EP2046706A1 - Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants - Google Patents

Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants

Info

Publication number
EP2046706A1
EP2046706A1 EP07787557A EP07787557A EP2046706A1 EP 2046706 A1 EP2046706 A1 EP 2046706A1 EP 07787557 A EP07787557 A EP 07787557A EP 07787557 A EP07787557 A EP 07787557A EP 2046706 A1 EP2046706 A1 EP 2046706A1
Authority
EP
European Patent Office
Prior art keywords
stream
water
product mixture
isopropanol
distillation column
Prior art date
Legal status (The legal status is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the status listed.)
Withdrawn
Application number
EP07787557A
Other languages
German (de)
English (en)
Inventor
Georg Degen
Sven Crone
Ralf Boehling
Ansgar Gereon Altenhoff
Wolfgang Rohde
Jochen Buerkle
Goetz-Peter Schindler
Thomas Holtmann
Markus Schmitt
Current Assignee (The listed assignees may be inaccurate. Google has not performed a legal analysis and makes no representation or warranty as to the accuracy of the list.)
BASF SE
Original Assignee
BASF SE
Priority date (The priority date is an assumption and is not a legal conclusion. Google has not performed a legal analysis and makes no representation as to the accuracy of the date listed.)
Filing date
Publication date
Application filed by BASF SE filed Critical BASF SE
Priority to EP07787557A priority Critical patent/EP2046706A1/fr
Priority to EP09153296A priority patent/EP2070900A3/fr
Publication of EP2046706A1 publication Critical patent/EP2046706A1/fr
Withdrawn legal-status Critical Current

Links

Classifications

    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C67/00—Preparation of carboxylic acid esters
    • C07C67/04—Preparation of carboxylic acid esters by reacting carboxylic acids or symmetrical anhydrides onto unsaturated carbon-to-carbon bonds
    • 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
    • C07C29/09—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by hydrolysis
    • 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
    • C07C29/09—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by hydrolysis
    • C07C29/095—Preparation of compounds having hydroxy or O-metal groups bound to a carbon atom not belonging to a six-membered aromatic ring by hydrolysis of esters of organic acids
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C31/00—Saturated compounds having hydroxy or O-metal groups bound to acyclic carbon atoms
    • C07C31/02—Monohydroxylic acyclic alcohols
    • C07C31/10—Monohydroxylic acyclic alcohols containing three carbon atoms
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C31/00—Saturated compounds having hydroxy or O-metal groups bound to acyclic carbon atoms
    • C07C31/02—Monohydroxylic acyclic alcohols
    • C07C31/12—Monohydroxylic acyclic alcohols containing four carbon atoms
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/10—Process efficiency

Definitions

  • the invention relates to a process for the preparation of isopropanol from propane and from 2-butanol from butane.
  • the hydration of alkenes to alcohols is well known and is operated on an industrial scale.
  • the two-stage hydration in which the alkene is reacted with sulfuric acid to alkyl sulfate and this is cleaved in the second step with water to the alcohol and the acid.
  • An advantage of this process is that the alkene can be used in the form of alkene crude mixtures, for example mixtures of the alkene, the corresponding alkane and further secondary constituents.
  • a disadvantage is the highly corrosive medium, the contamination of the product with sulfur-containing odoriferous substances, which may require additional purification steps, and the loss of the inert alkane content of the mixture, which is unreacted and discharged after the reaction.
  • a further disadvantage is that, after hydrolysis of the alkyl sulfate to the alcohol, the resulting dilute sulfuric acid must be concentrated in the esterification step prior to its reuse.
  • GB 2 238 539 discloses the two-stage hydration of 1-butene by reaction with trifluoroacetic acid to give 2-butyltrifluoroacetate in the presence of a strongly acidic ion exchange resin and subsequent hydrolysis of the ester to 2-butanol.
  • the direct, one-stage hydration of alkenes on strongly acidic catalysts for example ion exchangers, zeolites, heteropolyacids and mineral acids, is carried out on heterogeneous supports.
  • Direct hydration processes can be carried out in the gas phase, in the liquid phase or in the biphasic phase.
  • Disadvantages of the single-stage process are, in particular, the low conversions and the high demands on the purity of the alkene used.
  • propylene must be used as polymer grade propylene.
  • GB-A-2 147 290 discloses a process for producing isopropanol, 2-butanol and methyl tertiary butyl ether from an LPG gas mixture containing propane, n-butane and iso-butane.
  • the gas mixture is dehydrogenated to a mixture containing propene, n-butenes and isobutene and then passes through an etherification zone in which isobutene is etherified with methanol to give methyl tert-butyl ether.
  • Propene and n-butenes which under the chosen conditions essentially do not react with methanol, are simultaneously hydrated directly with water to give isopropanol or 2-butanol, respectively.
  • US 2004/0225165 A discloses a process for the preparation of alcohols having 3 or more carbon atoms from the corresponding alkanes, in which a stream containing propane or a longer alkane alkane is converted to an intermediate stream containing the corresponding olefin, and the intermediate stream by direct or indirect hydration to a product stream containing the corresponding alcohol is reacted.
  • various methods of direct hydration are described.
  • the remaining gas stream which consists essentially of propene, isobutene and isobutane, is fed into a hydration zone, where propene and isobutene are hydrated on an acidic ion exchanger resin directly to isopropanol and tert-butanol.
  • the remaining gas stream is separated on the one hand into an isobutane stream which is recycled to the dehydrogenation zone and on the other hand into a stream containing propane and propene, which is recycled to the propane separation before the hydration step.
  • the process is thus characterized in that unreacted propane is separated off prior to hydration and the hydration is carried out with a feed stream consisting essentially of the C 3 and C 4 olefins.
  • the object of the invention is to provide an economical process for the preparation of isopropanol and 2-butanol, which does not have the disadvantages of the prior art.
  • alkanols (I) selected from the group consisting of isopropanol and 2-butanol, from the corresponding alkanes (II) selected from the group consisting of propane and n-butane, with the steps :
  • the product gas stream b is at least compressed, optionally the product gas stream b is separated into an aqueous phase d, a phase c2 containing the alkene (III), the alkane (II) and optionally high boilers and a gas phase c3 comprising hydrogen and low boilers;
  • the product gas stream b or the alkene (III) and alkane (II) containing phase c2 is reacted in an esterification zone with an organic acid (IV), wherein a product mixture d containing the corresponding alkyl ester (V) of the organic acid and the unreacted alkane (II) is obtained;
  • the product mixture d is used to separate an alkane (II) -containing gas stream e1, which is optionally recycled to the dehydrogenation zone, and a product mixture e2 containing the alkyl ester is obtained;
  • the inventive method dispenses with the use of highly corrosive sulfuric acid. It is characterized by the fact that even in the ester-forming step D), even when using a feed gas stream containing the alkene (I) only in very dilute form in addition to other components (unreacted alkane, inert gases), high space / time yields and conversions be achieved. On a separation of non-reactive secondary components before carrying out the esterification step can therefore be dispensed with. Furthermore, since the ester cleavage step F) can be carried out with water in stoichiometric excess, the organic acid (IV) can be recovered in concentrated form and recycled directly to esterification step D) without further concentration.
  • an alkane (II), selected from propane and butane, containing feed gas stream a In the first process part A) is provided an alkane (II), selected from propane and butane, containing feed gas stream a.
  • propane this generally contains at least 80% by volume of propane, preferably 90% by volume of propane.
  • propane generally still contains butane (n-butane, iso-butane), butenes, ethane and ethene.
  • Typical compositions of the propane-containing feed gas stream are disclosed in DE-A 102 46 1 19 and DE-A 102 45 585.
  • the propane-containing feed gas stream a is obtained from liquid petroleum gas (LPG).
  • the feed gas stream generally contains at least 80% by volume of n-butane, preferably 90% by volume of n-butane. In addition, it generally still contains ethane, ethene, propane, propene, isobutane, butenes and C 5 - hydrocarbons.
  • the feed gas stream containing the alkane (II) is fed into a dehydrogenation zone and subjected to a generally catalytic dehydrogenation.
  • the alkane is partially dehydrogenated in a dehydrogenation reactor on a dehydrogenation catalyst to the alkene.
  • hydrogen and small amounts of low boilers and high boilers are referred to below as propene or 1-butene boiling hydrocarbons, high boilers higher than propane or 2-butene boiling hydrocarbons referred to.
  • the dehydrogenation is carried out in the presence of an oxygen-containing gas
  • carbon oxides (CO, CO 2 ), in particular CO 2 and water vapor and, if appropriate, to a lesser extent inert gases in the product gas mixture of the dehydrogenation generally fall.
  • the product gas stream of the dehydrogenation generally contains water vapor which has already been added to the dehydrogenation gas mixture and / or, upon dehydrogenation in the presence of oxygen (oxidative or non-oxidative), is formed during the dehydrogenation.
  • Inert gases nitrogen
  • unreacted alkane (II) is present in the reaction gas mixture.
  • the alkane dehydrogenation can in principle be carried out in all reactor types known from the prior art.
  • a comparatively comprehensive description of the invention suitable reactor types also includes "Catalytica® ® Studies Division, Oxidative Dehydrogenation and Alternative Dehydrogenation Proces- ses" (Study Number 4192 OD, 1993, 430 Ferguson Drive, Mountain View, California, 94043-5272, USA).
  • the dehydrogenation can be carried out as oxidative or non-oxidative dehydrogenation.
  • the dehydration can be performed isothermally or adiabatically.
  • the dehydrogenation can be carried out catalytically in a fixed bed, moving bed or fluidized bed reactor.
  • the non-oxidative catalytic alkane dehydrogenation is preferably carried out autothermally.
  • oxygen is added to the reaction gas mixture of the dehydrogenation in at least one reaction zone and at least partially burned the hydrogen and / or hydrocarbon contained in the reaction gas mixture, whereby at least a portion of the required Dehydriereben in the at least one reaction zone is generated directly in the reaction gas mixture.
  • One feature of non-oxidative driving versus oxidative driving is the at least intermediate formation of hydrogen, which is reflected in the presence of hydrogen in the dehydrogenation product gas. In oxidative dehydrogenation, there is no free hydrogen in the dehydrogenation product gas.
  • a suitable reactor form is the fixed bed tube or tube bundle reactor.
  • the catalyst dehydrogenation catalyst and optionally special oxidation catalyst
  • the catalyst is a fixed bed in a reaction tube or in a bundle of reaction tubes.
  • Typical reaction tube internal diameters are about 10 to 15 cm.
  • a typical Dehydrierrohrbündelreaktor comprises about 300 to 1000 reaction tubes.
  • the temperature in the inside of the reaction tube usually ranges from 300 to 1200 ° C., preferably in the range from 500 to 1000 ° C.
  • the working pressure is usually between 0.5 and 8 bar, frequently between 1 and 2 bar when using a low steam dilution.
  • GHSV high steam dilution
  • STAR process steam active reforming process
  • Linde process the dehydrogenation of propane or butane by Phillips Petroleum Co. loads
  • GHSV high steam dilution
  • the catalyst geometry may be for example, spherical or cylindrical (hollow or solid). It can be operated in parallel and a plurality of fixed bed or tubular reactors, one of which turns at least one is in the state of regeneration.
  • the non-oxidative catalytic, autothermal dehydrogenation can also be carried out in a heterogeneously catalyzed fluidized bed according to the Snamprogetti / Yarsintez-FBD process.
  • two fluidized beds are operated side by side, one of which is usually in the state of regeneration.
  • the working pressure is typically 1 to 2 bar
  • the dehydrogenation temperature is generally 550 to 600 ° C.
  • the heat required for the dehydrogenation can be introduced into the reaction system by preheating the dehydrogenation catalyst to the reaction temperature.
  • an oxygen-containing co-feed can be dispensed with the preheater, and the heat required is generated directly in the reactor system by combustion of hydrogen and / or hydrocarbons in the presence of oxygen.
  • a hydrogen-containing co-feed may additionally be admixed.
  • the non-oxidative catalytic, autothermal dehydrogenation is preferably carried out in a tray reactor.
  • This contains one or more successive catalyst beds.
  • the number of catalyst beds can be 1 to 20, advantageously 1 to 6, preferably 1 to 4 and in particular 1 to 3.
  • the catalyst beds are preferably flowed through radially or axially from the reaction gas.
  • such a tray reactor is operated with a fixed catalyst bed.
  • the fixed catalyst beds are arranged in a shaft furnace reactor axially or in the annular gaps of concentrically arranged cylindrical gratings.
  • a shaft furnace reactor corresponds to a horde.
  • the performance of dehydrogenation in a single shaft furnace reactor corresponds to one embodiment.
  • the dehydrogenation is carried out in a tray reactor with 3 catalyst beds.
  • the amount of the oxygen-containing gas added to the reaction gas mixture is selected such that the combustion of mixtures of hydrogen present in the reaction gas mixture and optionally of hydrocarbons present in the reaction gas mixture and / or coke present in the form of coke for the dehydrogenation of the alkane (propane or n-butane) required amount of heat is generated.
  • the total amount of oxygen supplied based on the total amount of propane, 0.001 to 0.5 mol / mol, preferably 0.005 to 0.25 mol / mol, particularly preferably 0.05 to 0.25 mol / mol.
  • Oxygen can be used either as pure oxygen or as an oxygen-containing gas containing inert gases.
  • the oxygen content of the oxygen-containing gas used is high and at least 50% by volume, preferably at least 80% by volume. %, more preferably at least 90 vol .-% is. It may be particularly advantageous if the oxygen-containing gas is technically pure oxygen with an O 2 content of about 99% by volume. In addition, a driving is possible in which air is fed as oxygen-containing gas.
  • the hydrogen burned to generate heat is the hydrogen formed during the catalytic alkane dehydrogenation and, if appropriate, the hydrogen gas additionally added to the reaction gas mixture.
  • sufficient hydrogen should be present so that the molar ratio H 2 IO 2 in the reaction gas mixture immediately after the introduction of oxygen is 1 to 10, preferably 2 to 5, mol / mol. This applies to multi-stage reactors for each intermediate feed of oxygen-containing and possibly hydrogen-containing gas.
  • the hydrogen combustion takes place catalytically.
  • the dehydrogenation catalyst used generally also catalyzes the combustion of the hydrocarbons and of hydrogen with oxygen, so that in principle no special oxidation catalyst different from this one is required.
  • the reaction is carried out in the presence of one or more oxidation catalysts which selectively catalyze the combustion of hydrogen to oxygen in the presence of hydrocarbons.
  • the combustion of these hydrocarbons with oxygen to CO, CO 2 and water is therefore only to a minor extent.
  • the dehydrogenation catalyst and the oxidation catalyst are present in different reaction zones.
  • the oxidation catalyst may be present in only one, in several or in all reaction zones.
  • the catalyst which selectively catalyzes the oxidation of hydrogen is disposed at the sites where higher oxygen partial pressures prevail than at other locations of the reactor, particularly near the oxygen-containing gas feed point.
  • the feeding of oxygen-containing gas and / or hydrogen-containing gas can take place at one or more points of the reactor.
  • an intermediate feed of oxygen-containing gas and of hydrogen-containing gas takes place before each tray of a tray reactor.
  • the feed of oxygen-containing gas and of hydrogen-containing gas takes place before each horde except the first horde.
  • behind each feed point is a layer of a special oxidation catalyst, followed by a layer of the dehydrogenation catalyst.
  • the dehydrogenation temperature is generally 400 to 1100 0 C
  • the pressure in the last Ka talysatorbett the tray reactor is generally from 0.2 to 5 bar, preferably 1 to 3 bar.
  • a preferred catalyst which selectively catalyzes the combustion of hydrogen contains oxides and / or phosphates selected from the group consisting of the oxides and / or phosphates of germanium, tin, lead, arsenic, antimony or bismuth.
  • Another preferred catalyst which catalyzes the combustion of hydrogen contains a noble metal of VIII. And / or I. Maury.
  • the dehydrogenation catalysts used generally have a carrier and an active composition.
  • the carrier is usually made of a heat-resistant oxide or mixed oxide.
  • the dehydrogenation catalysts contain a metal oxide selected from the group consisting of zirconium dioxide, zinc oxide, alumina, silica, titania, magnesia, lanthana, ceria and mixtures thereof as a carrier.
  • the mixtures may be physical mixtures or chemical mixed phases such as magnesium or zinc-aluminum oxide mixed oxides.
  • Preferred supports are zirconia and / or silica, particularly preferred are mixtures of zirconia and silica.
  • the active composition of the dehydrogenation catalysts generally contains one or more elements of subgroup VIII, preferably platinum and / or palladium, more preferably platinum.
  • the dehydrogenation catalysts may comprise one or more elements of main group I and / or II, preferably potassium and / or cesium.
  • the dehydrogenation catalysts may contain one or more elements of III. Subgroup including the lanthanides and actinides, preferably lanthanum and / or cerium.
  • the dehydrogenation catalysts may contain one or more elements of III. and / or IV. Main group, preferably one or more elements from the group consisting of boron, gallium, silicon, germanium, tin and lead, particularly preferably tin.
  • the dehydrogenation catalyst contains at least one element of the VIII. Subgroup, at least one element of the I. and / or II. Main group, at least one element of III. and / or IV. main group and at least one element of III. Subgroup including the lanthanides and actinides.
  • all dehydrogenation catalysts can be used which are described in WO 99/46039, US Pat. No. 4,788,371, EP-A 705,136, WO 99/29420, US Pat. No. 5,220,091, US Pat. No. 5,430,220, US Pat. No. 5,877,369, EP 0 1 17 146, DE-A 199 37 106, DE-A 199 37 105 and DE-A 199 37 107 are disclosed.
  • Particularly preferred catalysts for the above-described variants of the autothermal propane dehydrogenation are the catalysts according to Examples 1, 2, 3 and 4 of DE-A 199 37 107.
  • the autothermal alkane dehydrogenation is preferably carried out in the presence of steam.
  • the added water vapor serves as a heat carrier and supports the gasification of organic deposits on the catalysts, whereby the coking of the catalysts counteracted and the service life of the catalysts is increased.
  • the organic deposits are converted into carbon monoxide, carbon dioxide and possibly water.
  • the dehydrogenation catalyst can be regenerated in a manner known per se.
  • steam can be added to the reaction gas mixture or, from time to time, an oxygen-containing gas can be passed over the catalyst bed at elevated temperature and the deposited carbon burned off. Dilution with water vapor shifts the equilibrium to the products of dehydration.
  • the catalyst is reduced after regeneration with a hydrogen-containing gas.
  • a gas mixture which generally has the following composition: 10 to 45 vol .-% propane, 5 to 40 vol .-% propene, 0 to 5 vol .-% Methane, ethane, ethene and C 4 + hydrocarbons, 0 to 5 vol .-% carbon dioxide, 0 to 20 vol .-% of water vapor and 0 to 25 vol .-% hydrogen and 0 to 5 vol .-% inert gases.
  • a gas mixture which generally has the following composition: 5 to 40% by volume of butane, 10 to 60% by volume of 1-butene and 2-butene, 0 up to 10% by volume of methane, ethane, ethene, propane, propene and C 5 + Hydrocarbons, 0 to 5% by volume of carbon dioxide, 0 to 20% by volume of steam and 0 to 25% by volume of hydrogen and 0 to 5% by volume of inert gases.
  • the product gas stream b when leaving the dehydrogenation zone is generally under a pressure of 1 to 5 bar, preferably 1, 5 to 3 bar, and has a temperature in the range of 400 to 700 0 C.
  • the product gas stream b can be separated into two partial streams, with a partial stream being returned to the autothermal dehydrogenation, in accordance with the cycle gas procedure described in DE-A 102 11 275 and DE-A 100 28 582.
  • the product gas stream b is compressed.
  • the compression of the product gas stream b is carried out at pressures of 5 to 150 bar, preferably 15 to 100 bar, more preferably 20 to 60 bar.
  • the compression can take place in several stages with intermediate cooling, for example in three or four stages, preferably it takes place in several stages, for example in three stages.
  • the cooling can be done in several stages, preferably it is multi-stage.
  • the cooling agents are air in air coolers, cold or cold water and refrigerants such as ethene, propene and propane, which are cooled to pressures in the range of -40 0 C to -100 0 C by compressing to pressures up to 20 bar and then relaxing , for use.
  • the product gas stream b is separated into an aqueous phase d, a hydrocarbon phase c2 comprising the alkene (III) and the unreacted alkane (II) and a gas phase c3 comprising hydrogen and low-boiling components.
  • the separation step within the process part C) generally takes place when the product gas stream b contains water vapor. But it can also be done only a water separation (see below).
  • water can first be separated off.
  • the separation of water can be carried out by condensation by cooling and optionally compressing the product gas stream b and can be carried out in one or more cooling and optionally compression stages.
  • the product gas stream b is cooled to a temperature in this range of 30 to 80 0 C, preferably 40 to 65 ° C.
  • the condensation can be at intermediate compression prior to compression and / or in the compression stages.
  • the product gas stream b is passed through a cascade of heat exchangers and thus initially cooled to a temperature in the range from 50 to 200 ° C. and then in a stream
  • a water vapor-depleted product gas stream b is obtained.
  • drying by means of molecular sieve can be provided.
  • the autothermal alkane dehydrogenation is carried out with the introduction of pure oxygen or oxygen-enriched air as the oxygen-containing gas
  • the work-up of the product gas stream b and the recovery of the alkane and alkene-containing mixture c2 can also be carried out as described below.
  • the product gas stream b is cooled and a liquid hydrocarbon stream c2 containing propane and propene or n-butane and butenes can be separated by condensation, leaving a residual gas stream c3 containing hydrogen and low boilers.
  • the hydrocarbon stream c2 may additionally contain methane, ethane, ethene and C 4 + hydrocarbons; in general it contains at least small amounts of ethane and ethene.
  • the temperature and the pressure in the separation step within the compression stage C can also be selected so that a large part of the alkanes and alkenes contained in the product gas stream b are contained in the gas stream c3.
  • This gas stream c3 can either be fed into the esterification zone D or alternatively into an absorption stage (as described below), just like the hydrocarbon stream c2.
  • n-butane dehydrogenation the conditions are chosen so that predominantly n-butane and butenes condense out.
  • the residual gas stream c3 in addition to hydrogen as low boilers generally still methane and carbon monoxide.
  • it may also contain C 3 - and C 4 hydrocarbons.
  • the product gas stream b is generally compressed to a pressure in the range of 5 to 60 bar and cooled to a temperature in the range of -10 to -60 0 C.
  • an aqueous phase d can condense on cooling and compression and separated from the C3 hydrocarbon phase c2 by phase separation in a phase separator if complete removal of water from the product gas stream b has not occurred before the condensation step.
  • all accumulating condensate streams can be fed to the phase separator.
  • a prior removal of water from the product gas stream b before condensing the C 3 hydrocarbon phase c2 can also be dispensed with. Then, water condenses out as the aqueous phase d together with the alkane and alkene-containing hydrocarbon phase c2. Aqueous phase and hydrocarbon phase are then subsequently separated in a phase separator.
  • cooling of the product gas stream is accomplished by heat exchange with a coolant.
  • the cooling can be done in several stages using multiple cooling circuits.
  • the cooling can be carried out in a plurality of stages in a column, wherein the rising in the column gas is removed, cooled, (partially) condensed and returned to the column. At the bottom of the column, the condensate, at the top of the column, the non-condensed gas is removed, which is not condensed in the uppermost cooling circuit.
  • the alkane dehydrogenation is carried out as an autothermal dehydrogenation with simultaneous combustion of the hydrogen formed, the result is a low hydrogen content of the product gas stream b.
  • the separation step C if this is carried out, the C 3 or C 4 hydrocarbons can be condensed out predominantly and only a very small part of the C 3 or C 4 hydrocarbons is hydrogenated / Low boilers containing exhaust stream c3 discharged.
  • the carbon dioxide gas scrubber Before carrying out the hydrocarbon condensation, it is possible to separate off carbon dioxide from the product gas stream by washing with gas, a product gas stream b depleted in carbon dioxide being obtained.
  • the carbon dioxide gas scrubber may be preceded by a separate combustion stage in which carbon monoxide is selectively oxidized to carbon dioxide.
  • carbon monoxide is selectively oxidized to carbon dioxide.
  • For CO 2 separation generally sodium hydroxide solution, potassium hydroxide solution or an alkanolamine solution is used as the scrubbing liquid; preference is given to using an activated N-methyldiethanolamine solution.
  • the product gas stream c is compressed to a pressure in the range of 5 to 25 bar by single or multi-stage compression.
  • a carbon dioxide-depleted product gas stream b having a CO 2 content of generally ⁇ 100 ppm or even ⁇ 10 ppm can be obtained.
  • the liquid hydrocarbon condensate stream c2 obtained in the cooling and condensation step C) generally contains 20 to 60 mol% of alkane (II), 20 to 60 mol% of alkene (III), 0 to 20 mol% of low boilers and 0 to 5 mol% high boilers.
  • the liquid hydrocarbon condensate stream c2 obtained in the cooling and condensation step C) may contain 20 to 70 mol% propane, 20 to 60 mol% propene, 0 to 10 mol% methane, 0 to 10 mol% % Ethane and ethene and 0 to 5 mol% of C 4 + hydrocarbons.
  • the work-up of the product gas stream b and the recovery of the alkane and alkene-containing mixture c2 can also be carried out as described below.
  • steam is separated off by condensing by cooling the product gas stream b and, if appropriate, compressing it to obtain a water gas-depleted product gas stream b.
  • alkane and alkene are separated from non-condensable or low-boiling gas constituents by contacting the product gas stream b with an inert absorbent and then desorbing the alkane and alkene dissolved in the inert absorbent to obtain a gaseous C 3 or C 4 hydrocarbon stream, and the Exhaust stream c3 containing hydrogen and low boilers (in the case of the propane dehydrogenation methane, ethane, ethene, nitrogen, carbon monoxide, carbon dioxide, optionally oxygen and optionally inert gases, in the case of n-butane dehydrogenation additionally propane and propene) is separated ,
  • the described work-up of the product gas stream b can also be carried out correspondingly in the case of the autothermal alkane dehydrogenation with the introduction of pure oxygen or oxygen-enriched air as the oxygen-containing gas.
  • the gas stream b is brought into contact with an inert absorbent, the C 3 or C 4 hydrocarbons and also small amounts of C 2 -hydrocarbons are absorbed in the inert absorbent and a laden with C 3 - or C 4 -Koh hydrocarbons absorbent and the other gas components containing exhaust gas c3 are obtained.
  • These are mainly carbon oxides, hydrogen, inert gases and C 2 hydrocarbons and methane.
  • Certain amounts of propane and propene or C 4 - hydrocarbons may still be contained in the stream c3, since the separation is generally not completely complete.
  • the C 3 - or C 4 -hydrocarbons are released from the absorbent again.
  • Inert absorbent used in the absorption stage are generally high-boiling nonpolar solvent in which the separated C 3 - or C 4 hydrocarbon mixture has a significantly higher solubility than the other gas components to be separated.
  • the absorption can be carried out by simply passing the stream c through the absorbent. But it can also be done in columns. It can be used in cocurrent, countercurrent or cross flow. Suitable absorption columns include plate columns having bubble, VEN repayment and / or sieve trays, columns with structured packings, for example fabric packings or sheet-metal packings with a specific surface area of 100 to 1000 m 2 / m 3 as Mellapak ® 250 Y, and packed columns such as with balls, rings or saddles made of metal, plastic or ceramic as packing.
  • trickle and spray towers, graphite block absorbers, surface absorbers such as thick-film and thin-layer absorbers and bubble columns with and without internals are also possible.
  • the absorption column has an absorption part and a rectification part.
  • heat can then be introduced into the column bottom.
  • a stripping gas stream can be fed into the bottom of the column, for example from nitrogen, air, steam or propane / propene mixtures.
  • the laden absorbent is brought into contact with the stripping gas stream.
  • C 2 " hydrocarbons are stripped out of the loaded absorption medium.
  • a portion of the top product can be condensed and returned to the top of the column in order to limit solvent losses.
  • Suitable absorbents are comparatively nonpolar organic solvents, for example C 4 -C alkenes, naphtha or aromatic hydrocarbons. Substances such as the middle oil fractions from the Paraffederillation, or ethers with bulky groups, or mixtures of these solvents, which may be added to a polar solvent such as 1, 2-dimethyl phthalate.
  • Suitable absorbents are furthermore esters of benzoic acid and phthalic acid with straight-chain C 1 -C -alkanols, such as n-butyl benzoate, methyl benzoate, ethyl benzoate, dimethyl phthalate, diethyl phthalate, and so-called heat transfer oils, such as biphenyl and diphenyl ether, their chlorinated derivatives and triaryl alkenes.
  • a suitable absorbent is a mixture of biphenyl and diphenyl ether, preferably in the azeotropic composition, for example, the commercially available Diphyl ®. Frequently, this solvent mixture contains dimethyl phthalate in an amount of 0.1 to 25 wt .-%.
  • Suitable absorbers are also butanes, pentanes, hexanes, heptanes, octanes, nonanes, decanes, undecanes, dodecanes, tridecanes, tetradecanes, pentadecanes, hexadecanes, heptadecanes and octadecanes or fractions obtained from refinery streams, which contain the abovementioned linear alkanes as main components contain.
  • Preferred absorbents are Cs-C-io-hydrocarbons, particularly preferred are Cg-hydrocarbons, especially nonanes.
  • the loaded absorbent is heated and / or relaxed to a lower pressure.
  • the desorption may also be by stripping, usually with steam, or in a combination of relaxation, heating and stripping in one or more process steps.
  • the desorption can be carried out in two stages, wherein the second desorption stage is carried out at a lower pressure than the first desorption stage and the desorption gas of the second stage is returned to the absorption stage.
  • the absorbent regenerated in the desorption stage is returned to the absorption stage.
  • a portion of this absorbent stream which may contain C 4 + hydrocarbons, is discharged, recovered and recycled, or discarded.
  • the desorption step is carried out by relaxation and / or heating of the loaded desorbent.
  • the desorption step is additionally stripped with steam.
  • the separation is generally not completely complete, so that in the C 3 - or C 4 - hydrocarbon stream - depending on the type of separation - still small amounts or even traces of other gas constituents, in particular the low-boiling hydrocarbons may be present.
  • the desorbed C3 or C 4 -Koh hydrocarbon stream can be cooled, in which case it can be additionally compressed in one or more further compression stages.
  • the liquid condensate stream c2 is obtained from C 3 or C 4 hydrocarbons.
  • the stream c2 may still contain small amounts of C 2 hydrocarbons.
  • the aqueous condensate stream generally accumulates when it is stripped with water vapor to desorb the dissolved gases.
  • the compression can again be done in one or more stages.
  • a total pressure of from 1 to 29 bar, preferably from 1 to 10 bar is compressed to a pressure in the range from 12 to 30 bar.
  • a cooling in the gas stream to a temperature in the range of 15 to 8O 0 C, preferably 15 to 6O 0 C
  • the komp- rimATOR gas mixture is cooled to a temperature of -10 0 C to 6O 0 C, preferably from -10 0 C to 30 0 C.
  • An optional aqueous condensate stream can be separated from the liquid C 3 or C 4 hydrocarbon stream in a phase separation apparatus.
  • the absorbent used in the absorption column is the same organic acid which is reacted in the esterification zone with the corresponding alkene. In this case, it is then possible to dispense with the desorption step described above.
  • the absorbent laden with C 3 or C 4 hydrocarbons ie in this case the organic acid (for example acetic acid or formic acid), can then, if appropriate after further heating and / or compression, be conducted directly into the esterification zone.
  • the organic acid separated from the separation step (G) after the ester cleavage zone can be used as the absorbent feed into the absorption column. This is then not recycled directly into the esterification zone, but in the absorption stage and from there, loaded with the C 3 - or C 4 hydrocarbons, in the esterification zone.
  • the separation step within process part C) may or may not be performed. As described above, however, at least one compression of the product gas stream b is always carried out.
  • the resulting dehydrogenation gas mixture b which in the case of propane dehydrogenation consists essentially of propane, propene, hydrogen and Low boilers and in the case of n-butane dehydrogenation consists essentially of butane, 1-butene, 2-butene, hydrogen and low boilers, fed directly and without prior separation of C 3 - or C 4 hydrocarbons in the esterification zone and with the organic Acid can be brought into contact.
  • the product gas stream b contains steam, since the dehydrogenation was carried out with the introduction of oxygen and / or with the introduction of steam, separation of water vapor alone-for example by condensation as described above under C) -can be sufficient and the remaining, the C 3 - or C 4 hydrocarbons, hydrogen and low boilers containing mixture are reacted gaseous or liquid with the organic acid.
  • the presence of carbon oxides and other inert gases (atmospheric nitrogen) also does not interfere with the ester formation reaction.
  • the residual gas stream c3 can, preferably it is predominantly recycled to the dehydrogenation stage A).
  • a partial flow is separated and discharged from the process, in order to avoid an accumulation of secondary components.
  • This partial stream can be incinerated or fed to a process stage for the recovery of alkane / alkene contained therein.
  • the recovery can be carried out as adsorption or adsorption, as membrane separation or rectification.
  • a partial stream of the residual gas stream c3 can also be fed to the esterification step D).
  • the hydrocarbon phase c2 can be fed directly or after further pressure increase of the ester formation stage D).
  • the aqueous condensate stream d can be discharged from the process or fed into the ester cleavage stage (process part F)).
  • a process stage D the product gas stream b or, if the separation step in the process step C) is carried out, the phase c2 containing the alkane (II) and the alkene (III) and optionally the residual gas stream c3 or, if the absorption step as described above is carried out with the organic acid, the organic acid loaded with C 3 or C 4 hydrocarbons in an esterification zone is reacted with an organic acid (IV), a product mixture d containing the corresponding alkyl ester (V) (isopropyl ester or Butyl ester) of the organic acid (IV) and unreacted alkane (II).
  • the esterification can be carried out in liquid phase or two-phase (with respect to the reactants) as a gas / liquid reaction.
  • the esterification is generally carried out at a pressure of 10 to 100 bar and at a temperature of 50 to 250 0 C.
  • the organic acid based on alkene, in amounts of 0.5 to 50 moles per mole of alkene, preferably in stoichiometric excess, preferably in amounts of 1, 1 to 6, particularly preferably 1, 2 to 2.5 moles per mole of alkene used.
  • 50 to 90% of the alkene reacts to the corresponding alkyl ester.
  • unsaturated secondary components contained in the product gas stream b or in stream c2 can react to form alkyl esters. These can readily be removed by distillation in the subsequent work-up steps.
  • this stream may still contain small amounts of the ester formed, as well as water and the low-boiling components already contained in stream c2.
  • esterification step D may comprise two zones, a reaction zone and a backwash zone.
  • the aim of the reaction zone is the highest possible conversion of the acid with the alkene.
  • the gas stream d2 is largely freed from acid residues by the addition of water. As a result, corrosion problems in other parts of the system can be avoided when recirculating the flow d2).
  • the ester formation can be carried out in fixed-bed reactors in trickle-bed mode or in fluidized bed mode.
  • the organic acid and the alkene-containing stream can be conducted in cocurrent or countercurrent, but also in crossflow.
  • a cascading of several catalyst beds takes place, optionally with intermediate introduction of the acid and / or the alkene-containing stream.
  • a cascade of 2-5, more preferably 2-3 different reaction zones is used.
  • the released heat of reaction can be dissipated by internal heat exchanger surfaces.
  • a driving with an external circulation in which a heat exchanger is mounted to dissipate the heat of reaction is possible. By adjusting the circulation rate, the axial temperature profile in the reactor can be set almost arbitrarily.
  • esterification can also be carried out in bubble columns or jet loop reactors.
  • the tubular reactor in fluidized bed mode, which is operated selectively at the loosening point of the catalyst (so-called floating bed driving), with external heat exchanger and 2-fold cascading.
  • this zone corresponds to an absorption column, which can be designed separately as a separate apparatus or integrated into the fixed bed reactor. As internals soils, packs or random beds can be used.
  • the amount of washing water is chosen so that the acid radical in the host rome d2 ⁇ 50 ppm, preferably ⁇ 10 ppm particularly preferably ⁇ 1 ppm.
  • a typical ratio of wash water volume: the amount of acid to be removed to reach ⁇ 1 ppm acid is in the range of 1 kg / kg to 20 kg / kg.
  • the washed with acid residues wash water can be used in the process step F) for ester cleavage.
  • Suitable heterogeneous ester formation catalysts are acidic ion exchange resins, in particular those of sulfonated, crosslinked with divinylbenzene polystyrene. These catalysts can have different pore structures - one distinguishes microporous, gel-like and macroporous catalysts. In addition, may be bound to the aromatic rings of polystyrene electronegative radicals, such as chlorine or fluorine.
  • Suitable catalysts of this type are, for example, Amberlyst 15, 16, 36, 39, 40, 46, 48, 70, 119, 139 Lewatit K1 131, K1221, K1461, K2420, K2629, K2649, Purolite CT169, CT175, CT275, Diaion RCP145H.
  • catalysts which contain a high content of acidic groups such as. Amberlyst 35, 36, 40, 49, 19, Lewatit K2649, Purolite CT275.
  • the catalysts mentioned here are usually available in dry or hydrous form. Both forms are suitable, in the hydrous catalysts, the water is displaced by washing with the organic acid.
  • a group of catalysts related to the acidic ion exchange resins is derived from sulfonated polycondensed aromatics or graphitic carbon. Such materials are z. Example by sulfonation of polycyclic aromatics, such as naphthalene or anthracene, under conditions that lead to a condensation of aromatics formed.
  • a similar process is based on the charring of organic material, eg sugars, under anaerobic conditions. The corresponding residues are then sulfonated.
  • Another group of organic heterogeneous catalysts is derived from ionic liquids which are adsorbed on suitable support materials.
  • inorganic catalysts such as acidic metal oxide catalysts or acidic zeolites, are also suitable.
  • the acidic metal oxide catalysts are in particular sulfated Zirconium oxide and zirconium tungstate or titanium tungstate systems.
  • this group of catalysts include the water-insoluble acidic salts of the heteropolyacids, for example tungstophosphate or molybdophosphoric acid or tungstosilicic acid.
  • Such insoluble salts form from these acids with metal cations with large ionic radii, such as K + , Rb + , Cs + , NH 4 + or Ag + .
  • these salts usually 10 to 90 mol%, in particular 40 to 85 mol%, of the acidic centers are exchanged for cations.
  • Another group of catalysts is derived from heteropolyacids or their salts, which are adsorbed on an inert support material such as silica gel, alumina or activated carbon.
  • Suitable zeolite catalysts include those of the structural type beta zeolites, faujasites, mordenites, and ZSM-5 zeolites.
  • the ratio of the Si / Al atoms (the modulus) in the zeolite framework is crucial for the catalytic activity of the zeolite.
  • Suitable zeolites which have a modulus between 2 and 500, in particular between 3 and 200 and very particularly preferably between 5 and 100, are suitable for the process according to the invention.
  • the inorganic catalysts described here are usually thermally activated, ie the materials are heated at temperatures between 50 and 900 0 C, preferably calcined between 90 and 500 0 C.
  • the esterification reaction can also be carried out homogeneously catalyzed.
  • Mineral acids in particular sulfuric acid, sulfonic acids or the free heteropolyacids and their acidic soluble salts or acidic ionic liquids are suitable for this purpose.
  • the ester formation preferably takes place in the presence of heterogeneous catalysts.
  • Preferred heterogeneous catalysts are acidic ion exchange resins, the acidic K, Cs, or ammonium salts of the heteropolyacids, and also beta zeolites and faujasites. Particularly preferred are ion exchange resins.
  • Preferred organic acids (IV) are formic acid and acetic acid, which react with propene to form isopropyl or isopropyl acetate and butenes (1-butene and 2-butene) to 2-butyl or 2-butyl acetate.
  • the product mixture d of the ester formation is expanded, whereby a gas stream e1 containing propane or n-butane is separated off and a product mixture e2 containing the alkyl ester (V) is obtained.
  • the gas stream e1 can with the gas stream d2 of the backwashing zone in the process part D united and the dehydrogenation zone B are fed together.
  • the product mixture d is depressurized from a pressure in the range of 20 to 60 bar to a pressure in the range of 2 to 10 bar.
  • propane or n-butane optionally low boilers such as ethane, ethene, methane, carbon oxides and inert gases contained in the product mixture are d separated with.
  • the gas stream e1 containing the alkane (II) is preferably recycled to the alkane dehydrogenation.
  • the gas stream e1 containing the alkane (II) may still contain low-boiling components such as ethane, ethene, methane, carbon oxides and inert gases as well as hydrogen. In general, it contains low boilers and optionally hydrogen, if the above-described (optional) separation of low boilers and hydrogen (residual gas stream c3) has not already occurred in step C). From the gas stream e1, a partial stream can be separated and discharged from the process to avoid accumulation of secondary components. This partial stream can be incinerated or fed to a process stage for the recovery of alkane / alkene contained therein. The recovery can be carried out as adsorption or adsorption, as membrane separation or rectification.
  • Hydrogen contained in the stream can be recovered, for example, by pressure swing adsorption. Both the recovered alkane / alkene and the recovered hydrogen can be recycled to the dehydrogenation. It is also possible to remove the entire stream e1 from the process or to feed it to the process stage for the recovery of alkane / alkene or hydrogen contained therein.
  • distillation or rectification can additionally be carried out.
  • a reflux ratio between 0.2 and 1.5
  • tray columns e.g. Bell bottom columns as well as packed columns or packed columns can be used.
  • the product mixture d can also be expanded from a pressure in the range of 20 to 60 bar to a pressure of generally 2 to 45 bar, for example 10 to 40 bar, or, in a special variant, 25 to 32 bar.
  • the post-expansion streams e1 and e2 may be worked up as described below using two columns K1 and K2. Several variants of this work-up can be carried out. The variants
  • the gas stream e1 present after the expansion is fed to a first column K1.
  • the pressure in this column is equal to or only slightly lower than the pressure of the feed stream.
  • the column K1 is preferably the pure enrichment part of a conventional column, i. this has no evaporator, but a condenser, and the feed is preferably at the bottom of the column.
  • the liquid stream e2 present after the expansion is fed to the second column K2.
  • the pressure in K1 is 10 to 40 bar, preferably 25 to 32 bar, and in K2 1 to 5 bar, preferably 1, 3 to 2 bar.
  • the bottom draw of the column K1 is also fed to the column K2.
  • the two top takeoff streams of the columns K1 and K2 correspond to the stream e1 and can be used as described above or further worked up.
  • the top outlet of the second column K2 is preferably recycled directly into the dehydrogenation, optionally using a compressor.
  • the pressure in K1 is 10 to 40 bar, preferably 25 to 32 bar, and in K2 2.5 to 7 bar, preferably 4 to 6 bar.
  • This variant has the advantage that no compressor is needed.
  • the condenser temperature of the column K2 is higher (for example, about 40 0 C). In general, this temperature is 30 to 50 0 C, preferably 37 to 45 ° C.
  • this stream is compressed to the pressure of the column K1 and the Column K1 supplied.
  • the pressure in K1 is 10 to 40 bar, preferably 25 to 32 bar, and in K2 1 to 5 bar, preferably 1, 3 to 2 bar.
  • the procedure is as described under variant Ib, however, larger amounts of C3 components (propane and propene) in the bottom discharge of the column K2 (for example about 1- 2% by mass) are allowed, so that the sump temperature is reduced to approx 100-110 0 C is limited and increased material requirements on the evaporator and bottom part of the column K2 do not result.
  • the pressure in the column K1 is generally 10 to 40 bar, preferably 25 to 35 bar
  • the pressure in the column K2 is generally 2.5 to 7 bar, preferably 4 to 6 bar.
  • the recycle stream d2 can be treated, if it contains even larger amounts of ester.
  • a distillative purification of the stream is preferred. This can be done in a separate column, or alternatively together with the treatment of the stream d in a common column. It may be followed by a fine cleaning by adsorption, absorption, gas scrubbing or catalytic purification stages.
  • the alkyl ester (V) containing product mixture e2 (consisting essentially of the alkyl ester (V) and the organic acid (IV), for example from isopropyl acetate and acetic acid) can be further worked up before the stream e2 in the process step F) out becomes.
  • the organic acid (IV), for example acetic acid can be separated off from the alkyl ester (V), for example isopropyl acetate, in a distillation column.
  • the organic acid (IV), for example acetic acid can be recovered as a bottom product and recycled to the esterification stage.
  • the alkyl ester (V), for example isopropyl acetate can be obtained as top product and is further passed into process stage F).
  • Examples of suitable process parameters in the case of the separation of acetic acid and isopropyl acetate are a pressure of up to 2 bar (all pressures absolute) and a reflux ratio of 0 to 3.
  • the reaction equilibrium in the hydrolysis reactor of process step G) positively influenced.
  • the investment and energy requirements of, for example, the apparatus (2), (3), (4) and (5) described below are reduced.
  • addition stage D) and hydrolysis stage F) are additionally decoupled.
  • the product mixture e2 containing the alkyl ester (V) is reacted with water in an ester cleavage zone to give a product mixture f containing the alkanol (I) (isopropanol or 2-butanol) and the organic acid (IV).
  • the alkanol (I) is separated off from the product mixture f and the organic acid (IV) is recovered.
  • the organic acid (IV) is generally recycled to the esterification step D), or it is fed as an absorbent into the absorption step for the separation of alkene and alkane from the gas phase c3.
  • the product mixture f can be separated into a stream gl containing the acid and a stream g2 containing the alkanol (I) and the alkyl ester (V). From the alkanol stream g2, the alkyl ester can be separated by distillation and recycled to the ester cleavage zone.
  • the separation of the alkanol (I) from the product mixture f depending on the phase equilibria by simple distillation or by azeotropic rectification using entrainers (eg benzene, cyclohexane or diisopropyl ether in the case of isopropanol), by extractive distillation using a Extracting agent (eg of ionic liquids or acetic acid) and by membrane processes (pervaporation or vapor permeation) take place.
  • entrainers eg benzene, cyclohexane or diisopropyl ether in the case of isopropanol
  • Extracting agent eg of ionic liquids or acetic acid
  • membrane processes pervaporation or vapor permeation
  • the ester cleavage can be carried out both homogeneously or heterogeneously catalyzed.
  • the catalysts described above, which are also used in the ester formation reaction are suitable.
  • Preferred heterogeneous catalysts are ion exchange resins.
  • Preferred homogeneous catalysts are sulfuric acid or heteropolyacids.
  • the ester cleavage zone can be configured as a reactor or as a reactive distillation column. A combination of reactor and reactive distillation column is possible.
  • the ester cleavage is carried out with a substoichiometric amount of water.
  • the organic acid (IV) is recovered substantially in concentrated form and need not be further concentrated before being recycled to the esterification step D).
  • Ester cleavage and distillation can be carried out in separate process steps.
  • the mixture e2 containing the alkyl ester (V) is mixed with water in an ester cleavage reactor to give a product mixture f containing the alkanol (I), the organic acid (IV), the alkyl ester (V) and water, and this mixture then separated in at least two distillation columns connected in series, optionally in conjunction with a dividing wall column.
  • the ester cleavage reactor can be designed as a fixed bed, trickle bed, fluidized bed or slurry reactor.
  • the catalyst can be operated selectively at the loosening point (so-called floating bed mode).
  • the reactor can be designed, for example, as a stirred tank reactor or tubular reactor. If the ester cleavage is carried out homogeneously, the catalyst is generally separated with the organic acid in process step G), generally via the bottom draw stream of the first distillation column.
  • the catalyst prior to the recycling of the organic acid to the esterification step D), the catalyst can be separated off, and the separated catalyst can be recycled to the ester-separation reactor.
  • the separation of the catalyst can be carried out thermally in an evaporator or in a multistage distillation column and in a phase separator or else by a combination of phase separation and thermal separation. If the same homogeneous catalyst used for the ester cleavage as for the ester formation, it can be dispensed with a separate separation of the catalyst. Heteropolyacids or sulfuric acid are particularly suitable for this process variant.
  • acetic acid is used as the organic acid in a first specific embodiment of the process according to the invention, and an isopropyl acetate-containing product mixture e2 is obtained, where
  • step F) the isopropyl acetate-containing product mixture e2 is reacted in an ester cleavage reactor with water to give a product mixture f containing isopropanol, acetic acid, isopropyl acetate and water, and
  • step G the product mixture f in a first distillation column (1) into a stream gl consisting essentially of acetic acid and a stream g2 containing isopropyl acetate, isopropanol and water separated, and the stream g2 in a second distillation column (2) into a stream h1 containing isopropyl acetate and isopropanol and a stream h2 consisting essentially of isopropyl acetate, isopropanol and water is separated.
  • the stream gl can be recycled to the esterification zone. This variant is shown in FIG.
  • the separation of the product mixture f into a stream g consisting essentially of acetic acid and a stream g2 containing isopropyl acetate, isopropanol and water can be combined with an additional hydrolysis step.
  • the distillation column (1) can be equipped with a side reactor containing the hydrolysis catalyst, or the distillation column (1) can be carried out as a reactive distillation column by equipping at least one column section with trays or packs containing the hydrolysis catalyst, or After the distillation column (1), it is possible to pass a partial stream of the stream g2 via a reactor containing the hydrolysis catalyst and to recycle it into the distillation column (1).
  • the process parameters for the additional hydrolysis reactor are, for example, a pressure of 2 to 10 bar and a temperature of 60 to 130 ° C.
  • the additional hydrolysis reactor has the following advantages: since the reaction conversion in the hydrolysis reactor (process step F) is limited by the reaction equilibrium, the flow f is practically in the reaction equilibrium.
  • the acetic acid separation in distillation column (1) makes possible a further hydrolysis conversion which can be realized in the additional hydrolysis reactor.
  • the investment requirements and energy requirements of the apparatus (2), (3), (4) and (5) described below are reduced.
  • isopropanol can be separated as distillate in a further column (4) and thus a stream i1, which consists essentially of isopropyl acetate, can be obtained and returned to the ester cleavage reactor (10).
  • the water contained in stream g2 can be separated by azeotropic distillation comprising two distillation columns (hereinafter referred to as second (2) and third column (3)) and a phase separator (5) using an entrainer such as benzene or diisopropyl ether and recycled as stream h3 into the ester cracking reactor (10). If diisopropyl ether is obtained as a by-product in the ester cleavage, preference is given to using diisopropyl ether as entrainer in the removal of water.
  • the ester cleavage is generally carried out at a pressure of 1 to 10 bar, preferably 2 to 5 bar, and a temperature of 50 to 150 0 C, preferably 80 to 120 ° C, performed. Preference is given to working in the presence of acidic ion exchanger as a catalyst. Water can be used in stoichiometric excess or excess, generally in stoichiometric excess, based on isopropyl acetate, preferably in amounts of from 0.5 to 0.9 moles of water per mole mol of isopropyl acetate.
  • the ester cleavage reaction is preferably carried out in a fluidized bed reactor in a floating bed mode.
  • the resulting product mixture f contains, for example, 5 to 20% by weight of isopropanol, 10 to 50% by weight of acetic acid, 30 to 70% by weight of isopropyl acetate and 5 to 15% by weight of water.
  • the stream gl which is generally obtained as the bottom draw stream of the first column (1), contains, for example, 90 to 100% by weight of acetic acid and may additionally contain 0 to 5% by weight of isopropyl acetate and 0 to 5% by weight of isopropanol
  • the stream g2, which is generally obtained as the top draw stream of the first column (1) contains, for example, 40 to 70% by weight of isopropyl acetate, 10 to 80% by weight of isopropanol and 0 to 20% by weight of water.
  • the stream h1 which is generally obtained as the bottom draw stream of the second column (2), contains, for example, 60 to 90% by weight of isopropyl acetate and 5 to 30% by weight of isopropanol and 0 to 5% by weight of acetic acid
  • the stream h2 which is generally obtained as the top draw stream of the second column (2), contains, for example, 30 to 60% by weight of isopropyl acetate, 5 to 15% by weight of water and 10 to 30% by weight of isopropanol.
  • the stream h2 may still contain up to 50% by weight of the entrainer and also traces of acetic acid.
  • the stream h3 obtained as the bottom draw stream of the third column (3) preferably contains at least 95% by weight of water and the stream h4, which is generally obtained as the top draw stream of the third column (3), contains for example 5 to 15% by weight. % Isopropyl acetate, 15 to 50% by weight of water and 30 to 70% by weight of isopropanol. In addition, stream h4 may still contain up to 5% by weight of entrainer and traces of acetic acid.
  • the stream i1 which is generally obtained as a bottom draw stream of a further (fourth) column (4), preferably contains at least 90% by weight of isopropyl acetate, up to 5% by weight of acetic acid and up to 5% by weight of isopropanol
  • the stream i2 which is generally recovered as the top draw stream of the fourth column (4), preferably contains at least 98% by weight of isopropanol.
  • the stream i2 may still contain up to 2% by weight of isopropyl acetate and traces of entrainer and water.
  • the top hatch streams h4 and h2 are separated into an aqueous and an organic phase by adding an entraining agent (benzene).
  • the aqueous phase is added as reflux to the third column (3) and preferably contains at least 70% by weight of water, up to 10% by weight of isopropyl acetate, up to 20% by weight of isopropanol and traces of entrainer.
  • the organic phase is added as reflux to the second column (2) and preferably contains at least 40% by weight of isopropyl acetate, up to 10% by weight of water, 10 to 40% by weight of isopropanol and 10 to 50% by weight of entrainer and traces of acetic acid.
  • the first column (1) generally has 10 to 30 theoretical plates and is operated at a pressure of 0.5 to 2 bar.
  • the second column (2) generally has 25 to 50 theoretical plates and is operated at a pressure of 0.5 to 2 bar.
  • the third column (3) generally has 1 to 15 theoretical plates and is operated at a pressure of 0.5 to 2 bar.
  • the fourth column (4) generally has 40 to 70 theoretical plates and is operated at a pressure of 5 to 10 bar.
  • step F) the isopropyl formate-containing product mixture e2 in an ester-separation reactor (10) is reacted with water to give a product mixture f comprising isopropanol, formic acid, isopropyl formate and water, and
  • step G the product mixture f in a first distillation column (1) into a stream gl containing formic acid and water and a stream g2 containing isopropyl, water and isopropanol is separated, and the stream g2 in a second distillation column (2) in a Stream g4 containing isopropyl formate, water and isopropanol and a stream g3 consisting essentially of isopropanol is separated and the stream gl is returned to the esterification zone and stream g4 is recycled to the ester cracking reactor (10).
  • g3 is generally obtained as a bottom draw stream or gaseous side draw stream.
  • water can be separated from formic acid from stream 1 g and thus only the concentrated formic acid can be recycled back into the esterification zone, whereby higher space / time yields are achieved and the low-acidic water stream is fed back into the ester cleavage reactor.
  • This variant is shown in FIG.
  • the ester cleavage is generally carried out at a pressure of 1 to 10 bar, preferably 2 to 5 bar, and a temperature of 50 to 150 0 C, preferably 80 to 120 0 C, performed. Preference is given to working in the presence of acidic ion exchanger as a catalyst.
  • Water is generally employed in stoichiometric excess, based on isopropyl formate, preferably in amounts of from 0.5 to 0.9 mol of water per mole of isopropyl formate. If necessary, water can also be added in stoichiometric excess.
  • the product mixture f contains, for example, 10 to 25% by weight of isopropanol, 15 to 40% by weight of formic acid, 20 to 60% by weight of isopropyl formate and 5 to 20% by weight of water.
  • the stream gl which is generally obtained as the bottom draw stream of the first column (1), generally contains a mixture of, for example, 80 to 95% by weight of formic acid and 5 to 20% by weight of water.
  • the stream g4, which is generally obtained as the top draw stream of the second column (2) contains a mixture of, for example, 70 to 90% by weight of isopropyl formate, up to 15% by weight of water and 5 to 15% by weight of isopropanol.
  • the first column (1) generally has 30 to 60 theoretical plates and is operated at a pressure of 0.5 to 3 bar.
  • the second column (2) generally has 30 to 60 theoretical plates and is operated at a pressure of 0.5 to 1 bar.
  • step F) the isopropyl formate-containing product mixture e2 is reacted in an ester cleavage reactor (10) with water to give a product mixture f comprising isopropanol, formic acid, isopropyl formate and water, and
  • step G) the product mixture f this side of the partition in a dividing wall distillation column (1) led into a stream gl containing formic acid and water, a distillate stream g4 containing isopropyl, water and isopropanol and a side stream g3 beyond the partition substantially consisting of isopropanol separated and the bottom stream g1 is recycled to the esterification zone and the g4 stream is recycled to the ester decomposition reactor (10).
  • the composition of the obtained streams g1, g3 and g4 is preferably the same as in the case of the connection of two separate columns.
  • the dividing wall column (1) generally has 10 to 30 theoretical plates in the precolumn and 30 to 60 theoretical plates in the main column and is operated at a pressure of 0.5 to 3 bar. This variant is shown in FIG.
  • ester cleavage and distillation can also be carried out partly simultaneously in one and the same process step in a reactive distillation column, wherein the mixture formed by ester cleavage is simultaneously separated at least partially by distillation and the organic acid (optionally together with a homogeneous catalyst) is recovered.
  • the reactive distillation column may be preceded by a pre-reactor.
  • the reactive distillation can be carried out homogeneously or heterogeneously catalyzed.
  • the reactive distillation column may contain conventional internals (for example, packings, fillers, trays).
  • Heterogeneous catalysts can be present in the form of catalytic internals, for example as catalytic packings or random packings, or as a suspension.
  • a housing of the catalyst in external containers is possible, which are fed by a side stream of the reaction column.
  • the ester-containing mixture e2 is reacted with water in an ester cleavage pre-reactor to give a product mixture f containing the alkanol (I), the organic acid (IV), the alkyl ester (V) and water, and from the product mixture f in a reactive distillation column, a product mixture f2 from the alkanol (I) and the alkyl ester (V) and a stream f3 from the organic acid (IV), which is recycled to the esterification zone, recovered and the product mixture f2 in a downstream (further) distillation column in a stream gl from the alkanol (I) and the alkyl ester (V) and small amounts of water and a stream g2 from the alkanol (I) separated, wherein the current gl in the ester cleavage zone, preferably in the reactive distillation column, is recycled.
  • step F the isopropyl acetate-containing product mixture e2 is reacted in an ester cleavage prereactor with water to give a product mixture f containing isopropanol, acetic acid, isopropyl acetate and water, from the product mixture f in a Reactive distillation column (1) a product mixture f2 essentially consisting of isopropanol, isopropyl acetate and small amounts of water and a stream f3 essentially consisting of acetic acid recovered and the stream f3 returned to the esterification zone, and
  • step G the product mixture f2 in a distillation column (2) is separated into a stream gl consisting essentially of isopropanol, isopropyl acetate and small amounts of water and a stream g2 consisting essentially of isopropanol.
  • the stream gl can be recycled to the ester cleavage reactor (10) or preferably to the reactive distillation column (1).
  • This variant is shown in FIG.
  • the product mixture f which contains, for example, 5 to 20% by weight of isopropanol, 10 to 90% by weight of acetic acid, 10 to 80% by weight of isopropyl acetate and 1 to 20% by weight of water, is dissolved in the reactive distillation column (1).
  • a stream f2 which contains, for example, 30 to 80 wt .-% of isopropanol and 10 to 70 wt .-% isopropyl acetate and up to 10 wt .-% water and is generally obtained as a top draw stream
  • a stream f3 which preferably consists of at least 90 wt .-% of acetic acid, besides can still contain isopropyl acetate and isopropanol and is generally obtained as a bottom draw stream, separated.
  • the catalytically active internals can be arranged both below and above the feed inlet. In addition, at the top and in the bottom of the column further separation-effective internals can be located.
  • the feed feed is preferably located below the reaction zone.
  • the reaction column generally has 30 to 80 theoretical plates and is operated at a pressure of 1 to 6 bar.
  • water can be added above and below the reaction zone as well as in the reaction zone itself.
  • the stream gl which is generally obtained as the top draw stream of the distillation column (2), contains, for example, 10 to 50% by weight of isopropanol and 30 to 90% by weight of isopropyl acetate.
  • the distillation column (2) generally has 10 to 50 theoretical plates and is operated at a pressure of 0.01 to 1 bar.
  • the ester cleavage reactor (10) upstream of the reactive distillation column (1) can also be omitted.
  • the advantage of the pre-reactor is in the heterogeneously catalyzed reactive distillation that here the amount of catalyst in the column can be reduced and also the pre-reactor acts as a guard bed before catalyst poisons.
  • step F) the isopropyl acetate-containing product mixture e2 is fed with water directly into a reactive distillation column (1) and from the product mixture e2 in the reactive distillation column (1) a product mixture f2 essentially consisting of isopropanol , Isopropyl acetate and small amounts of water and a stream f3 im Essentially consisting of acetic acid recovered and the stream f3 returned to the esterification zone, and
  • step G the product mixture f2 in a distillation column (2) in a stream gl consisting essentially of isopropanol, isopropyl acetate and small amounts of water and a stream g2 essentially consisting of isopropanol separated and the stream gl returned to the reactive distillation column (1).
  • a stream gl consisting essentially of isopropanol, isopropyl acetate and small amounts of water
  • a stream g2 essentially consisting of isopropanol separated and the stream gl returned to the reactive distillation column (1).
  • composition of the streams f2, f3, gl and g2 essentially corresponds to the procedure with upstream ester cleavage reactor.
  • a homogeneous catalyst can also be used.
  • the isopropyl acetate-containing product mixture e2 is fed with water into a reactive distillation column (1) and from the product mixture e2 in the reactive distillation column (1) a product mixture f2 essentially consisting of isopropanol, small amounts Water and isopropyl acetate and a stream f3 consisting essentially of acetic acid and the homogeneous catalyst recovered and returned to the esterification zone, and
  • step G the product mixture f2 in a distillation column (2) into a stream gl essentially consisting of isopropyl acetate, water and isopropanol and a stream g2 essentially consisting of isopropanol separated and the stream gl returned to the reactive distillation column.
  • a stream gl essentially consisting of isopropyl acetate, water and isopropanol
  • a stream g2 essentially consisting of isopropanol separated and the stream gl returned to the reactive distillation column.
  • the separation of the catalyst from stream f3 can be carried out thermally in an evaporator or a multistage distillation column and in a phase separator or the combination of phase separator and thermal separation.
  • the concentrated catalyst stream is returned to the enrichment section of the reaction column and the concentrated acetic acid back to the esterification step.
  • the composition of the streams f2, gl and g2 corresponds essentially to the procedure with heterogeneously catalyzed reactive distillation.
  • the stream f3, preferably consisting of acetic acid, can contain up to 50% by weight of catalyst.
  • the reaction column (1) are conventional internals for distillation.
  • column trays eg, chimney trays, allow the setting of defined residence times.
  • defined residence times can also be achieved in an external container connected to the column by withdrawing a stream from the column by means of a side draw, passing it through the container and returning it to the column.
  • the reaction column generally has 30 to 100 theoretical plates and is operated at a pressure of 1 to 6 bar.
  • water can be added to the reaction zone.
  • step F it is preferred to use water in step F) in relation to isopropyl acetate-stoichiometric excess.
  • aqueous-organic azeotropes such as isopropanol-water, isopropanol-isopropyl acetate-water
  • ester cleavage can be carried out as a reactive distillation (with and without a pre-reactor).
  • step F the isopropyl formate-containing product mixture e2 is reacted in an ester cleavage reactor (10) with water to give a product mixture f containing isopropanol, formic acid, isopropyl formate and water and from the product mixture f in a reactive distillation column (1 ) a product mixture f2 essentially consisting of isopropanol and isopropyl formate and a stream f3 consisting essentially of formic acid and water are recovered and the stream f3 is returned to the esterification zone, and
  • step G the product mixture f2 in a distillation column (2) is separated into a stream consisting essentially of isopropanol and isopropyl formate and a stream g2 consisting essentially of isopropanol and the stream g1 in the ester cleavage reactor (10) or preferably in the reactive distillation column ( 1) returned.
  • This variant is also shown in FIG.
  • the product mixture f which contains, for example, 5 to 15% by weight of isopropanol, 5 to 50% by weight of formic acid, 20 to 70% by weight of isopropyl formate and 1 to 15% by weight of water, is dissolved in the reactive distillation column (1). reacted further and at the same time into a stream f2 which contains, for example, 10 to 50% by weight of isopropanol and 30 to 80% by weight of isopropyl formate and is generally obtained as a top draw stream.
  • NEN NEN
  • stream f3 which contains, for example, 60 to 85 wt .-% of formic acid and 5 to 20 wt .-% water, next can also contain isopropyl formate and is generally recovered as a bottom draw stream, separated.
  • the stream gl which is generally obtained as the top draw stream of the distillation column (2), generally contains an azeotrope-boiling mixture of, for example, from 5 to 30% by weight of isopropanol and from 60 to 95% by weight of isopropyl formate.
  • the distillation column (2) generally has 10 to 30 theoretical plates and is operated at a pressure of 0.1 to 1 bar.
  • step F it is likewise preferred to use water in step F) in relation to the isopropyl formate stoichiometric deficit.
  • aqueous-organic azeotropes such as isopropanol-water, isopropanol-Isopropylformiat-water
  • the catalytically active internals can be arranged both below and above the feed inlet. In addition, at the top and in the bottom of the column further separation-effective internals can be located.
  • the feed feed is preferably located below the reaction zone.
  • the reaction column generally has 30 to 80 theoretical plates and is operated at a pressure of 1 to 6 bar.
  • the ester cleavage reactor (10) upstream of the reactive distillation column (1) can likewise be dispensed with.
  • the downstream column (2) and the reaction column (1) can also be combined to form a reactive dividing wall column.
  • Example 1 By means of the commercial simulation program AspenPlus, an integrated overall process for the production of isopropanol from propane was calculated by way of example.
  • the method comprises steps A to N. These are shown in FIGS. 7 and 8 with the associated material flows.
  • the calculated compositions of the individual streams can be taken from the table.
  • the Simulation calculation is based on a state equation model based on measurements of vapor-liquid and liquid-liquid phase equilibria.
  • the propane dehydrogenation reactor A is supplied with a propane-rich, gaseous stream (4). This is composed of the recycle stream (16) from the top of the rectification column G, which contains essentially propane, and the inlets of fresh propane (1), water (2) and oxygen (3) together.
  • the 590 0 C leaving the reactor A hot dehydrogenation gas mixture (5) is cooled in stages in the cooling and condensation stage B, wherein an aqueous condensate stream (6) is obtained.
  • the product gas stream (7) is fed to the compression stage C, in which it is compressed to 40 bar.
  • the compressed stream (8) is separated in the phase separator D into three partial streams.
  • the gas stream (1 1) is the absorption column E, the liquid, propane and propene-rich stream (10) is fed to the esterification reactor F.
  • propane and propene are preferably dissolved in an absorbent and fed together with the absorbent as stream (12) to the esterification reactor F.
  • the absorbent used is the acetic acid-rich stream (22) recycled from the bottom draw off of the rectification column I, which stream is supplemented with fresh acetic acid (34).
  • the gas stream (13) can be discharged as exhaust gas or, if appropriate after an additional separation step, recycled to the dehydrogenation reactor A.
  • 90% of the propene fed with the streams (12) and (10) is reacted with acetic acid to form isopropyl acetate at 40 bar and 110 ° C.
  • a selectivity of 100% is assumed to be isopropyl acetate.
  • the two effluents from the esterification reactor F, the gaseous stream (14) and the liquid stream (15) are fed to the rectification column G.
  • the gaseous top draw stream (16) of the column G is fed into the propane dehydrogenation reactor A and the liquid bottom draw stream (17) into the ester cracking reactor H.
  • the ester hydrolysis reactor H is supplied with water (18) and the two bottom draw streams (30) and (32) of the rectification columns L and N, respectively.
  • 20% isopropyl acetate is reacted at 5 bar and 100 ° C.
  • the two remaining head withdrawal streams (25) or (31) are guided into the phase separator M.
  • the substantially water-containing bottom draw stream (30) of the stripping column L is recycled to the ester hydrolysis reactor H.
  • the bottom draw stream (24) of the stripping column K is fed to the rectification column N.
  • the bottom draw stream (32) of column N which essentially contains isopropyl acetate, is recycled to the ester cleavage reactor H.
  • the desired product stream (33) is obtained with a content of 99.7% by mass of isopropanol.

Landscapes

  • Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)

Abstract

L'invention concerne un procédé de fabrication d'alcanols (I), choisis dans le groupe constitué de l'isopropanol et du 2-butanol, à partir des alcanes (II) correspondants, choisis dans le groupe constitué du propane et du n-butane. Selon l'invention, le procédé comprend les étapes suivantes : A) un courant gazeux d'alimentation a contenant l'alcane (II) est utilisé ; B) le courant gazeux d'alimentation a contenant l'alcane (II) est introduit dans une zone de déshydrogénation et l'alcane (II) subit une déshydrogénation en alcène (III), un courant gazeux produit b contenant l'alcène (III), l'alcane (II) non réagi, éventuellement des substances de point d'ébullition élevé, de la vapeur d'eau, de l'hydrogène et des substances de point d'ébullition bas étant obtenu ; C) le courant gazeux produit b est au moins comprimé, éventuellement séparé en une phase aqueuse c1, une phase c2 contenant l'alcène (III), l'alcane (II) et éventuellement les substances de point d'ébullition élevé et une phase gazeuse c3 contenant l'hydrogène et les substances de point d'ébullition bas ; D) le courant gazeux produit b, et respectivement la phase c2 contenant l'alcène (III) et l'alcane (II), est mis en réaction avec un acide organique (IV) dans une zone de formation d'ester, un mélange de produits d contenant l'ester alkylique (V) correspondant de l'acide organique et l'alcane (II) non réagi étant obtenu ; E) du mélange de produits d est séparé un courant gazeux e1 contenant l'alcane (II), ledit courant e1 pouvant éventuellement être remis en circulation dans la zone de déshydrogénation, et est obtenu un mélange de produits e2 contenant l'ester alkylique ; F) le mélange de produits e2 contenant l'ester alkylique est transformé dans une zone de décomposition d'ester avec de l'eau en un mélange de produits f contenant l'alcanol (I) et l'acide organique (IV) ; G) du mélange de produits f sont séparés l'alcanol (I) et l'acide organique (IV) et l'acide organique est éventuellement remis en circulation dans la zone de formation d'ester.
EP07787557A 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants Withdrawn EP2046706A1 (fr)

Priority Applications (2)

Application Number Priority Date Filing Date Title
EP07787557A EP2046706A1 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants
EP09153296A EP2070900A3 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir d'alcanes correspondants

Applications Claiming Priority (3)

Application Number Priority Date Filing Date Title
EP06117568 2006-07-20
EP07787557A EP2046706A1 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants
PCT/EP2007/057289 WO2008009648A1 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants

Related Child Applications (1)

Application Number Title Priority Date Filing Date
EP09153296A Division EP2070900A3 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir d'alcanes correspondants

Publications (1)

Publication Number Publication Date
EP2046706A1 true EP2046706A1 (fr) 2009-04-15

Family

ID=38645759

Family Applications (2)

Application Number Title Priority Date Filing Date
EP09153296A Withdrawn EP2070900A3 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir d'alcanes correspondants
EP07787557A Withdrawn EP2046706A1 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants

Family Applications Before (1)

Application Number Title Priority Date Filing Date
EP09153296A Withdrawn EP2070900A3 (fr) 2006-07-20 2007-07-16 Procédé de fabrication d'isopropanol et de 2-butanol à partir d'alcanes correspondants

Country Status (10)

Country Link
US (1) US7750194B2 (fr)
EP (2) EP2070900A3 (fr)
JP (1) JP2009543845A (fr)
KR (1) KR20090031633A (fr)
CN (1) CN101516816A (fr)
CA (1) CA2657739A1 (fr)
MX (1) MX2009000560A (fr)
RU (1) RU2009105743A (fr)
TW (1) TW200812952A (fr)
WO (1) WO2008009648A1 (fr)

Families Citing this family (14)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
EP2220016A1 (fr) * 2007-12-10 2010-08-25 Basf Se Procédé pour produire de l'isopropanol et du 2-butanol à partir d'alcanes adéquats
JP5374921B2 (ja) * 2008-05-14 2013-12-25 住友大阪セメント株式会社 透明性微粒子分散液と微粒子含有透明性樹脂組成物及び微粒子含有透明性樹脂並びに光学部材
WO2010071011A1 (fr) * 2008-12-15 2010-06-24 新日本石油株式会社 Méthode de production d'un ester d'acide acétique
US7588690B1 (en) * 2009-02-10 2009-09-15 The Purolite Company Method of iodide removal
KR101582001B1 (ko) * 2012-06-28 2015-12-31 주식회사 엘지화학 이소프로필 알코올의 제조 방법
US9611191B2 (en) * 2012-12-12 2017-04-04 Basf Se Reactor for carrying out an autothermal gas-phase dehydrogenation
WO2015026160A1 (fr) * 2013-08-20 2015-02-26 주식회사 엘지화학 Procédé de purification d'alcool isopropylique
JP6181312B2 (ja) 2013-08-20 2017-08-16 エルジー・ケム・リミテッド イソプロピルアルコールの精製方法
WO2015026162A1 (fr) * 2013-08-20 2015-02-26 주식회사 엘지화학 Procédé de purification d'alcool isopropylique
WO2015026161A1 (fr) * 2013-08-20 2015-02-26 주식회사 엘지화학 Procédé de purification d'alcool isopropylique
US9919260B2 (en) * 2013-11-22 2018-03-20 Lg Chem, Ltd. Method of recovering absorption solvent in butadiene production process by oxidative dehydrogenation
CN106800509A (zh) * 2015-11-26 2017-06-06 中国石油化工股份有限公司 一种醋酸异丙酯的生产方法
US10570071B1 (en) 2018-12-12 2020-02-25 Saudi Arabian Oil Company Membrane-based process for butanols production from mixed butenes
DE102019213493A1 (de) 2019-09-05 2021-03-11 Thyssenkrupp Ag Verfahren zur Herstellung von Alkoholen

Family Cites Families (21)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
IT1110745B (it) 1979-01-30 1986-01-06 Anic Spa Procedimento per la sintesi di sec-butanolo da buteni lineari
US4384148A (en) 1982-02-26 1983-05-17 Uop Inc. Hydration of olefins
EP0117146B1 (fr) 1983-02-22 1986-12-30 The Halcon Sd Group, Inc. Conversion du propane en acide acrylique
GB2147290A (en) * 1983-09-29 1985-05-09 Humphreys & Glasgow Ltd Alcohols from liquid petroleum gas
US4484013A (en) 1983-12-30 1984-11-20 Uop Inc. Process for coproduction of isopropanol and tertiary butyl alcohol
US4788371A (en) 1987-12-30 1988-11-29 Uop Inc. Catalytic oxidative steam dehydrogenation process
GB2238539A (en) 1989-11-29 1991-06-05 Shell Int Research Preparation of alcohols from olefins
US5220091A (en) 1992-02-26 1993-06-15 Phillips Petroleum Company Alkane dehydrogenation
US5344805A (en) 1993-05-03 1994-09-06 Phillips Petroleum Company Platinum and tin-containing catalyst and use thereof in alkane dehydrogenation
NO179131C (no) 1993-06-14 1996-08-14 Statoil As Katalysator, fremgangsmåte for dens fremstilling og fremgangsmåte for dehydrogenering av lette paraffiner
US5877369A (en) 1997-10-23 1999-03-02 Phillips Petroleum Company Hydrocarbon conversion catalyst composition and processes therefor and therewith
DE19756292C2 (de) 1997-12-10 1999-10-28 Mannesmann Ag Katalysator zur Umwandlung paraffinischer Kohlenwasserstoffe in korrespondierende Olefine
NO307693B1 (no) 1998-03-13 2000-05-15 Statoil Asa Forbedret katalysatorbaerermateriale, fremgangsmate for fremstilling derav, dehydrogeneringskatalysator, og katalytisk dehydrogenering av lette alkaner
DE19937106A1 (de) 1999-08-06 2001-02-08 Basf Ag Multikomponenten-Katalysatoren
DE19937107A1 (de) 1999-08-06 2001-02-08 Basf Ag Katalysator mit bimodaler Porenradienverteilung
DE19937105A1 (de) 1999-08-06 2001-02-08 Basf Ag Oxidkatalysatoren
DE10028582A1 (de) 2000-06-14 2001-12-20 Basf Ag Verfahren zur Herstellung von Acrolein oder Acrylsäure oder deren Gemischen aus Propan
DE10211275A1 (de) 2002-03-13 2003-09-25 Basf Ag Verfahren der kontinuierlichen heterogen katalysierten partiellen Dehydrierung
DE10245585A1 (de) 2002-09-27 2004-04-08 Basf Ag Verfahren zur Herstellung von wenigstens einem partiellen Oxidations- und/oder Ammoxidationsprodukt des Propylens
DE10246119A1 (de) 2002-10-01 2004-04-15 Basf Ag Verfahren zur Herstellung von wenigstens einem partiellen Oxidations- und/oder Ammoxidationsprodukt des Propylens
US7057081B2 (en) 2003-05-09 2006-06-06 Conocophillips Company Method for treating alkanes

Non-Patent Citations (1)

* Cited by examiner, † Cited by third party
Title
See references of WO2008009648A1 *

Also Published As

Publication number Publication date
US20100048960A1 (en) 2010-02-25
WO2008009648A1 (fr) 2008-01-24
CA2657739A1 (fr) 2008-01-24
TW200812952A (en) 2008-03-16
CN101516816A (zh) 2009-08-26
RU2009105743A (ru) 2010-08-27
EP2070900A3 (fr) 2012-01-11
KR20090031633A (ko) 2009-03-26
US7750194B2 (en) 2010-07-06
MX2009000560A (es) 2009-02-18
JP2009543845A (ja) 2009-12-10
EP2070900A2 (fr) 2009-06-17

Similar Documents

Publication Publication Date Title
EP2046706A1 (fr) Procédé de fabrication d'isopropanol et de 2-butanol à partir des alcanes correspondants
EP1682468B1 (fr) Procede de fabrication de 1-butene
EP1708977B1 (fr) Procede de production de butadiene et 1-butene
EP1453778B1 (fr) Procede pour produire du 1,2-dichloroethane
EP1824803B1 (fr) Procede de fabrication de propene a partir de propane
EP3218334B1 (fr) Procédé de fabrication de 1,3 butadiène par la déshydratation de n-butènes par préparation d'un flux de matière contenant du butane et de 2-butène
WO2006069674A1 (fr) Procede de production de propene a partir de propane
WO2006094938A2 (fr) Procede de production de propene a partir de propane
EP2897927B1 (fr) Procédé de production de butadiène par élimination de l'oxygène de flux d'hydrocarbures en c4
WO2006066848A1 (fr) Procede de production de butadiene a partir de n-butane
EP1836146B1 (fr) Procede de production de propene a partir de propane
WO2015007698A1 (fr) Procédé de production de 1,3-butadiène à partir de n-butènes par déshydrogénation oxydative
EP1701928B1 (fr) Procede de production de butadiene
WO2018178005A1 (fr) Procédé pour l'arrêt et la régénération d'un réacteur pour la déshydrogénation oxydative de n-butènes
WO2009074574A1 (fr) Procédé pour produire de l'isopropanol et du 2-butanol à partir d'alcanes adéquats
DE102005000798A1 (de) Verfahren zur Herstellung von Propen aus Propan
WO2016151074A1 (fr) Procédé de préparation de 1,3-butadiène à partir de n-butènes par déshydrogénation oxydative
WO2016150940A1 (fr) Procédé de préparation de 1,3-butadiène à partir de n-butènes par déshydrogénation oxydative
EP1478610B1 (fr) Procede de production de 4-vinylcyclohexene, ethylbenzene et styrene
EP1678105B1 (fr) Procede de fabrication du 1-butene
WO2018095840A1 (fr) Procédé pour la préparation de 1,3-butadiène à partir de n-butènes par déshydrogénation oxydante comprenant une séparation de furanne lors du traitement
WO2003029171A1 (fr) Procede de production de composes aromatiques alkyles
WO2003089407A1 (fr) Procede de preparation de nitriles insatures
DE10217844A1 (de) Verfahren zur Herstellung von ungesättigten Nitrilen aus Alkanen

Legal Events

Date Code Title Description
PUAI Public reference made under article 153(3) epc to a published international application that has entered the european phase

Free format text: ORIGINAL CODE: 0009012

17P Request for examination filed

Effective date: 20090220

AK Designated contracting states

Kind code of ref document: A1

Designated state(s): AT BE BG CH CY CZ DE DK EE ES FI FR GB GR HU IE IS IT LI LT LU LV MC MT NL PL PT RO SE SI SK TR

AX Request for extension of the european patent

Extension state: AL BA HR MK RS

STAA Information on the status of an ep patent application or granted ep patent

Free format text: STATUS: THE APPLICATION HAS BEEN WITHDRAWN

18W Application withdrawn

Effective date: 20100914