US3652452A - Solvent recovery process - Google Patents

Solvent recovery process Download PDF

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Publication number
US3652452A
US3652452A US880656A US3652452DA US3652452A US 3652452 A US3652452 A US 3652452A US 880656 A US880656 A US 880656A US 3652452D A US3652452D A US 3652452DA US 3652452 A US3652452 A US 3652452A
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solvent
primary
raffinate
stream
extraction
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US880656A
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Volker W Eyermann
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Honeywell UOP LLC
Universal Oil Products Co
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Universal Oil Products Co
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G21/00—Refining of hydrocarbon oils, in the absence of hydrogen, by extraction with selective solvents
    • C10G21/28—Recovery of used solvent
    • B—PERFORMING OPERATIONS; TRANSPORTING
    • B01—PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01D—SEPARATION
    • B01D11/00—Solvent extraction
    • B01D11/04—Solvent extraction of solutions which are liquid

Definitions

  • the process comprises mixing and cooling the ralfinate stream with a stream of primary solvent in a conditioning zone under conditions sufiicient to produce a conditioned raffinate free of primary solvent mist and thereby render the rafiinate more readily extractable in the secondary extraction zone.
  • the process has particular application where the primary extraction process is an aromatics extraction process, the rai'finate comprises nonaromatic hydrocarbons, the secondary extraction zone is an aqueous extraction zone, and the secondary solvent comprises water.
  • Primary solvents recoverable within the scope of the inventive process include sulfolane-type chemicals, sulfolene chemicals, polyethylene glycols, polypropylene glycols, dimethyl sulfoxide, etc.
  • the present invention relates to the solvent extraction of aromatic hydrocarbons from a hydrocarbon charge stream. More particularly, the present invention relates to the recovery of solvent from the non-aromatic rafiinate produced by the solvent extraction of aromatic from a hydrocarbon charge stream. Most specifically, the present invention relates to an improved process for the recovery of solvent from the non-aromatic raffinate by means of a secondary extraction process.
  • the non-aromatic raffinate which leaves the extraction zone of an aromatic hydrocarbon extraction process contains solvent.
  • the solvent which is withdrawn in the raffinate stream must be recovered not only because it may interfere with subsequent rafiinate processing or ultimate raffinate use, but primarily because continual loss of solvent in the rafiinate stream is a prohibitive economic expense in the aromatic extraction process.
  • the recovery of the solvent from the raffinate stream may be accomplished by distillation, or adsorption, or by a secondary solvent extraction process.
  • a typical solvent which is utilized in commercial aromatics extraction and which may be recovered in accordance with the practice of this invention is a solvent of the sulfolane type.
  • the solvent possesses a five membered ring containing one atom of sulfur and four atoms of carbon, with two oxygen atoms bonded to the sulfur atom of the ring.
  • the sulfolane type solvents may be indicated as having the following structural formula:
  • R R R and R are independently selected from the group comprising a hydrogen atom, an alkyl group having from one to ten carbon atoms, an alkoxy radical having from one to eight carbon atoms, and an arylalkyl radical having from one to tweleve carbon atoms.
  • Other solvents which may be included within this process are the sulfolenes such as Z-sulfolene or 3-sulfolene which have the following structures:
  • CH OH C CHg lH Hg H CH 2-Suliolene 3-Suliolene
  • Other typical solvents which have a high selectivity for separating aromatics from non-aromatic hydrocarbons and which may be processed within the scope of the present invention are Z-methylsulfolane, 2,4-dimethylsulfolane, methyl 2-sulfonyl ether, n-aryl-3-sulfonyl amine, 2-sulfonyl acetate, diethylene glycol, various polyethylene glycols, dipropylene glycol, various polypropylene glycols, dimethyl sulfoxide, N-methyl pyrollidone, etc.
  • sulfolane having the following structural formula:
  • one embodiment comprises a process for removal of primary solvent from a rafiinate stream containing a first concentration of primary solvent, said raffinate being produced by a primary extraction process, which comprises: (a) mixing and cooling said solvent-containing rafilnate stream in a conditioning zone with a first stream of said primary solvent under conditions sufficient to produce a conditioned rafiinate containing a second concentration of primary solvent below said first concentration; and, (b) passing the resulting mixture into a separation zone maintained under conditions suflicient to produce a second stream of primary solvent and a stream of said conditioned rafiinate.
  • a further embodiment of this invention may be characterised as an improvement in a process for the recovery of primary solvent from a solvent-containing rafiinate stream produced by a primary extraction process, wherein cooling of said rafiinate stream inherently produces a mist of free primary solvent phase in said raffinate stream, and said primary solvent is recovered from said rafiinate stream by contact with a secondary solvent in a secondary extraction zone, the improvement which comprises: (a) mixing and cooling said solvent-containing rafiinate stream in a conditioning zone with a first stream of said primary solvent under conditions sufficient to produce a conditioned rafiinate substantially free of primary solvent mist and thereby render said raffinate stream more readily extractible in said secondary extraction zone; (b) passing the resulting mixture into a separation zone maintained under conditions sufiicient to produce a second stream of said primary solvent and a stream of said conditioned rafiinate; and, (c) passing said stream of conditioned raffinate into said secondary extraction zone wherein dis solved primary
  • the poor recovery of the primary solvent from the rafiinate by aqueous extraction may be the result of any number of influences, it is believed that the primary cause is the presence of an entrained free primary solvent phase within the rafiinate stream.
  • the non-aromatic raffinate leaves the aromatics extraction zone at an elevated temperature in the range of from 150350 F., but normally at about 210 F. when the extraction is conducted for recovery of benzene, toluene, and Xylene (BTX extraction).
  • the rafiinate stream must therefore be passed through a heat exchanger and cooled before it is passed into the aqueous extractor for the recovery of the sulfolane solvent.
  • the solubility of sulfolane in the raffinate is reduced from the range of 1.5 to 2.0 mole percent to the range of about 0.5 to 0.7 mole percent, the solubility of sulfolane being dependent not only upon the temperature of the rafiinate but also upon the mole percent of aromatic hydrocarbon contained therein.
  • a separate phase of sulfolane solvent should therefore appear in the rafiinate stream.
  • the rafiinate is not readily extractible until it has been freed of the haze of sulfolane microdroplets, and it is only then that the aqueous extractor will function as designed.
  • the aqueous extractor means having been designed on the basis of the 100 F. sulfolane solubility, cannot function as designed since the presence of the sulfolane phase microdroplets within the rafiinate phase effectively results in a higher sulfolane concentration which is equivalent to the original solubility of sulfolane in the BTX rafiinate at 210 F.
  • the net result is that the concentration of sulfolane in the raflinate leaving the aqueous extractor will be from five to eight times as great as the design concentration.
  • Both the higher temperature and higher aromatic content of the ratfinate therefore, cause the kerosine raffinate to contain a substantially greater amount of dissolved sulfolane solvent than is contained in a BTX rafiinate.
  • the haze or mist content of the kerosine rafiinate is greater upon cooling to 100 F., and the problems encountered 1n adequately extracting the primary sulfolane solvent in the aqueous extraction zone are thus aggravated.
  • solvent-rich rafiinate may be rendered readily extractible by contacting the hazy raflinate stream with a stream of free primary solvent phase under conditions of high turbulence. It is found in the laboratory that if a test tube of hazy rafiinate is violently agitated, the haze or mist does not disappear. The ratio of the hydrocarbon phase to the free solvent phase is so high, the sulfolane haze typically comprising only from 0.8 to 1.5 mole percent, that the mist of microdroplets of free primary solvent phase cannot coalesce despite the turbulence.
  • the tube of hazy raifinate is given one or two manual shakes with a large volume of free solvent phase, however, the haze or mist of solvent immediately disappears into the free solvent phase leaving a clear transparent supernatant raflinate phase which only contains dissolved primary solvent.
  • the presence of the free solvent phase in the turbulent zone reduces the ratio of the hydrocarbon phase to non-hydrocarbon phase sufficiently to allow the mist of solvent microdroplets to be coalesced into the free solvent phase. This phenomenon can be observed both in sulfolane solvent systems and in glycol solvent systems.
  • the equivalent effect is achieved in the inventive process by providing that the cooled non-aromatic raffinate is contacted With a stream of primary solvent phase in a conditioning zone of high turbulence before the solventcontaining rafiinate is introduced into the aqueous extraction zone.
  • the solvent-containing raflinate will be rendered most readily extractible when the ratio of hydrocarbon phase to nonhydrocarbon phase in the turbulent conditioning zone is reduced to a level in the range of 5:11: 1.
  • An aromatic solvent extraction unit was designed for removing most of the aromatics out of a kerosine fraction in order to improve the smoke point.
  • the kerosine fraction fed to the extraction unit had a boiling range, characterized by true boiling point out points, of 347 F. to 450 F.
  • the smoke point of this kerosine fraction was 24 mm., and the aromatic content was 27.2 weight percent.
  • the specification product was a raffinate fraction having a smoke point of 35 mm.
  • the kerosine fraction was solvent extracted with a lean solvent composition comprising chemical sulfolane at a kerosine feed rate of 1500 b.p.s.d. to produce a raffinate product meeting the specification smoke point at a rate of 1010 b.p.s.d.
  • This kerosine fraction enters the process of the present invention at a temperature of 220 F., and it is passed into an extractor column 2 via line 1 at a pressure of 50 p.s.i.g. Extractor column 2 is operated under extraction conditions sufficient to remove a substantial portion of the aromatic content from the kerosine feed stock.
  • the bottom of extractor 2 is maintained at a temperature of 280 F. and at a pressure of 50 p.s.i.g.
  • a lean primary solvent stream enters extractor 2 via line 3 at a temperature of 300 F. and at a rate of 645.0 mols./hr.
  • this lean primary solvent comprises chemical sulfolane.
  • a net rafiinate fraction leaves the top of extractor 2 via line 4 at a temperature of about 300 F. and a pressure of 15 p.s.i.g. This rafiinate fraction will be discussed more fully hereinafter.
  • a rich primary solvent leaves extractor 2 via line 5 at a rate of 682.5 mols./hr.
  • This rich primary solvent comprises 41.5 mols./hr. of aromatic hydrocarbon and 641.0 mols./hr. of sulfolane solvent.
  • the rich solvent is discharged from the bottom of extractor 2 at a temperature of 280 F. and a pressure of 50 p.s.i.g.
  • the rich primary solvent is subsequently combined in line 5 with an additional solvent stream which enters line 5 via line 13 at a rate of 36.8 mols./hr. and at a temperature of F. from a processing source to be disclosed hereinafter.
  • the resulting mixture of rich primary solvent passes via line 5 at a total rate of 719.3 mols./hr.
  • the solvent stream comprises 41.5 mols./hr. of an aromatic hydrocarbon fraction and 677.8 mols./hr. of the sulfolane solvent.
  • the mixture of rich primary solvent is heated, by means not shown, to a temperature of 370 F., and upon passing through a control valve in line 5, not shown, is flashed into a fractionation zone 6 at a temperature of 330 F. and a pressure of 50 mm. of mercury absolute.
  • Fractionation zone 6 may typically comprise any type of a distillation column or a plurality of distillation columns. In the instant example, however, fractionation zone 6 comprises a single fractionating column operated under fractionating conditions sufficient to provide an aromatic extract having substantial freedom from the primary solvent, and a lean solvent fraction having substantial freedom from aromatic hydrocarbons.
  • the operating conditions utilized within fractionator 6 are a reboiler temperature of 405 F. and a reboiler pressure of mm. of mercury absolute. This fractionation condition results in an aromatic extract which is withdrawn from fractionator 6 via line 7 at a rate of 41.5 mols./hr. or 490 b.p.s.d.
  • a lean primary solvent fraction comprising the sulfolane solvent, which is withdrawn at a rate of 677.8 mols./hr. at a temperature of 405 F.
  • the lean primary solvent passing via line 3 is combined with an additional lean solvent fraction entering line 3 via line 8 at a rate of 1.0 mols./ hr. and a temperature of 340 F., said fraction being obtained from a source to be described hereinafter.
  • Te resulting mixture of lean primary sulfolane solvent is then passed at a temperature of 405 F. into a heat exchanger or cooler, not shown, wherein the lean solvent mixture is cooled to 300 F.
  • the cooled lean primary solvent thereafter passes via line 3 at a rate of 678.8 mols./hr. and a temperature of 300 F. to the vicinity of the extractor 2.
  • the lean primary solvent is there divided into two portions. The first portion of 645.0 mols./hr. is fed into the extractor 2 via line 3 as the lean primary solvent noted hereinabove.
  • the second portion of the lean primary solvent comprising chemical sulfolane is withdrawn from line 3 via line 9 at a rate of 33.8 mols./ hr. and passed into line 4 wherein it is mixed with the kerosine rafiinate being discharged at a rate of 67.5 mols./hr. from extractor 2.
  • the kerosine raflinate comprises 63.5 mols./hr. of hydrocarbons and 4.0 mols./hr. of dissolved and entrained sulfolane solvent.
  • the resulting mixture of rafiinate from extractor 2 and solvent from line 9, passes via line 4 at a rate of 101.3 mols./hr. and a temperature of 300 F. into a heat exchanger 10 wherein it is cooled to 100 F. before passing into a separator 12 via line 11.
  • the resulting mixture thus produces a conditioned ralfinate having contained therein, only an amount of sulfolane dissolved within the kerosine railinate under equilibrium solubility conditions.
  • the conditioned raffinate is thereby rendered more readily extractible in subsequent aqueous extraction means to be discussed hereinafter.
  • the mixture of conditioned raffinate phase and primary sulfolane solvent phase thus enters separator 12 via line 11.
  • the two phases are separated therein into a lean solvent phase and a clear supernatant raflinate phase.
  • the lean solvent phase is withdrawn from separator 12 via line 13 at a temperature of 100 F. and at a rate of 36.8 mols./ hr.
  • the lean primary solvent passes via line 13 into line 5 wherein it is mixed with the rich primary solvent leaving extractor 2 as disclosed hereinabove.
  • the conditioned raffinate leaves separator 12 via line 14 at a rate of 64.5 mols./hr.
  • the conditioned raffinate comprises 63.5 mols./hr. of mist-free kerosine hydrocarbon and 1.0 mols./hr. of dissolved sulfolane solvent.
  • the conditioned raflinate passes via line 14 into a secondary extraction means 15 comprising a typical Water-wash tower.
  • Water-Wash tower 15 is maintained at a temperature of 100 F., and at a pressure of 80 p.s.i.g. on the bottom and 50 p.s.i.g. at the top.
  • the conditioned ralfinate is contacted therein with a lean aqueous solvent or water stream which enters water-wash tower 15 via line 16 at a rate of 81.0 mols./hr. and at a temperature of 100 F.
  • the conditioned raffinate is contacted under conditions sufiicient to extract substantially all of the dissolved sulfolane solvent contained therein.
  • a net kerosine raflinate stream leaves water-Wash tower 15 via line 17 at a rate of 63.5 mols./hr. and a temperature of 100 F.
  • This kerosine raffinate passes via line 17 into storage facilities, not shown, at a rate of 1010 barrels per stream day and has the required specification smoke point of 35 mm.
  • a rich aqueous stream leaves the bottom of Water-wash tower 15 via line 18 at a rate of 82.0 mols./hr.
  • This stream comprises 81.0 mols./hr. of water and 1.0 mols./ hr. of primary sulfolane solvent.
  • the stream is passed into a water still 19 wherein the secondary solvent comprising water is separated from the primary solvent comprising chemical sulfolane.
  • the water still is operated at a reboiler temperature of about 350 F. to produce a sulfolane solvent having substantial freedom from water.
  • This lean sulfolane solvent leaves the bottom of water still 19 via line 8 at a rate of 1.0 mols./hr. and a temperature of about 340 F., and it enters line 3 wherein it is combined with the lean sulfolane solvent leaving the bottom of fractionator 6 as disclosed hereinabove.
  • An overhead vapor leaves the top of water still 19 via line 20 at a temperature of 170 F. and a pressure of 310 mm. of mercury absolute.
  • This overhead vapor comprises water 'vapor having substantial freedom from the sulfolane solvent, and it passes into a condenser 21 at a rate of 108.5 mols./hr.
  • the water vapor is cooled and condensed therein to a temperature of 100 F., and it passes via line 22 into a receiver 23.
  • the condensed Water is separated into two portions at receiver 23. A first portion passes via line 24 at a rate of 27.5 mols./hr. into the top of water still 19 to provide the reflux required within the water still.
  • a second portion of the water is withdrawn from receiver 23 via line 16 at a rate of 81.0 mols./hr., and the second portion is sent into the top of water-wash tower 15 to provide the necessary lean aqueous solvent disclosed hereinabove.
  • the inventive process recovers a primary solvent from a kerosine raffinate, and wherein the primary solvent comprises chemical sulfolane.
  • the process illustrated hereinabove by the foregoing example is not limited to that specific hydrocarbon fraction or that specific solvent composition.
  • the inventive process will find application in any extraction system wherein the raflinate phase inherently produces a mist of free primary solvent phase upon cooling.
  • the present invention has application in aromatic extraction processes utilizing glycol solvents since the misting problem is also prevalent in glycol solvent systems.
  • the ratio of the hydrocarbon phase to the nonhydrocarbon phase in the conditioning zone was 2:1. That is to say, two moles of kerosine raflinate were contracted with one mole of lean solvent by combining the stream of line 9 with the stream in line 4. But as noted hereinabove the ratio of hydrocarbon phase to non-hydrocarbon phase may be in the range of from 5:1 to 1:1. This ratio will vary with each specific application of the inventive process since specific operating conditions will dietate the ratio required. It is obvious that the effectiveness of the present invention is influenced not only by the rate of circulation of the primary solvent stream of line 9 to give the desired ratio, but also by the degree of turbulence within the conditioning zone comprising line 4, heat exchanger 10, and line 11.
  • the effectiveness of the inventive process is influenced by the temperature of the raflinate stream entering the conditioning zone, the solvent content of the rafiinate stream entering the conditioning zone, and the degree of cooling occurring within the conditioning zone.
  • the solvent content of the non-aromatic raflinate will vary since it is dependent upon the temperature level of the preceding aromatic extraction process unit (extractor 2), the mole percent of aromatics remaining in the millnate, and the amount of solvent physically entrained in the rafiinate. It is, therefore, not possible to define specifically the rate of circulation of the lean solvent stream to the conditioning zone which will be required to etfectuate the proper elimination of solvent mist.
  • the inventive process provides for maximum recovery of primary solvent from a solvent-containing rafiinate with a minimum of capital cost and a minimum of utility expenditure.
  • the water-wash tower must be designed to recover the full solvent content of the raffinate. That is to say, the secondary extraction zone must recover both the solvent of solution and the free solvent mist contained within the rafiinate phase.
  • the secondary extraction zone must contain a substantially greater number of contact stages and it also must have a larger diameter due to the greater amount of water circulation required within water-wash tower 15. Similarly, water still 19 must be larger and the capital and utility costs will therefore be substantially greater than those incurred by the inventive process.
  • the inventive process requires that the raffinate and the primary solvent phase be contacted within the conditioning zone in a region of high turbulence.
  • suflicient turbulence was provided by line 4, heat exchanger 10, and line 9.
  • the passage of the solvent phase and raflinate phase across the cooling tubes within heat exchanger provides a high degree of turbulence.
  • a mixing device may be installed in line 4 immediately before heat exchanger 10.
  • the mixing device may comprise an in-line mixer having a motor-driven propeller or it may comprise a turbine mixer.
  • such a mixing device may comprise a series of mixing orifices, or any other suitable mixing apparatus may be provided in line 4 in order to obtain the high degree of turbulence which is required for contacting the raffinate hydrocarbon phase with the non-hydrocarbon solvent phase. It is also within the scope of this invention to provide such a mixing apparatus in line 11 after the two phases have been cooled in heat exchanger 10. However, the preferred embodiment when installing a mixing device in the conditioning zone is to provide the device immediately ahead of the heat exchanger 10.
  • the present invention comprises a process for recovery of an aromatic extract and recovery of a non-aromatic raffinate from a primary solvent extraction process which comprises: (a) passing a mixed hydrocarbon feed stock into a primary extraction zone, wherein said feed stock is contacted under extraction conditions with a hereinafter specified first stream of water soluble primary solvent; (b) passing a second primary solvent stream containing aromatic hydrocarbons from said primary extraction zone into a distillation zone maintained under conditions sufiicient to produce an aromatic extract, and to produce a lean primary solvent having substantial freedom from aromatic hydrocarbons; (c) passing a first portion of said lean primary solvent into said primary extraction zone as said first primary solvent stream specified; (d) withdrawing from said primary extraction zone a first non-aromatic rafiinate stream containing primary solvent at a first concentration; (e) mixing and cooling said first raffinate stream with a second portion of said lean primary solvent in a conditioning zone under conditions sufiicient to
  • Primary solvent extraction process for recovery of an aromatic extract and recovery of a non-aromatic raffinate which comprises:
  • said solvent comprises a sulfolene selected from the group consisting of 2 sulfolene and 3-sulfolene.
  • said solvent comprises at least one polyalkylene glycol.
  • said solvent comprises at least one of the group consisting of diethylene glycol, dipropylene glycol, and triethylene glycol.

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  • Chemical & Material Sciences (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
US880656A 1969-11-28 1969-11-28 Solvent recovery process Expired - Lifetime US3652452A (en)

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CA (1) CA944302A (de)
DE (1) DE2058446C3 (de)
FR (1) FR2072509A5 (de)
GB (1) GB1320948A (de)

Cited By (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4364821A (en) * 1980-07-30 1982-12-21 Uop Inc. Fluid contacting process
US4897182A (en) * 1988-10-03 1990-01-30 Uop Aromatic extraction process having improved water stripper
US5008004A (en) * 1988-10-03 1991-04-16 Uop Aromatics extraction process having improved water stripper

Cited By (3)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4364821A (en) * 1980-07-30 1982-12-21 Uop Inc. Fluid contacting process
US4897182A (en) * 1988-10-03 1990-01-30 Uop Aromatic extraction process having improved water stripper
US5008004A (en) * 1988-10-03 1991-04-16 Uop Aromatics extraction process having improved water stripper

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Publication number Publication date
DE2058446B2 (de) 1974-09-26
GB1320948A (en) 1973-06-20
DE2058446A1 (de) 1971-06-03
DE2058446C3 (de) 1975-05-15
FR2072509A5 (de) 1971-09-24
CA944302A (en) 1974-03-26

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