US3564067A - Process for treatment of coke oven light oil - Google Patents

Abstract

A CATALYTIC PROCESS IS PROVIDED FOR EFFECTING SELECTIVE HYDROGENATION AND HYDROCRACKING OF COKE OVEN LIGHT OIL CONTAINING 10 TO 50% PRIMARY OIL. HEATING THE LIGHT OIL WITH HOT RECYCLE AND THE HYDROGEN, AND IMMEDIATE INTRODUCTION INTO CONTACT WITH HYDROGENATION CATALYST AT UPFLOW LIQUID PHASE CONDITIONS SUBSTANTIALLY REDUCES THE COKE-FORMING TENDENCIES OF NONAROMATIC UNSATURATED COMPONENTS IN THE LIGHT OIL. SUBSEQUENT CATALYTIC HYDDRODEALKYLATION PRODUCES IMPROVED YIELDS OF HIGH PURITY AROMATICS. IN ONE EMBODIMENT PROVISION IS MADE FOR THE RECOVERY OF NAPHTHALENE.

Description

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' 7 ATTORNEYS United AStates Patent Office 3,564,067 PROCESS FOR TREATMENT OF COKE OVEN LIGHT OIL Walter Brenner, Wayne, and Louis C. Doelp, Jr., Glen Mills, Pa., assignors to Air Products and Chemicals, Inc., Philadelphia, Pa., a corporation of Delaware Filed May 6, 1968, Ser. No. 726,883

Int. Cl. Cg 9/16, 7/00; C07c 3/58 U.S. Cl. 260-672 10 Claims ABSTRACT OF THE DISCLOSURE BACKGROUND OF THE INVENTION The present invention is concerned with purification and conversion of coke oven light oils to obtain therefrom aromatic hydrocarbons of enhanced value and in substantially pure form.

The light oils produced as by-products in coke oven plants contain chiefly benzene and other benzene-type hydrocarbons in relatively high proportions; however, these light oils are contaminated with a wide variety of hydrocarbonaceous materials in small amounts including paraflins, olefins and naphthenes as well as significant quantities of sulfur-containing compounds. The aromatic hydrocarbons being of higher commercial value are preferably recovering in high yield and in high purity in order to realize the full value thereof. Commercial processes are in successful operation for the separation and recovery of purified aromatic portions from so-called secondary light oils.

In conventional practice the light oil separated from coke-oven gas is subjected to rectification to separate out an overhead fraction comprising the secondary light oil and there is also recovered a bottoms fraction of so-called primary light oil, sometimes referred to as middle oil. This higher-boiling fraction contains a portion of the xylenes, naphthalene With or Without some methyl naphthalenes, and various hydrocarbons boiling largely in the range of from xylene to about 250 C. including thiophenes, styrenes and other unsaturates together wtih indene, coumarones and other easily polymerizable gumforming materials. Because of the coking tendencies of the naphthalenes and other bicyclics present in the promary light oil as well as of the gum-forming unsaturated hydrocarbons, hydrogenative treatment of these primary light oils over catalyst for recovery of purified aromatic hydrocarbons presents operational difficulties. For these reasons, recovery of aromatic hydrocarbons from coal tar oils, to the extent that is being practiced, is largely carried out by chemical treatment or to a lesser degree by thermal (non-catalytic) hydrotreating methods under conditions resulting in comparatively low yields of the desired aromatic hydrocarbons due to extensive cracking of hydrocarbons in the charge to gases and other low molecular weight compounds. Only incomplete removal of sulfur compounds is achieved by these methods.

3,564,067 Patented Feb. 16, 1971 SUMMARY OF THE INVENTION An object of the present invention, accordingly, is to provide an improved process for handling of the primary light oil components to obtain valuable aromatic hydrocarbons therefrom in economically attractive yields.

In accordance with the invention coke oven light oils containing at least 10% by volume of primary light oil including naphthalene and other bicyclic compounds as well as easily polymerizable gum-forming materials are pretreated with hydrogen over hydrogenative catalyst at processing conditions effecting selective hydrogenation and selective hydrocracking whereby subsequent processing to high yields of aromatics is effected with substantially reduced coke formation. Particularly beneficial results are obtained with charge stocks containing l0 to 50% primary light oil and 90 to 50% secondary light oil. Under the selected conditions some of the unsaturates, particularly readily polymerizable components, are hydrogenated. In one embodiment, the process conditions are controlled, particularly as to temperature and hydrogen partial pressure, so as to favor saturation of only one of the rings of the bicyclic components present in the feed. The hydrogenated eluent is distilled to remove (I) as bottoms, a higher boiling fraction containing substantially all unconverted naphthalene, and (II) a lower boiling overhead fraction, including the monocyclic hydrocarbons. At least a portion of the higher boiling fraction (I) is reheated and recycled to the bottom of the hydrogenation reactor, thus supplying heat to the entering fresh light oil charge and bringing the total charge (fresh plus recycle) to desired reaction temperature. The monocyclic fraction (Il) is treated to effect dealkylation of alkyl aromatics, hydrodesulferization and selective catalytic hydrocracking for recovery of enhanced amounts of extremely pure benzene with or without separate recovery of toluene and/or Xylenes.

In another embodiment of the present invention, process conditions are such as to favor retention of naphthalene such that the higher boiling fraction (I) is fractionated to obtain a naphthalene heart cut. The remaining portion of fraction (I) is then reheated and recycled to the bottom of the hydrogenation reactor.

An important aspect of the present invention lies in the novel arrangement for handling of the fresh feed containing unsaturates which readily polymerize to gums and tars on heating or prolonged storage. To avoid polymer deposition in the lines and on hot Contact surfaces the hydrogenative pretreatment is effected at conditions such that the fresh feed is in substantially liquid phase and moves in upward -flow through the bed of hydrogenation catalyst, the charge being brought from ambient temperature to desired inlet temperature by contact with preheated recycle bottoms stream (I) from fractionation of the hydrogenated efliuent, such recycle stream being in a ratio in excess of one volume per volume of fresh feed. The maintenance of the normally liquid portion of the charge in substantially liquid phase coupled with upow operation appear as essential features for successful extended performance. Nominally similar amounts of reactants at similar conditions provide a similar degree of liquid phase in downow operation; however, the percent liquid occupancy at the described conditions is effectively higher at upflow conditions. It is desirable, therefore, to select conditions related to the specific charge stock which will provide the lowest practical reaction temperature, the highest practical pressure consistent with the desired result and sound economics, and the lowest effective quantity of added hydrogen.

The quantity of hydrogen supplied to the pretreating reactor should be in excess of the stoichiometric requirement for saturating oleins and oleiinic side chains in the feed. In preferred practice, the amount of hydrogen supplied is less than one mole of hydrogen per mole of fresh feed and is generally in the order of about 0.5 mole of hydrogen per mole of fresh feed for a charge comprising to 50 percent by volume of primary light oil. When benzene is the main desired product, the operating conditions are relatively severe, involving temperatures above about 285 C. (285-370 C.) where hydrogenation of monocyclic aromatics tends to occur, and the added hydrogen should be below one mole per mole of fresh feed. When less severe operating conditions are employed, such as in the embodiment concerned with the separation and recovery of significant quantities of naphthalene and particularly when higher concentrations of primary light oil are present in the fresh feed, e.g., with hydrogenation temperatures between 250 C. and 285 C., it is possible to add as much as up to 2 moles of hydrogen per mole of fresh feed. In this latter instance the driving force to complete hydrogenation may be less; but larger quantities of hydrogen-consuming components, e.g., up to 40% indenes, naturally can account for larger amounts of added hydrogen.

BRIEF DESCRIPTION OF THE DRAWINGS The operation of the invention and certain of the advantages thereof will be understood and appreciated from the description which follows read in connection with the accompanying drawings in which:

FIG. 1 is a `flow diagram illustrating the treatment of coke oven light oil for the recovery of puried monocyclic aromatic hydrocarbons: and

FIG. 2 is a flow diagram of an embodiment illustrating the treatment of coke oven light oil wherein naphthalene is recovered in addition to other purified aromatic hydrocarbons.

DESCRIPTION OF THE PREFERRED EMBODIMENTS Referring now to FIG. l, the fresh feed of coke oven light oil at ambient temperature is introduced through conduit 10 leading to the bottom of hydrogenation reactor 11. This fresh feed includes the usual secondary light oil and up to about 50% by volume of higher boiling components, including naphthalenes, in the primary light oil boiling range. A heavier hot oil containing added hydrogen is introduced through conduit 12-the source of this heavier oil will be hereinafter explained-for intimate admixture with the fresh feed in conduit 10; and the admixture of fresh feed, heavier hot oil and hydrogen is introduced into the bottom of reactor 11. Thus, dilution of the fresh charge, establishment of the proper reaction temperature and commencement of the reaction occur substantially simultaneously.

Reactor 11 contains sulfided cobalt molybdate catalyst supported on alumina. The preferred catalyst is one containing, prior to sulfidation, 10 to 20%' by weight of the oxides of cobalt and molybdenum; the M003 being from 3 to 5 times that of the COO by weight. Suldation of the catalyst can be effected by pretreatment with H28. However, common practice with the sulfur-bearing stocks employed in this invention involves simply allowing the sulfur in the charge to effect the sulfidation of the catalyst. A typical cobalt molybdate catalyst for use in the hydrogenation reactor is that described in Example I of U.S. Pat. No. 3,207,802. Catalysts containing small amounts of other metal oxides or sulfides of the iron group, particularly nickel, in addition to cobalt may also be employed, such as catalysts of the type described in U.S. Pat. No. 2,880,171.

In reactor 11, under the selected process conditions, hydrogenation of polymerizable unsaturates, such as styrene, indene, and dicyclopentadienes, is effected. Saturation of one of the rings of the naphthalenes present to form tetrahydronaphthalene-type compounds may also occur. In addition, 50 to 80% of the ring sulfur compounds (thiophenes for example) are hydrogenated, facilitating subsequent desulfurization. A portion of these sulfur compounds may also be hydrocracked in this operation. The conditions employed in reactor 11 include temperatures in the range of about Z50-370 C. and pressures of about 50 to 100 atmospheres gauge. Since the desired selective hydrogenation requires only a short residence time, high oil throughput rates can be employed including fresh feed volume space rates (LHSV) in the order of 0.5 to 5.0. Preferred conditions are oil inlet temperatures of about 300 C. and pressure of 70 atmospheres at an hourly oil space rate (fresh feed only) of about l volume per volume of catalyst in the reactor. A volume space rate in the range of 0.8 to 20 yLHSV for the total oil feed is consistent with the described operation.

Under the conditions employed, a major portion of the oil will be in liquid phase during hydrogenation in the reactor 11 and the still liquid product as well as gases and vaporized products will be discharged at the top of the reactor through conduit 13. The mixed-phase effluent, preferably after cooling to about 60 to 90 C. as indicated at 14 is sent to liquid-gas separator or flash drum 15. From the separator 15 there is withdrawn through conduit 16 a vapor overhead composed of hydrogen and light hydrocarbon gases up to about C5. The high pressure liquid from high pressure liquid-gas separator 15 passes through conduit 17 and pressure reducing valve 18 into a low pressure liquid-gas separator 19 from which low pressure gas is vented through conduit 20 and the gasfree liquid passes through conduit 21 to fractionator 22. Separator 19 is operated at any appropriate pressure of less than about 7 atmospheres gauge and preferably in the range of about atmospheric to about 3 atmospheres pressure. Suitable valving, liquid level controllers and related instrumentalities are employed as required.

IFractionator 22 is operated under conditions and at a. cut point to sparate out a liquid fraction, including naphthalenes, from the vapor overhead which includes monocyclic aromatcs. If the fractionator 22 is operated at substantially atmospheric pressure, this cut point will be at a level in the range of about C. to 120 C., so that at least the major portion of th indane and tetralin will be in the liquid fraction.

The vapor overhead from fractionator 22 is withdrawn through conduit 23 and passes through condenser 24 in which essentially all of the C5+ components are condensed to liquid form. The effluent from condenser 24 is separated in a liquid-gas separator 25 with uncondensed vapors vented, or otherwise disposed of, through conduit 26. A portion of the condensate can be returned from the separator through conduit 27 as liquid reflux to fractionator 22 while the remaining portion of the condensate is passed from the separator through conduit 28 to pressurizing pump 29. The pressurized liquid from pump 29 is then passed through conduit 30 and combined with the vapor overhead from separator 15 in conduit 16 for further processing, as will be further described below.

The liquid fraction from fractionator 22 is withdrawn through conduit 31 and passes through pump 32 into conduit 33. Any excess liquid fraction beyond the requirements for recycle use may be withdrawn from the system `as from line 31 by way of valved outlet 50. To assist in the operation of fractionator 22, a reboiler may be provided at the bottom thereof. Hydrogen-containing gas is added via conduit 34 to the liquid fraction in conduit 33, and the mixture of oil and gas is then heated to a required temperature at 3S before being recycled through conduit 12 to the bottom of the pretreat reactor 11. Only a small amount of hydrogen need be added to the recycle oil stream, i.e., sufficient hydrogen to effect the saturation of polymerizable olefins and no more than one of the rings of naphthalene compounds contained in the fresh light oil feed admitted through conduit 10. As indicated above, as little as about 0.5 mole of hydrogen per mole of fresh oil charged is normally adequate. However, larger amounts of hydrogen can be employed, up to that amount which would tend to favor undesired hydrogenation of monocyclic aromatics and/or interfere with the substantially liquid phase operation under the conditions employed in reactor 11.

The' recycled oil admitted to reactor 11 provides the principal source of the heat required for initiation of the hydrogenation reaction in reactor 11. Accordingly, the stream in conduit 12 is brought to a temperature sul'licient to provide in the mixture thereof with the fresh feed from conduit the desired inlet temperature at the bottom of reactor 11. The fresh feed charged through conduit 10 contains unsaturated hydrocarbons which tend to polymerize at elevated temperature (e.g., at about 150 C.) and cause fouling by deposit of tarry and other polymerized substances on catalyst and other surfaces contacted thereby. This problem of fouling equipment and catalyst is overcome by heating the fresh feed directly and only with the hot stream from conduit 12 which is admixed therewith. Since the hot oil in conduit 12 has already Ibeen subjected to hydrogenation it is relatively free of troublesome unsaturates and provides an excellent diluent and wash oil for the fresh feed. Relatively long catalyst life and substantial freedom from fouling of equipment are obtained when the recycled hot oil in conduit 12 is employed at a ratio of at least 1 and preferably at about 3 volumes per volume of fresh feed. Higher recycle oil ratios of course can be employed but have no particular addedadvantage and tend unnecessarily to reduce the useful capacity of the reactor and other equipment. The use ofhot refractory oils in another relation for supplying heat directly to an oil charge containing polymerizable unsaturates is described in U.S. Letters Patent No. 3,216,924.

The vapor overhead from separator 15 passes through conduit 16 and is combined with the net overhead from fractionator 22 which passes through conduit 30. The combined material, free of polymerizable contaminants, passes through conduit 36 to the hydrodealkylation system shown diagrammatically at 37. Eiiiuent from the hydrodealkylation system 37 is withdrawn through line 49 for separation and recovery of hydrogen-containing gas and product aromatics by means not shown. The hydrogen-containing gas necessary to the hydrodealkylation reaction is introduced through conduit 38 and cornbined with the material in conduit 36. Typically, this hydrogen-containing gas comprises high pressure recycle gas, from the hydrodealkylation system 37, supplemented by such make-up hydrogen gas as may be desired or required. For example, the hydrogen-containing gas in conduit 38 should contain at least 70% free hydrogen in an amount in the range of 4 moles to 10 moles of hydrogen, usually about 6 moles, per mole of aromatic hydrocarbons in conduit 36.

In practice, the hydrodealkylation methods described in U.S. Letters Patent No. 3,081,259 may be utilized at 37. As therein described, the charge is introduced into at least one reactor together with hydrogen at an inlet temperature of about S90-630 C. and contacted at 30 to 70 atmospheres with high activity chromia-alumna catalyst so that during reaction 'temperatures of at least 630 C. are reached. Space rates are preferably employed such that under the reaction conditions the nominal residencetime of the material is less than three minutes. At these conditions sulfur-containing compounds are substantially completely converted as by hydrocracking, so that any sulfur is in the form of hydrogen sulfide; non-aromatic compounds are hydrocracked to light hydrocarbons; and a substantial portion of alkyl aromatics is hydrodealkylated to benzene. The effluent from the hydrodealkylation step is normally flashed to remove HZS, H2 and the lighter hydrocarbons. Such gaseous portion may be freed of HZS and then utilized with or without other purication as at least a portion of the hydrogen-containing gas 1ntroduced with the charge of hydrodealkylation. The nongaseous portion of the flashed-effluent is distilled and otherwise treated to effect separation and recovery of purified mononuclear aromatics, chiefly benzene.

The embodiment shown in FIG. 2 for the recovery of naphthalenc, as well as other puried aromatic hydrocarbons, is similar to that illustrated in FIG. 1. For the sake of simplicity, items which are identical structurally in FIGS. 1 and 2 have been designated with the same numeral. With the following exceptions the process steps and conditions employed in the embodiment illustrated by FIG. 2 are identical with those for the embodiment illustrated in FIG. 1.

Since naphthalene is recovered in the embodiment shown in FIG. 2, less severe process conditions are utilized in hydrogenation reactor 11 than in connection with the embodiment shown in FIG. 1. The desired selective hydrogenation of polymerizable unsaturates, without extensive saturation of naphthalene rings, can be effected by operating reactor 11 at a temperature in the range of 250285 C. and at a pressure of from about 50 to 100 atmospheres. Preferred conditions are oil inlet temperatures of about 260 C. and a pressure of 70 atmospheres at an hourly oil space rate (total oil including recycle) of 3 volumes per volume of catalyst in the reactor.

The actual recovery of naphthalene is effected by passing the liquid fraction withdrawn from fractionator 22 through conduit 39 to another fractionator 40. A naphthalene heart cut, e.g., boiling between 217-219 C., is recovered from fractionator 40 in conduit 41 and further processed, by means such as a stripper tower, not shown, in a manner suitable for the separation and recovery of pure naphthalene. The overhead fraction from fractionator 40 is passed through line 42 to overhead condenser 43. Uncondensed vapor may be vented through line 44 and condensate removed through line 45 from which a suitable recycle stream passes through line 46 for re-entry to fractionator 40. The remaining portion of the condensate from line 45 passes through line 47 for combination with the 'bottoms fraction, from fractionator 40, passed through line 48. The combined streams from lines `47 and 48 are passed through line 31 for return as recycle to reactor 11. Valved outlet 50 from line 31 is provided to permit withdrawal of any liquid in excess of that required for recycle purposes.

A fuller understanding of the invention will be had from the following examples depicting the application thereof in several preferred embodiments. It is to be understood that these examples are for illustrative purposes only and are not intended as limiting.

lExample I Fresh coke oven light oil feed containing 20% by weight primary light oil and by weight secondary light oil was introduced at ambient temperature into the bottom of a hydrogenation reactor. The composition of this fresh feed is shown in Table l.

A heavier hot oil having the composition shown in Table 2 was also added in an amount of 3.3 volumes per volume of fresh coke oven light oil feed to the bottom of the hydrogenation reactor together with 0.55 mole of hydrogen per mole of fresh feed.

TABLE 2 Material: Wt. percent C11C12+ aromatics 91.23 Naphthalene 7.37 Indanes 0.50 Indenes 0.44 Benzene 0.30 Toluene 0.14 Xylenes 0.02

The fresh feed together with the added heavier hot oil and hydrogen were contacted with alumina-supported presulded cobalt molybdate catalyst, having a composition of about 82% by weight alumina, 3% by weight COO, 15% by weight M003 and a bulk density of 0.7, at an average temperature of 313 C., a pressure of 69 atmosphere absolute and a fresh feed liquid hourly space velocity of 0.9.

The mixed-phase effluent from the hydrogenation reactor was then sent to a liquid-gas separator. The gaseous overhead fraction was passed directly to a chromia-alumina catalyst reactor for hydrodealkylation as indicated below.

The liquid fraction from the liquid-gas separator was reduced in pressure to 2 atmospheres absolute and sent to a fractionating column, containing theoretical plates. The composition of the condensed overhead from the fractionating column is shown in Table 3.

TABLE 3 Material: Wt. percent Benzene 62.13 Toluene 15.02 Indanes 9.81 Xylenes 4.16 C9 to C11 aromatics 3.17 Ethyl benzene 1.95 Naphthalenes 1.33 C2-C6 parafns 1.17 Tetralin 1.04 Diphenyl 0.22

The liquid fraction from the fractionating column, having a composition shown in Table 2, was recycled and added to the fresh feed introduced to the hydrogenation reactor.

The condensed overhead from the fractionating column together with the gas from the liquid-gas separator were l combined with fresh hydrogen in a ratio of 1.4 moles of hydrogen per mole of aromatics, and 7.8 moles of recycle gas, containing a minimum of 70 mole percent hydrogen, (from the hydrodealkylation reactor) per mole of aromatics. The combined mixture was then contacted with a chromia on alumina catalyst at a temperature of 635 C. for a residence time of 36.9 seconds. The chromia on alumina catalyst was a commercially available catalyst containing a nominal 20% by weight Cr2O3. Similar catalysts having 15 to 25% by weight of Cr203 impregnated on alumina which has a surface area of 100 to 200 square meters per gram before impregnation may be ernployed with generally similar results.

Analysis of the resulting liquid product from the hydrodealkylation shows 91.50 wt. percent benzene, 6.37 wt. percent toluene and minor amounts of the materials shown in Table 4.

TABLE 4 Material: Wt. percent Diphenyl 1.05 Ethyl benzene 0.59 Naphthalene 0.24 Xylenes 0.12 C2 to C5 parains 0.10

C9 to C10 aromatics 0.03

Benzene was separated from non-benzene portions of the liquid product by fractionation to obtain 99.9 mole percent benzene.

As seen by this example very little naphthalene is in the vapor overhead obtained from the fractionating column and practically no benzene is recycled.

Example II Fresh coke oven light oil feed containing approximately 50% by weight primary light oil and 50% by weight secondary light oil constituted the fresh feed portion of the charge to the hydrogenation reactor. This portion of the charge, boiling in the range of 60.5 C. to 260 C., had the composition shown in Table 5.

TABLE 5 Material: Wt. percent Benzene 38.39

Toluene 7.79 Xylenes 2.81 Ethyl benzene 0.07 Styrene 1.14 C9+ aromatics 8.34 Coumarone 1.50

Indanes 0.37

Indenes 16.28

Tetralin 1.04 Naphthalene 19.96 Methyl naphthalenes 1.19 Benzothiophenes 0.79 Thiophene 0.29 Other 0.04

To this fresh feed portion of the charge there is added, in a ratio of 1.2 to 1 volume of fresh feed, a stream of recycle oil, from the naphthalene separator, to which is added fresh hydrogen in an amount equivalent to 1.65 moles per mole of fresh feed. The recycle oil has the composition shown in Table 6.

TABLE 6 Material: Wt. percent C6C8 aromatics 2.07 09+ aromatics 20.43 Indanes 67.52

Tetralin 4.99 Methyl naphthalene 4.42 Benzothiophenes 0.57

The fresh feed portion of the charge was introduced at ambient temperature and raised to the reactor inlet temperature of about 260 C. by admixture with the heated recycle oil-hydrogen stream. The admixed stream was directly introduced to the upllow reactor into contact with an alumina-supported sulded cobalt molybdate catalyst similar to that described in Example I.

The reaction conditions included an operating pressure of 69 atmospheres absolute, a liquid hourly space velocity of 0.9, based on fresh feed, and the relatively low average temperature of about 264 C.

The effluent from the hydrogenation reactor was separated in a flashing operation into a vapor overhead, forming part of the charge to a subsequent hydrodealkylation stage or other use, and a liquid bottoms. The liquid portion amounted to 98.25% by weight of the total charge to the hydrogenation reactor. After pressure reduction to about 33 p.s.i.a. this liquid portion was distilled in a fractionation column containing 25 theoretical plates into an overhead fraction amounting to 39.11% by weight of total charge suitable as charge to a subsequent hydrodealkylation stage, and a bottoms fraction amounting to 59.14% by weight of the total charge to the hydrogenation stage.

The bottoms fraction was distilled in a second fractionation in a high eiciency fractionation column. An overhead fraction, comprising mainly C7-C9 aromatics, and a bottoms fraction were removed and recombined to -form the above-mentioned recycle oil used as a portion of the charge to the hydrogenation reaction. There was also separated and recovered an intermediate fraction containing substantially only the naphthalene in an amount equivalent to 12.54 Weight percent of the fresh feed to the hydrogenation reactor.

- The condensed overhead from the first fractionator is combined with the gas from the high pressure flash, hydrogen is added and the stream is then processed in the hydrodealkylation operation as described in Example I. Benzene, toluene and xylene ,(BTX) are recovered as low sulfur product having purity better than 99.9 mole percent.

In accordance with operation of this example desired aromatics are recovered in amounts, based on the fresh feed, of 12.54 :weight percent naphthalene and 76.96 percent by weight as BTX.

Thus, the present invention provides a process for the purification of coke oven light oil contaminated with nonaromatic hydrocarbons. The present invention is not only able to handle charge stocks which cannot be processed by other catalytic systems for any practical period of on stream time, but is able to obtain an aromatics recovery which is far in excess of that obtained from similar stocks treated by either chemical or thermal means. The recovery of low sulfur containing products is another unique feature of the present invention. This process utilizes partial product recycle as diluent for, and the preheating of, fresh coke oven light oil immediately prior to subjecting the charge stock to hydrogenation in an upflow reactor operated under substantially liquid phase conditions. Selective fractionation in the process prevents any significant amount of bicyclic aromatics from reaching the final hydrodealkylation system. Selective hydrogenation and hydrocracking of naphthalene and other bicyclics is achieved in one embodiment of the process. In another embodiment, naphthalene is recovered as a product by employing somewhat milder operating conditions for hydrogenating any unsaturates with substantially less hydrogenation and hydrocracking of the naphthalene and using a second fractionator to obtain a naphthalene heart cut.

What is claimed is:

1. The process for the recovery of aromatic hydrocarbons from coke oven light oil containing at least by volume of primary light oil, which comprises:

(a) admixing the coke oven light oil at ambient temperature and a hot stream comprising a recycle liquid fraction having an initial boiling point in the range of 100-120 C. in an amount in the range of 1 to 3 volumes per volume of light oil and hydrogen in an amount in the range of 0.5 to 2.0 mols of hydrogen per mole of light oil and obtaining an admixture having a temperature in the range of from about 250-300 C.

(b) introducing said admixture under substantially liquid phase conditions into an upilow hydrogenation reactor into contact rwith hydrogenation catalyst and effecting hydrogenation of readily polymerizable components,

(c) subjecting a mixed phase eiuent from the hydrogenation reactor to liquid-gas separation into a vapor overhead fraction and a liquid fraction,

(d) fractionating the liquid fraction into a lower boiling fraction having an end boiling point in the range of about 100-120 C. and a higher boiling liquid fraction,

`(e) recycling at least a portion of the higher boiling liquid fraction as said recycle liquid fraction,

(f) admixing at least a portion of the vapor overhead fraction from (c), at least a major portion of the lower boiling fraction from (d), and hydrogen to form an hydrodealkylation charge stream,

|( g) subjecting the hydrodealkylation charge stream to hydrodealkylation at hydrodealkylation conditions 10 including contact with chromia-alumina catalyst, and

(h) recovering substantially pure aromatic hydrocar- =bons from the effluent from hydrodealkylation.

2. The process of claim 1 in which the coke oven light oil contains l0 to 50 percent by Volume of primary light oil.

3. The process of claim 1 in which the coke oven light oil, recycle liquid fraction and hydrogen-containing gas admixture is contacted with cobalt molybdate catalyst in the hydrogenation reactor at a temperature in the range of about 250 to about 370 C., a pressure of 50 to atmospheres gauge and a space rate of between 0.5 and 5.() LHSV, based on the coke oven light oil.

4. The process of claim 3 in which the cobalt molybdate catalyst is a sulfided cobalt molybdate catalyst supported on alumina containing, prior to sultidation, 10 to 20% by weight of the oxides of cobalt and molybdenum and in which M003 is present in an amount 3 to 5 times that of the COO by weight.

5. The process of claim 1 in which the mixed phase eluent from the hydrogenation reactor is cooled to a temperature between 60 and 90 C., before being subjected to liquid-gas separation.

6. The process of claim 1 in which the fractionation in (d) is effected at a pressure in the range of 1 to 3 atmospheres and with a cut point in the range of about 100 to about 120 C. at atmospheric pressure.

7. The process of claim 1, in which the hydrodealkylation conditions include a temperature in the range of 590 to 630 C. and a pressure in the range of 30 to 70 atmospheres gauge with resulting dealkylation of alkyl aromatics to aromatic hydrocarbons; and fractionating the effluent from hydrodealkylation to separate and recover the aromatic hydrocarbons, including benzene.

8. The process of claim 1 in which naphthalene is recovered as product by taking a heart cut of a fraction boiling in the range of between 217 and 219 C. from a second fractionation of the higher boiling fraction in (d).

9. The process for the recovery of high purity aromatic hydrocarbons, including; benzene and naphthalene, from coke oven light oil containing 10 to 50 percent by volume of primary light oil, which comprises: (a) admixing the coke oven light oil having an ambient temperature with a hot stream of recycle liquid and hydrogen-containing gas, said recycle liquid being employed in an amount in the range of 1-3 volumes per volume of coke oven light oil and the hydrogen in said hydrogen-containing gas being employed in an amount in the range of 0.5 to 2.0 moles of hydrogen per mole of coke oven light oil, said admixture having a temperature in the range of 250 to 285 C. obtained with said hot stream being the direct and only source of heating for the coke oven light oil;

(b) introducing the admixture into an upow hydrogenation reactor into contact with hydrogenation catalyst at a temperature between 250 and 285 C., a pressure between 50 and 100 atmospheres gauge and a space rate based on the coke oven light oil of between 0.5 and 5.0 LHSV, said contacting being effected with at least said coke oven light oil being in substantially liquid phase;

(c) subjecting the mixed phase effluent from the hydrogenation reactor to liquid-gas separation to obtain a vapor overhead fraction and a liquid fraction;

(d) fractionating the liquid fraction into a lower boiling fraction and a higher boiling fraction, said fractionating having a cut point equivalent to between about 100 and about 120 C. at atmospheric pressure;

(e) subjecting the higher boiling fraction to further fractionationto separate out and recover a heart cut boiling in the range of between 217 and 219 C. and comprising mostly naphthalene;

(f) recombining the fractions, minus said heart cut,

11 from said further fractionation and forming a recycle liquid;

(g) admixing said recycle liquid and hydrogen-containing gas and heating said last mentioned admixture to provide said hot stream in (a);

(h) combining said vapor overhead fraction of (c) and said lower boiling fraction of (d) with hydrogen to form a hydrodealkylation charge stream;

(i) contacting the hydrodealkylation charge stream with a dealkylation catalyst in a hydrodealkylation 1 reactor at hydrodealkylation conditions including a temperature between about 590 and 630 C. and a pressure between 30 and 70 atmospheres gauge; and (j) recovering aromatics, including high purity benzene, from the hydrodealkylation reactor eflluent. 10. The process of claim 1 in which said dealkylation catalyst is chromia on alumina and benzene is recovered from the reactor effluent by fractionation.

References Cited UNITED STATES PATENTS Donovan et al. 208-216 Butler et al. 208-255 Butler et al. 208-89X Maerker et al 260-674 Watkins 20S-143 McKinney et al. 208-143 Assistant Examiner U.S. Cl. X.R.

P34050 UNITED STATES PATENT OFFICE CERTIFICATE OF CORRECTION Patent No. 3,56#,O67 Dated February 16, 1971 lnventods) Walter Brenner and Louis C Doelp, Jr.

It is certified that erro: appears in the above-identified" patent and that said Letters Patent are hereby corrected as shown below:

r- Column l, line bf2, "recovering" should read --recovered Column 1, line 59, "promary" should read --primary Column 2, line 314, "hydrodesulferization" should read --hydrodesulfurization Column il, line 37, l'sparate" should read separate column u, line u2, "th" should read the-- Column 5, line T5, "of" should read to Column 7, line 19, "atmosphere:l should read atmospheres-- Column lO, line lll, delete semi-colon Column lO, line 68, "fraotonating" should read --fractiona Column l1, line 16, "claim l" should read --claim 9 Signed and sealed this 18th day of April 1972.

(SEAL) Attest:

EDWARDFLFLETCHERJR. ROBERT GOTTSCHALK At testing Officer Commissioner of Patents

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