WO1986003189A1 - Procede de production de butane-1,4-diol - Google Patents

Procede de production de butane-1,4-diol Download PDF

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Publication number
WO1986003189A1
WO1986003189A1 PCT/GB1985/000524 GB8500524W WO8603189A1 WO 1986003189 A1 WO1986003189 A1 WO 1986003189A1 GB 8500524 W GB8500524 W GB 8500524W WO 8603189 A1 WO8603189 A1 WO 8603189A1
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WO
WIPO (PCT)
Prior art keywords
ester
hydrogenolysis
process according
vaporous
butane
Prior art date
Application number
PCT/GB1985/000524
Other languages
English (en)
Inventor
Keith Turner
Mohammad Sharif
Colin Rathmell
John Wilson Kippax
Anthony Benjamin Carter
John Scarlett
Arthur James Reason
Norman Harris
Original Assignee
Davy Mckee (London) Limited
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Priority claimed from US06/673,797 external-priority patent/US4584419A/en
Priority claimed from GB858514002A external-priority patent/GB8514002D0/en
Application filed by Davy Mckee (London) Limited filed Critical Davy Mckee (London) Limited
Priority to JP61500083A priority Critical patent/JPH0655684B2/ja
Priority to BR8507068A priority patent/BR8507068A/pt
Publication of WO1986003189A1 publication Critical patent/WO1986003189A1/fr
Priority to KR1019860700459A priority patent/KR870700588A/ko

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    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/132—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group
    • C07C29/136—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH
    • C07C29/147—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof
    • C07C29/149—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by reduction of an oxygen containing functional group of >C=O containing groups, e.g. —COOH of carboxylic acids or derivatives thereof with hydrogen or hydrogen-containing gases
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C29/00—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring
    • C07C29/17—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by hydrogenation of carbon-to-carbon double or triple bonds
    • C07C29/177—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by hydrogenation of carbon-to-carbon double or triple bonds with simultaneous reduction of a carboxy group
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07D—HETEROCYCLIC COMPOUNDS
    • C07D307/00—Heterocyclic compounds containing five-membered rings having one oxygen atom as the only ring hetero atom
    • C07D307/02—Heterocyclic compounds containing five-membered rings having one oxygen atom as the only ring hetero atom not condensed with other rings
    • C07D307/04—Heterocyclic compounds containing five-membered rings having one oxygen atom as the only ring hetero atom not condensed with other rings having no double bonds between ring members or between ring members and non-ring members
    • C07D307/06—Heterocyclic compounds containing five-membered rings having one oxygen atom as the only ring hetero atom not condensed with other rings having no double bonds between ring members or between ring members and non-ring members with only hydrogen atoms or radicals containing only hydrogen and carbon atoms, directly attached to ring carbon atoms
    • C07D307/08—Preparation of tetrahydrofuran

Definitions

  • This invention relates to a process for the production of butane-l,4-diol, more particularly to a process for the production of butane-l,4-diol by hydrogenolysis of a dialkyl ester, usually a di-(C* ] _ to C ⁇ alkyl) ester, of a C_ j dicarboxylic acid, such as maleic acid, fumaric acid, or succinic acid.
  • Butane-l,4-diol is used as a monomer in the production of plastics such as polybutylene terephthalate. It is also used as an intermediate for manufacture of butyrolactone and of the important solvent, tetrahydrof ran.
  • butane-l,4-diol involves reacting acetylene and formaldehyde by the Reppe reaction to give but-2-yne- 1,4-diol which is then hydrogenated to form butane-1,4- diol.
  • US-A-4001282 describes a process for production of butyrolactone by passing vaporised maleic acid, maleic anhydride, or a mixture thereof together with water and hydrogen over a metallic catalyst capable of hydrogenolysing a carboxylic group to a hydroxymethyl group.
  • Typical catalysts include copper-zinc catalysts (such as Girdler G-66 ARS and G-66-BRS) and copper chromite catalysts (such as Girdler G-13).
  • the reported products include succinic acid
  • US-A-4048196 teaches production of butane-1,4- diol and/or tetrahydrofuran by multi-stage catalytic hydrogenation of maleic anhydride or succinic anhydride.
  • maleic anhydride or succinic anhydride is hydrogenated over a nickel catalyst to give butyrolactone. This is then hydrogenated in the liquid phase over a copper/zinc oxide or hydroxide catalyst to give butane-l,4-diol and
  • US-A-2079414 describes use of copper chromite as a catalyst for effecting hydrogenation of esters. It is recommended that, in operating in the vapour phase, temperatures within the range of 300°C to 400°C should be
  • US-A-2040944 recommends use of temperatures of 230°C to 400°C for hydrogenation of esters of non-aromatic polybasic acids with a monohydric aliphatic alcohol containing at least four carbon atoms. It recommends
  • 35 copper chromite as catalyst and teaches that the catalyst can be prepared by ignition of a copper ammonium chromate precipitate and used without further treatment or after reduction by hydrogen at a temperature of 500°C or higher. It goes on to mention that either the liquid phase or vapour phase can be used, depending largely upon the ester to be hydrogenated. Pressures of 100 to 250 bar are recommended, as well as use of about 5 to 20 moles of hydrogen per mole of ester.
  • Example 1 describes a liquid phase batch reaction in which crude butyl succinate is hydrogenated at 3000 p.s.i.g. (207 bar) at 255°C using a copper chromite catalyst.
  • US-A-4172961 describes in Example 1 experiments in which a mixture of dibutyl butoxysuccinate, dibutyl maleate and dibutyl fumarate is hydrogenated using a copper chromite catalyst at 2000 psig to 4000 psig (141.65 bar to 282.26 bar and at a temperature of 250°C to yield butane-1,4-diol.
  • Copper chromite is suggested as the catalyst for use in both hydrogenation zones; use of temperatures of about 100°C to about 200°C and hydrogen pressures of about 2000 psig to about 3500 psig (about 141.65 bar to about 247.11 bar) in the first hydrogenation is recommended, whilst use of temperatures of about 225°C to about 300°C and pressures of about 3000 psig to about 4000 psig (about 241.95 bar to 282.26 bar) in the second hydrogenation zone is said to provide the necessary severity of operating conditions required to convert substantially all of the dialkyl esters to a product comprising butane-l,4-diol and monohydric alkanol.
  • Butyrolactone is produced, according to GB-A- 1168220, by vapour phase hydrogenation of maleic anhydride, succinic acid, an ester of maleic acid, an ester of succinic acid, or an ester of fumaric acid in the presence of a copper-zinc catalyst to which may be added small amounts of one or more promoters other than chromium.
  • a copper-zinc catalyst to which may be added small amounts of one or more promoters other than chromium.
  • reaction material should be fed to the catalyst at low speed. It is also difficult to reactivate the
  • butane-l,4-diol using as a starting material a precursor that can be produced from maleic anhydride and hence ultimately from butane or benzene as feedstock. According to the present invention there is provided a process for the production of butane-l,4-diol
  • a vaporous feed stream comprising a dialkyl ester of a C 4 dicarboxylic acid in vapour form c and excess hydrogen; allowing the ester to undergo hydrogenolysis in the first hydrogenolysis zone under substantially adiabatic reaction conditions thereby to form a vaporous first reaction mixture that is substantially free from the
  • the hydrogenolysis catalyst comprises copper chromite. It is especially preferred to use a
  • the ester is a di-(C* j _ to C 4 alkyl) ester of a but-2-en-l,4-dioic acid or of succinic acid. .,,-
  • the first feed temperature preferably lies in the range of from about 150°C to about 200°C, and even more preferably in the range of from about 170°C to about 190°C.
  • the two hydrogenolysis zones are each operated substantially adiabatically. The first reaction mixture thus exits the first hydrogenolysis zone at a temperature above the first inlet temperature.
  • the second inlet temperature at which the cooled vaporous first reaction mixture is supplied as second vaporous feed mixture to the second hydrogenolysis zone is preferably at least about 5°C lower than the temperature at which the vaporous first reaction mixture exits the first hydrogenolysis zone.
  • the second inlet temperature is at least about 10°C lower, e.g. about
  • the operating pressure is preferably at least about 25 bar but not more than about 70 bar, and is most preferably in the range of from about 35 bar to about 45 bar. Usually it is at least about 30 bar.
  • the two hydrogenolysis zones may each comprise a separate reactor which is operated under substantially adiabatic conditions.
  • the two hydrogenolysis zones may comprise separate beds of catalyst in the same reactor vessel.
  • the dialkyl ester of a C. dicarboxylic acid used in the process of the invention is preferably derived from an alkyl alcohol containing from 1 to 4 carbon atoms.
  • esters include diethyl maleate, diethyl fumarate, diethyl succinate, and mixtures of two or more thereof.
  • suitable esters include the dimethyl, di- n-propyl, di-i-propyl, di-n-butyl, di-i_-butyl, and di-sec- butyl esters of maleic, fumaric and succinic acids, as well as mixtures thereof.
  • the ester is selected from diethyl maleate, diethyl fumarate, diethyl succinate, and mixtures of two or more thereof.
  • ester Besides using an undiluted ester as feedstock it is also possible to use a solution of the ester in a suitable inert solvent, e.g. methanol, ethanol, or n- or iso-propanol.
  • a suitable inert solvent e.g. methanol, ethanol, or n- or iso-propanol.
  • the ester or ester solution feed can be admixed with recycled ester recovered in the product recovery section of the plant. If a di-(C* j _ to C 4 alkyl) maleate or fumarate is used as starting material, then the product stream from the second hydrogenolysis zone may include a minor amount of the corresponding dialkyl succinate; this can be recycled from the product recovery section of the plant for admixture with the di-(C, to C 4 alkyl) maleate or fumarate or solution thereof used as fresh feed supplied to the first hydrogenolysis zone.
  • the sole desired product is butane-1,4-diol
  • all of any co- product gamma-butyrolactone can be recycled from the product recovery section for admixture with fresh feed ester or ester solution.
  • a ready market can be found for most if not all of the co-product gamma- butyrolactone.
  • the plant operator can in this case adjust the output of the plant to meet market requirements either by adjusting the operating temperatures of the hydrogenolysis zones or by recycling a part of the butane- 1,4-diol product or a part of the byproduct gam a- butyrolactone from the product recovery section for admixture with the ester or ester solution feed.
  • the process requires that the starting ester and any other condensible component present be in the vapour phase in the first and second hydrogenolysis zones.
  • the composition of the vaporous mixture must be controlled so that, under the selected operating conditions, the temperature of the mixture in contact with the catalyst is always above the dew point of the ester and of any other condensible component present.
  • the temperature of the mixture in contact with the catalyst is at all times preferably at least about 5°C, more preferably at least about 10°C, and even more preferably at least about 15°C, higher than the dew point of the mixture. This can normally be achieved by selecting an appropriate gas:ester ratio in the vaporous mixture.
  • a convenient method of forming a vaporous mixture for use in the process of the invention is to spray the liquid ester or ester solution into a stream of hot hydrogen-containing gas so as to form a saturated or partially saturated vaporous mixture.
  • a vaporous mixture can be obtained by bubbling a hot hydrogen- containing gas-through a body of the liquid ester or ester solution. If a saturated vapour mixture is formed it should then be heated further or diluted with more gas so as to produce a partially saturated vaporous mixture prior to contact with the catalyst.
  • the vaporous mixture will normally contain excess hydrogen. It may additionally contain a minor amount of carbon oxides.
  • the vaporous mixture may further include vaporised inert solvent (if used) and one or more inert gases (e.g. N , A, CH etc) which may be present in the hydrogen supply in a major or minor amount. It may also include vaporous material recycled from the product recovery section.
  • the hydrogen supply is substantially free from sulphur compounds, from halogens such as Cl 2 , and from halogen-containing compounds such as HC1.
  • the H 2 :ester molar ratio is typically at least about 100:1 up to about 800:1 or more.
  • Preferabl*' it is at least about 200:1 up to about 500:1.
  • the presence of the excess hydrogen and of any inert gases that may be present helps to moderate the temperature rise in the first hydrogenolysis zone.
  • the preferred catalyst used in the first and/or second hydrogenolysis zone is a reduced copper chromite catalyst.
  • the catalyst is reduced at a temperature of not more than about 200°C, for an extended period using a mixture of H 2 and an inert gas, such as nitrogen, methane or argon.
  • a typical gas used for reduction of the catalyst is an H 2 in N 2 mixture containing, for example, from about 1% up to about 15% by volume of H .
  • the catalyst is reduced for at least about 24 hours prior to use.
  • copper chromite may be written as CuCr 0 .
  • copper chromite is non-stoichiometric and some authors have, for example, described a copper chromite catalyst as copper chromium oxide of the formula CuO.CuCr 2 0 4 .
  • the catalyst may contain an excess of copper oxide. It may further or alternatively contain a minor amount of at least one stabilizer, such as barium or manganese.
  • the catalyst contains, before reduction, from about 25 to about 45% by weight of copper and from about 20 to about 35% by weight of chromium.
  • the most preferred catalysts are those containing from about 32 to about 38% by weight of copper and from about 22 to about 30% by weight of chromium.
  • Such catalysts preferably contain no more than about 15% by weight of a stabilizer or stabilizers, if present. It may be supported on a suitable inert carrier. Desirably the catalyst is in finely divided form having an internal surface area, as measured by the well-known BET method, of at least about 30 sq. m. per gram and preferably at least about 60 sq. m. per gram. Preferably it is formed into cylindrical pellets or into other conventional catalyst shapes, such as rings, saddles, or the like.
  • the ester is preferably supplied to the first hydrogenolysis zone at a rate corresponding to a liquid hourly space velocity in the range of from about 0.1 hr up to about 0.6 hr--- or higher, for example up to about 1.5 hr -1 or even up-, to about 3.0 hr -1 .
  • liquid hourly space velocity we mean the number of unit volumes of the liquid ester supplied to the vaporization zone per unit volume of catalyst per hour.
  • gaseous hourly space velocity in the range of from about 2500 hr -1 up to about 160000 hr -1 , for example up to about 85000 hr "1 , most preferably ' in the range of from about 8000 hr "1 to about 30000 hr -1 .
  • gaseous hourly space velocity we mean the number of unit volumes of vaporous mixture measured at 1 bar and 0°C passed over a unit volume of catalyst per hour.
  • the feed temperature to the second hydrogenolysis zone is not more than about 175°C and even - more preferably lies in the range of from about 160°C to about 175°C.
  • further gas and/or ester can be admixed with the product stream from the first hydrogenolysis zone prior to admission to the second hydrogenolysis zone in order to adjust the temperature or the H :ester molar ratio.
  • one or more materials recovered from the product recovery section of the plant e.g. dialkyl succinate, unreacted dialkyl maleate or fumarate, gamma-butyrolactone and/or butane-1,4-diol
  • dialkyl succinate, unreacted dialkyl maleate or fumarate, gamma-butyrolactone and/or butane-1,4-diol can be admixed with the product stream from the first hydrogenolysis zone prior to admission to the second hydrogenolysis zone, instead of or in addition to recycling such material to the inlet end of the first hydrogenolysis zone.
  • a di-(C ⁇ to C 4 alkyl) maleate or fumarate such as diethyl maleate or fumarate
  • a catalyst which either has little or no ester hydrogenolysis activity or which is maintained under conditions such that ester hydrogenolysis and formation of butane-l,4-diol are minimised.
  • the di-(C j _ to C 4 alkyl) succinate-containing reaction mixture from the upstream hydrogenation zone is cooled, if necessary, prior to entry to the first hydrogenolysis zone in order to remove the heat of hydrogenation of the C:C bond of the maleate or fumarate ester prior to entry to the first hydrogenolysis zone.
  • the temperature rise in the first hydrogenolysis zone is kept as small as possible and the production of by-product tetrahydrofuran is thereby minimised.
  • the di-(C* ⁇ to C 4 alkyl) maleate or fumarate to the upstream hydrogenation zone at a relatively high rate, for example at a rate corresponding to a liquid hourly space velocity of at least about 1.0 hr "1 , preferably from about 3.0 hr "1 to about 6.0 hr "1 .
  • any upstream hydrogenation zone for conversion of dialkyl maleate and/or fumarate to dialkyl succinate at substantially the same pressure as is used in the first and second hydrogenolysis zones.
  • typical operating pressures in the upstream hydrogenation zone range from about 25 bar to about 70 bar and preferably lie in the range of from about 35 bar to about 45 bar.
  • Vapour phase hydrogenation conditions are usually selected in the upstream hydrogenation zone.
  • the inlet temperature to the upstream hydrogenation is preferably kept as low as is practicable consistent with use of vapour phase conditions and is typically in the range of from about 160°C to about 180°C.
  • the volume of catalyst in the upstream hydrogenation zone is selected to provide the desired high rate of ester throughput.
  • the reaction mixture exiting the upstream hydrogenation zone is supplied as feed mixture to the first hydrogenolysis zone.
  • the reaction mixture exiting the upstream hydrogenation zone Prior to introduction to the first hydrogenolysis zone the reaction mixture exiting the upstream hydrogenation zone is cooled somewhat so as to enable as low an inlet temperature as possible to be used in the first hydrogenolysis zone and thereby limit the maximum temperature achieved therein.
  • Further hydrogen and/or fresh feed ester and/or one or more materials recycled from the product recovery zone may be admixed with the reaction mixture from the upstream hydrogenation zone prior to its introduction into the first hydrogenolysis zone.
  • the product mixture exiting the second hydrogenolysis zone contains, in addition to unreacted hydrogen and possibly other gases, a mixture of condensible materials including butane-l,4-diol and alkyl alcohol (e.g. a C* j _ to C 4 alkyl alcohol) derived from the alkyl moiety of the dialkyl ester starting material.
  • the condensible materials may further include butyrolactone, dialkyl succinate and possibly also a small amount of unreacted ester and minor amounts of by-products. including n-butanol and tetrahydrofuran.
  • These condensibl-e materials are preferably condensed from the product mixture and separated in any suitable fashion, e.g.
  • dialkyl maleate and/or intermediate ester e.g. dialkyl succinate
  • dialkyl maleate and/or intermediate ester can be recycled for admixture with the ester or ester solution feed.
  • some of the butane-1,4-diol, and/or gamma-butyrolactone byproduct can be recycled for admixture with the product stream from the first hydrogenolysis zone.
  • the unreacted hydrogen-containing gases from the product recovery step can be recycled.
  • Purge lines can be used to control the level of inerts and/or byproducts in the circulating gas stream and in any liquid recycle line.
  • diethyl maleate is fed in line 1 and admixed with a liquid recycle stream in line 2 containing diethyl succinate, and possibly also gamma- butyrolactone and/or butane-1,4-diol.
  • the combined liquid stream is fed by pump 3 to feed heater 4 in which it is heated to 210°C by steam supplied in line 5.
  • the resulting hot liquid stream is passed to a spray nozzle 6 in feed saturator 7, in which the resulting spray encounters an ascending stream of hot hydrogen-containing gas supplied in line 8 at 42 bar.
  • Liquid is withdrawn from the bottom of feed saturator 7 in line 9 and is pumped by a feed saturator pump 10 to circulation heater 11 before being sprayed back into feed saturator 7 through nozzle 12.
  • Reference numeral 13 indicates a .spray eliminator pad in the top of feed saturator 7.
  • the ester vapour laden stream exits feed saturator 7 in line 14 at a temperature of 166°C and passes via heat exchanger 15 to a steam heater 16 in which its temperature is raised to 170°C.
  • the H 2 :ester molar ratio is approximately 300:1.
  • This mixture is then passed under substantially adiabatic reaction conditions through a first bed 17 of copper chromite catalyst containing 25% by weight of copper and 35% by weight of chromium and having a surface area of 85 m 2 /g.
  • the vaporous reaction mixture exits first bed 17 at an exit temperature of about 185°C.
  • Analysis of this first reaction mixture indicates the absence of diethyl maleate and shows that, besides hydrogen and inert gases (e.g.
  • diethyl maleate is converted smoothly and substantially quantitatively to diethyl succinate, 95.5 mol% of which is then converted to products with a selectivity to ethanol of substantially 100%, a selectivity to tetrahydrofuran of 4.3 mol%, a selectivity to i-butanol of 0.2 mol%, a selectivity to gamma-butyrolactone of 16.0 mol%, and a selectivity to butane-1,4-diol of 79.3 mol%, the balance being minor byproducts.
  • the butane-1,4-diol:gamma-butyrolactone molar ratio in this first product mixture from the first hydrogenolysis zone 17 is 4.96:1.
  • the size of the catalyst charge in first bed 17 is preferably selected in relation to the rate of supply of vaporous diethyl maleate thereto so that reaction can proceed substantially to equilibrium in bed 17.
  • the rate of supply of vaporous diethyl maleate to bed 17 corresponds to a liquid hourly space velocity of about 0.45 hr "1 .
  • the vaporous first reaction mixture exits first bed 17 at about 185°C and is passed by way of line 18 to heat exchanger 15, in which it is cooled to 170°C, and is then passed to a second bed 19 of the same copper chromite catalyst.
  • the second reaction mixture passes in line 20 to a heat exchanger 21 and then to a product cooler 22, in which it is cooled by means of cooling water supplied in line 23, to product catchpot 24.
  • the liquid condensate is recovered in line 25, whilst the gases exit in line 26.
  • the liquid condensate is passed through a pressure let ⁇ down valve 27 to a pressure let-down catchpot 28 and thence to a product recovery section 29 in which the product butane-1,4-diol is separated from gamma- butyrolactone, from tetrahydrofuran, from ethanol, from n- butanol, from diethyl succinate, and from any other minor components present in the condensate.
  • Separation of the condensate in product recovery section 29 can be achieved, for example, by distillation in several stages, including a "light ends" distillation stage to remove tetrahydrofuran, ethanol, iv-butanol and other low boiling byproducts, and then distillation of the resulting bottoms product to yield an overhead product comprising an azeotrope of gamma-butyrolactone and diethyl succinate and a bottom product comprising butane-l,4-diol.
  • Tetrahydrofuran is recovered in line 30, gam a- butyrolactone in line 31, and butane-1,4-diol in line 32. Diethyl succinate, and possibly also some gamma- butyrolactone and/or butane-l,4-diol, is recycled in line 2.
  • Fresh hydrogen is supplied to the plant in line 33 and is fed by way of compressor 34 and cooler 35 for admixture with the recycled gas in line 26.
  • the combined gas stream is compressed by recycle compressor 36 and fed to line 8.
  • a gas purge stream is taken in line 37 after passage through pressure let-down valve 38 and is combined with the vent gas in line 39 from catch pot 28; the combined stream in line 40 passes to gas purge condenser 41 which is fed with refrigerant in line 42 from refrigeration unit 43.
  • the purge gas exits in line 44 whilst any condensate is recovered in line 45 and fed to product recovery section 29.
  • a purge can be taken from the bottom of feed saturator 7 in line 46 as necessary.
  • Reference numeral 47 indicates a normally closed valve which can be opened at start up of the plant so as to cause the vaporous mixture in line 14 to bypass heat exchanger 15.
  • butane-1,4-diol can be passed from the purified butane-1,4-diol product line 32 via line 48 to a dehydration zone 49 which contains a charge of a dehydration catalyst such as gamma-alumina, aluminium phosphate, silica-alumina, a molecular sieve, an acidic clay or similar dehydration catalyst so as to convert at least a proportion of the butane-l,4-diol to tetrahydrofuran.
  • This zone is maintained at a temperature in the range of from about 200°C to about 300°C.
  • dehydration zone 49 can be supplied with a crude butane-1,4-diol stream from product recovery section 29 in line 50.
  • the tetrahydrofuran-containing product stream from dehydration section 49 is passed to product recovery section 29 in line 51, thereby increasing the amount of tetrahydrofuran appearing in line 30.
  • the illustrated plant can, with relatively little modification, be operated using another dialkyl ester of a C 4 dicarboxylic acid as feedstock, such as diethyl succinate or a mixture of diethyl maleate and diethyl succinate, in place of diethyl maleate.
  • diethyl succinate or a mixture thereof with diethyl maleate may, for example, be produced from diethyl maleate by hydrogenation in an upstream hydrogenation zone (not shown), the product of which is fed to the illustrated plant in line 1.
  • diethyl maleate supplied in line 1 is hydrogenated in the vapour phase to diethyl succinate in a hydrogenation zone (not shown) in line 14 upstream from the first hydrogenolysis zone.
  • a hydrogenation zone contains, for example, a relatively small charge of copper chromite catalyst and the rate of passage of ester therethrough preferably corresponds to a liquid hourly space velocity of at least about 3.0 hr "1 .
  • the resulting hydrogenated ester containing reaction mixture Prior to entry to first hydrogenolysis zone 17 the resulting hydrogenated ester containing reaction mixture, which now contains a minor amount only each of diethyl maleate and butane-l,4-diol and a major amount of diethyl succinate, is cooled to remove the heat of hydrogenation of the C:C bond of the unsaturated ester starting material.

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  • Organic Chemistry (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

Du butane-1,4-diol est produit par hydrogénolyse en phase vapeur d'un ester d'alkyle d'un acide C4 dicarboxylique, p.ex. du maléate de diéthyle, sur un catalyseur d'oxyde mélangé de Cu-Cr ou Cu-Zn réduit. Deux zones d'hydrogénolyse adiabatiques sont utilisées. Le mélange sortant de la première de ces zones est refroidi (par exemple de 5oC) et le mélange refroidi résultant est amené vers la seconde zone dans laquelle il se rééquilibre à une température inférieure pour augmenter le rendement de la production de butane-1,4-diol aux dépens de gamma-butyrolactone. Des conditions caractéristiques de réaction consistent en des températures comprises entre 150oC et 200oC, des pressions de 25 à 70 bars, et un rapport molaire H2:ester compris entre 100:1 et 800:1. Lorsque l'on utilise un ester de maléate, il est souvent avantageux de l'hydrogéner en son ester de succinate correspondant dans une zone d'hydrogénation amont avant de le faire entrer dans la première zone d'hydrogénolyse.
PCT/GB1985/000524 1984-11-21 1985-11-18 Procede de production de butane-1,4-diol WO1986003189A1 (fr)

Priority Applications (3)

Application Number Priority Date Filing Date Title
JP61500083A JPH0655684B2 (ja) 1984-11-21 1985-11-18 ブタン−１，４−ジオ−ルの生成方法
BR8507068A BR8507068A (pt) 1984-11-21 1985-11-18 Processo para a producao de butano-1,4-diol
KR1019860700459A KR870700588A (ko) 1984-11-21 1986-07-12 부탄-1,4-디올의 제조방법

Applications Claiming Priority (4)

Application Number Priority Date Filing Date Title
US06/673,797 US4584419A (en) 1983-11-29 1984-11-21 Process for the production of butane-1,4-diol
US673,797 1984-11-21
GB8514002 1985-06-04
GB858514002A GB8514002D0 (en) 1985-06-04 1985-06-04 Process

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WO1986003189A1 true WO1986003189A1 (fr) 1986-06-05

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EP (1) EP0204730A1 (fr)
JP (1) JPH0655684B2 (fr)
BR (1) BR8507068A (fr)
CA (1) CA1249837A (fr)
ES (1) ES8701703A1 (fr)
WO (1) WO1986003189A1 (fr)

Cited By (38)

* Cited by examiner, † Cited by third party
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WO1988000937A1 (fr) * 1986-08-01 1988-02-11 Davy Mckee (London) Limited Procede pour la coproduction de butane-1,4-diol et de gamma-butyrolactone
US4765869A (en) * 1986-08-01 1988-08-23 Davy Mckee (London) Limited Process for the production of a dialkyl maleate
US4810807A (en) * 1987-10-13 1989-03-07 The Standard Oil Company Hydrogenation of maleic anhydride to tetrahydrofuran
US4919765A (en) * 1987-07-29 1990-04-24 Davy Mckee (London) Limited Process for the purification of tetrahydrofuran
WO1991001960A1 (fr) * 1989-08-04 1991-02-21 Davy Mckee (London) Limited Procede
US5030609A (en) * 1987-07-29 1991-07-09 Davy Mckee (London) Limited Process for producing a hydrogenation catalyst of copper chromite
EP0443392A1 (fr) * 1990-02-20 1991-08-28 BASF Aktiengesellschaft Procédé de fabrication de tétrahydrofuranne et alphabutyrolactone
US5254758A (en) * 1989-08-04 1993-10-19 Davy Mckee (London) Limited Hydrogenation process
EP0589314A1 (fr) * 1992-09-23 1994-03-30 BASF Aktiengesellschaft Procédé de préparation de 2-méthyl-1,4-butanediol et 3-méthyltétrahydrofurane
US5310954A (en) * 1989-08-04 1994-05-10 Davy Mckee (London) Limited Process for preparing tetrahydrofuran
US5387752A (en) * 1993-12-02 1995-02-07 Eastman Chemical Company Process for the production of cyclohexanedimethanol
US5387753A (en) * 1993-12-02 1995-02-07 Eastman Chemical Company Process for the preparation of alcohols and diols
US5395987A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Preparation of cyclohexanedimethanol with a particular ratio
US5395991A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Process for the production of alcohols and diols
US5395990A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Process for the production of alcohols and diols
US5395986A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Process for the production of cyclohexanedimethanol
US5406004A (en) * 1993-12-02 1995-04-11 Eastman Chemical Company Hydrogenation process for the preparation of alcohols and diols utilizing gas and liquid phases
US5414159A (en) * 1993-12-02 1995-05-09 Eastman Chemical Company Process
EP0699652A2 (fr) 1994-09-02 1996-03-06 Bayer Ag Procédé pour la préparation de butane-1,4-diol à partir de l'anhydride maléique
US5536854A (en) * 1993-09-21 1996-07-16 Basf Aktiengesellschaft Preparation of 2-methyl-1,4-butanediol and 3-methyltetrahydrofuran
US5698713A (en) * 1994-02-22 1997-12-16 Lonza S.P.A. Process for the production of gamma-butyrolactone
WO1999035113A2 (fr) * 1998-01-08 1999-07-15 Pantochim S.A. Procede de production de tetrahydrofurane, de gamma-butyrolactone, et de butanediol
US6077964A (en) * 1996-05-15 2000-06-20 Basf Aktiengesellschaft Process for preparing gamma-butyrolactone, butane-1, 4-diol and tetrahydrofuran
US6100410A (en) * 1996-05-14 2000-08-08 Basf Aktiengesellschaft Process for the production of 1,4-butanediol, γ-butyrolactone and tetrahydrofuran
US6204395B1 (en) 1997-11-14 2001-03-20 Basf Aktiengesellschaft Process for the preparation of butane-1,4-diol, γ-butyrolactone and tetrahydrofuran
US6239292B1 (en) 1997-11-13 2001-05-29 Basf Aktiengesellschaft Process for preparing gamma-butyrolactone, butane-1,4-diol and tetrahydrofuran
EP1108702A1 (fr) * 1999-12-13 2001-06-20 Kvaerner Process Technology Limited Procédé pour la coproduction de diols aliphatiques et d'éthers cycliques
US6274743B1 (en) 1998-03-23 2001-08-14 Basf Aktiengesellschaft Process for the preparation of butanediol, butyrolactone and tetrahydrofuran
US6297389B1 (en) 1998-01-09 2001-10-02 Lonza S.P.A. Process for the production of gamma-butyrolactone
US6492535B1 (en) 1998-02-02 2002-12-10 Lonza S.P.A. Process for the production of gamma-butyrolactone
WO2003006446A1 (fr) * 2001-07-12 2003-01-23 Davy Process Technology Limited Procede de production d'ethers, et notamment de thf
WO2005058855A1 (fr) * 2003-12-16 2005-06-30 Davy Process Technology Limited Procede de production d'ethers
US8816104B2 (en) 2008-02-28 2014-08-26 Davy Process Technology Limited Process
US9527796B2 (en) 2013-12-06 2016-12-27 Johnson Matthey Davy Technologies Limited Process for the preparation of succinic acid ester
US9776948B2 (en) 2013-12-06 2017-10-03 Johnson Matthey Davy Technologies Limited Process for the preparation of succinic acid ester
US9776947B2 (en) 2013-10-14 2017-10-03 Johnson Matthey Davy Technologies Limited Process for the production of dialkyl succinate from maleic anyhdride
US10584091B2 (en) 2015-04-28 2020-03-10 Johnson Matthey Davy Technologies Limited Process for the recovery of dialkyl succinate or dialkyl maleate
WO2021115813A1 (fr) * 2019-12-10 2021-06-17 Basf Se Procédé de production de 1,4-butanediol, de gamma-butyrolactone et de tétrahydrofurane en phase gazeuse tout en évitant les dépôts de polymère

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DE3503485A1 (de) * 1985-02-01 1986-08-07 Lentia GmbH Chem. u. pharm. Erzeugnisse - Industriebedarf, 8000 München Verfahren zur herstellung reiner dialkylsuccinate
JPH02233674A (ja) * 1989-03-07 1990-09-17 Mitsubishi Kasei Corp ラクトン類の製造法
DE19842847A1 (de) * 1998-09-18 2000-03-23 Basf Ag Verfahren zur Herstellung von Tetrahydrofuran

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EP0143634A2 (fr) * 1983-11-29 1985-06-05 DAVY McKEE (LONDON) LIMITED Procédépour la préparation de butane-1,4-diol

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EP0143634A2 (fr) * 1983-11-29 1985-06-05 DAVY McKEE (LONDON) LIMITED Procédépour la préparation de butane-1,4-diol

Cited By (53)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
WO1988000937A1 (fr) * 1986-08-01 1988-02-11 Davy Mckee (London) Limited Procede pour la coproduction de butane-1,4-diol et de gamma-butyrolactone
US4765869A (en) * 1986-08-01 1988-08-23 Davy Mckee (London) Limited Process for the production of a dialkyl maleate
US4919765A (en) * 1987-07-29 1990-04-24 Davy Mckee (London) Limited Process for the purification of tetrahydrofuran
US5030609A (en) * 1987-07-29 1991-07-09 Davy Mckee (London) Limited Process for producing a hydrogenation catalyst of copper chromite
US4810807A (en) * 1987-10-13 1989-03-07 The Standard Oil Company Hydrogenation of maleic anhydride to tetrahydrofuran
WO1991001960A1 (fr) * 1989-08-04 1991-02-21 Davy Mckee (London) Limited Procede
US5254758A (en) * 1989-08-04 1993-10-19 Davy Mckee (London) Limited Hydrogenation process
US5310954A (en) * 1989-08-04 1994-05-10 Davy Mckee (London) Limited Process for preparing tetrahydrofuran
EP0443392A1 (fr) * 1990-02-20 1991-08-28 BASF Aktiengesellschaft Procédé de fabrication de tétrahydrofuranne et alphabutyrolactone
EP0589314A1 (fr) * 1992-09-23 1994-03-30 BASF Aktiengesellschaft Procédé de préparation de 2-méthyl-1,4-butanediol et 3-méthyltétrahydrofurane
US5536854A (en) * 1993-09-21 1996-07-16 Basf Aktiengesellschaft Preparation of 2-methyl-1,4-butanediol and 3-methyltetrahydrofuran
US5395987A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Preparation of cyclohexanedimethanol with a particular ratio
EP0656338A1 (fr) * 1993-12-02 1995-06-07 Eastman Chemical Company Hydrogénation en phase vapeur d'esters ou lactones en composés hydroxys
US5395991A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Process for the production of alcohols and diols
US5395990A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Process for the production of alcohols and diols
US5395986A (en) * 1993-12-02 1995-03-07 Eastman Chemical Company Process for the production of cyclohexanedimethanol
US5406004A (en) * 1993-12-02 1995-04-11 Eastman Chemical Company Hydrogenation process for the preparation of alcohols and diols utilizing gas and liquid phases
US5414159A (en) * 1993-12-02 1995-05-09 Eastman Chemical Company Process
US5387753A (en) * 1993-12-02 1995-02-07 Eastman Chemical Company Process for the preparation of alcohols and diols
US5387752A (en) * 1993-12-02 1995-02-07 Eastman Chemical Company Process for the production of cyclohexanedimethanol
US5698713A (en) * 1994-02-22 1997-12-16 Lonza S.P.A. Process for the production of gamma-butyrolactone
EP0699652A3 (fr) * 1994-09-02 1996-03-27 Bayer Ag
EP0699652A2 (fr) 1994-09-02 1996-03-06 Bayer Ag Procédé pour la préparation de butane-1,4-diol à partir de l'anhydride maléique
US5705715A (en) * 1994-09-02 1998-01-06 Bayer Aktiengesellschaft Process for preparing 1,4-butanediol from maleic anhydride
US6100410A (en) * 1996-05-14 2000-08-08 Basf Aktiengesellschaft Process for the production of 1,4-butanediol, γ-butyrolactone and tetrahydrofuran
US6077964A (en) * 1996-05-15 2000-06-20 Basf Aktiengesellschaft Process for preparing gamma-butyrolactone, butane-1, 4-diol and tetrahydrofuran
US6239292B1 (en) 1997-11-13 2001-05-29 Basf Aktiengesellschaft Process for preparing gamma-butyrolactone, butane-1,4-diol and tetrahydrofuran
US6204395B1 (en) 1997-11-14 2001-03-20 Basf Aktiengesellschaft Process for the preparation of butane-1,4-diol, γ-butyrolactone and tetrahydrofuran
WO1999035113A2 (fr) * 1998-01-08 1999-07-15 Pantochim S.A. Procede de production de tetrahydrofurane, de gamma-butyrolactone, et de butanediol
US6248906B1 (en) 1998-01-08 2001-06-19 Aldo Bertola Process for the production of tetrahydrofuran, gammabutyrolactone and butanediol
WO1999035113A3 (fr) * 1998-01-08 1999-09-16 Pantochim Sa Procede de production de tetrahydrofurane, de gamma-butyrolactone, et de butanediol
US6297389B1 (en) 1998-01-09 2001-10-02 Lonza S.P.A. Process for the production of gamma-butyrolactone
US6492535B1 (en) 1998-02-02 2002-12-10 Lonza S.P.A. Process for the production of gamma-butyrolactone
US6274743B1 (en) 1998-03-23 2001-08-14 Basf Aktiengesellschaft Process for the preparation of butanediol, butyrolactone and tetrahydrofuran
US6844452B2 (en) 1999-12-13 2005-01-18 Davy Process Technology Ltd. Process for the co-production of aliphatic diols and cyclic ethers
EP1108702A1 (fr) * 1999-12-13 2001-06-20 Kvaerner Process Technology Limited Procédé pour la coproduction de diols aliphatiques et d'éthers cycliques
WO2001044148A1 (fr) * 1999-12-13 2001-06-21 Davy Process Technology Ltd Procede permettant de produire a la fois des diols aliphatiques et des ethers cycliques
EA005339B1 (ru) * 2001-07-12 2005-02-24 Дэйви Проусесс Текнолоджи Лимитед Способ получения простых эфиров, в типичном случае тетрагидрофурана
CN1309714C (zh) * 2001-07-12 2007-04-11 大卫技术有限公司 醚的生产方法
WO2003006446A1 (fr) * 2001-07-12 2003-01-23 Davy Process Technology Limited Procede de production d'ethers, et notamment de thf
US6936727B2 (en) 2001-07-12 2005-08-30 Davy Process Technology Limited Process for the production of ethers, typically thf
AU2004299317B2 (en) * 2003-12-16 2010-07-22 Davy Process Technology Limited Process for the production of ethers
EA010589B1 (ru) * 2003-12-16 2008-10-30 Дэйви Проусесс Текнолоджи Лимитед Способ получения простых эфиров
US7598404B2 (en) 2003-12-16 2009-10-06 Davy Process Technology Limited Process for the production of ethers
WO2005058855A1 (fr) * 2003-12-16 2005-06-30 Davy Process Technology Limited Procede de production d'ethers
KR101144761B1 (ko) 2003-12-16 2012-05-10 데이비 프로세스 테크놀로지 리미티드 에테르 제조 방법
US8816104B2 (en) 2008-02-28 2014-08-26 Davy Process Technology Limited Process
US9776947B2 (en) 2013-10-14 2017-10-03 Johnson Matthey Davy Technologies Limited Process for the production of dialkyl succinate from maleic anyhdride
US9527796B2 (en) 2013-12-06 2016-12-27 Johnson Matthey Davy Technologies Limited Process for the preparation of succinic acid ester
US9776948B2 (en) 2013-12-06 2017-10-03 Johnson Matthey Davy Technologies Limited Process for the preparation of succinic acid ester
US10584091B2 (en) 2015-04-28 2020-03-10 Johnson Matthey Davy Technologies Limited Process for the recovery of dialkyl succinate or dialkyl maleate
WO2021115813A1 (fr) * 2019-12-10 2021-06-17 Basf Se Procédé de production de 1,4-butanediol, de gamma-butyrolactone et de tétrahydrofurane en phase gazeuse tout en évitant les dépôts de polymère
CN114829347A (zh) * 2019-12-10 2022-07-29 巴斯夫欧洲公司 在避免聚合物沉积的同时在气相中生产1,4-丁二醇、γ-丁内酯和四氢呋喃的方法

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JPS62501702A (ja) 1987-07-09
ES8701703A1 (es) 1986-12-01
BR8507068A (pt) 1987-07-14
CA1249837A (fr) 1989-02-07
ES549746A0 (es) 1986-12-01
JPH0655684B2 (ja) 1994-07-27
EP0204730A1 (fr) 1986-12-17

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