of the liquid polymer bottoms withdrawn from the bottom of the fractionator 65 by line 68 through line 69 connected,to reboiler 31 where it is heated and thereafter returned to fractionator 65. The remaining portion of the polymer bottoms is withdrawn by way of line 93 and either removed from the system by way of line 94, or alternatively, all or a portion of the polymer bottoms can be drawn off by way of line for recycle to auxiliary reactor 85, as described below.

Fractionator tower 65 is designed to withdraw a benzene product stream from the upper portion thereof by line 70 which is provided with cooler 71 for cooling the benzene product stream to a temperature of about F. To assure recovery of a high purity benzene product stream from the fractionator, the benzene stream is withdrawn from the fractionator at about the fifth tray and any lower boiling materials are withdrawn from the top of the tower by line 72, cooled in cooler 73 to a temperature of about F. and then passed to separator 74. All or a portion of this material is employed as a cool reflux stream and is withdrawn from separator 74 and returned to the top portion of the fractionator above the point of withdrawal of benzene product material by line 75. Line 76, which is connected to line 75, is provided for withdrawing any excess reflux material from the fractionation system. Any lighter than benzene material is withdrawn from separator 74 by line 98.

Unconverted alkyl aromatics are removed from a lower portion of the fractionator by line 77 provided with cooler 78 for reducing the temperature of the stream to about 100 F. The cooled liquid stream is recycled by line 79 to the feed in line 11. A portion of the entire part of the cooled liquid stream from line 79 can be employed as additional feed to the auxiliary reactor 85 by passing it through line 91 to line 81.

A polyphenyl-rich fraction which boils between the liquid recycle and the polymer bottoms is removed from the fractionator 65 by line 80 and is pumped through lines 92 and 81 wherein it is combined with make-up hydrogen-containing gas and/or recycle gas introduced by line 16. The fraction removed by line 80 comprises polyphenyls, such as diphenyl, methyl diphenyl, C14 diphenyls and triphenyls, and fused ring constituents such as naphthalene, anthracene, phenanthrene, pyrene, and tribenzene. Sufficient hydrogen-containing gas is introduced through line 16 to give a hydrogen-to-hydrocarbon mol ratio in line 81 within the range of from about 1.5 to 20.0 and, preferably, from about 4 to 10.

A variation of this procedure can also be used in which all or part of the polymer bottoms removed by line 93 is passed through line 95 to line 92 and used as feed to the auxiliary reactor 85. When such a procedure is employed, line 80 either can be eliminated entirely, or can be closed off by valve means not shown. Any remaining part of the polymer bottoms not passed through line 95 thus passes through line 94 and is available for other purposes.

Furthermore, when all of the material entering line 77 is to be passed through line 91 for recycle to auxiliary reactor 85,

lines

77, 79 and 91 and exchanger 78 can be eliminated or closed off by valve means not shown because the operation of fractionator 65 can be modified t0 obtain almost the same flow rate and composition of material in line 91 using only lines 68, 70, 72 and 80 to withdraw material from the fractionator.

An alternative operation is such that no flow occurs through line v80 and only lines 68, 70, 72 and 77 are used to remove material from the fractionator.

A third possible operation of fractionator 65 is such that only lines 68, 70 and 72 are used to remove material from the fractionator. In this operation, lines 77, 79, 80 and 91 and exchanger 78 are unnecessary and need not be present.

The stream comprising the combined polyphenyl-rich v fraction and make-up hydrogen-containing gas is passed from line l81 to heat exchanger 82 wherein it is indirectly heated to about 900 F. without hot effluent from the auxiliary reactor as hereinafter described. Thereafter, the stream is passed by line 83 to furnace 22 wherein it is heated to a temperature of about 1200D F. The stream is then passed from furnace 22 by line 84 to auxiliary reactor 85 wherein at least a portion of the polyphenyl-rich fraction is converted to benzene. The overall reaction which takes place in the auxiliary reactor as exemplified by C12 diphenyl is represented by the following equation:

The reaction in the auxiliary reactor 85 can be conducted thermally at temperatures within the range of from about 900 to 1500 F. and, preferably, between 1000 to 1350 F., at a 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 to 200 seconds and, preferably, from to 120 seconds. Especially good results are obtained with contact times in excess of 60 seconds when using temperatures in the preferred range. Advantageously, the temperature is sufficiently low and the contact time is suiciently long to provide a conversion of at least 75 percent of the benzene precursors in the feed and a benzene selectivity of at least 94 percent. Thus, operating conditions in the auxiliary reactor need not be the same as, indeed, as elsewhere explained, they are advantageously different from, those employed in the primary reactors. While it is not necessary to use a catalyst in the auxiliary reactor, nonhydrogenation-dehydrogenation catalysts, such as natural clays, aluminas, silicaaluminas and sulfided metal (e.g., Ni) supported catalysts may 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 85 is preferably sized to provide a sucient reactant holding time to obtain a maximum temperature of about l300 F. at about 400 p.s.i.g. An euent is recovered from auxiliary reactor 85 by line 86. The effluent in line 86 is quenched with recycle gas from line 45. The quenched eluent is then passed to indirect heat exchanger 82, wherein it gives up a portion of its heat to preheat the stream in line `81. The euent is thereby cooled to a temperature of about 500 F. The effluent is passed from heat exchanger 82 through line 87 and is introduced into line 34 wherein it is combined with the product effluent stream from primary reactors 24 and 27, and passed to the hydrogenation chamber 35. While the broad range of hydrogenating conditions described in copending applications, Ser. Nos. 466,255 or "466,256, led June 23, 1965, and now Pat. Nos. 3,310,593 and 3,310,594, is satisfactory for purposes of the benzene product quality, a relatively severe combination of conditions is desirable for purposes of the hydrogenation of the polymer bottoms. Thus, temperatures in the range of from 575 to 625 F., pressures of 400 to 600 p.s.i.g., and liquid hourly space velocities of 3 to 6 are preferred. The particular combination of conditions selected should, however, not be so severe as to result in significant proportions of cyclohexane in the benzene product.

The auxiliary reactor is advantageously operated at a lower temperature and so as to permit a longer contact time than in the casewith the main reactors, since the selectivity to benzene for a feed stock containing chiefly diphenyls is improved at temperatures lower than those favoring the selectivity to benzene of a feed stock containing chiefly toluene or similar alkylbenzenes.

As previously mentioned, it is preferred that the hydrogen-containing gas introduced through line 16 which is combined with the polyphenyl-rich fraction in line 81 be of relatively high purity, e.g., percent or more by volume of hydrogen. Therefore, make-up hydrogen gas is preferred instead of recycle hydrogen-containing gas which may 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, and 16 can be used, and if necessary, recycle gas alone introduced through lines 17, 90 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 85. 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 85, 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 make-up 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 88 and 89 to avoid freezing of any benzene material, cooled in refrigeration exchanger 48.

An important feature of this invention is the manner in which the reactants are brought to and maintained at reaction temperature. In the practice of this invention, the reactants, i.e. hydrogen, feed and recycle, are blended at temperatures below reaction temperature and allowed to become thoroughly admixed while the temperature is raised in a series of steps over a period of time to one at which the reaction will proceed spontaneously. The reaction then is allowed to proceed solely by means of autogenous heat.

This system is relatively temperature stable, and is unlikely, even in the absence of quenching or automatic controls, to reach temperatures that promote undesired side reactions involving degeneration of the hydrocarbons present in the system to methane and ethane. These side reactions are promoted at higher temperatures, and lea'd to reduced benzene selectivity.

More particularly, by providing a well mixed system that is substantially homogeneous as to composition and temperature when incipient reaction temperature is reached, we avoid any fluctuations in reactor inlet temperature that might result from incomplete mixing of hydrogen and alkyl aromatic charge stock, as well as any variations in the hydrogen-to-hydrocarbon ratio at the reactor inlet that might result from incomplete mixing, either of which will produce much larger temperature fluctuations at the reactor outlet. By causing the reaction to proceed solely by means of autogenous heat, we avoid, without depending on sophisticated and expensive controls, the runaway temperatures that are likely to occur when external heating is employed and there is a substantial fluctuation in the amount of exothermic heat generated as a result of a change in temperature or feed composition at the reactor inlet.

This manner of preheating is employed for the feed to the auxiliary reactor as well as for the primary reactor, but it is more important in the case of the auxiliary reactor, as it is much easier to see to it that all of the toluene is in vapor form before reaction temperature is achieved, by whatever method of preheating is employed, than is the case with polymer bottoms. The presence of any liquid in the feed to either the primary or auxiliary reactors at reaction temperature is undesirable, since liquid phase material is more susceptible to coking, due to slower movement of the liquid through the reaction zone or settling on conduit or vessel walls. However, in the case of polymer bottoms feed, which cokes more readily than the relatively pure alkyl aromatic feed to the primary reactors, the method of preheating the feed is rvery important in order to minimize coking. Thus, in the case of the polymer bottoms, which contains coke precursors, homogeneous mixture With hydrogen is especially important when reaction temperature is reached, since to avoid or minimize coke formation, it is important to disperse coke-forming molecules as much as possible to reduce their activity, or probability of contact with other coke-forming molecules.

According to the practice of this invention, the polymer feed to the auxiliary reactor is intermixed with hydrogen at below reaction temperature, and thorough intermixture is allowed to take place before heating to reaction temperature. Then, further heating to just below reaction temperature is effected, and again a period of time is provided to allow uniform distribution of heat. Next, the mixture is raised to the reaction temperature in the red furnace 22 and introduced into the reactor, where the reaction is caused to proceed solely With autogenous heat. The gradual mixing and heating in increments, with pauses between to allow homogeneity of temperature and mixture to occur, tends to minimize the possibility of any liquid material being present when reaction temperatures are reached, thus minimizing the possibility of coking. The subsequent conduct of the reaction solely by means of autogenous heat results in a relatively thermally stable, easily controlled system, and avoids the runaway temperatures that lead to poor benzene selectivity which are likely to occur with thermally unstable, diiculty controlled, externally heated reactor systems.

An additional feature of this invention is the prolonged coke-free operation resulting from the use of hydrogenation (i.e. tower 35) to clean up the eluent from the primary and auxiliary reactors. It has been found that in operations involving recycle of unreacted polyphenylrich polymer bottoms to the auxiliary reactor 85, that significantly reduced coking tendencies result from the hydrogenation step.

.. ,The practice of this invention 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 alkylbenzenes 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 EXAMPLE 1 In a continuous process, a toluene feed containing about 98 percent by volume toluene and 2 percent by volume xylene is mixed with make-up and recycle gases in a hydrogen-to-hydrocarbon mol ratio of 4.6 to 1. These reactants are preheated in a furnace to a temperature of about ll F. at about 460 p.s.i.g., and are charged to the first primary reactor where the temperature increases to about 1285 F. The first primary reactor eluent is quenched with recycle gas to a temperature of about 12.30 F., and is fed to the second primary reactor where the temperature rises to about 1245 F. The nominal residence time of the feed in the primary reactors is about 46 seconds. After quenching with recycle gas, the effluent is passed through a series of exchangers wherein the ternperature is lowered to about 500 F. before entering the hydrogen treater for production purification. The treated effluent then is flashed in a high pressure separator. The liquid is charged to a stripper and the gas returned in part to a recycle gas compressor, and part rejected to fuel. The liquid is stabilized in the stripper and charged to a fractionator where it is separated into three fractions, one fraction comprising 80.6 mol percent of 99.9 percent benzene, a second fraction containing 17.0 mol percent of unconverted toluene and a third total fractionator bottoms fraction containing 2.4 mol percent of polymer containing about 41 mol percent toluene, 4 mol percent xylene, 1 mol percent naphthalene, 38 mol percent C12 diphenyl, 6 mol percent C13 diphenyl, 5 mol percent C13 fluorene, l mol percent C14 diphenyl, l mol percent C14 fluorene, l mol percent C1., anthracene-phenanthrene and 2 mol percent C161L aromatic. The conversion in the main reactors is 83.5 percent and the selectivity of the conversion of toluene and xylene to benzene in these reactors is 95.2 mol percent to benzene, 4.5 mol percent to C11,+ aromatics, and 0.3 mol percent to C1-C2 gas. The polymer fraction recovered from the fractionator is mixed with 14.1 percent by weight, based on the weight of the polymer of a gas containing 95.6 percent by volume of hydrogen. The mixture is preheated to about 1200 F. and is charged to an auxiliary flow reactor where the nominal residence time or holding time is about 78 seconds at an average reactor temperature of 1250 F. and a pressure of about 460 p.s.i.g. The effluent from this reactor is determined to be comprised of about 3836 parts by weight of a liquid and 921.5 parts by weight of a gas. The liquid contains 2716 parts by weight of benzene, 288 parts by weight of toluene, 7 parts by weight of xylene, 73 parts by weight of naphthalene, 441 parts by weight of C12 diphenyl, 31 parts by weight of C13 diphenyl, 165 parts by weight of C13 fluorene, 4 parts by weight of C1,= fluorene, 84 parts by weight of C14 anthracene-phenanthrene, and 27 parts by weight of C16Jr aromatics. The gas contains 85.1 percent by volume of hydrogen. The conversion of Abenzene precursors in the auxiliary reactor is 76.6 percent and the selectivity to benzene in this reactor is 96.7 percent. The overall selectivity to benzene for the combined reactions (i.e., for the reactions which take place in the primary reactors and the reactions which take place in the auxiliary reactor) based on the fresh feed to the entire unit, is 98.1 mol percent. From a comparison of the results obtained from the primary reactors aloneI and from the combination of the primary reactors and the auxiliary reactor, it is seen that a higher benzene selectivity is obtained by the use of the auxiliary reactor which converts polyphenyls to benzene.

EXAMPLE 2 The process of Example 1 is repeated except that the effluent from the auxiliaryireactor is combined with the the auxiliary reactor for converting the fractionator bottoms to benzene.

EXAMPLE 3 The process of Example 2 is repeated, except that a side stream containing chiefly diphenyl and close boiling polyphenyls also is removed from the fractionator and charged to the auxiliary reactor, along with hydrogen gas in the proportion and under the reaction conditions specied in Example 2. The overall selectivity to benzene for the main and auxiliary reactors is greater than that in the main reactors alone.

EXAMPLE 4 The process of Example 3 is repeated, except that toluene from the toluene side stream is combined with the diphenyl side stream of Example 3 in the proportion of 40 mol percent, and the mixture is charged, with hydrogen, as described in Example 3, to the auxiliary reactor. The Overall selectivity to benzene for the main and auxiliary reactors is greater than that in the main reactors alone.

EXAMPLE 5 The process of Example 1 is repeated, except that the eilluent from the auxiliary reactor, containing recycle polymer bottoms, is passed to a fractionator, without being subjected to hydrogenation. The recycle polymer bottoms is combined in the fractionator with fresh polymer bottoms obtained from the primary reactors as described in Example 1, so as to recover a mixed polymer bottoms stream containin g fresh polymer bottoms formed in the primary reactors, and unhydrogenated recycle polymer bottoms from the auxiliary reactor. This mixed polymer bottoms is mixed with a hydrogen-containing gas and heated as described in Example 1, and the mixture is passed into the auxiliary reactor. The mixed polymer bottoms contains recycle and fresh polymer in a ratio of 177 parts recycle polymer to 265 parts fresh polymer. The composition of the feed to the auxiliary reactor, and the product effluent stream from the auxiliary reactor, are set forth in Table I, below. The conversion of benzene precursors is 75.5 percent and the benzene selectivity is 94.5 percent. The nominal residence time in the auxiliary reactor is 80 seconds, at an average reactor temperature of 1250 F., and a pressure of about 460 p.s.i.g. The hydrogen-to-hydrocarbon mol ratio is maintained at 6.36 and 11.8 percent of the polymer bottoms is bled from the system during operation. Coking, i.e. plugging, is encountered in the preheater system for the auxiliary reactor after approximately 100 hours. The tendency of the total feed to the auxiliary reactor to coke may be monitored by periodic observation of the bromine index of the feed. The bromine index is a measure of unsaturation, including unsaturation due to aromatics. The bromine index is determined in accordance with ASTM procedure D-1491. At startup, the bromine index of the feed to the auxiliary reaction is found to be 71. After 4 hours, the bromine index is 105, and at the time coking occurs, the bromine index of the feed is 146.

TABLE I [Feed and product compositions for auxiliary reactor with polymer recycle, without hydrogenation] Fresh Recycle Total feed polymer polymer to auxiliary bottoms bottoms reactor Product Component, percent by weight:

Benzene 0. 1 1. 2 0. 5 62. 4 28. 3 5. 1 19. 0 4. 8 2. 9 0. 1 l. 8 0. 1 1. 3 13. 5 6. 2 5. 8 51. 9 25. 5 41. 3 13. 7 C13 diphenyl 3. 4 0. 4 2. 2 0. 2 Cia Iluorene 7.1 29. 2 15. 9 12.2 C14 diphenyl 1.0 0.2 0. 7 0. l C11 antliracene- 1. 5 14. (i ti. 8 5. 8 Cia-iaromatics. 2. 5 10. 2 5. 6 4. U

EXAMPLE 6 The process of Example 5 is repeated, except that the effluent from the auxiliary reactor, as well as the eflluent from the primary reactor, is subjected to hydrogenation as described herein. Thus, in this example, the recycle polymer bottoms obtained from the auxiliary reactor is subjected to hydrogenation, whereas in Example 5, there was no hydrogenation of this material. The ratio of fresh polymer to recycle polymer in the mixed polymer introduced to the auxiliary reactor is 261 parts polymer to parts recycle, with a nominal residence time in the reactor of 81 seconds. The composition of the mixed polymer feed to the auxiliary reactor, and the product obtained therefrom are set forth in Table 1I, below. The conversion of benzene precursors is 81.7 percent and the benzene selectivity is 95.3 percent. An average reactor ternperature of about 1256 F. is maintained at a pressure of about 460 p.s.i.g. The hydrogen-to-hydrocarbon ratio is 6.43, with 11.5 percent of the polymer bottoms being bled from the system. The hydrogen treater is run at a temperature of 500 F. initially, at a pressure of about 460 p.s.i.g. and a hydrogen-to-hydrocarbon mol ratio of 6.43. Space velocity through the hydrogenation tower is 4.18. After 72 hours, the bromine index of the auxiliary reactor feed increases from an initial value of 71 to 135, ndicating that the severity of the treatment, i.e. the combined effect of the temperature, pressure, hydrogen oil ratio, and space velocity is not sufficient to obtain the desired degree of hydrogenation. Therefore, the hydrogenation tower temperature is raised to an average of 600 F. at `which temperature a bromine index of 99 is found after 118 hours of operation. The bromine index decreases further, to 92, after 484 hours, indicating that the presence of the hydrogenation tower at effective operating conditions will greatly prolong coke-free operation.

TABLE lI [Feed and product compositions for auxiliary reactor with polymer recycle, with hydrogenation] Fresh Recycle Total feed polymer polymer to auxiliary bottoms bottoms reactor Product Component, percent by weight:

Benzene 0. 1 5. 4 2. 2 59. 9 28. 3 3.3 18. 4 1. 4 2. J 0. 2 1. 8 0. 1 l. 3 16. 1 7. 2 6. 8 51.9 30. 2 43. 4 12. 7 C13 diphenyl- 3. 4 0. 7 2. 3 0.3 C13 fluorene- 7. 1 27. 9 15. 3 11. 7 C14 diphenyl 1. 0 0. 2 0. 7 0. 1 C14 anthracene. 1. 5 10. 2 4. 9 4. 3 C15-|- aromatics 2. 5 5. 8 3.8 2. 7

Obviously, many modifications and variations of the invention as hereinabove set forth can be made without departing from the spirit and scope thereof. For example, the liquid recycle stream withdrawn by line 77 from the fractionator may be divided into two portions, one portion being passed as described previously to the feed in line 11, and the other portion being combined with the polyphenyls in line 81 by way of line 91. Since the conversion of diphenyl to benzene is much less exothermic than the conversion-of toluene to benzene, the proper proportioning of liquid recycle to polyphenyls can be made to maintain a temperature prole in the auxiliary reactor which substantially increases asthe dealkylation reactions proceed. Based on a given quantity of benzene produced, the dealkylation of toluene is between .3.5 to 4.0 times more exothermic than the reaction producing benzene from diphenyl and hydrogen. Hence, the addition of some alkylbenzenes to the feed of the auxiliary reactor 85 has the advantage that, while it slightly increases the volume required for the auxiliary reactor, it decreases the amount by which the feed to said reactor has to be heated to obtain a given conversion in said reactor. In addition, the inclusion of an alkyl benzene component in the polyphenyl feed to the auxiliary reactor reduces the possibility of coking by reducing the partial pressure required for the polyphenyls to form a vapor mixture and by reducing the activity of the polyphenyl molecules for coke formation. In general, the inclusion of about 30 to 60 mol percent and preferably lessthan 50 mol percent toluene, or equivalent amounts of other alkylbenzenes, is suicient to provide the heat required to maintain the desired temperature profile in the auxiliary reactor. Small amounts of fresh feed also could be injected into lines 80 or 81 and combined with the polyphenyls to produce the desired temperature profile in the auxiliary reactor.

We claim:

1. A process for the 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 within the range of about 10007 to 1800 F.;

(b) recovering an effluent comprising unreacted hydrogen and alkyl aromatic, benzene product, polyphenyls and fused ring constituents from said primary reaction zone;

(c) passing said effluent from said primary reaction zone to a catalytic hydrogenation zone wherein the hydrocarbon portion of said eluent from said primary reaction zone is subjected to partial hydrogenation without significant saturation of benzene;

(d) removing the thus hydrogen-treated eluent from said hydrogenation zone and passing said hydrogentreated eluent to a separation zone wherein a polyphenyl-rich stream is removed;

(e) mixing said polyphenyl-rich stream with hydrogencontaining gas and subjecting said mixture in an auxiliary reaction zone to temperatures within the range of from about 900 to about 1500 F., whereby at least a portion of said polyphenyl-rich lstream is converted to benzene;

(f) combining the efluent stream from said auxiliary reaction zone with the eluent from the primary reaction zone before said eluent from said primary reaction zone enters said hydrogenation zone; and

(g) thereafter recovering the benzene product from the combined effluents.

2. The process of claim 1 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-hydrocarbon mol ratio is within the range of from about 1.5 to 20.0; wherein the reaction in said hydrogenation zone is conducted at a pressure of from about 100 to 1000 p.s.i.g. at a liquid hourly space velocity of from 1 to 20, and at a temperature of from about 150 to 625 F., and wherein the reaction in said auxililary 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. The process of claim 1 wherein the reaction in said primary reaction zone is conducted at temperatures within 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-hydrocarbon mol ratio is within the range of from about 3 to 10; wherein the reaction in said hydrogenation zone is conducted at temperatures within the range of from 575 to 625 F., and a pressure of from 400 to 600 p.s.i.g. and liquid hourly space velocities of from 3 to 6; and wherein the reaction in said auxiliary reaction zone is conducted at temperatures in the range of from about 1000o to l350 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-hydrocarbon mol ratio is within the range of from about 4 to l0.

4. The process of claim 1 wherein the conversion of said alkyl aromatic in said primary reaction zone is in excess of mol percent.

5. The process of claim 1 wherein the conversion of said alkyl aromatic in said primary reaction zone is from to 95 percent.

6. The process of claim 1 wherein said hydrogen-containing gas which is mixed with said polyphenyl-rich stream comprises at least 80 percent by volume of hydrogen.

7. The process of claim 1 wherein said alkyl aromatic is toluene.

8. The process of claim 1 wherein the said removed polyphenyl-rich stream is mixed with hydrogen-containing gas and subjected in admixture with 30 to `60 percent in alkyl aromatic in an auxiliary reaction zone to temperatures in the range of from about 900 to 15 00 F., lwhereby at least a portion of said polyphenyl-rich stream is converted to benzene.

9. A process for the 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 within the range of about 1000 to 1800 F.;

(b) recovering an effluent comprising unreacted hydrogen, an alkyl aromatic, benzene product, polyphenyls and fused ring constituents from said primary reaction zone;

(c) fractionally distilling said eiuent from the primary reaction zone to obtain a polyphenyl-rich stream;

(d) mixing said polyphenyl-rich stream with hydrogencontaining gas and subjecting said mixture in an auxiliary reaction zone to temperatures within the range of about 900 to about l500 F., whereby at least a portion of said polyphenyl-rich stream is converted to benzene;

(e) recovering an efuent comprising benzene product and unreacted hydrogen, polyphenyls and fused ring constituents from said auxiliary reaction zone;

(f) passing at least the effluent from the auxiliary reaction zone to a catalytic hydrogenation zone wherein the hydrocarbon portion of said effluent from the auxiliary reaction zone is subjected to partial hydrogenation without substantial saturation of aromatic, and recovering an effluent from said hydrogenation zone;

(g) passing the eluent from the hydrogenation zone to a separation zone wherein a polyphenyl-rich stream is removed;

(h) mixing at least the polyphenyl-rich stream obtained from the efuent from the hydrogenation zone with hydrogen-containing gas and recycling said mixture to the auxiliary reaction zone at temperatures within the range of about 900 to about 1500 F., whereby a further portion of said polyphenyl-rich stream is converted to benzene;

(i) recovering an effluent comprising benzene product,

unreacted hydrogen and recycled polyphenyls and fused ring constituents; and (j) recovering benzene product from said effluents.

References Cited UNITED STATES PATENTS Coonradt et al. 208-78 Chapman et al 260-670 Richter 260--668 Weidenhammer 260-621 Feigelman 260-672 Payne et al. 260-672 King et al 260--672 Myers et al. 260-672 Nelson et al 260-672 Peterson et al. 260-672 Hoertz et al. 260--672 Carr et al. 260-672 Carr et al. 260-672 Sze 260-672 Hill 260-672 Engelbrecht et a1. 260-672 Juhl et al. 260- 672 OTHER REFERENCES Fowle & Pitts, Thermal Hydrodealkylation, Chem. Eng. Progress 58 (4), 37-40 (April 1962).

Masamune et al., MHCz Newest in Hydrodealkylation," Hydrocarbon Processing 46 (2), 155-158 (February 1967).

DELBERT E. GANTZ, Primary Examiner l5 G. E. SCHMITKONS, Assistant Examiner U.S. Cl. X.R.

260-668F, 672NC, 674H, 675