US3647900A - Process for concentration and conversion of hydrocarbons - Google Patents

Abstract

NAPHTHALENE IS PRODUCED FROM HYDROCARBON FRACTIONS CONTAINING SUBSTANTIAL VOLUMES OF POLYCYCLIC AROMATICS, SUCH AS LIGHT CYCLE OILS FROM A CATALYTIC CRACKING UNIT, BY CONTACTING THE HYDROCARBON FEED WITH AN ORGANIC SULFOXIDE, SUCH AS DIMETHYLSULFOXIDE; SEPARATING AN EXTRACT PHASE AND A FIRST RAFFINATE PHASE FROM ONE ANOTHER; PASSING THE EXTRACT PHASE TO A VACUUM DISTALLATION COLUMN UNDER CONDITIONS TO FORM AZEOTROPIC MIXTURE OF A SECOND RAFFINATE PHASE AND THE SOLVENT AS AN OVERHEAD AND THE EXTRACT PHASE AS A BOTTOMS; SEPARATING THE AZEOTROPIC MIXTURE INTO A SECOND RAFFINATE PHASE AND A SOLVENT PHASE; COMBINING THE FIRST RAFFINATE PHASE AND THE SECOND RAFFINATE PHASE AND FROM A WATER-SOLVENT MIXTURE; SUBJECTING THE WATER-SOLVENT MIXTURE TO VACUUM DISTILLATION TO SEPARATE WATER AS AN OVERHEAD AND SOLVENT AS A BOTTOMS; RECYCLING THE SOLVENT FROM THE VACUUM DISTILLATION AND THE SOLVENT SEPA-

RATED FROM THE AZEOTROPE BACK TO THE EXTRACTION UNIT; AND SUBJECTING THE EXTRACT FROM THE AZEOTROPIC DISTILLATION UNIT TO HYDRODEALKLYATION TO PRODUCE A DEALKYLATED PRODUCT. IF DESIRED, A THIRD RAFFINATE PHASE CAN BE SEPARATED FROM THE ORIGINAL EXTRACT PHASE OF THE EXTRACTION STEP BY COOLING THE EXTRACT PHAST TO ESSENTIALLY AMBIENT TEMPERATURE. WHEN A FEED MATERIAL IS SUFFICIENTLY CONCENTRATED IN POLYCYCLIC AROMATICS, THE EXTRACTION STEP MAY BE ELIMINATED AND THE FEED CAN BE COMBINED WITH SOLVENT JUST PRIOR TO THE AZEOTROPIC DISTILLATION STAGE, OR A FEED OF LOW POLYCYCLIC AROMATICS MAY BE SUBJECTED TO EXTRACTION AND A FEED OF HIGHER POLYCYCLIC AROMATIC CONTENT MAY BE ADDED TO THE EXTRACT JUST PRIOR TO THE AZEOTROPIC DISTILLATION STAGE.

Description

March 7, 1972 G. J. ROZMAN PROCESS FOR CONCENTRATION AND CONVERSION OF HYDROCARBONS Filed July 17, 1968 5 Sheets-Sheet 1 March 7, 1972 G, J, ROZMAN v3,647,900

PROCESS FOR CONCENTRATION AND CONVERSION OF HYDROCARBONS Filed July 17, 1968 3 Sheets-Sheet 2 WLS 5 0f i 1 E D a 1 OO f 'i i 8 N 4 Lo J Lf) U D HONMMWBGOHGAH f rf) A mmm/mov f N Q Q m 0. 8 D LL no rf) E 2 3 E Q Lu m m 0 v LL 9 n :r 0

` Holm/un@ 5 s Q 5 E INVENTOR g N GEORGEJROZMAN G. J. RozMAN 3,647,900

PROCESS FOR-CONCENTRATION AND CONVERSION OF HYDROCARBONS March 7, 1972 5 Sheets-Sheet 3 Filed July 17, 1968 United States Patent O1 heel 3,647,900 Patented Mar. 7, 1972 3,647,900 PROCESS FOR CONCENTRATION AND CONVERSION OF HYDROCARBONS George J. Rozman, Ashland, Ky., assigner to Ashland Oil & Refining Company, Houston, Tex. Filed July 17, 1968, Ser. No. 745,516 Int. Cl. C07c 7/06, 7/10, 3/58 U.S. Cl. 260--674 R Claims ABSTRACT OF THE DISCLOSURE Naphthalene is produced from hydrocarbon fractions containing substantial volumes of polycyclic aromatics, such as light cycle oils from a catalytic cracking unit, by contacting the hydrocarbon feed with an organic sulfoxide, such as dimethylsulfoxide; separating an extract phase and a first rainate phase from one another; passing the extract phase to a vacuum distillation column under conditions to form an azeotropic mixture of a second rafnate phase and the solvent as an overhead and the extract phase as a bottoms; separating the azeotropic mixture into a second rainate phase and a solvent phase; combining the first raffinate phase and the second raffinate phase and mixing the same with water; separating the raiiinate phase from a water-solvent mixture; subjecting the water-solvent mixture to vacuum distillation to separate Water as an overhead and solvent as a bottoms; recycling the solvent from the vacuum distillation and the solvent separated from the azeotrope back to the extraction unit; and subjecting the extract from the azeotropic distillation unit to hydrodealkylation to produce a dealkylated product. If desired, a third rainate phase can be separated from the original extract phase of the extraction step by cooling the extract phase to essentially ambient temperature. When a feed material is suliciently concentrated in polycyclic aromatics, the extraction step may be eliminated and the feed can be combined with solvent just prior to the azeotropic distillation stage, or a feed of low polycyclic aromatics may be subjected to extraction and a feed of higher polycyclic aromatic content may be added to the extract just prior to the azeotropic distillation stage.

BACKGROUND OF THE INVENTION The present invention relates to the production of naphthalene from hydrocarbon feedstocks by forming a concentrate of polycyclic aromatics and thereafter hydrodealkylating the same. More specifically, the present invention relates to the production of a concentrate of polycyclic aromatics from petroleum feedstocks containing substantial volumes of polycyclic hydrocarbons by azeotropic distillation of the hydrocarbon feedstock with an organic sulfoxide and, if desired, the polycyclic hydrocarbons can then be subjected to hydrodealkylation to produce a dealkylated product.

DESCRIPTION OF THE PRIOR ART The commercial importance of naphthalene resides primarily in its use as an intermediate in the production of phthalic anhydride. For example, roughly 80% of all of the naphthalene produced domestically is consumed in the production of phthalic anhydride.

Heretofore, the primary source of naphthalene has been coal tar fractions. However, the uncertainties and fluctuations in the production of coal tars makes it undesirable to tie the production of naphthalene to variable sources of coal tars, particularly since the major portion of coal tar is produced as a by-product of the manufacture of coke for the production of steel.

Naphthalene does not exist in any great volumes in crude petroleum hydrocarbons, While the amounts of naphthalene in crude oils vary to some extent, the total of all bicyclic aromatic hydrocarbons in petroleum is usually about 5% to 30%. Accordingly, it is impractical to separate such naphthalene from the crude by simple distillation, since a number of other contaminating materials boil in the same boiling range. However, certain refined petroleum fractions, such as fractions obtained as products of catalytic reforming, catalytic cracking and thermal cracking, do contain significant quantities of naphthalene and alkyl-substituted naphthalenes to be of interest as feedstocks for further processing. Some feedstocks, and, particularly, products of reforming, which boil above 420 F., contain large quantities of naphthalenes and alkyl-substituted naphthalenes and only minor quantities of monocyclic aromatics and parailins. These feeds may, therefore, be directly processed in a hydrodealkylation unit to convert the alkyl naphthalenes to naphthalene. However, when attempting to process fractions containing smaller amounts of naphthalene precursors, it becomes impractical and prohibitively expensive, to directly hydrodealkylate this material since the consumption of hydrogen is extremely high due to the hydrogenation of large quantities of monocyclic aromatics and the hydrodealkylation catalysts are subject to rapid deactivation when utilizing a catalytic hydrodealkylation system.

While a number of techniques for concentrating naphthalene precursors have been tested in the past, and found useful to a greater or lesser extent, none have been found to economically produce substantial quantities of naphthalene precursors which, in turn, produce economic yields of naphthalene. However, a highly effective technique in accordance Awith the present invention has been the utilization of organic sulfoxides and, particularly dimethylsulfoxide, as a solvent for the separation of parains and monocyclic hydrocarbons from dicyclic of polycyclic hydrocarbons in the feed material. While organic sulfoxides have heretofore been suggested as a general solvent for separations of hydrocarbons, oxyorganic mixtures and the like, the commercial use of such solvents has been practically nil. There are believed to be two basic reasons for such lack of utility and these are primarily directed to the cost of the solvent. First, there has been no really effective means developed for recovery of the solvent and its re-use. Secondly, when the desired component is present in substantial amounts, solvent extraction in and of itself utilizes substantial volumes of solvent for only slight increases in the concentration of the desired component. Thus it would be most advantageous if the recovery of solvent could be improved while also improving the purity of the extract or the solvent extraction step could be eliminated completely.

SUMMARY OF THE INVENTION It is therefore an object of the present invention to provide an improved method for the production of naphthalene. Another object of the present invention is to provide an improved method for the production of naphthalene from petroleum hydrocarbon mixtures. A still further object of the present invention is to provide an improved method for the production of naphthalene from catalytic light cycle oils obtained from the catalytic cracking of petroleum hydrocarbons. Another and further object of the present invention is to provide an improved method for the production of naphthalene by solvent extracting a petroleum hydrocarbon mixture containing saturated hydrocarbons, monocyclic aromatics, and dicyclic aromatics to separate dicyclic aromatics therefrom and thereafter subjecting the dicyclic aromatic extract to catalytic hydrodealkylation. A still further object of the present invention is to provide an improved method for the production of naphthalene wherein a catalytic light cycle oil is subjected to solvent extraction with an organic sulfoxide, a raffinate phase and an extract phase containing the solvent are separated, the extract phase is separated into an extract and the solvent, the solvent is recycled to the extraction step and the extract is subjected to catalytic hydrodealkylation. Another object of the present invention is to provide an improved method for the dealkylated polycyclic hydrocarbons from mixtures of polycyclic, monocyclic and saturated hydrocarbons, wherein the mixture is subjected to solvent extraction with an organic sulfoxide, a raffinate phase and an extract phase containing the solvent are separated from one another, the extract phase is subjected to azeotropic distillation conditions to form an azeotropic overhead of solvent and a further raffinate phase, the solvent is separated from the second raffinate phase and recycled to the extraction step and the extract phase, separated as a bottoms in the azeotropic distillation, is subjected to catalytic hydrodealkylation. A further object of the present invention is to provide an improved method for the separation of organic cornpounds. Another and further object of the present invention is to provide an improved method for the separation of organic compounds utilizing an organic sulfoxide as an azeotrope former and subjecting the mixture to azeotropic distillation. Another and further object of the present invention is to provide an improved method for the recovery of polycyclic hydrocarbons from mixtures of polycyclic hydrocarbons, monocyclic hydrocarbons and saturated hydrocarbons, wherein the mixture is contacted with an organic sulfoxide as an azeotrope former, and the final mixture is then subjected to azeotropic distillation to form an azeotrope of the organic sulfoxide and non-polycyclic hydrocarbon and a bottoms fraction of polycyclic hydrocarbons. Yet another object of the present invention is to provide an improved technique for the separation of dealkylated polycyclic aromatics from mixtures of polycyclic aromatics, monocyclic aromatics and saturated hydrocarbons wherein the mixture is subjected to azeotropic distillation in the presence of an organic sulfoxide as an azeotrope former, and the polycyclic aromatics are collected as a bottoms fraction and subjected to hydrodealkylation to produce a dealkylated product.

These and other objects and advantages of the present invention will be apparent from the following description.

Briefly, in accordance with the present invention, organic compounds are subjected to azeotropic distillation in the presence of an organic sulfoxide as an azeotrope former.- More specifically, polycyclic aromatics are separated from mixtures containing polycyclic aromatics, monocyclic aromatics and paraflins by azeotropic distillation of the mixture in the presence of an organic sulfoxide, as an azeotrope former; the monocyclic aromatics and paraffins are recovered as an overhead azeotrope with the organic sulfoxide and polycyclic aromatics are recovered as a substantially pure bottoms product. The feed mixture to the azeotropic distillation step may be obtained by first extracting the organic mixture with an organic sulfoxide in a solvent extraction operation and utilizing the extract phase as a feed material. The bottoms product from the `azeotropic distillation step comprising substantial quantities of polycyclic aromatics may be subjected to hydrodealkylation in order to produce a polycyclic aromatic product, such as naphthalene, in substantially improved quantities.

The present invention will be more clearly understood by reference to the drawings when read in conjunction with the following detailed description.

4 BRIEF DESCRIPTION OF THE DRAWINGS In accordance with the present invention, the details of the present invention may be :best understood by reference to the drawings, wherein:

FIG. 1 is a flow diagram illustrating the operation of the present method;

FIG. 2 is a ow diagram illustrating a variation of the present method; and

FIG. 3 is a ternary phase diagram illustrating the selection of optimum operating conditions in the extractor of FIGS. l and 2.

DESCRIPTION OF THE PREFERED EMBODIMENTS In accordance with FIG. l, a hydrocarbon feed is introduced to the system through line 10 and a solvent through line 12. The hydrocarbon feed and solvent are contacted in extractor 14 to form a raffinate phase which is discharged through line 16 and an extract phase which is discharged through line 18. Extractor 14 is a conventional liquid-liquid solvent extraction unit. This extractor may comprise a single stage or a multiple stage unit. The extract phase in line 18, may, preferably, be cooled to essentially ambient temperature in cooler 20 and then passed to accumulator 22 where a second raffinate phase is separated and discharged through line 24 and the remaining extract phase containing solvent is discharged through line 26. The solvent and extract are heated in heater 28. The heated extract phase is passed through line 30 to azeotropic distillation unit 32. In azeotropic distillation unit 32, the extract is separated into an overhead fraction which is discharged through line 34. The overhead through line 34 is essentially pure, single-phase solvent since the bulk of the raffinate type materials were removed through cooling and are discharged through line 24. Hence the solvent in line 34, containing only a very small amount of oil, can be recycled directly to the extraction unit without further purification or separation. The extract, freed of solvent is discharged through line 36 as a bottoms fraction from unit 32. The extract is then mixed with hydrogen, supplied through line 38, in a hydrodealkylation unit 40. In hydrodealkylation unit 40, the materials are dealkylated to produce the corresponding non- `alkyl compounds which are discharged through line 42. The raffinate fraction from line 16 and the raffinate fraction from line 24 are combined and introduced to mixer 44. Water is introduced through line 46 and mixed with the raffinate. The raffinate-water mixture passes through line 48 to clarifier 50. Clarifier 50 is preferably a centrifuge operated under pressure in order to minimize solvent loss due to evaporation. The claried raffinate is discharged through line 52, and may be passed to dryer 54 for the removal of residual water before passage through line 56 to storage or other use. The Water-solvent mixture from clarifier 5t) is passed through line 58 to vacuum distillation unit 60. In vacuum distillation unit 60, the Water is separated as an overhead and discharged through line 62. Solvent is discharged as a bottoms fraction through line 64 and is thence recycled to the solvent extractor.

Where the hydrocarbon feed is more concentrated in polycyclic aromatics, the extraction phase may be eliminated completely. In this case, the hydrocarbon feed would be introduced through line 66, and the solvent material through line 68. In this instance, the overhead from vacuum distillation unit 32 will contain raffinate material in addition to solvent. However, a two-phase separation of these two materials can be effected by passing the overhead from line 34 through fvalve 70 and line 72 to an accumulator 74. In accumulator 74, the ratinate phase is separated from the solvent. The raflinate phase is discharged through line 76 and may be treated in the same manner as was previously described with respect to the ranates from lines 16 and 24. The solvent separated in accumulator 74 is discharged through line 78 and may then be recycled to line 68 for further use. As an intermediate-type operation, the extractor may be eliminated and a preliminary separation effected by cooling as previously described. In this instance, the hydrocarbon feed will be introduced through line 80 and solvent through line 82. The operation would be substantially the same with respect to the recovery of raffinate through line 24 and the processing of the extract and solvent through vacuum distillation unit 32. The only variation would, of course, be the passage of recovered solvent from line 78 through line 84 for re-use as an azeotrope former.

FIG. 2 of the drawings illustrates a variation of the process in which the cooling cycle is eliminated. This can be done if a fairly low temperature, approximately ambient, is produced at the extractor. In accordance with FIG. 2, hydrocarbon feed is introduced through line 110 and solvent through line 112 to an extraction unit 114. Rafnate is discharged through line 116 and extract through line 118. The extract is preheated in heater 120 and passed through line 122 to azeotropic distillation unit 124. In azeotropic distillation unit 124, the mixture is separated into an overhead fraction which is discharged through line 126 to an accumulator 128. In accumulator 128, a phase separation is effected to produce a rainate phase, which is discharged through line 130, and a solvent phase, which is discharged through line 132. The solvent phase is of sufficient purity that it may be recycled back to the extraction unit. The raffinate in line 130 is combined with the raiiinate from the extractor in line 116 and charged to a mixer 134. Water is mixed with the raffinate through line 136. The water-raffinate mixture is discharged through line 138 to a clarifier 140. Clarifier 140 may, of course, be a centrifuge, as previously suggested. In any event, a raffinate phase is separated and discharged through line 142 which raffinate phase may be dried in dryer 144 and thence passed to storage through line `146. The mixture of water and solvent is discharged from clarifier 140 through line 148. This mixture then passes to vacuum distillation unit 150. Vacuum distillation 150 separates the mixture into an overhead fraction comprising water which is discharged through line 152, and a bottoms fraction comprising solvent which is discharged through Iline 154 for recycle to the extraction unit. The extract from unit 124 is fed to reactor 156 with hydrogen from line 158 to hydrodealkylate the material and the dealkylated product may be discharged through line 160.

In a manner similar to that discussed with respect to FIG. 1, the extractor 114 may be eliminated where the hydrocarbon feed is of sutiicient concentration (extract component) to 4make Isuch elimination practical. The hydrocarbon feed would be introduced to the system through line 162 and the solvent through line 164. Processing through azeotropic distillation unit 124 would be the same except that recovered solvent in line 132 and in line 154 would be recycled through line 166 for reuse.

The hydrocarbon feed material, in accordance with the present invention, is a highly aromatic hydrocarbon liquid containing saturated hydrocarbons, monocyclic aromatics and polycyclic aromatics. More specifically, the feed material is a hydrocarbon mixture containing between about 20 to 40% or more of naphthalene precursors. Even though these naphthalene precursors appear to be present in rather substantial amounts, it has been found, in accordance with the present invention, that such feedstocks cannot be dealkylyated to naphthalene directly but must be Isubjected to purification or concentration of the polycyclic aromatics prior to dealkylation. By such purification, it has been found possible to produce a highly concentrated dealkylation feedstock and thereby substantially reduce the hydrogen requirements in the dealkylation unit and produce higher concentrations of naphthalene per unit capacity of the dealkylation unit. A specific feedstock falling within this definition is a light cycle oil obtained by the catalyic conversion of petroleum hydrocarbon oils through contact with an acidic catalyst, for instance, a

silica-alumina catalyst. A heart-cut of light cycle oil, boiling between about 400 and y600 F. can be utilized as a feedstock for the present process. However, this heart-cut should preferably boil between about 430 and 530 iF. The properties of such a light cycle oil heart-cut are illustrated inthe following Table I.

It should also be recognized that other organic mixtures such as naphthenes and parailins; oxyorganic compounds, such as orthoand paranitrophenol; organic acids, such as acetic, propionic, lactic, etc.; organic substances containing water, such as aqueous alcohols; etc. are separable by the azeotropic distillation taught herein. In short, any mixture selectively separable by solvent extraction with an organic sulfoxide is separable by the present method. It is to be observed, however, that the azeotrope forms with the component or components generally rejected as raffinate in a conventional solvent extraction.

TABLE L-PROPERTIES OF JIIJQFHT CYCLE OIL-HEART- ASTH distillation COC Fire, F FIAP I1STM Dl3l9) LV, percent:

By selecting a heart-cut, it is obvious that substantial amounts of paraflins and materials boiling below methyl naphthalene are removed. These materials are primarily saturates and monocyclic aromatics and are not desired. Materials boiling above dimethylnaphthalene, such as acenaphthene and other tricyclic fused ring aromatics, as well as tri- `and higher alkylated naphthalenes are also minimized by heart-cutting. The' light Icycle oil in Table I shows a typical 35.4% acenaphthene and higher boiling content, but this could obviously be reduced to an 8-l2% level or lower by better fractionation which would make the feedstock more valuable.

The solvent utilized in accordance with the present invention is an organic sulfoxide of the general formula wherein R1 and R2 are organic radicals, such as aliphatic, alicyclic, Varomatic or mixed hydrocarbon radicals or organic radicals containing a polar grouping, specifically oxygen, sulfur, nitrogen, halide and/or possibly allied atoms. R1 and R2 may be joined together to form a heterocyclic ring, such as tetrahydro and dihydro-l-thiophene oxide, and their derivatives wherein one or more of the hydrogen `atoms on the ring may be replaced by an organic radical of any of the types mentioned. The preferred organic sulfoxide from the standpoint of availability and cost is dimethylsulfoxide, often referred to as DMSO. Dimethylsulfoxide is a water-white, highly polar solvent which is completely miscible with water and very hygroscopic. Accordingly, care should be taken to prevent contamination of the dimethylsulfoxide by water from the Iatmosphere. Accordingly, the dimethylsulfoxide charged to the extraction unit should be essentially anhydrous dimethylsulfoxide. In this instance, the term anhydrous is meant to include dimethylsulfoxide containing less than about 1% water.

The ratio of dimethylsulfoxide to hydrocarbon feed to be utilized in accordance with the present invention should be between about 0.03 and 5.0 to 1. A preferred range is between 0.5 and 2.0 to 1 and a ratio of 1.5 to 1 has been found optimum. Since the constituents of the light cycle oil or other highly aromatic feed material may differ, the relative ratio of solvent to hydrocarbon feed cannot be precisely set forth in raw numbers. Therefore, the operable as well as the optimum ratio can be readily determined by the use of ternary phase diagrams such as that illustrated in FIG. 3 of the drawings.

A ternary phase diagram is a binodal curve plotted on triangular coordinate paper, showing the relative miscibilities of a three-component system. Percentage concentrations of all three components can be read directly from the plot. The area under the curve is the area of immiscibility and mixtures within this area will result in two-phase separation producing a raflinate phase and an extract phase. Mixtures above the curve will be miscible or homogeneous in lall proportions. The shape and height of the curve shoiw the maximum purification which can be obtained for one of the two components of the system, excluding solvent.

Any mixture under the curve, and which is therefore in the immiscibility range, can be split into a raffinate phase and an extract phase and each phase analyzed for the three components of the system. The analysis of the rathnate produces one point on the graph and the analysis of the extract provides a second point on the graph. A straight line connecting these two points Iand having the original mixture somewhere along the line is called a tie-line and the two points form a point on either side of the curve. The distances from the starting mixture to the terminal points are measures (inverse) of the volumes of raiinate and extract, respectively which will be produced.

The curve can be established by a titration technique. For example, mixtures of the two hydrocarbons may be made in increments of and each mixture titrated with dimethylsulfoxide at `a constant temperature until a blackout or cloud point is detected. Separation into an extract and a ratlnate and the analysis of the rainate is then carried out, as previously mentioned. This point forms a point on the right side of the curve. Conversely, mixtures of incremental volumes of solvent and the hydrocarbon of the top apex are then made and the hydrocarbon at the right apex is then titrated in the same man- Iier. The phases are separated and analyzed and an appropriate point obtained to form the right side of the curve is obtained.

While the ternary diagram of FIG. 2 is not thermodynamically rigorous, it is used so long as one recognizes that a given temperature produces a single curve, a single plot must be at a single pressure for various temperatures and, when multicomponent systems, such as a light cycle oil is treated, it is necessary to arbitrarily choose a definition of polycyclics and non-polycyclics to be used as two of the three components, with the other component being the solvent utilized. As an aid, the refractive index of various polycyclic and non-polycyclic mixtures can be placed along the line connecting the apex C and the apex B of the plot. The line from the light cycle oil composition point to the apex A is a constant composition line of component C and The distance Cl-Rl represents the volume of extract phase produced and the length El-Cl, the Volume of rafiinate phase produced. The dashed lines drawn from the apex A to border line C B, that is lines A-E1-E3 and A-Eg-Eb illustrate at line B-C, the refractive index and composition which each extract phase will have.

The E4 composition point represents the best extract that can be made by liquid phase extraction, at the temperature of interest (80 F), that is 90% polycyclics and 10% nonpolycyclics. The E3 composition point shows au 86% polycyclics composition, which differs only slightly from the E4 product. Further, the amount of extract produced at the 1:3 (point C2) ratio would be less than about half of that produced at the 1:1 (point C1) solvent to oil ratio. Hence, it is obvious that the solvent to oil ratios may be optimized for economic purposes and to facilitate handling in the recovery systems in addition to determining workable ratios by the use of the ternary phase diagram.

FIG. 3 also illustrates curves prepared for different temperatures than the 80 F. curve just discussed.

The extraction temperature may vary to a certain extent except that it cannot be below the freezing point of dirnethylsulfoxide. Accordingly, a temperature range of about to 200 F. may be utilized, and, preferably between and 140 F. A temperature of about 140 F. has been found optimum.

It should also be recognized that the extraction may be carried out either in a single or multi-stage extraction.

It has also been found that in the recovery of residual solvent from the rainate phase, the oil or ranate can be practically quantitatively displaced from the solvent by a ratio of water-to-solvent between about i114 and 10:1, preferably a ratio of 4:1 should be used.

The following examples will illustrate the principles and advantages of the present invention.

The variables and advantages of the present invention will be apparent from the following examples.

Example 1 A series of runs utilizing the light cycle oil heart-cut previously described was conducted in a single-stage extraction unit. The feed material had an API gravity at 60 F. of 26.5 and contained about 35.4% of acenaphthene and higher boiling materials. `It has been found that the solvent capacity can be substantially increased up to a certain point (where a single phase results) by increasing the temperature of extraction above ambient temperature conditions. This series of tests therefore shows not only the results of using various solvent-to-oil ratios but the affect of a series of higher temperatures on the extraction. Runs 1 through 9 were conducted utilizing dimethyl sulfoxide as the solvent, whereas, for comparicomponent B. Point C1 represents a 1: 1 ratio of solvent to 55 son, the extraction in Run 10 was made with furfuml /lllixge pmstplcltzs riestlotsplste rtorgflltegge 01g; containing 2% water at the optimum conditions previous- 1 a and the extract phase El, which are plotted after analyzing 1y freund for this partlculal: Solvent The change. m API their Components as Set forth above. R1 fans on the right gravity between the extraction` and the rainate illustrate hand Side of the curve and E1 on the 1eft hand Side' Simi G0 quite effectively the comparatively poor results obtained lady, the mixture C2 produces the right hand point R2 and with furfural. As a matter of fact, the furfural extraction the lefphand point E2, Lines E1 C1 R1 and E2 C2 R2 is almost as bad as a Run 8 utilizing a solvent-to-oil represent two tie-lines, as previously discussed. ratio of only .39.

TABLE II.-EXTRACTION Test 1 2 3 4 5 (i 7 8 9 l0 S l .i 0.37

Total tlii'uput S9 Ratt. yield Raft. gravity Gravity oil iii extract Example 2 The extract of Run No. 6 of the previous example was subjected to cooling to ambient temperature to effect a TABLE IV.-AZE0&1ROPIC DISTILLATIONS OF EXTRACTS Test No 13 14 15 16 17 1S 8:0 ratio 0.71 1.46 2. 52 0.88 1. 39 1.39 Temp., F 103 111 118 143 137 137 Wt. percent extract- 49 57 62 56. 4 64 64 API gr. (oil) 13. 0 14. 4 14. 5 15.7 15. 8 15.8 D (oil) 1. 5723 1, 5710 1. 5704 Azeotrope overhead 0i Vol. percent 29. 7 58. 7 77. 0 39 1 37 36. 5

In HC In DMSO In HC In DMSO In HC In DMSO layer layer layer layer layer layer qm2 1. 5579 l. 5587 1. 5633 1. 4825 l. 5202 1. 4850 1. 5285 1. 4850 1. 5267 API gr 16. 8 16. 8 15. 8 32. 8 22. 2 14. 5 9. 1 7. 9 15. 5 Bottoms oil:

70. 3 41.3 23.0 61 1 63 63. 5 1. 5783 1. 5903 1. 5944 1. 6017 1. 6030 1. 6008 API gr. 1l. 4 11. 2 10. 5 9. 6 0. 7 10. 1 1 Wt. percent.

Example 4 further separation of reinate from extract as explained 3 in connection with FIG. 1 of the drawings. The results of this cooling and separation are illustrated by the tests set forth in Table III. It is to be noted that the refractive index of 1.5210 for Run 12 indicates a very low level of polycyclic aromatics in the oil separated from the azeotrope overhead. The table also shows a refractive index of 1.5804 for extract recovered as a bottoms of the cooling step after removal of dimethylsulfoxide. This value compares with the corresponding refractive index of 1.5633 for the extract used as a feed to the cooling 1 EP. 2 Recovered extract from 11 and 12.

3 Original extract.

EXAMPLE 3 In this particular series of tests, extracts from Example 1 were subjected to azeotropic distillation in the presence of dimethylsulfoxide. The extract was taken directly from the extraction unit and was not subjected to the cooling illustrated in the previous example. Table IV below sets forth the results of these distillation tests. In Table IV, Runs 13, 14 and l5 were made on extracts from previous Runs l, 2 and 3 of Example 1, Run 16 was on an extract from previous Run 6, and Runs 17 and 18 were on portions of the extract from previous Run 4. A reux ratio of 5:1 was utilized in all tests, except Run 17 and this was at a reflux ratio of 10 to 1. It should also be noted that Runs 13, 14 and l5 were made in a Hempel unit which is essentially a single pass operation; whereas, 16, 17, and 18 were made in a Todd column The bottoms product from the previously described azeotropic distillation Runs 16 and 17 were combined and dealkylated in the following Run No. 19 and the bottoms product from the previous azeotropic distillation Run 18 was dealkylated as Run No. 20 below. The operating conditions of the dealkylation test are set forth in Table V below. The catalyst employed in the dealkylation was a conventional dealkylation catalyst comprising chromium oxide on alumina.

TABLE V.DEALKYLATION DATA [Bottoms from azeotropic distillation of extract-test] Example 5 In order to illustrate the applicability of the azeotropic distillation of the present application, utilizing dimethylsulfoxide as an azeotrope former, on materials other than solvent extracts, the following Runs were made. In Run 21, a synthetic light cycle oil mixture having a refractive index of 1.5268 was made up. This mixture was subjected to azeotropic distillation with internal reflux only. In Run 22, a light cycle oil was subjected to azeotropic distillation at a 1:1 dimethylsulfoxide-to-oil ratio and utilizing a 5:1 reflux ratio. Finally, in Run 23, a light cycle oil was subjected to azeotropic distillation at a 2:1 dimethylsulfoxide-to-oil ratio utilizing a 5:1 reiiux ratio.

TABLE 1V Overhead Vohlrie Hydrocarbon layer DMSO layer mls. Vol. (ml.) Percent oil nn2 Vol. (ml.) Percent oil m32 47. 5 9. 7 92. 8 1. 4686 37. 3 3. 2 1. 5310 49. 3 8. 8 93. 3 1. 4680 40. 5 3. 5 1. 5468 49.8 11. 6 86. 2 1. 4712 38. 2 6. 5 1. 5568 49. 8 14. 4 91. 6 1. 4750 35. 0 8. 6 1. 5468 49. 7 15. 2 94. 1 1. 4763 34. 5 9. 3 1. 5624 48. O 13. (l 94. 7 1. 4761 35.0 10. 0 1. 5613 49. 7 14. 2 90. 3 1. 4764 35. 5 7. 9 1. 5483 49. 6 10. 1 89. 2 1. 4739 39; 5' 6. 8 1. 5520 50. 2 9. 2 92. 5 1. 4723 41. 0 8. 0 1. 5560 50. 1 9. 6 94. 8 1. 4751. 40. 5 7. 4 1. 5538 5l. 0 6. 0 90. 0 1. 4648 45. 0 6. 0 1. 5383 50. 7 6. 2 91. 9 1. 4651 44. 5 5. 2 1. 5512 50. 1 6. 1 90. 3 1. 4661 44.0 5. 8 1. 5580 49. 7 31. 2 92. 0 1. 5115 18. 5 24. 1. 5780 25. 25.0 97. 6 1. 5284 No separation 25. 0 25. 0 98. 9 l. 5299 No separation 22 1 240 74 94. 6 1. 4584 166 1. 3 1. 5158 2 240 51 96. 1 1. 4584 189 1.3 1. 5158 3 120 25 96. 0 1. 4596 95 1. 3 1. 5158 4 240 47 97. 8 1. 4617 193 2. 8 1. 5214 5 240 35 97. 1 1. 4578 205 2. 2 1. 5249 6 170 7 98. 0 1. 4951 163 25. 4 1. 5847 Bottoms un 1.5713

It is obvious from the previous example that a dimethylsulfoxide-to-oil ratio higher than 1:1 is necessary for effective separation.

Example 6 The bottoms product of Run 22 and Run 23 were cornposited and subjected to dealkylation as set forth in Table VII below.

TABLE VII Dealkylation data.

Bottoms from azeotropic distillation of light cycle oil Still another azeotropic distillation of a light cycle oil heart-cut was carried out utilizing a 3.5 to 1 dimethylsulfoxide-to-oil ratio and a rellux ratio of 5.1. The results of this Run are set forth in Table VIII below.

TABLE VIII.HEARTCUT AZEOTROPIC DISTILLATION Run 25; Charge:1

I-10 gm. NaOH Overhead: 2,120 mls. HC layer (1.3% DMSO by Water wash) 13, 805 mls. DMSO layer (97.5% DMSO by water wash) 1y 800 mls. bottoms 275 mls. loss (=1.97% based on DMSO) Bottoms charge to column: Charge 1,800

Bottoms 1,360 Overhead 410 Column holdup nils.

[Properties of oil in fractions] In order to illustrate the advantages of azeotropic distillation of a feed mixture containing substantial volumes of polycyclic hydrocarbons, yet another test was run in which a heavy reformate fraction from a catalytic reforming operation was subjected to azeotropic distillation, using a 1:1 dimethylsulfoxide-to-oil ratio in Run 26 and a 2:1 ratio in Run 27. A reflux ratio of 2:1 was used in Run 26 and a reflux ratio of 5:1 was used in Run 27. Finally, Run 28 shows a dealkylation of the bottoms from the azeotropic distillation Runs 26 and 27. Table IX shows the properties of the feed material; Table X shows the results of the azeotropic distillation; and Table XI shows the dealkylation data.

TABLE IX Typical properties of heavy reformate ASTM Distillation:

IBP F-- 426 5 F-- 436 10 F 441 20 F 446 30 F-- 450 40 F 454 50 F 460 F 465 F 474 F-- 483 F 500 F 522 EP F 532 API Gravity 12.0

Fomrez GC Analysis:

N1 21.6 N 17.9 MN 34.4 DMN-I-BP 17.2 ACN 8.9

1 Saturates and poly-benzenes.

TABLE X.DMSO AZEOTROPE TESTS WITH HEAVY REFORMATE Run Chage: mls 2, 000 1, 000 DMSO, mls 2, 000 2, 000

10h .fr n Totivns..i1 c fil 2, 620 2, 530 tVol. percent. 65. 5 84. 3 o toms' 1, 300 485 34.0 16.2

20 -15 0.5 43.2 t t ff d 08.0 Bo tom, percen o ee 0.4 ggso t' l. ere t 74.2 8 Ohd composl ion vo p en 25.4K() MASHC Bottoms composition, vol. percent 1.8 DMSO 100 HC 08. 2 HC Bottoms 071320 1,6045 1. 6053 Bottoms, API gravity 8.3 8.7

HC HC Fomrez GC Feed Ohd Btms Ohd. Btms 21. 6 66. 9 1. 2 24. 6 Trace 17. 9 29. 4 11. 9 64. 0 1. 34.4 3.7 53.8 11.4 47.3 17. 2 18. 4 Trace 32. 7 8.9 14.7 18.4

TABLE XI Dealkylation data 68% Bottoms from azeotropic distillation of heavy reformate Run 28 Feed properties:

API gravity at 60 F 8.3 Refractive index 1.6045 Dealkylation operating conditions:

WHSV 1 Temperature, F. 1325 Pressure, p.s.i.g 400 Hz/HC 15/1 Liquid yield, wt. percent 86.5

Product analysis, wt. percent on feed:

Benzene 1.4 Toluene 2.5 Xylene 1.4 X-N 1.0 Naphthalene 58.9 M-Napthalene 25.3 a M-Napthalene 5.9 DiMN-l-Biph. 2.9 ACN and 0.7

I claim:

1. A method for separating a mixture of light cycle oil obtained by cracking a hydrocarbon mixture at least one component of which forms an azeotropic mixture with an organic sulfoxide, comprising; mixing a substantially anhydrous organic sulfoxide with said light cycle oil in a ratio of sulfoxide to light cycle oil between about 0.3 and 5.0 to 1 to form an azeotropic mixture of said one component and subjecting the resultant mixture to azeotropic distillation under conditions to separate an overhead product comprising said azeotropic mixture from a substantially pure, liquid bottoms product comprising the non-azeotrope forming components of said resultant mixture.

2. A method in accordance with claim 1, wherein the organic sulfoxide is dimethylsulfoxide.

3. A method in accordance with claim 1, wherein the remainder of the mixture is subjected to hydrodealkylation.

4. A method in accordance with claim 1, wherein `the azeotropic mixture is separated into organic sulfoxide and the azeotropic component.

S. A method in accordance with claim 4, wherein the azeotropic component is contacted with water and the azeotropic component is separated from a residual organic sulfoxide-water mixture.

6. A method in accordance with claim 5, wherein the organic sulfoxide-water mixture is distilled to separate water from the organic sulfoxide.

7. A method in accordance with claim 6, wherein the organic sulfoxide is recycled to the azeotropic distillation step.

8. A method in accordance with claim 4, wherein the organic sulfoxide is recycled to the azeotropic distillation step.

9. A method of separating light cycle oil obtained by cracking a hydrocarbon mixture, comprising; contacting said light cycle oil with a substantially anhydrous organic sulfoxide, in a ratio between about 0.3 and 5 to 1 sulfoxide to light cycle oil and at a temperature of about 65 to 200 F., separating a raffinate phase from an extract phase containing a substantial volume of residual raffinate, and subjecting the extract phase to azeotropic distillation under conditions to separate an overhead product comprising an azeotropic mixture of sulfoxide land said residual raflinate from a substantially pure, liquid bottoms product comprising the non-azeotrope-forming components of said extract phase.

10. A method in accordance with claim 9, wherein the extract phase is cooled and the cooled extract phase is separated into a further raiiinate phase and an extract phase.

11. A method in accordance with claim 10, wherein the further raiiinate phase is combined with the iirst raiinate phase.

12. A method in accordance with claim 10, wherein the raffinate phase is treated to remove residual organic sulfoxide by contacting the raffinate phase with water and separating the rainate from a mixture of water and organic sulfoxide.

13. A method in accordance with claim 12, wherein the Water-organic sulfoxide mixture is distilled to separate the water from dimethylsulfoxide.

14. A method in accordance with claim 13, wherein the organic sulfoxide is recycled to the initial contacting step.

15. A method in accordance with claim 9, wherein the organic sulfoxide is dimethylsulfoxide.

16. A method in accordance with claim 9, wherein the remainder of the extract phase is subjected to hydrodealkylation.

17. A method in accordance with claim 9, wherein the azeotropic mixture is separated into organic sulfoxide and the azeotropic component.

18. A method in accordance with claim 17, wherein at least a part of the organic sulfoxide is recycled to the initial contacting step.

19. A method in accordance with claim 17, wherein at least a part of the organic sulfoxide is recycled to the azeotropic distillation step.

20. A method in accordance with claim 17, wherein the azeotropic component is recycled back to the azeotropic distillation step.

References Cited UNITED STATES PATENTS 2,365,898 12/1944 Morris et al. 260-674 2,773,006 12/1956 Carver et al 260--674 3,005,032 10/1961 Makin 260-674 3,244,759 4/ 1966 Schaeier et al. 260-674 PAUL M. COUGHLAN, IR., Primary Examiner C. E. SPRESSER, JR., Assistant Examiner U.S. C1. X.R.

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