US3316316A

US3316316A - Benzene-naphtha reforming process

Info

Publication number: US3316316A
Application number: US367334A
Authority: US
Inventors: Jr Walker F Johnston
Original assignee: Standard Oil Co
Current assignee: Standard Oil Co
Priority date: 1964-05-14
Filing date: 1964-05-14
Publication date: 1967-04-25
Anticipated expiration: 1984-04-25

Description

Apri 25, w67 w. F. JQHNSTON, JR Ew BENZENE-NAPHTHA REFORMING PROCESS Filed May 14, 1964 1 Sm t United States Patent O 3,316,316 BENZENE-NAPHTHA REFRMIN G PRUCESS Walker F. Johnston, Jr., Flossmoor, Ill., assignor to Standard Gil Company, Chicago, Ill., a corporation of Indiana Filed May 14, 1964, Ser. No. 367,334 9 Claims. (Cl. 260-668) My invention relates to the production of ultra high octane number blending stocks from naphtha feed stocks and oleiinic benzene-containing material boiling in the gasoline boiling range. More particularly my invention relates to a combined process of catalytic reforming, hydrogenation of the olefinic benzene-containing material and introduction of this hydrogenated material into the reforming process at a selected point to produce a reformate product from which high purity benzene and ultra high octane gasoline blending components may be easily recovered in high yield.

The primary object of the invention is to provide means for economically processing olefinic benzene-containing materials, such as steam-cracked naphthas and dripo` lene, and naphtha to produce in high yield a high octane reformate suitable for use as a gasoline blending component. Another object is a combination process for the recovery of high quality, low sulfur benzene from dripolene without destroying an appreciable amount of the benzene contained therein and without resort to elaborate and costly treating and purification processes. The present invention provides means whereby substantial savings in catalyst requirements for producing high quality benzene and 10U-{- Research octane number gasoline blending stocks can be achieved and also provides means for maximizing the recoverable liquid yield of benzene and high octane product.

It is well known that the high temperature pyrolysis of gaseous hydrocarbons to prepare ethylene results in the by-production of a normally liquid mixture of hydrocarbons through reactions such as polymerization, alkylation, aromatization, dehydrogenation, and the like. The mixture is commonly termed dripolene, and while it contains virtually all classes of hydrocarbons it predominates in oleiinic and aromatic hydrocarbons, mainly benzene. As the demand for ethylene for the production of polyethylene and plastics and other petrochemicals rises, increasingly large supplies of dripolene are becoming available. Because of its benzene content, dripolene represents an exceedingly valuable material, and one which is steadily becoming more obtainable.

Thus far lthree large-volume usages have developed for dripolene. It is blended into motor gasolines, where the high octane numbers of its aromatic and olenic components render dripolene a desirable blending stock. Dripolene may be fed to aromatics extraction units for the recovery `of benzene. Finally, exceedingly valuable resins have been made by thermal or catalytic polymerization of high-boiling dripolene fractions.

Dripolene however contains cylic di-oleiins and other reactive olelinic compounds which give rise to problems which have severely restricted the quantity of dripolene blended into motor fuels or fed to aromatics extraction units. Cyclic di-olefins tend to form gum-like polymers in the presence of air or upon heating, and for this reason only relatively small amounts, usually less than about 0.5 volume percent, of dripolene can be blended into premium motor fuels. And in aromatics extraction units, it is r'ce found that cyclic di-olefins tend to concentrate in the aromatic extract, thereby complicating the preparation of pure aromatic compounds, such as benzene.

A catalytic hydrogenation process has previously been discovered whereby dripolene may be hydrogenated successfully lto remove these di-olens and other gum-formers essentially completely by converting them to saturated or less olefinic compounds, This process is `described in U.S. Patent Number 2,953,612. However, the benzene fraction, boiling in the range of about -190 F., of hydrogenated dripolene still contains an appreciable amount of olenic and sulfur compounds. Many commercial processes utilizing benzene as a starting material require high purity benzene containing essentially no oleiinic or sulfur impurities, therefore the benzene recovered from hydrogenated dripolene must be further purified before it is marketed. Also the F. plus fraction of hydrogenated dripolene has a clear Research octane number of about 98.5 which, by modern standards, is relatively low. Since current premium motor gasolines are being marketed having octane numbers above 100, it is desirable that the research octane number of the hydrogenerated dripolene components other than benzene be increased in order to enhance their usefulness as premium gasoline blending components.

Previous attempts to recover benzene and high octane reformate from dripolene by including even small concentrations of dripolene with the normal naphtha feed to catalytic reformers have been unsuccessful because of fouling of pre-heat exchangers and furnaces, as well as the catalyst beds themselves, due to polymerization and condensation of the olenic materials contained in the dripolene. These fouling problems were overcome by hydrogenating the -dripolene prior to including it in the hydroformer feed. While this mode of operation is generally successful, much of the benzene present in the hydrogenated dripolene, i.e. up to about 30%, is lost in the reforming process by conversion to other compounds.

In the catalytic reforming of naphthas several reactions occur, principally dehydrogenation, isomerization, dehydrocyclization and cracking. The dehydrogenation reaction is the easiest and fastest of the reactions, therefore it takes place in the lirst portion of the reforming catalyst. An example of this reaction is dehydrogenation of cyclohexane to produce benzene and hydrogen. Since the dehydrogenation reaction is highly endothermic, and some of the other reactions are also mildly endothermic, catalytic reforming processes normally employ a plurality of catalyst beds or zones, termed reactors herein, with provisions for reheating the process stream between the beds or zones. In such multi-reactor catalytic reformers the dehydrogenation reaction is essentially complete in the first one or two reactors. My invention contemplates treating the -rst reactor or reactors to be contacted by the naphtha feed as the dehydrogenation reactors which I will designate as the dehydrogenation stage or simply as the first stage. In the subsequent stage, which comprises the remaining reactors or zones, the isomerization, dehydrocyclization and cracking reactions predominate since the dehydrogenation reaction is essentially completed in the first stage.

I have found that when hydrogenated dripolene is introduced into the process stream of a naphtha reformer subsequent to the dehydrogenation stage no loss of benzene occurs. The components of hydrogenated dripolene, other than benzene, are up-graded in research octane rating and he sulfur and olefin impurities present in the benzene fracion of the hydrogenated dripolene are simultaneously renoved so that high purity benzene may be easily recovered rom the reformate product with essentially no loss of the tenzene present in the hydrogenated dripolene. Thus, the )rocess of my invention results in the production of high luantities of high quality benzene and high octane reormate from naphtha and hydrogenated dripolene with ess processing equipment than in any other known process.

I prefer to limit the amount of hydrogenated dripolene which is introduced into the reformer below a 1:1 volume .'atio of hydrogenated dripolene to naphtha charge. A :onvenient amount of dripolene overhead charge is that amount which will result in the net hydrogen produced in the reformer being consumed in the hydrogenator so that the combination process is operated in hydrogen balance. This results in maximum utilization of the chemically combined hydrogen present in the naphtha feed. lt is desirable to desulfurize the naphtha feed to the reformer in order to reduce the load on the H28 removal facilities and minimize corrosion of process equipment.

Although the amount of hydrogen consumed in the hydrogenation of -the dripolene overhead varies with the properties of the particular dripolene used, the hydrogen consumption is normally in the range of 500 to 750 standard cubic feet per 4barrel of dripolene. Likewise the yield of hydrogen from the reformer will vary depending upon the properties of the particular naphtha feed. Reforming of South Texas heavy naphtha typically yield-s 1,000 standard cubic feet (s.c.f.) of hydrogen/ bbl. of feed. In this situation, a dripolene overhead/ naptha volume ratio of about 1.3:1 to 21:1 results in operation of the process of the invention in hydrogen balance. Of course, if Asome of the reformer hydrogen is used to hydrodesulfurize the naphltha feed, this ratio is lowered correspondingly.

Example To illustrate my invention and the advantages thereof, l have chosen for dripolene hydrogenation the process of U.S. Patent 2,953,612 and a 5-reactor catalytic reformer utilizing platinum-alumina-halogen catalyst. The reformer catalyst has a platinum content in the range of 0.1 to 1 weight percent and is disposed in 5 reactors with furnaces provided prior to each reactor for heating the process stream. The catalytic reforming process operates at a pressure in the range of 50 to 500 p.s.i.g., average reactor temperatures in the range of 750 to 50 F., and a weight hourly space velocity (Wo/hn/Wc) in the range of 0.5 to 5. Hydrogen-containing gas is recycled in an amount to provide about 2500 to 6000 standard cubic feet of hydrogen per barrel of naphtha feed.

Three modes of operation are compared:

(l) Separately reforming naphtha and recovery of ben zene from hydrogenated dripolene,

(2) Reforming naphtha containing 15 vol. percent of hydrogenated dripolene,

(3) And reforming naphtha with introduction of hydrogenated dripolene into the reforming process stream subsequent to the first stage, which in this 5-reactor system is subsequent -to the second reforming reactor.

The quantities of naphtha and hydrogenated dripolene utilized in each of the three modes of operation and the reforming severity in each case are the same to permit ready comparison. Benzene recovery in each case is effected by separating from the reformate a narrow-boiling fraction of benzene concentrate and recovering the benzene therefrom by extractive distillation with phenol. The quantity of high quality benzene and the amount and octane of the reformate produced by each of the three modes of operation are compared in Table I. These data show that operation according to the present invention, Mode 3, produces more high octane reformate than either Mode 1 or Mode 2. Also the quality of the benzene produced in Mode 3 is higher than that produced 4 according to Mode 1 and more of it is produced than by Mode 2. Although these differences might at rst glance appear small, these are valuable products which are produced in great quantities, therefore the monetary advantage of operating according to the present invention is indeed substantial.

i Hydrogenated Dripolene overhead, not. reformed.

2 lyllixtlre of 85 bbls. paphtlia and 15 bbls. hydrogenated dripolene over rea 3 85 bbls. naphtha [ed to reformer, 15 bbls, hydrogenated dripolene overhead introduced to reformer subsequent the dehydrogenatiou stage.

4 C-lreformate, less benzene. HDO not reformed in Mode 1.

According to the invention, the naphtha feed is charged to a reforming Zone in which at least three reactors connected in series are provided. 1t is preferred that the naphtha feed be desulfurized. Any desulfurization process known to the art may be employed. The reactors are equipped in the usual manner for inter-heating between reactors in order to compensate for endothermic temperature drop and to introduce additional heat into the later stages of the reaction. The reactors contain a platinum type reforming catalyst in the form of pelleted, pilled, extruded or beaded particles. In the reforming zone, separate dehydrogenation and isomeriZation-dehydrocyclization-cracking stages are readily identified by the relatively large endothermic temperature drop across the catalyst beds of the first one or two reactors encountering the naptha charge.

The efuent from the reforming zone is separated conventionally into a recycle hydrogen gas fraction and a liquid reformate fraction. The latter, if separate benzene recovery is desired, is further fractionated so as to separate light hydrocarbons boiling in the C6 range and to recover a C74- fraction having an ultra high research octane rating. With a commercially available platinumalumina catalyst as used in Ul-traforming, for example, under `reforming conditions including a pressure in the range of about 15 0` to 400 p.s.i.g., the severity level should be sufficient to produce at least -a 95 research octane C5| reformate from a Mid-Continent heavy naphtha charge, and more advantageously, such as to produce a `100-1- research octane C54- reformate.

Dripolene normally boils in the range of about 100 to 400 F. as determined by ASTM distillation, and may contain appreciable amounts of light, normally gaseous hydrocarbons. My invention is particularly concerned with the fraction of dripolene boiling in the range of about 100-375 F., although it is not essential that the dripolene boil entirely over this range or that all of the dripolene fraction boil within the range. My dripolene charging stock is obtained as an overhead or heartcut fraction in the distillation of total dripolene to obtain about to of the charge fraction, which I term dripolene overhead, while the bottoms 4may be used to prepare resins by processes well known -to the art.

Dripolene overhead is hydrogenated in the presence of a platinum-alumina catalyst at elevated tem-peratures and pressures and in the presence of hydrogen gas to selectively hydrogenate monoolens and diolens without simultaneously hydrogenating the aromatic compounds to naphthenes.

The dripolene fraction and a hydrogen-containing gas stream initially contact the platinum-alumina hydrogenation catalyst at conditions of temperature and pressure `such that substantially -all (ie. about 80% or more, and preferably at least 90%) of the dripolene remains in the liquid phase, while maintaining a critically low hydrogen sulfide concentration in the hydrogen-containing gas stre-am.

The a-mount of hydrogen sulfide that may be present in the process is quite critical, and accordingly it is essential to provide to the reaction zone a hydrogen-containing gas stream containing less than 12 grains of hydrogen sulfide per 100 standard cubic feet of gas. If for example the level of sulfur in the hydrogen-containing gas exceeds 12 grains, both the catalyst activity `and the catalyst life diminish rapidly, and if the level increases to as much as 57 grains per 100 s.c.f., the catalyst becornes completely deactivated in a matter of minutes. Fortunately however, the effect of either hydrogen sulfide gas or merca-ptan sulfur in the dripolene feed on the hydrogenation process -appears to be temporary with respect to its effect on product quality although periods of high sulfur do materially increase the amount of coke deposition. It is convenient to use the hydrogen-rich gas produced by the catalytic reformer as the source of hydrogen for the hydrogenation process.

Suitable platinum-alumina hydrogenation catalysts are conveniently those catalysts which have been found eminently suitable for use in naphtha reforming processes. Generally, these catalysts contain from about 0.01 to about by weight of platinum and may optionally from -about `0.05 to about 3% by weight of a halogen, preferably chlorine and/or fluorine, on a high surface area alumina support such as the alumina described in Heard Reissue Patent Number 212,196. The catalysts may be in the form Iof pills, pellets, extrudates, spheres or the like, and conventionally have a size between about y1@ to 1A in maximum dimension. A particularly suitable catalyst is one which has been partially deactivated by continued use in a naphtha reforming process, since it appears that catalysts previously used for reforming are more stable and have less tendency to hydrogenate aromatic compounds than fresh platinum-alumina catalysts. Furthermore, used catalysts exhibit less tendency to cause wasteful 'hydrocracking of hydrocarbons and thus result in higher yields of recoverable liquid product.

The conditions of pressure, temperature, liquid hourly space velocity and hydrogen-containing gas rate which are employed are interrelated such that the commingled feedstock and hydrogen-containing gas, as it initially contacts the catalyst bed, consists of la gas phase and a liquid phase wherein the liquid phase comprises substantially all of the charge stock. It is desired that at least 80 mol percent, and preferably at least 90 mol percent of the dripolene charge contact the catalyst `as a liquid. Pressures within the range of 100 to 1000 p.s.i.g. are desired, with pressures fr-om 300 to 500 p.s.i.g. preferred from a commercial standpoint as this latter range favors conditions at which the hydrogenation reaction occurs rapidly. `Within the broad presure range the bed inlet tem-perature may be between 50 and 200 F., most desirably between 100 and 150 F., typically 115 F. With most dripolene stocks the temperature rise through an adiabatic bed, for complete oleflnic saturation, is on the order of 350-450 F. and provides an average reactor temperature of about 280-340 F. This average temperature may be increased by providing more catalyst or may be decreased by increasing the proportion of hydrogen-containing gas to charge stock. With respect to the hydrogen-containing gas, it is desirably employed in a proportion of 500 to 10,000 standard cubic feet per barrel `(s.c.f./b.) of charge stock, preferably from 1000 to 4000 s.c.f./b., e.g. 1500 s.c.f./b. This gas preferably is cornposed of at least 70% hydrogen as derived from the naphtha hydroforming operation. Although the experi- .mentally observed consumption of hydrogen usually varies between about 5 00 and 750 s.c.f./b., it is preferred to maintain a substantially large amount in the reaction zone.` This may be accomplished economically be relcycling the excess hydrogen. The hydrogen-containing gas, if recycled, must be chemically treated to maintain the critically low hydrogen sulfide level therein.

The effluent from the hydrogenation reactor is introduced into the reforming process stream subsequent to the first, or hydrogenation, stage. The entire eflluent stream may be so introduced, or a hydrogen-rich gas may 'be recovered from the effluent for recycle to the hydrogenation reactor, and only the hydrocarbon portion of the effluent introduced into the reformer. A convenient way to recover :a hydrogen-rich gas for recycle is to cool the eflluent to condense the normally liquid hydrocarbons therein and pass the cooled stream to a gas-liquid separa- Itor from which hydrogen-rich gas and liquid hydrogenated dripolene may be withdrawn.

Simultaneous reforming o-f the effluent from the reformer hydrogenation stage and the introduced hydrogenated dripolene is then effected in the subsequent stage of the reformer. Then, if desired, the C6 reformate is separated by fractionation into a fraction boiling in the range of about to 190 F. and a heavy reformate fraction having an initial boiling point in the range of about to 230 F. which contains at least about 85 yvolume percent aromatics. The latter constitutes a product stream having a clear research octane number significantly in excess of 100. The 140 to 190 F. cutis subjected to solvent extraction or extractive distillation to recover high purity benzene. Alternatively, the total reformate may be used as a high octane gasoline blending component having an ususually high front-end octane because of the increased benzene content.

The invention will be further described by reference to the accompanying drawing which is a flow plan of a preferred embodiment of the invention in simplified diagrammatic form.

The feed, constituting a 200 to 400 F. mixture of South Texas naphthas, is charged to the system through line 10. The feed is preheated in fired heater 11 and is mixed in line 12 with recycle hydrogen gas from line 13. The mixture is charged to reactor 14, which is the first of a train of ve serially connected reactors, each of which contains a bed of platinum-alumina catalyst in pellet form. The reaction mixture is flowed from reactor 14 via line 14a to interheater 16 and from thence via connection 17a to reactor 17. The effluent from reactor 17 is passed by means of l-ine 18, into which hydrogenated dripolene from line 66 is charged, through interheater 19 and from thence by means of connection 20a to reactor 20. From reactor 20, the reaction mixture is passed by means of line 21, interheater 22 and connection 23 to reactor'24. From reactor 24, the reaction mixture is passed by means of line 25, interheater 26 and connection 27 to the last reactor 28. Hydrogenated dripolene may also be introduced into the reformer process stream via valved line 67 :and/or 68.

The eflluent from the final reactor 282 is flowed through line `29 and cooler 30 to high pressure gas separator 31. In separator 31, a recycle gas rich in hydrogen is recovered by line 32 for recompression and recycle through line 33, heater 34 and line 13. Hydrogen-rich make gas is Withdrawn from line 32 via line 3S for use in hydrogenating the dripolene. Excess make gas may be vented Ifrom the system through valved connection 52.

Turning noW to the dripolene hydrogenation, dripolene liquid is withdrawn from external storage tanks and conducted through line 36 to fractionator 37 which is provided with corrosion resistant distillation trays or perforated pans, wherein an overhead dripolene charge stock fraction comprising about 80% of the total dripolene is separated `by distillation from about 20% of high boiling bottoms, which latter is sent via line 38 to the resins plant, not shown. The total dripolene fed to fractionator 37 has an analysis approximating the typical dripolene described 7 reviously. The 80% fractionator 37 overhead which is vken through line 39 has the following compositi-on:

TABLE I1 'harge analysis:

Gravity, A.P.I 32.4 R.V.P.,p.s.i.a 6.18 LSTM distillation, F.:

I.B.P. 134 163 179 50% 1189 70% `202 90% 280 F.B.P. 356

Light hydrocarbons anlysis )omponentz C3 liquid vol. percent 0.1 fc4: dO-.... 1C4,= do 0.2 210,: do 0.2 nC4= do 0.2 C., diolen do 1.6 C5 diolen do 6.6 C5 monoolen do 3.1 C5 paraffin do 0.2 Cs-ido 87.5 Benzene do 53 C64-gravity, A.P.I 28 0 The 80% dripolene charge fraction contains 70 parts per million sulfur and 29 parts per million organic chlorides, and has a bromine number of 48 (indicative of total olens) and a maleic anhydride value (MAV, representing conjugated diolefins) of 47 Ing/g. The bottoms withdrawn through line 38 has an ASTM distillation boiling range between about 200` and 420 F., preferably between about 230 and 375 F.

The dripolene charge is conducted through line 39, cooler 40, and line 49 to a charge pump, not shown, which may be a multistage centrifugal pump adapted to pump the dripolene charge to the reactor system operating at a pressure of 325 pounds per square inch gage. The cooler 40V outlet temperature is about 80 F. The charge stock from the pump is sent through line 41 to junction 42, where it is met by a stream of recycle hydrogen-containing gas from line 43 in the amount of 1350 standard cubic feet of total hydrogen-containing gas per barrel of charge. The gas has a composition of approXimately 80% hydrogen, with the balance consisting primarily `of methane, ethane, and some propane and propylene, together with less than the critical limit of 12 grains of H28 per 100 `standard cubic feet of total gas. It is highly preferred that this gas contain, if possible, less than 3 grains per 100 cubic feet of HZS. The temperature of the commingled liquid and gas stream is 115 F., and at this temperature the commingled stream passes via line 44 into reactor 45 shown symbolically as a single bed or chamber, although it may comprise a plurality of serially lor parallel-connected reaction chambers. At these operating conditions, 94 mol percent of the dripolene is in the liquid phase when the commingled stream initially contacts the catalyst.

The reaction zone 45 operates essentially adiabatically, that is the commingled dripolene charge and hydrogencontaining gas stream are permitted to increase in temperature by the eXothermic heat of monoolen and diolefin hydrogenation on passage through the catalyst bed. The catalyst employed is spent Ultraforming catalyst obtained after more than one years use in a regenerative naphtha reforming unit and has an activity for reforming of substantially less than that of fresh Ultraforming catalyst, but is very nearly as active for hydrogenation as is fresh catalyst. The catalyst in chamber 45 is in the form of pellets having an average length and diameter approximating 1/8 and is disposed so as to permit downllow passage of the commingled stream. A Weight hourly space velocity of 2 is employed. In passage through the reaction zone the dripolene plus hydrogen stream temperature is increased to 625 F., which provides an average reaction temperature of 370 F. In this zone, 625 standard cubic feet per barrel of hydrogen is consumed by olefin hydrogenation, a quantity which compares closely with the theoretical hydrogen consumption based on the observed experimental heat of reaction, 280 B.t.u./ lb. The quantity of catalyst in reaction zone 45 is that which provides a weight hourly space velocity of 2.0, i.e. 2.0 pounds of dripolene charged per hour for each pound of catalyst in Zone 45. The hydrogenated stream leaving chamber 45 passes through line 46, valved line 47, to cooler 48, and then through line 49 to gas-liquid separator 50.

The hydrogenated product stream comprising hydrogenated dripolene in vapor form together with excess hydrogen-containing gas is cooled in the cooler 48 wherein the hydrogenated dripolene condenses as a liquid which is sent, along with the non-condensible hydrogen-containing gas, to the gas-liquid separator 50. At the gas-liquid separator 50, the hydrogen-containing gas is separated and withdrawn through line 51 and conducted to amine scrubber 53, where a descending stream of diethanolamine or other agent, from line 60, effective to absorb HZS is employed to remove hydrogen sulfide gas formed by the destructive hydrogenation of sulfur compounds in the dripolene charge or in the naphtha reformer feed and carried in the reformer make gas stream feed to the hydrogenation system via

lines

35 and 54. The amine is withdrawn via line 55, heated in a stripper, not shown, for the purpose of releasing absorbed H28 and recycled to the amine scrubber 53.

Depending upon the hydrogen sulfide concentration of the reformer hydrogen-containing gas, it may be added at either valved line 54, or valved line 43. Briefly, if the reformer gas is relatively low in hydrogen sulfide, it may be added to the system through valved line 43. The cornposition -of reformer gas varies with the operation of the reformer and may range for example from 70-95% hydrogen, the balance being saturated light hydrocarbons such as methane, ethane and propane. If this gas is of a purity below about it may be desirable 'to vent a portion of the gas from gas-liquid separator 50 through valved vent line 57 so as to prevent a build-up of noncondensible methane, ethane and propane within the recycle gas system.

Where large quantities of reformer gas are available, the present recycle gas system may be eliminated in favor of a once-through hydrogen ow. In this case, the entire hydrogenation reactor efuent may be passed directly via line 46, and valved line 69 into line 65 for introduction into the reformer process stream.

After treatment in amine scrubber 53, the essentially hydrogen-sulfide-free hydrogen-containing gas, .e. containing less than 12 grains of HZS per 100 standard cubic feet, is conducted via line 58 to water -scrubber 59 where a descending stream of water from line 61 scrubs entrained or vaporized amine from the gas. The rich water stream is withdrawn through line 62 and is concentrated for amine recovery in a distillation column, not shown. If desired, water vapor removal facilities such as a glycol scrubbing tower or a silica gel or alumina drier may follow water scrubber 59 in line 63.

The treated gas passes from water scrubber 59 through

lines

63 and 43 back to the juncture 42 with dripolene charge line 41 and thence to reaction Zone 45.

Returning now to receiver 50, hydrogenated dripolene as a liquid condensate passes through

lines

64, 65 and any one or more of

valved lines

66, 67, and 68 into the reformer process stream in one or more of

lines

18, 21 and 25.

The average pressure on the reforming system is 300 p.s.i.g., the space velocity is 1.0 weight of fresh feed per hour per weight of catalyst, and the recycle rate is 5,000 s.c.f. of hydrogen per barrel of feed. The feed is preheated to a temperature of 900 F. in heater 11, and the recycle hydrogen is heated to a temperature of 1025 F. in heater 34, providing a reactor inlet temperature of 940 F. The outlet temperature of reactor 14 is 800 F. The reaction mixture is reheated in interheater 16 to a temperature to provide an inlet temperature to reactor 17 of 940 F. The outlet temperature is 880 F. In similar fashion, the feed streams to the remaining reactors are reheated to obtain inlet temperatures of 940 F. The outlet temperatures for

reactors

20, 24 and 28 are, respectively, `910", 925 and 930 F.

The effluent from reactor 28 is cooled to obtain a temperature of 100-120 F. in gas liquid separator 31 at 280 p.s.i.g. The octane number of the usual (35+ reformate obtained from the South Texas charge naphtha is 100 research octane clear under the above reaction conditions. This reformate may be withdrawn via valved line 70 for utilization as a high octane gasoline blending component. Alternatively, when high purity benzene is to be recovered, the reformate from the separator 31 is fed to fractionator 71 wherein a heavy reformate having an initial boiling point of about 230 F. and having a research octane number of about 108 research octane clear, comprising about 54 volume percent of the product is recovered as bottoms from fractionator 71 via line 72. The benzene-containing overhead fraction from fractionator 71 is passed via line 73 to a solvent extraction tower 74 wherein the benzene-containing overhead fraction is contacted with diethyleneglycol solvent at a solvent to oil ratio of 6:1, 300 F. and 150 p.s.i.g. The rich solvent stream containing the benzene is withdrawn from the bottom of the extraction tower 74 and passed via line 75 into solvent stripper 76 wherein benzene is distilled overhead from the solvent and withdrawn via line 77. A lean solvent is withdrawn from the bottom of stripper 76 via line 78 and recycled to the extraction tower 74 via line 79. Raflinate is withdrawn from the top of extraction tower 74 via line 80. The raffinate may be withdrawn from the system via valved line 81, or alternatively, the rafnate may be recycled to the reforming process via valved line 82, line 65 and any one or more of

valved lines

66, 67 and 68.

In the operation of the invention, three or more reactors can be used in the reforming system. In the case of three reactors, the first reactor constitutes the dehydrogenation stage. In the `case of a greater number of reactors, the first two reactors will usually constitute the dehydrogenation stage. The conditions in the dehydrogenation stage approximate 750 to 900 F. average -temperature, with a space velocity in this first stage, based on naphtha feed, of about 3 to 10 WHSV. The temperature in the subsequent stage will approximate 900 to l,000 F. The pressure may be in the range of 100 to 500 p.s.i.g., preferably 15 0-300, and the hydrogen recycle rate in the range of about 2,000 to 10,000 s.c.f. per barrel.

The reforming catalyst may comprise any of the platinum-type reforming catalysts, preferably on an alumina type base, although other supports, such as deactivated silica-alumina, alumina-titania, and the like, may he used. The presence of chlorine or iluorine, in known manner, may be desired in the reforming zone.

Although the drawing illustrates only the hydrocarbon flow, it will be understood that the system should be equipped for catalyst regeneration and/or rejuvenation. In the regeneration step, carbon is burned off the partially deactivated catalyst with -a dilute oxygen containing gas. Higher oxygen partial pressures and severities are used in rejuvenation of more severely deactivated catalysts. The regeneration may be effected periodically in blockedout operations, or it may be effected in the manner of ultraforming by use of a swing reactor as has been described in the technical literature.

The cut point `in the reformate splitter depends sorne what upon the severity of reforming, the feed stock and the desired heavy reformate octane. With C7+ charge naphthas at severity levels producing octane product, the C7 aromatics should be included in the heavy reformate. An initial boiling point in the range of about 225 to 250 F. lis recommended. At lower severities or with more refractory feeds, the initial of the heavy reformate may be in the range of about 250 to 275 F., however, this will require an additional distillation `step to separate benzene and toluene recovered from the extract.

A variety of extractive agents can he used inthe solvent extraction for treating the light reformate. Polyhydroxy solvents such as diethyleneglycol, dipropyleneglycol, triethyleneglycol, or mixtures thereof, `advantageously promoted in selectivity by the addition of water, `are particularly suitable. Other useful solvents are described in U.S. Patent 2,365,517. A newer solvent, butyrolacetone, has certain advantages for processing reformates. Sulfur dioxide also is feasible although it requires added facilities for refrigerated handling of the solvent. Usually, the feed and solvent will be contacted countercurrently in one or more extraction columns of the number of theoretical extraction stages required to effect the degree of separation desired. The conditions of extraction will be determined vby the nature of the solvent and its selectivity for aromatics at various temperature conditions. Usually, selectivity is improved with decreasing temperatures, and temperatures in the range of say -40 F. to 300 F. or more may be used, with adjustment of pressure to obtain the desired phase separation, at solvent to feed ratios advantageously in the range of l/ 1 to 25/1. Various methods may be used to separate extract and solvent, but, in general, distillation is most satisfactory. Traces of solvent can he removed from the separated rainate and extract phases by washing with water or other solvents, or by stripping.

The invention has the advantage of improving reformate yield at any given severity yby introducing greater selectivity into the conduct of the various reforming reactions and of producing high quality, lowrsulfur benzene. The most difficult reforming reaction, cyclization of parans, is promoted in rate fby decreasing the concentration of naphthenes in the subsequent reaction zone. Thus, injection of the hydrogenated dripolene has the additional advantage of reducing the effective concentration of naphthenes to a very low level. Also, more selective handling of the hydrocarbons in the feed is possible since greater advantage is taken of the different conditions obtaining through a series of reforming reactors by injecting the hydrogenated dripolene and, if desired, the paraifnic raffinate to the latter reactors.

I claim:

1. A method of producing high quality benzene from dripolene which comprises hydrogenating dripolene, thereafter charging hydrogenated dripolene to a naphtha reforming process comprising dehydrogenation, isomerization, dehydrocyclization and cracking reactions subsequent the dehydrogenation reaction, withdrawing reformate, and recovering high quality benzene therefrom.

2. A method Iof producing benzene and reformate from olerinic benzene-containing material boiling in the gasoline boiling range and naphtha which comprises hydrogenating said benzene-containing material to reduce the olefin and diolefin content thereof, charging said naphtha and hydrogen to the dehydrogenation stage of a reformer, charging said hydrogenated benzene-containing material and eluent from said dehydrogenation stage to a subsequent stage of said reformer and recovering benzene and reformate from the effluent of said. reformer.

3. A method of producing benzene and reformate from oleflnic benzene-containing material boiling in the gasoline boiling range and naphtha which comprises hydrogenating said benzene-containing material to reduce the fouling tendency thereof, charging said naphtha and hydrogen to the dehydrogenation stage of a reformer, charging said ydrogenated benzene-containing material and efuent rom said dehydrogenation stage to a subsequent stage of aid reformer, recovering benzene, reformate and hydroen-rich gas from the eiuent from said reformer, recycling portion of said recovered hydrogen-rich gas to the deiydrogenation stage of said reformer to provide said hylrogen, and employing another portion of said recovered iydrogen-rich gas in said hydrogenation.

4. An improved method of concurrently reforming iaphtha and hydrogenated dripolene which comprises hylrogenating said dripolene to produce hydrogenatcd drip- )lene, charging said naphtha and hydrogen to the first stage of a reformer, charging both of said hydrogenated :lripolene and hydrocarbon eiuent from said rst stage to :he second stage of said reformer, and recovering reformate from the efuent of said second stage, said reformate having a higher yield and/or octane rating than if said dripolene were fed to the rst stage of said reformer in the Conventional manner.

5. An improved method of concurrently reforming naphtha and hydrogenated dripolene which comprises separately desulfurizing said naphtha to produce desulfurized naphtha and hydrogenating said dripolene to produce hydrogenated dripolene, charging said desulfurized naphtha to the dehydrogenation stage of a multi-stage reformer, charging both of said hydrogenated dripolene and effluent from said dehydrogenation stage to a subsequent stage of said reformer, and recovering reformate from the effluent of said subsequent stage.

6. An improved process for producing benzene and reformate from naphtha and dripolene which comprises hydrogenating said dripolene to produce hydrogenated dripolene, charging said naphtha to the first stage of a multi-stage reformer, charging said hydrogenated dripolene and effluent from said rst stage to the second stage of said reformer, recovering benzene from the efiluent of said second stage and recovering at least a portion of the remainder of said second stage efuent as reformate, whereby the yield of benzene is higher than when said hydrogenated dripolene is charged to said reforming rst stage.

7. The method of claim 1 wherein said reforming process employs a platinum-alumina catalyst.

8. The method of claim 1 wherein said hydrogenating is carried out in the presence of a platinum-alumina catalyst.

9. The method of claim 1 wherein the hydrogen consumed by said hydrogenating is produced in said reformer.

References Cited by the Examiner UNITED STATES PATENTS 2,865,837 12/1958 Holcomb et al. 208-138 2,953,612 9/1960 Haxton et al 208-144 DELBERT E. GANTZ, Primary Examiner.

A. RIMENS, Assistant Examiner'.