CA1249837A - Process for the production of butane-1,4-diol - Google Patents
Process for the production of butane-1,4-diolInfo
- Publication number
- CA1249837A CA1249837A CA000495906A CA495906A CA1249837A CA 1249837 A CA1249837 A CA 1249837A CA 000495906 A CA000495906 A CA 000495906A CA 495906 A CA495906 A CA 495906A CA 1249837 A CA1249837 A CA 1249837A
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- Prior art keywords
- hydrogenolysis
- ester
- vaporous
- butane
- diol
- Prior art date
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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
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- Organic Chemistry (AREA)
- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
Abstract
ABSTRACT
PROCESS FOR THE PRODUCTION OF BUTANE-1,4-DIOL
Butane-1,4-diol is produced by vapour phase hydrogenolysis of an alkyl ester of a C4 dicarboxylic acid, e.g. diethyl maleate, over a reduced Cu-Cr or Cu-Zn mixed oxide catalyst. Two adiabatic hydrogenolysis zones are used. The mixture exiting the first of these zones is cooled (by, for example, 5°C) and the resulting cooled mixture is fed to the second zone in which it re-equilibriates at a lower temperature to increase the butane-1,4-diol yield at the expense of gamma-butyrolactone. Typical reaction conditions include use of temperatures of 150°C to 200°C, pressures of 25 to 70 bar, and an H2:ester molar ratio of 100:1 to 800:1. When using a maleate ester it is often advantageous to hydrogenate this to the corresponding succinate ester in an upstream hydrogenation zone prior to entry to the first hydrogenolysis zone.
PROCESS FOR THE PRODUCTION OF BUTANE-1,4-DIOL
Butane-1,4-diol is produced by vapour phase hydrogenolysis of an alkyl ester of a C4 dicarboxylic acid, e.g. diethyl maleate, over a reduced Cu-Cr or Cu-Zn mixed oxide catalyst. Two adiabatic hydrogenolysis zones are used. The mixture exiting the first of these zones is cooled (by, for example, 5°C) and the resulting cooled mixture is fed to the second zone in which it re-equilibriates at a lower temperature to increase the butane-1,4-diol yield at the expense of gamma-butyrolactone. Typical reaction conditions include use of temperatures of 150°C to 200°C, pressures of 25 to 70 bar, and an H2:ester molar ratio of 100:1 to 800:1. When using a maleate ester it is often advantageous to hydrogenate this to the corresponding succinate ester in an upstream hydrogenation zone prior to entry to the first hydrogenolysis zone.
Description
PROCESS_FOR THE PRODUCTIO~ OF BUTAN~-1,4-DIOL~
This invention relates to a process for the production of butane-1,4-diol, more particularly to a process for the production of butane-1,4-diol by hydroyenolysis of a dialkyl ester, usually a di-(Cl to C4 alkyl) ester, of a C4 dicarboxylic acid, such as maleic acid, fumaric acid, or succinic acid.
Butane-1,4-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, tetrahydrofuran.
The most commonly adopted present method of manufacturing butane-1,4-diol involves reacting acetylene and formaldehyde by the Reppe reaction to give but-2-yne-1l4-diol which is then hydrogenated to form butane-1,4-diol.
Alternatively it has been proposed in EP-B-0018163 to react allyl alcohol, which can be produced frompropylene, with iso-butylene to form allyl t-butyl ether.
This compound is then hydroformylated using, for example a rhodium complex hydroformylation catalyst, to give 4-t-butoxybutyraldehyde, which is then hydrogenated and cleaved under mild conditions with the aid of an acid catalyst to give butane-1,4-diol and so-butylene which is recycled for reaction with further allyl alcohol.
There have also been a number of proposals to produce butane-1,4-diol from maleic anhydride. According to these proposals maleic anhydride, which is produced by oxidation of butane or benzene, is esterified to give a diester of maleic acid which is then hydrogenated in one or more stages to give butane-1,4-diol. Alternatively it is proposed that, maleic acid or anhydride should be directly hydrogenated. In some of these proposals ,~
butyrolactone is an intermediate product.
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). Besides butyrolactone the reported products include succinic acid anhydride, propionic acid, butyric acid, propanol and n-butanol, but no mention is made of butane-1,4-diol.
US-A-4048196 teaches production of butane-1,4-diol and/or tetrahydrofuran by multi-stage catalytic hydrogenation of maleic anhydride or succinic anhydride.
In a first liquid phase hydrogenation step maleic anhydride or succinic anhydride is hydroqenated 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-1,4-diol and tetrahydrofuran.
In US-A-4083809, US-A-4105674 and GB-A-1534136 there is described a process for producing butyrolactone using a Cu-Pd catalyst for vapour phase hydrogenation of maleic acid, succinic acid, their anhydrides, and mix-tures of two or more thereof.
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 300C to 400C should be used. Diethyl succinate is mentioned.
US-A-2040944 recommends use of temperatures of 230C to 400C for hydrogenation of esters of non-aromatic polybasic acids with a monohydric aliphatic alcohol containing at least four carbon atoms. It recommends ~r~ k 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 500C or higher.
It goes on to mention that either the liql-id 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. Rxamp]e 1 describes a liquid phase batch reaction in which crude butyl succinate is hydrogenated at 3000 p.s.i.g. (207 bar) at 255C using a copper chromite catalyst.
A discussion of the use of copper chromite as a catalyst for hydrogenation of esters is to be found in "Organic Reactions", Vol. 8, published in 1954 by J.
Wiley and Sons, Inc.. Chapter 1 of this reference book is by Horner Adkins and is entitled "Catalytic Hydrogenation of ~sters to Alcohols". Table II on page 15 lists two experiments in which diethyl succinate is reacted at 5000 p.s.i. (345 bar) and 150C for 4 hours and at 3300 p.s.i.
(227.5 bar) and 250C for 6.5 hours respectively. This reference suggests that the "copper chromite" catalyst is more correctly described as an approximately equimolecular combination of cupric oxide and cupric chromite, i.e. CuO, CuCr204.
Production of copper chromite catalysts for use in hydrogenation of esters is described in ~R-A-1276722.
This recommends use of ester hydrogenation conditions including use of temperatures of between 100C and 350C, 0 preferably between 200C and 300C, and hydrogen pressures between 'l50hpz and 500 hpz" (by which is presumably meant between 50 bar and 500 bar).
Production of butane-1,4-diol and tetrahydrofuran by a process in which a dialkyl maleate is - subjected to hydrogenolysis in the liquid phase in the - q -presence of a copper chromite catalyst is described in GB-A-1454440 and GB-.~-1464263. A similar liquid phase process using nickel-based catalysts is described in GB-A-15~7198.
Hydrogenation of a di-(Cl to C7 alkyl) succinate at 100 to 300 atmospheres and 200C to 260C over a copper chromite catalyst to produce butane-1,4-diol is described in DE-A-2719867.
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 250C to yield butane-1,4-diol.
A two stage hydrogenation procedure in which a dialkyl maleate is hydrogenated first to the corresponding dialkyl succinate in a first hydrogenation zone and then the resulting dialkyl succinate is hydrogenated to yield butane-1,4-diol in a second hydrogenation zone is 20 described in US-A-4032458. Copper chromite is suggested as the catalyst for use in both hydrogenation zones; use of temperatures of about 100C to about 200C 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 225C
to about 300C 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-1,4-diol and monohydrie alkanol.
Butyrolactone is produced, according to GB-A-1168220, by vapour phase hydrogenation of rnaleic anhydride, succinic acid, an ester of maleic acid, an -35 ester of succinie acid, or an ester of fumaric acid in the -- 5 ~
presence of a copper-zinc catalyst to which may be added small amounts of one or rnore promoters other than chromium. This specification mentions that the preparation of butyrolactone by hydrogenation of the chosen starting materials was already known and states (see page 1, lines 23 to 25j:
"It is also possible to carry out the hydrogenation in the vapour phase when the preferred catalyst is copper chromite."
GB-A-1168220 continues (see page 1, lines 29 to 39):
"~urthermore hitherto known hydrogenation processes in the vapour phase suffer from the disadvantage of having to be carried out at a relatively high temperature, for example~ about 300~C, and moreover in order to obtain a good conversior.
rate, the reaction material should be fed to the catalyst at low speed. It is also difficult to reactivate the ~0 copper-chromite catalyst, when the activity has been lowered by use for a period of time."
It is an object of the present invention to provide a novel improved process for the production of butane-1,4-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-1,4-diol which comprises:
providing first and second hydrogenolysis zones, each containing a charge of a heterogeneous ester hydrogenolysis catalyst;
supplying to the first hydrogenolysis zone at an elevated pressure and at an elevated first feed temperature in excess of the threshold temperature for the hydrogenolysis reaction a vaporous feed stream comprising a dialkyl ester of a C4 dicarboxylic acid in vapour form 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 starting ester and contains, in addition to unreacted hydrogen, butane-1,4-diol and ~ a-butyrolactone in a first molar ratio;
cooling the vaporous first reaction mixture;
supplying a second vaporous feed stream comprising resulting cooled vaporous first reaction mixture to the second hydrogenolysis zone at a second feed temperature in excess of the threshold temperature for the hydrogenolysis reaction;
allowing the second vaporous feed stream to react further and to equilibriate under substantially adiabatic reaction conditions in the second hydrogenolysis zone; and recovering from the second hydrogenolysis zone a vaporous second reaction mixture containing, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a second molar ratio that is greater than the first molar ratio.
Preferably the hydrogenolysis catalyst comprises copper chromite. It is especially preferred to use a reduced copper chromite catalyst which contains, before reduction, from about 25 to about 45~ by weight of copper and from about 20 to about 35% by weight of chromium.
Preferably the ester is a di-(Cl to C4 alkyl~
ester oE a but-2-en-1,4-dioic acid or of succinic acid.
-35 In operating the process of the invention with a copper chromite catalyst it is preferred to operate at a temperature no higher than about 200C. Hence the first feed temperature preferably lies in the range of from about 150C to about 200C, and even more preferably in the range of from about 170C to about 190C. The two hydrogenolysis zones are each operated substan-tially 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 5C lower than the temperature at which the vaporous first reaction mixture exits the first hydrogenolysis zone. Typically the second inlet temperature is at least about 10C lower, e.g. about 15C lower, than the exit temperature from the first hydrogenolysis zone. The extent of cooling between the two hydrogenolysis zones will in no case be so much as to cause condensation of condensible products upon entry to the second hydrogenolysis zone. In most cases it is unlikely that the second inlet temperature will be more than about 25C lower than the exit temperature from the first hydrogenolysis zone.
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. Alternatively the two hydrogenolysis zones may comprise separate beds of catalyst in the same reactor vessel.
The dialkyl ester of a C4 dicarboxylic acid used in the process of the invention is preferably derived from -3 an alkyl alcohol containing from 1 to 4 carbon atoms.
~_7~ ?~
Examples of such es~ers include diethyl maleate, diethyl fumarate, diethyl succinate, and mixtures of two or more thereof. Other 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, aswell as mixtures thereof. Preferably the ester is selected from diethyl maleate, diethyl fumarate, diethyl succinate, and mixtures of two or more thereof.
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.
It is desirable to pretreat the ester feedstock, e.g. by distiliation, so as to remove substantially all of any sulphurous or halogenated impurities present in the feedstock. It is also preferred to remove substantially all water therefrom The ester or ester solution feed can be admixed with recycled ester recovered in the product recovery section of the plant. If a di-(Cl to C4 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-(Cl to C4 alkyl) maleate or fumarate or solution thereof used as fresh feed supplied to the first hydrogenolysis zone. If the sole desired product is butane-1,4-diol, then all of any co-product ga_ma-butyrolactone can be recycled from the product recovery section for adrnixture with fresh feed ester or ester solution. Often, however, a ready market can be found for most if not all of the co-product gam_a-butyrolactone. The plant operator can in this case adjust the output of the plant to meet marke-t requirements either -35 by adjusting the operating temperatures of the - 9 ~
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. This means that 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 10C, and even more preferably at ]east about 15C, 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. Alternatively such 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.
Reduction of a ma~eate or fumarate ester to butane-1,4-diol involves reaction of 5 moles of H2 with each mole of ester, according to the following equation:
R-O-CO-CH=CH-CO-O-R + 5H2 = 2ROH + HO(CH2)4OH (I) where R is an alkyl radical containing from 1 to 4 carbon atoms.
However, when a succinate ester is hydrogenolysed, only 4 moles of H2 are consumed:
R-O-CO-CH2-CH2-CO-O-R + 4H2 = 2ROH ~ HO(CH2)4OH (II) where R is as defined above.
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. N2, A, CH4 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. Preferably the hydrogen supply is substantially free from sulphur compounds, from halogens such as C12, and from halogen-containing compounds such as HCl.
In the vaporous mixture the H2:ester molar ratio is typically at least about 100:1 up to about 800:1 or more. Preferably it is at ]east 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.
In practice the reduction of a maleate ester, such as diethyl maleate, is more complex than is suggested by equation ~I) above and results in production of variable amounts of by-products, including tetrahydrofuran, gamma- butyrolactone and n-butanol. Although the reaction mechanism has not been fully elucidated yet, the currently available evidence is consistent with the following sequence:-3~ d CH - CO2Et +H2 CH2 - Co2Et CH - CO2Et CH2 - CO2Et +2H2 -EtOH
~0 ~ ~ .
CH2 - C ~ +EtOH CH2 - CO2Et / CH2 - CH2 ~ -EtOHCH2 CH2OH
+332 ~+232 ~\ ~t E2
This invention relates to a process for the production of butane-1,4-diol, more particularly to a process for the production of butane-1,4-diol by hydroyenolysis of a dialkyl ester, usually a di-(Cl to C4 alkyl) ester, of a C4 dicarboxylic acid, such as maleic acid, fumaric acid, or succinic acid.
Butane-1,4-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, tetrahydrofuran.
The most commonly adopted present method of manufacturing butane-1,4-diol involves reacting acetylene and formaldehyde by the Reppe reaction to give but-2-yne-1l4-diol which is then hydrogenated to form butane-1,4-diol.
Alternatively it has been proposed in EP-B-0018163 to react allyl alcohol, which can be produced frompropylene, with iso-butylene to form allyl t-butyl ether.
This compound is then hydroformylated using, for example a rhodium complex hydroformylation catalyst, to give 4-t-butoxybutyraldehyde, which is then hydrogenated and cleaved under mild conditions with the aid of an acid catalyst to give butane-1,4-diol and so-butylene which is recycled for reaction with further allyl alcohol.
There have also been a number of proposals to produce butane-1,4-diol from maleic anhydride. According to these proposals maleic anhydride, which is produced by oxidation of butane or benzene, is esterified to give a diester of maleic acid which is then hydrogenated in one or more stages to give butane-1,4-diol. Alternatively it is proposed that, maleic acid or anhydride should be directly hydrogenated. In some of these proposals ,~
butyrolactone is an intermediate product.
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). Besides butyrolactone the reported products include succinic acid anhydride, propionic acid, butyric acid, propanol and n-butanol, but no mention is made of butane-1,4-diol.
US-A-4048196 teaches production of butane-1,4-diol and/or tetrahydrofuran by multi-stage catalytic hydrogenation of maleic anhydride or succinic anhydride.
In a first liquid phase hydrogenation step maleic anhydride or succinic anhydride is hydroqenated 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-1,4-diol and tetrahydrofuran.
In US-A-4083809, US-A-4105674 and GB-A-1534136 there is described a process for producing butyrolactone using a Cu-Pd catalyst for vapour phase hydrogenation of maleic acid, succinic acid, their anhydrides, and mix-tures of two or more thereof.
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 300C to 400C should be used. Diethyl succinate is mentioned.
US-A-2040944 recommends use of temperatures of 230C to 400C for hydrogenation of esters of non-aromatic polybasic acids with a monohydric aliphatic alcohol containing at least four carbon atoms. It recommends ~r~ k 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 500C or higher.
It goes on to mention that either the liql-id 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. Rxamp]e 1 describes a liquid phase batch reaction in which crude butyl succinate is hydrogenated at 3000 p.s.i.g. (207 bar) at 255C using a copper chromite catalyst.
A discussion of the use of copper chromite as a catalyst for hydrogenation of esters is to be found in "Organic Reactions", Vol. 8, published in 1954 by J.
Wiley and Sons, Inc.. Chapter 1 of this reference book is by Horner Adkins and is entitled "Catalytic Hydrogenation of ~sters to Alcohols". Table II on page 15 lists two experiments in which diethyl succinate is reacted at 5000 p.s.i. (345 bar) and 150C for 4 hours and at 3300 p.s.i.
(227.5 bar) and 250C for 6.5 hours respectively. This reference suggests that the "copper chromite" catalyst is more correctly described as an approximately equimolecular combination of cupric oxide and cupric chromite, i.e. CuO, CuCr204.
Production of copper chromite catalysts for use in hydrogenation of esters is described in ~R-A-1276722.
This recommends use of ester hydrogenation conditions including use of temperatures of between 100C and 350C, 0 preferably between 200C and 300C, and hydrogen pressures between 'l50hpz and 500 hpz" (by which is presumably meant between 50 bar and 500 bar).
Production of butane-1,4-diol and tetrahydrofuran by a process in which a dialkyl maleate is - subjected to hydrogenolysis in the liquid phase in the - q -presence of a copper chromite catalyst is described in GB-A-1454440 and GB-.~-1464263. A similar liquid phase process using nickel-based catalysts is described in GB-A-15~7198.
Hydrogenation of a di-(Cl to C7 alkyl) succinate at 100 to 300 atmospheres and 200C to 260C over a copper chromite catalyst to produce butane-1,4-diol is described in DE-A-2719867.
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 250C to yield butane-1,4-diol.
A two stage hydrogenation procedure in which a dialkyl maleate is hydrogenated first to the corresponding dialkyl succinate in a first hydrogenation zone and then the resulting dialkyl succinate is hydrogenated to yield butane-1,4-diol in a second hydrogenation zone is 20 described in US-A-4032458. Copper chromite is suggested as the catalyst for use in both hydrogenation zones; use of temperatures of about 100C to about 200C 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 225C
to about 300C 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-1,4-diol and monohydrie alkanol.
Butyrolactone is produced, according to GB-A-1168220, by vapour phase hydrogenation of rnaleic anhydride, succinic acid, an ester of maleic acid, an -35 ester of succinie acid, or an ester of fumaric acid in the -- 5 ~
presence of a copper-zinc catalyst to which may be added small amounts of one or rnore promoters other than chromium. This specification mentions that the preparation of butyrolactone by hydrogenation of the chosen starting materials was already known and states (see page 1, lines 23 to 25j:
"It is also possible to carry out the hydrogenation in the vapour phase when the preferred catalyst is copper chromite."
GB-A-1168220 continues (see page 1, lines 29 to 39):
"~urthermore hitherto known hydrogenation processes in the vapour phase suffer from the disadvantage of having to be carried out at a relatively high temperature, for example~ about 300~C, and moreover in order to obtain a good conversior.
rate, the reaction material should be fed to the catalyst at low speed. It is also difficult to reactivate the ~0 copper-chromite catalyst, when the activity has been lowered by use for a period of time."
It is an object of the present invention to provide a novel improved process for the production of butane-1,4-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-1,4-diol which comprises:
providing first and second hydrogenolysis zones, each containing a charge of a heterogeneous ester hydrogenolysis catalyst;
supplying to the first hydrogenolysis zone at an elevated pressure and at an elevated first feed temperature in excess of the threshold temperature for the hydrogenolysis reaction a vaporous feed stream comprising a dialkyl ester of a C4 dicarboxylic acid in vapour form 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 starting ester and contains, in addition to unreacted hydrogen, butane-1,4-diol and ~ a-butyrolactone in a first molar ratio;
cooling the vaporous first reaction mixture;
supplying a second vaporous feed stream comprising resulting cooled vaporous first reaction mixture to the second hydrogenolysis zone at a second feed temperature in excess of the threshold temperature for the hydrogenolysis reaction;
allowing the second vaporous feed stream to react further and to equilibriate under substantially adiabatic reaction conditions in the second hydrogenolysis zone; and recovering from the second hydrogenolysis zone a vaporous second reaction mixture containing, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a second molar ratio that is greater than the first molar ratio.
Preferably the hydrogenolysis catalyst comprises copper chromite. It is especially preferred to use a reduced copper chromite catalyst which contains, before reduction, from about 25 to about 45~ by weight of copper and from about 20 to about 35% by weight of chromium.
Preferably the ester is a di-(Cl to C4 alkyl~
ester oE a but-2-en-1,4-dioic acid or of succinic acid.
-35 In operating the process of the invention with a copper chromite catalyst it is preferred to operate at a temperature no higher than about 200C. Hence the first feed temperature preferably lies in the range of from about 150C to about 200C, and even more preferably in the range of from about 170C to about 190C. The two hydrogenolysis zones are each operated substan-tially 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 5C lower than the temperature at which the vaporous first reaction mixture exits the first hydrogenolysis zone. Typically the second inlet temperature is at least about 10C lower, e.g. about 15C lower, than the exit temperature from the first hydrogenolysis zone. The extent of cooling between the two hydrogenolysis zones will in no case be so much as to cause condensation of condensible products upon entry to the second hydrogenolysis zone. In most cases it is unlikely that the second inlet temperature will be more than about 25C lower than the exit temperature from the first hydrogenolysis zone.
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. Alternatively the two hydrogenolysis zones may comprise separate beds of catalyst in the same reactor vessel.
The dialkyl ester of a C4 dicarboxylic acid used in the process of the invention is preferably derived from -3 an alkyl alcohol containing from 1 to 4 carbon atoms.
~_7~ ?~
Examples of such es~ers include diethyl maleate, diethyl fumarate, diethyl succinate, and mixtures of two or more thereof. Other 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, aswell as mixtures thereof. Preferably the ester is selected from diethyl maleate, diethyl fumarate, diethyl succinate, and mixtures of two or more thereof.
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.
It is desirable to pretreat the ester feedstock, e.g. by distiliation, so as to remove substantially all of any sulphurous or halogenated impurities present in the feedstock. It is also preferred to remove substantially all water therefrom The ester or ester solution feed can be admixed with recycled ester recovered in the product recovery section of the plant. If a di-(Cl to C4 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-(Cl to C4 alkyl) maleate or fumarate or solution thereof used as fresh feed supplied to the first hydrogenolysis zone. If the sole desired product is butane-1,4-diol, then all of any co-product ga_ma-butyrolactone can be recycled from the product recovery section for adrnixture with fresh feed ester or ester solution. Often, however, a ready market can be found for most if not all of the co-product gam_a-butyrolactone. The plant operator can in this case adjust the output of the plant to meet marke-t requirements either -35 by adjusting the operating temperatures of the - 9 ~
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. This means that 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 10C, and even more preferably at ]east about 15C, 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. Alternatively such 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.
Reduction of a ma~eate or fumarate ester to butane-1,4-diol involves reaction of 5 moles of H2 with each mole of ester, according to the following equation:
R-O-CO-CH=CH-CO-O-R + 5H2 = 2ROH + HO(CH2)4OH (I) where R is an alkyl radical containing from 1 to 4 carbon atoms.
However, when a succinate ester is hydrogenolysed, only 4 moles of H2 are consumed:
R-O-CO-CH2-CH2-CO-O-R + 4H2 = 2ROH ~ HO(CH2)4OH (II) where R is as defined above.
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. N2, A, CH4 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. Preferably the hydrogen supply is substantially free from sulphur compounds, from halogens such as C12, and from halogen-containing compounds such as HCl.
In the vaporous mixture the H2:ester molar ratio is typically at least about 100:1 up to about 800:1 or more. Preferably it is at ]east 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.
In practice the reduction of a maleate ester, such as diethyl maleate, is more complex than is suggested by equation ~I) above and results in production of variable amounts of by-products, including tetrahydrofuran, gamma- butyrolactone and n-butanol. Although the reaction mechanism has not been fully elucidated yet, the currently available evidence is consistent with the following sequence:-3~ d CH - CO2Et +H2 CH2 - Co2Et CH - CO2Et CH2 - CO2Et +2H2 -EtOH
~0 ~ ~ .
CH2 - C ~ +EtOH CH2 - CO2Et / CH2 - CH2 ~ -EtOHCH2 CH2OH
+332 ~+232 ~\ ~t E2
2.0 / / CH2 - - CH2 - OH
CH3cH2cH2cH2oll~ / -H20 CH2 -- CH2 +H2 l l -H2O
O ~ ~
r~7 The preferred catalyst used in the first and/or second hydrogenolysis zone is a reduced copper chromite catalyst. This should be prepared by careful reduction of copper chromite prior to use. Preferably the catalyst is reduced at a temperature of not more than about 200C, for an extended period using a mixture of H2 and an inert gas, such as nitrogen, methane or argon. A typical gas used for reduction of the catalystis an H2 in N2 mixture containing, for example, from about 1% up to about 15~ by volume of H2. Usually the catalyst is reduced for at least about 24 hours prior to use. Best results are obtained when reduction is effected for several days at a temperature of from about 120C to about 180C prior to use in the process of the invention. It will usually be unnecessary to exceed about 10 days reduction pretreatment. If the catalyst is reduced at a temperature in excess of about 200C, the activity is noticeably inferior to the activity obtained by reduction at lower temperatures. If the catalyst is supplied in pre-reduced form then the period of reduction can be shorter. In the later stages of pre-treatment higher H2 concentrations can be used; thus H2 can replace the H2/N2 mixture towards the end of the reduction pre-treatment. It is best to use an elevated pressure during this pre-treatment period; for example, pressures of from 1 bar up to about 50 bar or higher can be used. After the reduction treatment the catalyst should be maintained under an inert gas, a hydrogen/inert gas mixture or hydrogen until use.
The formula of copper chromite may be written as CuCr2O4. However, it is known that copper chromite is non-stoichiometric and some authors have, for example, described a copper chromite catalyst as copper ^hromium oxide of the formula CuO.CuCr2O4. Thus the catalyst may - contain an excess of copper oxide. It may ~urther 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 sultable 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.
Other catalysts that may be used in the first and/or second hydrogenolysis zone of the process of the invention include a reduced copper oxide/zinc oxide catalyst of the type disclosed in ~O-A-82/03854.
2G 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~l up to about 0.6 hr~l or higher, for example up to about 1.5 hr~l or even up to about 3.0 hr~l. By the term "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. This normally corresponds to a gaseous hourly space velocity in the range of from about 2500 hr 1 up to about 160000 hr~l, for example up to about 30 85000 hr~l, most preferably in the range of from about 8000 hr~l to about 30000 hr~l. By the term "gaseous hourly space velocity" we mean the number of unit volumes of vaporous mixture measured at 1 bar and 0C passed over a unit volume of catalyst per hour.
Usually the feed temperature to the second }, ? 7 hydrogenolysis zone is not more than about 175C and even more preferably lies in the range of from about 160C to about 175C.
If desired 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 H2:ester molar ratio. It is also contemplated that one or more materials recovered from the product recovery section of the plant (e.g. dialkyl succinate, unreacted dialkyl ma]eate or fumarate, ga_ma-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.
The reduction of a di-(Cl to C4 alkyl) maleate to the corresponding di-(Cl to C4 alkyl) succinate according to equation (III) above is an exothermic reaction. As the yield of by-product tetrahydrofuran is found to be dependent upon the maximum temperature to which the reaction mixture is exposed in the hydrogenolysis zones and upon the residence time at elevated temperatures, it is preferred to use a di-(Cl to C4 alkyl) succinate as the ester supplied to the first hydrogenolysis zone or as a major component thereo~, rather than the corresponding di-(Cl to C4 alkyl) maleate or fumarate, since this results in production of a correspondingly smaller temperature rise in the first hydrogenolysis zone and hence in production of a smaller proportion of by-product tetrahydrofuran. Hence, in a particularly preferred process according to the invention a di-(Cl to C4 alkyl) maleate or fumarate, such as diethyl maleate or fumarate, is hydrogenated to the corresponding di-(Cl to C4 alkyl) succinate, such as diethyl succinate, in a hydrogenation zone upstream from the first hydrogenolysis zone using 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-1,4-diol are minimised. Prior to entry to the first hydrogenolysis zone, the di-(Cl to C4 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. In this way the temperature rise in the first hydrogenolysis zone is kept as small as possible and the production of by-product tetrahydrofuran is thereby minimised.
As examples of catalysts that can be used in the upstream hydrogenation zone when using a di-(Cl to C4 alkyl) maleate or fumarate, there can be mentioned, for example, copper oxide/zinc oxide catalysts of the type disclosed in WO-A-82/03854, supported nickel, palladium, ruthenium, and cobalt hydrogenation catalysts, zinc oxide, copper chromite catalysts, and copper chromite catalysts which have been deliberately deactivated to reduce their ester hydrogenolysis characteristics. When using copper chromite catalysts, which may or may not have been deactivated, it is expedient to supply the di-(Cl to C4 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~l, preferably from about 3.0 hr~l to about 6.0 hr~l.
For convenience and ease of construction and operation of the plant it is preferred to operate any upstream hydrogenation zone for conversion of dialkyl maleate and/or fumarate to dialkyl succinate at 9~3~
substantially the same pressure as is used in the first and second hydrogenolysis zones. Hence 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 160C to about 180C. 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. 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 O 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 mater~als 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-1,4-diol and alkyl alcohol (e.g. a C1 to C4 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, - 17 ~
including n-butanol and tetrahydrofuran. These condensible materials are preferably condensed from the product mixture and separated in any suitable fashion, e.g. by distillation in one or more stages under normalr elevated or reduced pressure. In designing a suitable product recovery system, it should be borne in mind that some of the components present in the product mixture are capable of forming azeotropic mixtures with one or more other components of the product mixture. The liquid butane-1,4-diol product and any butyrolactone formed can be passed forward for purification whilst any minor byproducts can be used as fuel for the process. The alkyl alcohol can be recycled for reaction with further maleic or succinic anhydride or with further maleic acid, fumaric acid, or succinic acid to form fresh dialkyl C4 dicarboxylic acid ester for use in the process of the invention. Any unreacted starting ester (e.g. dialkyl maleate) and/or intermediate ester (e.g. dialkyl succinate~
can be recycled for admixture with the ester or ester solution feed. If desired 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.
In order that the invention may be clearly understood and readily carried into effect a preferred form of plant designed to operate a process in accordance with the invention will now be described, by way of example only, with reference to the accompanying drawing which is a flow diagram of a plant for effecting hydrogenolysis of diethyl maleate.
It will be appreciated by those skilled in the art that, as the drawing is diagrammatic, many further - ]8 ~ d ~ 3 ~
items of equipmentr such as temperature and pressure sensors, valves, control equipment, and the like, will be required in a commercial plant. The provision of such ancillary items of equipment forms no part of the present invention and will be in accordance with conventional chemical engineering practice. Moreover the provisions for heat exchange within the illustrated plant are merely one way of achieving the desired temperature levels and any other equivalent system of heat exchange can be used instead.
Referring to the drawing, 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 210C 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 166C and passes via heat exchanger 15 to a steam heater 16 in which its temperature is raised to 170C. The H2: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 m2/g The vaporous reaction mixture exits first bed 17 at an exit temperature of about 185C. Analysis of this first reaction mixture indicates the absence of diethyl a-.,P,~7 _ ~9 _ maleate and shows that, besides hydrogen and inert gases (e.g. CO, CO2, methane, propane, N2, A and any other inert gases which may be present in the hydrogen supply to the plant), it contains also significant amounts of diethyl succinate, ethanol, tetrahydrofuran, n-butanol, garn a-butyrolactone, and butane-1,4-diol. In passage through first bed 17 under the substantially adiabatic reaction conditions employed in the process, 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 n-butanol of 0.2 mol%, a selectivity to qam_a-butyrolactone of 16.0 mol~, and a selectivity to butane-1,4-diol of 79.3 mol%, the balance being minor byproducts. Hence the butane-1,4-diol:qamma-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 oE
about 0.45 hr~l.
The vaporous first reaction mixture exits first bed 17 at about 1~5C and is passed by way of line 18 to heat exchanger 15, in which it is cooled to 170C, and is then passed to a second bed 19 of the sarne eopper chromite eatalyst. In passage through bed 19 Further hydrogenation reaetions take p]aee and the reaction mixture re-equilibriates to yield a second reaetion mixture whieh exits the second hydrogenolysis ~one ~ormed by catalyst bed 19 at a temperature of 171-172C. The butane-1,4-diol:gamma-butyrolactone rnolar ratio in this second _ _ reaction mixture is 9:1. The volume of catalyst in the
CH3cH2cH2cH2oll~ / -H20 CH2 -- CH2 +H2 l l -H2O
O ~ ~
r~7 The preferred catalyst used in the first and/or second hydrogenolysis zone is a reduced copper chromite catalyst. This should be prepared by careful reduction of copper chromite prior to use. Preferably the catalyst is reduced at a temperature of not more than about 200C, for an extended period using a mixture of H2 and an inert gas, such as nitrogen, methane or argon. A typical gas used for reduction of the catalystis an H2 in N2 mixture containing, for example, from about 1% up to about 15~ by volume of H2. Usually the catalyst is reduced for at least about 24 hours prior to use. Best results are obtained when reduction is effected for several days at a temperature of from about 120C to about 180C prior to use in the process of the invention. It will usually be unnecessary to exceed about 10 days reduction pretreatment. If the catalyst is reduced at a temperature in excess of about 200C, the activity is noticeably inferior to the activity obtained by reduction at lower temperatures. If the catalyst is supplied in pre-reduced form then the period of reduction can be shorter. In the later stages of pre-treatment higher H2 concentrations can be used; thus H2 can replace the H2/N2 mixture towards the end of the reduction pre-treatment. It is best to use an elevated pressure during this pre-treatment period; for example, pressures of from 1 bar up to about 50 bar or higher can be used. After the reduction treatment the catalyst should be maintained under an inert gas, a hydrogen/inert gas mixture or hydrogen until use.
The formula of copper chromite may be written as CuCr2O4. However, it is known that copper chromite is non-stoichiometric and some authors have, for example, described a copper chromite catalyst as copper ^hromium oxide of the formula CuO.CuCr2O4. Thus the catalyst may - contain an excess of copper oxide. It may ~urther 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 sultable 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.
Other catalysts that may be used in the first and/or second hydrogenolysis zone of the process of the invention include a reduced copper oxide/zinc oxide catalyst of the type disclosed in ~O-A-82/03854.
2G 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~l up to about 0.6 hr~l or higher, for example up to about 1.5 hr~l or even up to about 3.0 hr~l. By the term "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. This normally corresponds to a gaseous hourly space velocity in the range of from about 2500 hr 1 up to about 160000 hr~l, for example up to about 30 85000 hr~l, most preferably in the range of from about 8000 hr~l to about 30000 hr~l. By the term "gaseous hourly space velocity" we mean the number of unit volumes of vaporous mixture measured at 1 bar and 0C passed over a unit volume of catalyst per hour.
Usually the feed temperature to the second }, ? 7 hydrogenolysis zone is not more than about 175C and even more preferably lies in the range of from about 160C to about 175C.
If desired 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 H2:ester molar ratio. It is also contemplated that one or more materials recovered from the product recovery section of the plant (e.g. dialkyl succinate, unreacted dialkyl ma]eate or fumarate, ga_ma-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.
The reduction of a di-(Cl to C4 alkyl) maleate to the corresponding di-(Cl to C4 alkyl) succinate according to equation (III) above is an exothermic reaction. As the yield of by-product tetrahydrofuran is found to be dependent upon the maximum temperature to which the reaction mixture is exposed in the hydrogenolysis zones and upon the residence time at elevated temperatures, it is preferred to use a di-(Cl to C4 alkyl) succinate as the ester supplied to the first hydrogenolysis zone or as a major component thereo~, rather than the corresponding di-(Cl to C4 alkyl) maleate or fumarate, since this results in production of a correspondingly smaller temperature rise in the first hydrogenolysis zone and hence in production of a smaller proportion of by-product tetrahydrofuran. Hence, in a particularly preferred process according to the invention a di-(Cl to C4 alkyl) maleate or fumarate, such as diethyl maleate or fumarate, is hydrogenated to the corresponding di-(Cl to C4 alkyl) succinate, such as diethyl succinate, in a hydrogenation zone upstream from the first hydrogenolysis zone using 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-1,4-diol are minimised. Prior to entry to the first hydrogenolysis zone, the di-(Cl to C4 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. In this way the temperature rise in the first hydrogenolysis zone is kept as small as possible and the production of by-product tetrahydrofuran is thereby minimised.
As examples of catalysts that can be used in the upstream hydrogenation zone when using a di-(Cl to C4 alkyl) maleate or fumarate, there can be mentioned, for example, copper oxide/zinc oxide catalysts of the type disclosed in WO-A-82/03854, supported nickel, palladium, ruthenium, and cobalt hydrogenation catalysts, zinc oxide, copper chromite catalysts, and copper chromite catalysts which have been deliberately deactivated to reduce their ester hydrogenolysis characteristics. When using copper chromite catalysts, which may or may not have been deactivated, it is expedient to supply the di-(Cl to C4 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~l, preferably from about 3.0 hr~l to about 6.0 hr~l.
For convenience and ease of construction and operation of the plant it is preferred to operate any upstream hydrogenation zone for conversion of dialkyl maleate and/or fumarate to dialkyl succinate at 9~3~
substantially the same pressure as is used in the first and second hydrogenolysis zones. Hence 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 160C to about 180C. 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. 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 O 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 mater~als 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-1,4-diol and alkyl alcohol (e.g. a C1 to C4 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, - 17 ~
including n-butanol and tetrahydrofuran. These condensible materials are preferably condensed from the product mixture and separated in any suitable fashion, e.g. by distillation in one or more stages under normalr elevated or reduced pressure. In designing a suitable product recovery system, it should be borne in mind that some of the components present in the product mixture are capable of forming azeotropic mixtures with one or more other components of the product mixture. The liquid butane-1,4-diol product and any butyrolactone formed can be passed forward for purification whilst any minor byproducts can be used as fuel for the process. The alkyl alcohol can be recycled for reaction with further maleic or succinic anhydride or with further maleic acid, fumaric acid, or succinic acid to form fresh dialkyl C4 dicarboxylic acid ester for use in the process of the invention. Any unreacted starting ester (e.g. dialkyl maleate) and/or intermediate ester (e.g. dialkyl succinate~
can be recycled for admixture with the ester or ester solution feed. If desired 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.
In order that the invention may be clearly understood and readily carried into effect a preferred form of plant designed to operate a process in accordance with the invention will now be described, by way of example only, with reference to the accompanying drawing which is a flow diagram of a plant for effecting hydrogenolysis of diethyl maleate.
It will be appreciated by those skilled in the art that, as the drawing is diagrammatic, many further - ]8 ~ d ~ 3 ~
items of equipmentr such as temperature and pressure sensors, valves, control equipment, and the like, will be required in a commercial plant. The provision of such ancillary items of equipment forms no part of the present invention and will be in accordance with conventional chemical engineering practice. Moreover the provisions for heat exchange within the illustrated plant are merely one way of achieving the desired temperature levels and any other equivalent system of heat exchange can be used instead.
Referring to the drawing, 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 210C 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 166C and passes via heat exchanger 15 to a steam heater 16 in which its temperature is raised to 170C. The H2: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 m2/g The vaporous reaction mixture exits first bed 17 at an exit temperature of about 185C. Analysis of this first reaction mixture indicates the absence of diethyl a-.,P,~7 _ ~9 _ maleate and shows that, besides hydrogen and inert gases (e.g. CO, CO2, methane, propane, N2, A and any other inert gases which may be present in the hydrogen supply to the plant), it contains also significant amounts of diethyl succinate, ethanol, tetrahydrofuran, n-butanol, garn a-butyrolactone, and butane-1,4-diol. In passage through first bed 17 under the substantially adiabatic reaction conditions employed in the process, 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 n-butanol of 0.2 mol%, a selectivity to qam_a-butyrolactone of 16.0 mol~, and a selectivity to butane-1,4-diol of 79.3 mol%, the balance being minor byproducts. Hence the butane-1,4-diol:qamma-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 oE
about 0.45 hr~l.
The vaporous first reaction mixture exits first bed 17 at about 1~5C and is passed by way of line 18 to heat exchanger 15, in which it is cooled to 170C, and is then passed to a second bed 19 of the sarne eopper chromite eatalyst. In passage through bed 19 Further hydrogenation reaetions take p]aee and the reaction mixture re-equilibriates to yield a second reaetion mixture whieh exits the second hydrogenolysis ~one ~ormed by catalyst bed 19 at a temperature of 171-172C. The butane-1,4-diol:gamma-butyrolactone rnolar ratio in this second _ _ reaction mixture is 9:1. The volume of catalyst in the
3~
~ 20 -second bed 19 is approximately twice that in the first bed 17; hence the rate at which vaporous diethyl maleate is supplied to bed 17 corresponds to a liquid hourly space velocity, taken over both beds 17 and 19, of about 0.15 hr 1.
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 ga_ma-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, n-butanol and other low boiling byproducts, and then distillation of the resulting bottoms product to yield an overhead product comprising an azeotrope of ~amma-butyrolactone and diethyl succinate and a bottom product comprising butane-1,4-diol.
Tetrahydrofuran is recovered in line 30, _a__a-butyrolactone in line 31, and butane-1,4-diol in line 32.
Diethyl succinate, and possibly also some g_mma-butyrolactone and/or butane-1,4-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.
~J~
-- 21 ~
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.
If desired, 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 ga__a-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-1,4-diol to tetrahydrofuran. This zone is maintained at a temperature in the range of from about 200C to about 300C. (Alternatively 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.
It will readily be appreciated by those skilled in the art that the illustrated plant can, with relatively little modification, be operated using another dialkyl ester of a C4 dicarboxylic acid as feedstock, such as diethyl succinate or a mixture of diethyl maleate and diethyl succinate, in place of diethyl maleate. Some changes to the design of the product recovery section 29 - 22 ~ 3~ .J~
may, however, be required. Such 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.
In an alternative modification of the p].ant of the drawing diethyl maleate supplied in line 1 is hydrogenated in the vapour phase to diethyl succinate in a hydrogenation zone tnot shown) in line 14 upstream from the first hydrogenolysis zone~ Such upstream 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~l.
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-1,4-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.
~ 20 -second bed 19 is approximately twice that in the first bed 17; hence the rate at which vaporous diethyl maleate is supplied to bed 17 corresponds to a liquid hourly space velocity, taken over both beds 17 and 19, of about 0.15 hr 1.
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 ga_ma-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, n-butanol and other low boiling byproducts, and then distillation of the resulting bottoms product to yield an overhead product comprising an azeotrope of ~amma-butyrolactone and diethyl succinate and a bottom product comprising butane-1,4-diol.
Tetrahydrofuran is recovered in line 30, _a__a-butyrolactone in line 31, and butane-1,4-diol in line 32.
Diethyl succinate, and possibly also some g_mma-butyrolactone and/or butane-1,4-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.
~J~
-- 21 ~
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.
If desired, 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 ga__a-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-1,4-diol to tetrahydrofuran. This zone is maintained at a temperature in the range of from about 200C to about 300C. (Alternatively 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.
It will readily be appreciated by those skilled in the art that the illustrated plant can, with relatively little modification, be operated using another dialkyl ester of a C4 dicarboxylic acid as feedstock, such as diethyl succinate or a mixture of diethyl maleate and diethyl succinate, in place of diethyl maleate. Some changes to the design of the product recovery section 29 - 22 ~ 3~ .J~
may, however, be required. Such 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.
In an alternative modification of the p].ant of the drawing diethyl maleate supplied in line 1 is hydrogenated in the vapour phase to diethyl succinate in a hydrogenation zone tnot shown) in line 14 upstream from the first hydrogenolysis zone~ Such upstream 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~l.
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-1,4-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.
Claims (17)
1. A process for the production of butane-1,4-diol by hydrogenolysis of a dialkyl ester of a C4 dicarboxylic acid which comprises:
providing first and second hydrogenolysis zones, each containing a charge of a heterogeneous ester hydrogenolysis catalyst;
supplying to the first hydrogenolysis zone at an elevated pressure and at an elevated first feed temperature in excess of the threshold temperature for the hydrogenolysis reaction a vaporous feed stream comprising a dialkyl ester of a C4 dicarboxylic acid in vapour form and excess hydrogen;
allowing the ester to undergo hydrogenolysis in the first hydrogenolysis zone under substanttially adiabatic reaction conditions thereby to form a vaporous first reaction mixture that is substantially free from the starting ester and contains, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a first molar ratio;
cooling the vaporous first reaction mixture;
supplying a second vaporous feed stream comprising resulting cooled vaporous first reaction mixture to the second hydrogenolysis zone at a second feed temperature in excess of the threshold temperature for the hydrogenolysis reaction;
allowing the second vaporous feed stream to react further and to equilibriate under substantially adiabatic reaction conditions in the second hydrogenolysis zone; and recovering from the second hydrogenolysis zone a vaporous second reaction mixture containing, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a second molar ratio that is greater than the first molar ratio.
providing first and second hydrogenolysis zones, each containing a charge of a heterogeneous ester hydrogenolysis catalyst;
supplying to the first hydrogenolysis zone at an elevated pressure and at an elevated first feed temperature in excess of the threshold temperature for the hydrogenolysis reaction a vaporous feed stream comprising a dialkyl ester of a C4 dicarboxylic acid in vapour form and excess hydrogen;
allowing the ester to undergo hydrogenolysis in the first hydrogenolysis zone under substanttially adiabatic reaction conditions thereby to form a vaporous first reaction mixture that is substantially free from the starting ester and contains, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a first molar ratio;
cooling the vaporous first reaction mixture;
supplying a second vaporous feed stream comprising resulting cooled vaporous first reaction mixture to the second hydrogenolysis zone at a second feed temperature in excess of the threshold temperature for the hydrogenolysis reaction;
allowing the second vaporous feed stream to react further and to equilibriate under substantially adiabatic reaction conditions in the second hydrogenolysis zone; and recovering from the second hydrogenolysis zone a vaporous second reaction mixture containing, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a second molar ratio that is greater than the first molar ratio.
2. A process according to claim 1, in which the heterogeneous hydrogenolysis catalyst comprises copper chromite.
3. A process according to claim 2, in which the copper chromite 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.
4. A process according to claim 1, in which the first feed temperature ranges from about 150°C to about 200°C.
5. A process according to claim 4, in which the first feed temperature is from about 170°C to about 190°C.
6. A process according to claim 1, in which the vaporous first reaction mixture is recovered from the first hydrogenolysis zone at a temperature in the range of from about 170°C to about 200°C.
7. A process according to claim 1, in which the second feed temperature is at least about 5°C lower than the temperature at which the vaporous first reaction mixture exits the first hydrogenolysis zone.
8. A process according to claim 1, in which the second feed temperature lies in the range of from about 150°C to about 190°C.
9. A process according to claim 1, in which the hydrogen:ester molar ratio in the first vaporous feed stream is in the range of from about 100:1 to about 800:1.
10. A process according to claim 1, in which the first feed temperature is at least about 5°C above the dew point of the first vaporous feed stream.
11. A process according to claim 1, in which the second feed temperature is at least about 5°C above the dew point of the second vaporous feed stream.
12. A process according to claim 1, in which the pressure lies in the range of from about 25 bar to about 70 bar.
13. A process according to claim 1, in which the ester is fed to the first hydrogenolysis zone at a rate corresponding to a liquid hourly space velocity of from about 0.1 hr-1 to about 3.0 hr-1.
14. A process according to claim 1, in which the ester is a di-(C1 to C4 alkyl) ester of a C4 dicarboxylic acid.
15. A process according to claim 14, in which the ester is selected from diethyl maleate, dialkyl fumarate and a mixture thereof.
16. A process according to claim 14, in which the ester is selected from diethyl succinate, and mixtures thereof with one or both of diethyl maleate and diethyl fumarate.
17. In a process for the production of tetrahydrofuran by dehydration of butane-1,4-diol the improvement comprising producing butane-1,4-diol by the steps comprising:
providing first and second hydrogenolysis zones, each containing a charge of a heterogeneous ester hydrogenolysis catalyst;
supplying to the first hydrogenolysis zone at an elevated pressure and at an elevated first feed temperature in excess of the threshold temperature for the hydrogenolysis reaction a vaporous feed stream comprising a dialkyl ester of a C4 dicarboxylic acid in vapour form 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 starting ester and contains, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a first molar ratio;
cooling the vaporous first reaction mixture;
supplying a second vaporous feed stream comprising resulting cooled vaporous first reaction mixture to the second hydrogenolysis zone at a second feed temperature in excess of the threshold temperature for the hydrogenolysis reaction;
allowing the second vaporous feed stream to react further and to equilibriate under substantially adiabatic reaction conditions in the second hydrogenolysis zone; and recovering from the second hydrogenolysis zone a vaporous second reaction mixture containing, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a second molar ratio that is greater than the first molar ratio.
providing first and second hydrogenolysis zones, each containing a charge of a heterogeneous ester hydrogenolysis catalyst;
supplying to the first hydrogenolysis zone at an elevated pressure and at an elevated first feed temperature in excess of the threshold temperature for the hydrogenolysis reaction a vaporous feed stream comprising a dialkyl ester of a C4 dicarboxylic acid in vapour form 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 starting ester and contains, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a first molar ratio;
cooling the vaporous first reaction mixture;
supplying a second vaporous feed stream comprising resulting cooled vaporous first reaction mixture to the second hydrogenolysis zone at a second feed temperature in excess of the threshold temperature for the hydrogenolysis reaction;
allowing the second vaporous feed stream to react further and to equilibriate under substantially adiabatic reaction conditions in the second hydrogenolysis zone; and recovering from the second hydrogenolysis zone a vaporous second reaction mixture containing, in addition to unreacted hydrogen, butane-1,4-diol and gamma-butyrolactone in a second molar ratio that is greater than the first molar ratio.
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 |
Publications (1)
| Publication Number | Publication Date |
|---|---|
| CA1249837A true CA1249837A (en) | 1989-02-07 |
Family
ID=26289326
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA000495906A Expired CA1249837A (en) | 1984-11-21 | 1985-11-21 | Process for the production of butane-1,4-diol |
Country Status (6)
| Country | Link |
|---|---|
| EP (1) | EP0204730A1 (en) |
| JP (1) | JPH0655684B2 (en) |
| BR (1) | BR8507068A (en) |
| CA (1) | CA1249837A (en) |
| ES (1) | ES8701703A1 (en) |
| WO (1) | WO1986003189A1 (en) |
Families Citing this family (41)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| DE3503485A1 (en) * | 1985-02-01 | 1986-08-07 | Lentia GmbH Chem. u. pharm. Erzeugnisse - Industriebedarf, 8000 München | METHOD FOR PRODUCING PURE DIALKYLSUCCINATES |
| WO1988000937A1 (en) * | 1986-08-01 | 1988-02-11 | Davy Mckee (London) Limited | Process for the co-production of butane-1,4-diol and gamma-butyrolactone |
| GB8618892D0 (en) * | 1986-08-01 | 1986-09-10 | Davy Mckee Ltd | Process |
| GB8717989D0 (en) * | 1987-07-29 | 1987-09-03 | Davy Mckee Ltd | Catalyst |
| GB8717993D0 (en) * | 1987-07-29 | 1987-09-03 | Davy Mckee Ltd | Process |
| US4810807A (en) * | 1987-10-13 | 1989-03-07 | The Standard Oil Company | Hydrogenation of maleic anhydride to tetrahydrofuran |
| JPH02233674A (en) * | 1989-03-07 | 1990-09-17 | Mitsubishi Kasei Corp | Production of lactones |
| GB8917862D0 (en) * | 1989-08-04 | 1989-09-20 | Davy Mckee London | Process |
| GB8917864D0 (en) * | 1989-08-04 | 1989-09-20 | Davy Mckee London | Process |
| GB8917859D0 (en) * | 1989-08-04 | 1989-09-20 | Davy Mckee London | Process |
| DE4005293A1 (en) * | 1990-02-20 | 1991-08-22 | Basf Ag | PROCESS FOR THE MANUFACTURING OF TETRAHYDROFURAN AND BUTYROLACTONE |
| DE4231782A1 (en) * | 1992-09-23 | 1994-03-24 | Basf Ag | Process for the preparation of 2-methyl-1,4-butanediol and 3-methyltetrahydrofuran |
| US5536854A (en) * | 1993-09-21 | 1996-07-16 | Basf Aktiengesellschaft | Preparation of 2-methyl-1,4-butanediol and 3-methyltetrahydrofuran |
| GB9324782D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| GB9324786D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| GB9324784D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| GB9324823D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| GB9324752D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| GB9324785D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| GB9324753D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| GB9324783D0 (en) * | 1993-12-02 | 1994-01-19 | Davy Mckee London | Process |
| IT1273320B (en) * | 1994-02-22 | 1997-07-08 | Alusuisse Italia Spa | PROCEDURE FOR THE PRODUCTION OF RANGE-BUTYROLACTONE |
| DE4431220A1 (en) | 1994-09-02 | 1996-03-07 | Bayer Ag | Process for the preparation of 1,4-butanediol from maleic anhydride |
| ZA973972B (en) * | 1996-05-14 | 1998-03-23 | Kvaerner Process Tech Ltd | A process for the production of at least one C4 compound selected from butane-1,4-diol, gamma-butyrolactone and tetrahydrofuran. |
| ZA973971B (en) * | 1996-05-15 | 1998-03-23 | Kvaerner Process Tech Ltd | A process for the production of at least one C4 compound selected from butane-1,4-diol, gamma-butyrolactone and tetrahydrofuran. |
| GB9724004D0 (en) | 1997-11-13 | 1998-10-21 | Kvaerner Process Tech Ltd | Process |
| GB9724195D0 (en) | 1997-11-14 | 1998-01-14 | Kvaerner Process Tech Ltd | Process |
| BE1011698A6 (en) | 1998-01-08 | 1999-12-07 | Pantochim Sa | Method for producing tetrahydrofuran, gamma-butyrolactone and butanediol. |
| IT1298096B1 (en) | 1998-01-09 | 1999-12-20 | Lonza Spa | PROCEDURE FOR THE PRODUCTION OF RANGE-BUTYROLACTONE |
| IT1298535B1 (en) | 1998-02-02 | 2000-01-12 | Lonza Spa | PROCEDURE FOR THE PRODUCTION OF RANGE-BUTYROLACTONE |
| ES2158645T3 (en) | 1998-03-23 | 2001-09-01 | Basf Ag | PROCEDURE FOR THE PREPARATION OF 1,4-BUTANODIOL, BUTIROLACTONE AND TETRAHYDROFURAN. |
| DE19842847A1 (en) * | 1998-09-18 | 2000-03-23 | Basf Ag | Production of tetrahydrofuran, useful as solvent and intermediate for polytetramethylene glycol, involves reacting anhydrous mixture of butan-1,4-diol and other alcohol on acid catalyst |
| EP1108702A1 (en) | 1999-12-13 | 2001-06-20 | Kvaerner Process Technology Limited | Process for the co-production of aliphatic diols and cyclic ethers |
| GB0117090D0 (en) * | 2001-07-12 | 2001-09-05 | Kvaerner Process Tech Ltd | Process |
| GB0329152D0 (en) * | 2003-12-16 | 2004-01-21 | Davy Process Techn Ltd | Process |
| GB0803663D0 (en) | 2008-02-28 | 2008-04-09 | Davy Process Techn Ltd | Process |
| GB201318175D0 (en) | 2013-10-14 | 2013-11-27 | Johnson Matthey Davy Technologies Ltd | Process |
| GB201321627D0 (en) | 2013-12-06 | 2014-01-22 | Johnson Matthey Davy Technologies Ltd | Process |
| GB201321611D0 (en) | 2013-12-06 | 2014-01-22 | Johnson Matthey Davy Technologies Ltd | Process |
| GB201507234D0 (en) | 2015-04-28 | 2015-06-10 | Johnson Matthey Davy Technologies Ltd | Process |
| KR20220110820A (en) * | 2019-12-10 | 2022-08-09 | 바스프 에스이 | Process for gas phase production of 1,4-butanediol, gamma-butyrolactone and tetrahydrofuran while avoiding polymer deposits |
Family Cites Families (2)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US4032458A (en) * | 1975-08-08 | 1977-06-28 | Petro-Tex Chemical Corporation | Production of 1,4-butanediol |
| GB8331793D0 (en) * | 1983-11-29 | 1984-01-04 | Davy Mckee Ltd | Process |
-
1985
- 1985-11-18 JP JP61500083A patent/JPH0655684B2/en not_active Expired - Lifetime
- 1985-11-18 EP EP19850905648 patent/EP0204730A1/en not_active Withdrawn
- 1985-11-18 WO PCT/GB1985/000524 patent/WO1986003189A1/en not_active Application Discontinuation
- 1985-11-18 BR BR8507068A patent/BR8507068A/en unknown
- 1985-11-21 ES ES549746A patent/ES8701703A1/en not_active Expired
- 1985-11-21 CA CA000495906A patent/CA1249837A/en not_active Expired
Also Published As
| Publication number | Publication date |
|---|---|
| ES8701703A1 (en) | 1986-12-01 |
| WO1986003189A1 (en) | 1986-06-05 |
| BR8507068A (en) | 1987-07-14 |
| ES549746A0 (en) | 1986-12-01 |
| JPS62501702A (en) | 1987-07-09 |
| EP0204730A1 (en) | 1986-12-17 |
| JPH0655684B2 (en) | 1994-07-27 |
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