US3065167A

US3065167A - Process for separating aromatic hydrocarbons

Info

Publication number: US3065167A
Application number: US848418A
Authority: US
Inventors: Frederick J Zuiderweg; Gerrit H Reman; Schaafsma Albert
Original assignee: Shell Oil Co
Current assignee: Shell USA Inc
Priority date: 1959-03-24
Filing date: 1959-10-23
Publication date: 1962-11-20
Anticipated expiration: 1979-11-20
Also published as: DE1165187B; GB899281A; BE588894A; NL249677A

Description

DISTILLATION STRIPPER EXTRACTOR COLUMN Nov. 20, 1962 F J. ZUIDERWEG ET A1.

PROCESS FOR SEPARATING AROMATIC HYDROCARBONS Filed 0G13. 23, 1959 THEIR ATTORNEY United States Patent Oiice 33,065,167 Patented Nov. 20, 1962 ware Filed Get. 23, 1959, Ser. No. 843,412 Ciairns priority, appiication Great Britain Mar. 24, 1959 9 Claims. (Cl. 20S- 311) This invention relates to a process for the extraction and recovery of aromatic hydrocarbons from a liquid hydrocarbon mixture.

Several processes for the extraction and recovery of aromatic hydrocarbons from liquid hydrocarbon mixtures have already been proposed. Various high boiling, aromatic selective solvents have been suggested, among these are diethylene glycol, dipropylene glycol, and sulfolane. Because of the light-heavy selectivity of these solvents, the hydrocarbons in the mixture are extracted in the following order: light aromatics, heavy aromatics, light paraflns, and heavy paraflins. It has developed in the use of these high boiling solvents that when the process is operated to obtain a high yield of aromatics, contamination with light paraihnic material is usually experienced. A procedure for increasing the purity is de? scribed in British Patent No. 739,200 wherein a hydrocarbon mixture containing aromatic hydrocarbons is introduced into a multi-stage extractor at an intermediate point and a glycolic, water-containing selective solvent for aromatic hydrocarbons is introduced into one end of the extractor. At the same end a ratlinate poor in aromatics and in solvent is withdrawn, while at the other end an aromatic-rich extract phase is withdrawn and introduced into a column which is at a lower pressure than the extractor and wherein only part of the most volatile aromatics, together with non-aromatics of equivalent volatility, are removed as the top products and after condensation are returned as reflux to the extractor at the end from which the aromatic-rich extract is withdrawn. From the remainder of the extract phase the aromatics are separated from the solvent by distillation in a distilling column at atmospheric or subatmospheric pressure, direct steam being injected near the bottom of the distilling column. lt is stated in the example of the British patent that by using this process the purity of the recovered aromatics may amount to 97.5% and that this value may be increased to about 98% by using a more complex recovery system for the extract phase comprising three columns instead of two, and using two diiierent reflux streams instead of one single redux stream. This addition of a further column together with all its auxiliary equipment to the recovery system which is obviously a very costly measure, clearly demonstrates the extreme rdilliculties encountered in attempting to increase the products purity in extraction processes at the present type.

In this connection it should be noted that the speciiications for nitration grade toluene require a very high aromatics content, viz. of at least 98.5% by volume (cf. ASTM specification 13841-).

It has now been found that by introducing certain changes in the process as described in British specification No. 739,200 and by adhering to certain specic operating conditions, it is possible to obtain aromatics having a purity of 97.5-99% or even higher, while using a simple recovery system that involves two columns only.

The invention will now be illustrated with reference to the accompanying schematic drawing, wherein the sole gure is a process ilow diagram of a preferred embodiment of the improved process.

The combination of the following conditions has been found to give the aforesaid improved results:

(1) The solvent employed is a high boiling aromatic selective material having an atmospheric (normal) boiling temperature within the range of 225 to 295 C., and should contain not more than 2% by weight of water. Suitable solvents include diethylene glycol, dipropylene glycol, sulfolane, and mixtures thereof.

(2) The temperature in the extraction zone should be between and 155 C.

(3) The extract-phase from the extraction system should enter the stripping zone with a temperature between 140 and 155 C.

(4) The pressure in the stripping zone should be lower than the pressure in the extraction system, but not so low as to cause appreciable flash vaporization (i.e. vaporization by pressure reduction without heating) and should be at least 1.5 atm. abs.

(5) The top temperature in this stripping zone `should be between 140 and 155 C. and the bottom temperature should be at least 190 C.

(6) The condensed top vapors from the stripping zone should be freed from substantially all water present as a second liquid phase before being returned to the extraction system.

(7) The bottom product from the stripping zone should be cooled before expansion to distilling zone pressure by at least 15 C. to a temperature between 170 and 180 C.

(8) The pressure in the distilling zone should be below 0.5 atm. abs.

(9) The bottom temperature in the distilling zone should be at least C., but should be at least 30 C. below the bottom temperature in the stripping zone.

(10) Direct steam should be injected into said distilling zone in order to enable the hydrocarbons to be distilled off at not too high bottom temperatures.

Thus, according to the present invention there is provided a process for the extraction and recovery of aromatic hydrocarbons from a liquid hydrocarbon mixture containing one or more aromatic hydrocarbons, by introducing the mixture into a counterflow multistage extraction system, introducing, at one end of the system, a high boiling, aromatic selective solvent having an atmospheric boiling temperature within the range of 225 C. to 295 C. and containing dissolved water; maintaining the extraction system under pressure suicient to keep the flowing contents liquid and maintaining the owing contents at an elevated temperature; withdrawing a railinate poor in aromatics and in solvent from the same end of the system as that at which the solvent is introduced; withdrawing an aromatic-rich extract phase from the other end of the system, stripping said extract phase in a stripping zone to liberate a mixture of aromatic and nonaromatic hydrocarbons, said stripping being effected at a pressure lower than that prevailing in the extraction system, condensing the so liberated vapors and returning the condensate to the extraction zone; conducting the remainder of the extract phase to a distilling zone operating at a sub-atmospheric pressure, in which distilling zone separation is etfected between hydrocarbons a/nd solvent, with direct steam being introduced into said zone; and returning substantially hydrocarbon-free solvent containing dissolved water to the extraction system. The temperature in the extraction system is maintained between 140 and 155 C. The solvent supplied to one end of the extraction system contains not more than 2% by weight of water. The extract-phase withdrawn from the extraction system is introduced into the stripping zone,`

at a temperature between 140 C. and 155 IC. without the occurrence of appreciable flash vaporization. The stripping zone is operated at a pressure of at least 1.5

atm. abs. at a top temperature between 140 C. and 155 C. and at a bottom temperature of at least 190 C. The vapors from said stripping zone are condensed and recycled to the extraction system after removing substantially all water present as a second liquid phase. The condensed vapors are introduced into the extraction system at the end at which the extract phase is withdrawn and/or at an intermediate point of the extraction zone the bottom product of the stripping zone is cooled before expansion to the distilling zone pressure, by at least C. to a temperature between 170 and 180 C. The distilling zone is operated at a pressure of less than 0.5 atm. abs. and at a bottom temperature of at least 150 C., with the bottom temperature being at least 30 C. lower than the bottom temperaure in the stripping zone.

The present process can be applied to feedstocks having a wide or a narrow boiling range. It is especially suitable for separating aromatics from catalytically reformed gasolines, such as hydroformates and platformates, or from fractions thereof. rihe feedstock should have an ASTM nal boiling point of not higher than 220 C. Preferably the upper cutting point should be not higher than 160 C. When using a reformate fraction as starting material the upper cutting point should advantageously be approximately the same as the upper cutting point of the feedstock for the reforming operation. The lower cutting point =of the feedstock should preferably be about 100 C.

When carrying out the process according to the invention the volumetric ratio of solvent to hydrocarbon feed should be between 3:1 and 8:1, preferably between 4:1 and 6:1, whereas the ratio `of the amount of top product returned from the stripping zone to the extraction system to the amount of hydrocarbon feed should be between 0.2:1 and 0.7:1, preferably between 0.3:1 and 0.5:1.

Various high boiling solvents maybe used in the present process. Among these are diethylene glycol, dipropylene glycol and sulfolane. Diethylene glycol is preferred. The suitable solvents boil within the temperature range of 225 C. to 295 C. The solvent used in the extraction systeml may contain a small amount, not exceeding 2% by Weight, of water.

The extraction system should be a countercurrent multistage extraction system, e.g. a column containing packing material, or sieve plates, a rotating disc contactor, a multiplicity of mixer-settler combinations, and the like. The number of theoretical stages should preferably be at least 5.

The feed to the extraction system may be introduced as an intermediate point, but it is generally preferred to introduce it at or at least near that end of the extraction system at which the aromatic rich extract phase is withdrawn, because in that case the best compromise between product purity and recovery (yield) is usually realized. Suitable inlet points are at said end of the extraction system (viz. at the first theoretical stage) or at the second theoretical stage, the latter embodiment being preferred because in that case the purity is distinctly higher whereas the recovery is `only slightly decreased. Feed introduction at a plurality of points between the middle and the extract phase end of the system may be useful under specific circumstances.

The stream of hydrocarbons and solvent that is obtained as top product from the stripping zone and (after condensation and after removing substantially all water separating as a second liquid phase in the condensing operation) is returned to the extraction system, should also be introduced at one or more points at or near the end of the extraction system at which the aromatic rich extract phase is withdrawn. It is preferred to introduce this stream at `or near the end of the system (i.e. at the rst or second theoretical stage of the system) because in these cases the best compromise between product purity and recovery is realized. The recycle material may 4 in some instances be advantageously returned to the ex.- traction system in a plurality of streams between the middle and the extract phase end of the system. If necessary, the recycle stream may be vheated before entering the extraction system.

The temperature in the extraction system should be between and 155 C. There may be a certain temperature gradient over the system, provided that the above limits are adhered to.

The extract phase leaving the extraction system `is introduced in the stripping zone, at or near its top, with a temperature between 140 and 155 C., which implies that there will be no or substantially no intentional cooling of this stream between extraction system and stripping zone.

The stripping zone is operated at a pressure that is lower than that prevailing in the extraction system but that is at least 1.5 atm. abs. The diierence in pressure in the extraction system and in the stripping zone should be not so large as to cause appreciable flash vaporization. In fact, if there is any vaporous material present in the extract phase when entering the stripping zone, the gravimetric ratio of this vaporous material to the fresh feed to the extraction system should not exceed 1:10. Preferably this ratio should be below 5:100, and in the most preferred embodiment no vaporous material will be present at all. In this zone the top temperature should be between 140 and 155 C., and the bottom temperature should be at least 190, the difference between top andV bottom temperature being at least 40 C. The top product contains substantially all water and non-aromatic hydrocarbons present in the extract phase, and, in addition, some solvent and aromatic hydrocarbons.

In order to substantially reduce or prevent deterioration of the solvent due to lthe high bottom temperature in the 4stripping zone it is generally advisable that the solvent should contain a minor proportion, e.g. between 0.05 and 1% by weight of phenothiazine or of a substituted phenothiazine. The use of these compounds for the purpose referred to is described in -greater detail in the co-pending application Serial No. 862,761 of Philip I. Garner, led December 30, 1959.

The bottom product leaving the stripping zone is cooled by at least 15 C. to a temperature between 170 and 180 C., and then passed to the distilling zone operating at a pressure below 0.5 atm. abs., preferably 0.2 atm. abs. and at a bottom temperature of at least C., which temperature should always be at least 30 C. and preferably at least 40 C. lower than the bottom temperature in the stripping zone.

Whereas the stripping zone will be operated without introduction of direct steam near or in the bottom part, direct steam should be introduced into the distilling zone because otherwise the temperatures required to obtain a substantially hydrocarbon-free solvent as the bottom product would be so high as to cause decomposition of the solvent.

As a consequence of this introduction 4of steam in the ydistilling zone the solvent leaving the distilling zone contains a certain amount (less than 2% by weight) of dissolved water.

lf necessary, there may be cooling of the solvent stream that is recycled to the extraction system.

In the preferred embodiment of the present process the following operating conditions are adhered to:

Temperature in extraction system C, 150 Temperature of extract phase on entering stripping zone Top temperature in stripping zone C 145 Bottom temperature in stripping zone C 195 Temperature of bottom stream from stripping zone after cooling but before expansion to distilling zone pressure C 170 Top temperature in distilling zone C 85 Bottom temperature in distilling zone C-.. 155

Pressure in extraction system atm. abs 5 Pressure in stripping zone atm. abs-- 1.8 Pressure in distilling zone atm. abs 0.2

Under these conditions the water content of the solvent leaving the distilling column will be about 0.6% by weight.

It should be realized that the present process, in which there is substantially no iiash vaporization between extraction system and stripping zone, a rather large temperature drop over the stripping zone, a distinctly higher bottom temperature in the stripping zone than in the distilling zone, and a cooling between stripping zone and distilling zone, is fundamentally different from the approach of UK. patent specification No. 739,200, already mentioned above, in which the temperature drop over the stripping zone is rather small (as reflected by statement that it is usually not necessary to add heat to the stripping zone by means of a reboiler), in which the main heat input in recovery system is obviously in the bottom of the distilling zone and not in that of the stripping zone, and in which there is obviously no external cooling between stripping zone and distilling zone.

As a consequence of this essentially different situation the present process has the advantage that a greater proportion of the non-aromatics still present in the extract phase is removed in the overhead of the stripping zone, which results in a higher purity of the final extract as compared with the previous process.

A liquid hydrocarbon mixture containing both aromatic and non-aromatic hydrocarbons, is extracted in a multi-stage countercurrent extractor 1, which operates under pressure and at a temperature between 140 and 155 C. For this purpose the feed is introduced into the extractor through one or more of the feed inlet lines 2 to 5, whereas the selective solvent, containing dissolved water, is introduced into the extractor 1 at or near its top through line 6.

The rainate phase, which contains only relatively small amounts of water, solvent and aromatic hydrocarbons, is withdrawn from the top of the extractor 1 through line 7 and further processed to remove substantially all solvent present therein.

The aromatic-rich extract phase is withdrawn from the bottom of the extractor through line 8 and passes through a reducing valve (not shown) to the stripper column 10, the cooled extract phase being introduced at or near the top of the stripper. The pressure in column 10 is at least 1.5 atm. abs., but lower than the pressure in extractor 1.

The stripper 10 is provided with a reboiler 11 .and is operated with a fairly high temperature drop over the column. ln the stripper, separation is eliected into a top product containing some solvent, part of the aromatic hydrocarbons and most of the Water and non-aromatic hydrocarbons present in the extract phase leaving the extractor 1, and a bottom product that contains only a small amount of water and contains solvent and aromatic hydrocarbons and at most a minor amount of nonaromatic hydrocarbons. The top vapors are passed through line 12 to the condenser 13 and the resulting liquid passed to the settler 14, wherein separation in two layers, viz. a water-rich layer and a layer consisting mainly of solvent and hydrocarbons, takes place. The water layer is removed via line 15, whereas the solventhydrocarbon layer is passed through line 16 and heater 17, to one or more of the lines 18 to 20, through which it enters the extractor in its lower part.

The bottoms product from the stripper is passed through line 21, provided with a cooler 33 and with a subsequent reducing valve (not shown) to distilling column 22. In column 22, which operates at subatmospheric pressure, separation is effected into a top product containing aromatic hydrocarbons and substantially free 6 from solvent, .and into a bottom product, that is substantially free from aromatic hydrocarbons.

The top vapors are withdrawn through line 23, provided with condenser 24 and the condensate is partly Withdrawn via line 25, as the aromatic-rich product, and partly passed through line 26 equipped with the settler 26A, wherein separation in two layers, viz. a lower watererich layer and an upper layer consisting mainly of hydrocarbons, takes place. The upper layer is recycled as reliux to column 22. The bottom product from this column is withdrawn through line 27 and returned to the extractor through line 6, if necessary after cooling in cooler 28. If desired, part of the bottom product may be withdrawn through line 29 for purification or rejection, and fresh orpuried solvent may be introduced through line 30. Distilling column 22 is operated without a reboiler, hot steam being introduced into the bottom of this column via line 31 and heater 32 to strip dissolved hydrocarbons from the descending liquid.

lt should be realized that the drawing and its above description are schematic and that many auxiliary features such as valves and pumps, have not been represented or discussed.

We claim as our invention:

1. A process for the extraction and recovery of aromatic hydrocarbons from a liquid hydrocarbon mixture containing aromatic and non-aromatic hydrocarbons cornprising (l) introducing the mixture into a counterflow multi-stage extraction system wherein the temperature is maintained between and 155 C.; (2) introducing, at one end of the extraction system, a high boiling, aromatic selective solvent having a normal boiling temperature within the range of 225 C. to 295 C. and which solvent contains dissolved water in the proportion of not more than 2% by weight; (3) passing the solvent in countercurrent flow to the hydrocarbon mixture to obtain a solvent extract phase enriched in aromatics; (4) removing the extract phase from the extraction system and introducing said extract phase at a temperature between 140 and 155 C. without appreciable flash vaporization into a stripping zone, which is operated at a pressure of at least 1.5 atmospheres absolute and at a top temperature between 140 C. and 155 C. and with a bottom temperature of at least 190 C. to separate overhead a vapor phase enriched in non-aromatic hydrocarbons; (5) condensing the vapors from said zone and forming a first liquid phase consisting essentially of hydrocarbons along With a second liquid phase `and recycling them to the extraction system after removing substantially all water present as the second liquid phase, with the condensed vapors being introduced into the extraction system between the intermediate point of the system and the end at which the extract phase is withdrawn; (6) withdrawing the liquid remainder of the extract phase from the stripping zone and cooling the withdrawn extract phase before reduction in pressure by at least 15 C. to a temperature between 170 and 180 C., and (7) distilling the cooled extract phase in a distillation zone operated at a pressure of less than 0.5 atmosphere `absolute and a bottom temperature of at least C. which is at least 30 C. below the bottom temperature in the stripping zone whereby an aromatic-rich fraction is separated as an overhead product.

2. A process in accordance with claim 1 wherein the aromatic selective solvent is selected from the group consisting of diethylene glycol, dipropylene glycol, sulfolane, and mixtures thereof.

3. A process according to claim l in which the feed mixture is a catalytically reformed gasoline, or a fraction thereof.

4. A process according to claim 1 in which the feed stock has an upper cutting point not higher than C.

5. A process according to claim 1 in which the feed to the extraction system is introduced near that end of e' the extraction system at which the extract phase is withdrawn.

6. A process according to claim 1 in Which the condensed vapors substantially free from Water are introduced near the end of the extraction systern at which the extract phase is Withdrawn.

7. A process according to claim 1 in which the extract phase is introduced in the stripping zone near the top of said Zone.

8. A process according to claim 1 in which the Weight ratio of the vaporous material present in the extract phase on entering the stripping Zone to the fresh feed to the extraction system is less than 5 100.

9. A process according to claim 8 in which no vapors are present in the extract phase on entering the stripping zone.

References Cited in the file of this patent UNITED STATES PATENTS Schumacker Dec. 15, Durrum Sept. 17, Weedrnan Oct. 9, Poienberger Aug. 20, Findlay Sept. 10, Hawkins et a1. Oct. 8, Broughton Mar. 17, Scoeld Sept. 22,

FORETGN PATENTS Great Britain Oct. 26,

Claims

1. A PROCESS FOR THE EXTRACTION AND RECOVERING OF AROMATIC HYDROCARBONS FROM A LIQUID HYDROCARBON MIXTURE CONTAINING AROMATIC AND NON-AROMATIC HYDROCARBONS COMPRISING (1) INTRODUCING THE MIXTURE INTO A COUNTERFLOW MULTI-STAGE EXTRACTION SYSTEM WHEREIN THE TEMPERATURE IS MAINTAINED BETWEEN 140 AND 155*C.; (2) INTRODUCING, AT ONE END OF THE EXTRACTION SYSTEM, A HIGH BOILING, AROMATIC SELECTIVE SOLVENT HAVING A NORMAL BOILING TEMPERATURE WITHIN THE RANGE OF 225*C. TO 295*C. AND WHICH SOLVENT CONTAINS DISSOLVED WATER IN THE PROPORTION OF NOT MORE THAN 2% BY WEIGHT; (3) PASSING THE SOLVENT IN COUNTERCURRENT FLOW TO THE HYDROCARBON MIXTURE TO OBTAIN A SOLVENT EXTRACT PHASE ENRICHED IN AROMATICS; (4) REMOVING THE EXTRACT PHASE FROM THE EXTRACTION SYSTEM AND INTRODUCING SAID EXTRACT PHASE AT A TEMPERATUE BETWEEN 140* AND 155*C. WITHOUT APPRECIABLE FLASH VAPORIZATION INTO A STRIPPING ZONE, WHICH IS OPERATED AT A PRESSURE OF AT LEAST 1.5 ATMOSPHERES ABSOLUTE AND AT A TOP TEMPERATURE BETWEEN 140*C. AND 155*C. AND WITH A BOTTOM TEMPERATURE OF AT LEAST 190*C. TO SEPARATE OVERHEAD A VAPOR PHASE ENRICHED IN NON-AROMATIC HYDROCARBONS; (5) CONDENSING THE VAPORS FROM SAID ZONE AND FORMING A FIRST LIQUID PHASE CONSISTING ESSENTIALLY OF HYDROCARBONS ALONG WITH A SECOND LIQUID PHASE AND RECYCLING THEM TO THE EXTRACTION SYSTEM AFTER REMOVING SUBSTANTIALLY ALL WATER PRESENT AS THE SECOND LIQUID PHASE, WITH THE CONDENSED VAPORS BEING INTRODUCED INTO THE EXTRACTION SYSTEM BETWEEN THE INTERMEDIATE POINT OF THE SYSTEM AND THE END AT WHICH THE EXTRACTION PHASE IS WITHDRAWN; (6) WITHDRAWING THE LIQUID REMAINDER OF THE EXTRACT PHASE FROM THE STRIPPING ZONE AND COOLING THE WITHDRAWN EXTRACT PHASE BEFORE REDUCTION IN PRESSURE BY AT LEAST 15* C. TO A TEMPERATURE BETWEEN 170* AND 180*C., AND (7) DISTILLING THE COOLED EXTRACT PHASE IN A DISTILLATION ZONE OPERATED AT A PRESSURE OF LESS THAN 0.5 ATMOSPHERE ABSOLUTE AND A BOTTOM TEMPERATURE OF AT LEAST 150*C. WHICH IS AT LEAST 30*C. BELOW THE BOTTOM TEMPERATURE IN THE STRIPPING ZONE WHEREBY AN AROMATIC-RICH FRACTION IS SEPARATED AS AN OVERHEAD PRODUCT.