WO2009074574A1 - Method for producing isopropanol and 2-butanol from the corresponding alkanes - Google Patents

Abstract

The invention relates to a method for producing alcohols (I), selected from isopropanol and 2-butanol, from the corresponding alkanes (II), selected from propane and n-butane, comprising the steps: A) a feed stream a, comprising the alkane (II) is prepared, B) the feed stream a, comprising the alkane (II) is introduced into a dehydrogenation zone and the alkane (II) subjected to a dehydrogenation to give the alkene (III), wherein a product gas stream b containing the alkene (III), unreacted alkane (II), optional high-boilers, steam, hydrogen and low-boilers is obtained, C) the product gas stream b is at least compressed, the product gas stream b is optionally separated into an aqueous phase c1, a phase c2, comprising the alkene (III), the alkane (II) and optionally high-boilers and a gas phase c3, comprising hydrogen and low-boilers, D) the product gas stream b and/or the phase c2, comprising the alkene (III) and alkane (II) is reacted in an esterification zone with an alkanoic acid (IV) having 3 or more carbon atoms, wherein a product mixture d, comprising the corresponding alkyl ester (V) of the alkanoic acid and the unreacted alkane (II) is obtained, E) a gas stream e1 comprising an alkane (II) is separated off from the product mixture d, optionally recycled back to the dehydrogenation zone and a product mixture e2, comprising the alkyl ester (V) is obtained; F) the product mixture e2, comprising the alkyl ester (V) is reacted with water in an ester cleavage zone to give a product mixture f containing the alcohol (I) and the alkanoic acid (IV), G) the alcohol (I) and the alkanoic acid (IV) are separated off from the product mixture f and the alkanoic acid optionally recycled to the esterification zone.

Description

 Process for the preparation of isopropanol and 2-butanol from the corresponding alkanes

description

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. Technically, 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 before its reuse in the esterification step.

In addition, the esterification of olefins with carboxylic acids and subsequent ester cleavage of the alkyl esters formed to give the corresponding alcohol with recovery of the acid is known. Thus, US 4,384,148 describes the reaction of ethene with acetic acid to ethyl acetate in an autoclave. The subsequent hydrolysis of the distillatively separated ethyl acetate with water in a molar ratio of 1: 5 in an autoclave yields a mixture containing ethyl acetate, ethanol and diethyl ether. 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.

Furthermore, 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. A disadvantage of the single-stage process are in particular the low sales and the high demands on the purity of the alkene used. For example, 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 passed through an etherification zone in which isobutene is etherified with methanol to 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 comprising propane or a longer alkane alkane is converted to an intermediate product stream containing the corresponding olefin, and the intermediate product stream direct or indirect hydration to a product stream containing the corresponding alcohol is reacted. The document mentions as an indirect process the two-stage hydration of propene with concentrated sulfuric acid with intermediate formation of the sulfuric acid ester and subsequent hydrolysis of the sulfuric acid ester to the alcohol. In addition, various methods of direct hydration are described.

From US 4,484,013 a process for the preparation of isopropanol and tert-butanol is known, in which a feed gas stream of propane and isobutane fed into a dehydrogenation zone and non-oxidative, ie, in a product gas mixture containing propene, isobutene, and unreacted propane and isobutane Absence of air or oxygen, being dehydrated. From the product gas stream of the dehydrogenation first low boilers (hydrogen) and then propane are separated by distillation, the latter being recycled to the dehydrogenation zone. 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.

The object is achieved by a process for the preparation of 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 :

A) a feed gas stream a containing the alkane (II) is provided;

B) the feed gas stream a containing the alkane (II) is fed into a dehydrogenation zone and the alkane (II) is subjected to dehydrogenation to the alkene (III), a product gas stream b comprising the alkene (III), unreacted alkane (II), optionally high boilers, water vapor, hydrogen and low boilers are obtained;

C) the product gas stream b is at least compressed, optionally the product gas stream b is separated into an aqueous phase c1, an alkene (III), the alkane (II) and optionally high boilers containing phase c2 and a gas phase c3 containing hydrogen and low boilers;

D) the product gas stream b or the alkene (III) and the alkane (II) containing phase c2 is reacted in an esterification zone with an alkanoic acid having 3 or more carbon atoms (IV), wherein a product mixture d containing the corresponding alkyl ester (V ) of the alkanoic acid (IV) and the unreacted alkane (II) is obtained;

E) is separated from the product mixture d an alkane (II) containing gas stream e1, which is optionally recycled to the dehydrogenation, and a

Product ester e2 containing alkyl ester (V) obtained;

F) the product mixture e2 containing the alkyl ester is reacted with water in an ester cleavage zone to give a product mixture f containing the alkanol (I) and the alkanoic acid (IV);

G) from the product mixture f, the alkanol (I) and the alkanoic acid (IV) are separated and the alkanoic acid is optionally recycled to the esterification zone. 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 (III) only in very dilute form in addition to other components (unreacted alkane, inert gases), high space / time yields and sales are achieved. Separation of non-reactive secondary components before the esterification step can therefore be dispensed with.

In the first process part A) is provided an alkane (II), selected from propane and butane, containing feed gas stream a. In the case of propane this contains in the

Generally at least 80% by volume of propane, preferably 90% by volume of propane. In addition, it generally still contains butane (n-butane, iso-butane), butenes, ethane and ethene.

Typical compositions of the propane-containing feed gas stream are described in DE-A 102

46 1 19 and DE-A 102 45 585 discloses. The propane-containing feed gas stream a is usually obtained from liquefied petroleum gas (LPG).

In the case of n-butane as alkane (II), 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.

In a process part B), the feed gas stream containing the alkane (II) is fed into a dehydrogenation zone and subjected to a generally catalytic dehydrogenation. In this case, the alkane is partially dehydrogenated in a dehydrogenation reactor on a dehydrogenation catalyst to the alkene. In addition, hydrogen and small amounts of low boilers and high boilers. Low-boiling hydrocarbons are referred to herein as lower hydrocarbons than propene or 1-butene, and higher-boiling hydrocarbons referred to as propene or 2-butene, respectively. Thus, for example, in the propane dehydrogenation as low boilers methane, ethane and ethene and high boilers C ₄ ⁺ hydrocarbons (n-butane, isobutane, butenes, butadiene) incurred. In n-butane dehydrogenation, for example, methane, ethane and ethene, propane and propene can be obtained as low boilers and C ₅ ⁺ hydrocarbons as high boilers. In addition, if 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 optionally, to a lesser extent, inert gases in the product gas mixture generally fall to the dehydrogenation. The dehydrogenation product gas stream 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 on dehydrogenation. Inert gases (nitrogen) become when carrying out the dehydrogenation in the presence of oxygen introduced with the fed oxygen-containing gas stream in the dehydrogenation zone, if no pure oxygen is fed. In addition, unreacted alkane (II) (propane or n-butane) 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. For this purpose, 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 Dehydrierwärme 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. In these, the catalyst (dehydrogenation catalyst and optionally special oxidation 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. but also between 3 and 8 bar when using a high steam dilution (according to the so-called "steam active reforming process" (STAR process) or the Linde process) for the dehydrogenation of propane or butane by Phillips Petroleum Co. loads (GHSV) are from 500 to 2000 h ^"1, based on the hydrocarbon used. 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. Conveniently, 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 dehydrated temperature usually 550 to 600 ⁰ C. The heat required for the dehydrogenation can be introduced into the reaction system in that the dehydrogenation catalyst is preheated to the reaction temperature. By adding 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. Optionally, 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 may 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. In general, such a tray reactor is operated with a fixed catalyst bed. In the simplest case, 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. In a further preferred embodiment, the dehydrogenation is carried out in a tray reactor with 3 catalyst beds.

In general, 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. In general, the total amount of oxygen introduced, based on the total amount of propane, is 0.001 to 0.5 mol / mol, preferably 0.005 to 0.25 mol / mol, particularly preferably 0.05 to 0.25 mol / mol. Angry- The substance can be used either as pure oxygen or as an oxygen-containing gas containing inert gases. In order to avoid high propane and propene losses in the work-up (see below), it may be advantageous if 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% by volume. 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. Preferably, so much hydrogen should be present 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. In one embodiment, 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. Preferably, the dehydrogenation catalyst and the oxidation catalyst are present in different reaction zones.

In the case of a multistage reaction, the oxidation catalyst can be present in only one, in several or in all reaction zones.

Preferably, 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.

In one embodiment of the process according to the invention, an intermediate feed of oxygen-containing gas and of hydrogen-containing gas takes place before each tray of a tray reactor. In a further embodiment of the inventive In the process, oxygen-containing gas and hydrogen-containing gas are introduced to each horde except the first horde. In one embodiment, behind each feed point is a layer of a special oxidation catalyst, followed by a layer of the dehydrogenation catalyst. In another embodiment, no special oxidation catalyst is present. The dehydrogenation temperature is generally 400 to 1100 ⁰ C, the pressure in the last catalyst bed of the tray reactor generally 0.2 to 5 bar, preferably 1 to 3 bar. The load (GHSV = gasous hourly space velocity) is generally from 500 to 2000 h ^-1 , in the case of high-load operation also up to 100 000 h ^-1 , preferably 4000 to 16 000 h ^-1 .

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. Nebengruppe.

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. Preferably, 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 zirconium dioxide and / or silicon dioxide, particular preference is given to mixtures of zirconium dioxide and silicon dioxide.

The active composition of the dehydrogenation catalysts generally contains one or more elements of subgroup VIII, preferably platinum and / or palladium, more preferably platinum. In addition, the dehydrogenation catalysts may comprise one or more elements of main group I and / or II, preferably potassium and / or cesium. Furthermore, the dehydrogenation catalysts may contain one or more elements of III. Subgroup including the lanthanides and actinides, preferably lanthanum and / or cerium. Finally, 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.

In a preferred embodiment, 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 min. at least one element of the III. Subgroup including lanthanides and actinides.

For example, according to the invention, 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. Thus, 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. Optionally, the catalyst is reduced after regeneration with a hydrogen-containing gas.

In the autothermal propane dehydrogenation under feeding of substantially pure oxygen, a gas mixture is obtained, 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.

In the autothermal butane dehydrogenation with introduction of substantially pure oxygen, a gas mixture is obtained, 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 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 Hydrogen and 0 to 5 vol .-% 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.

In process part C), 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. In one embodiment of the method according to the invention, a one- or two-stage compression of the product gas stream b to a pressure in the range of 5 to 12 bar and then a one or two stage compression to a pressure in the range of 10 to 25 bar. The cooling can be done in several stages, preferably it is multi-stage. As the coolant air coming pan in air coolers, flow or cold water, as well as refrigerants, such as ethene, propene and Pro, which by compacting at pressures up to 20 bar and subsequent decision-span temperatures in the range from -40 ⁰ C to -100 ⁰ C. be cooled, used.

Optionally, 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).

From the product gas stream b, 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. In general, 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 done before the compression and / or in the compression stages as intermediate cooling. In one embodiment of the method according to the invention, the product gas stream b is passed through a cascade of heat exchangers and thus first cooled to a temperature in the range of 50 to 200 ⁰ C and then in a quench tower with water to a temperature of 40 to 80 ⁰ C, for example 55 ° C, further cooled. In this case, the major part of the water vapor condenses, but also a part of the high boilers contained in the product gas stream b. In the case of propane dehydrogenation, these may be C ₄ ⁺ hydrocarbons, in particular the C ₅ ⁺ hydrocarbons.

In this case, a water vapor-depleted product gas stream b is obtained. This generally contains up to 5 vol .-% water vapor. For practically complete removal of water from the product gas stream b, drying by means of molecular sieve can be provided.

If 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.

Subsequently, 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. In the case of propane dehydrogenation, 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 chosen 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, either as well as the hydrocarbon stream c2, can be fed into the esterification zone D or alternatively into an absorption stage (as described below).

In the case of 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. In addition, it may also ethane and ethene and - if the autothermal dehydrogenation is not carried out under the supply of pure oxygen - in particular nitrogen and noble gases (mainly argon) included. In addition, it may also contain C ₃ and C ₄ hydrocarbons. For this purpose, 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. In addition to the hydrocarbon stream c2, 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 no complete removal of water from the product gas stream b has taken place before the condensation step. In multi-stage cooling and compression all accumulating condensate streams can be fed to the phase separator.

Prior removal of water from the product gas stream b before condensing the C ₃ -hydrocarbon phase c 2 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.

In general, 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 uncondensed gas is removed, which is not condensed in the uppermost cooling circuit.

If the alkane dehydrogenation is carried out as autothermal dehydrogenation with simultaneous combustion of the hydrogen formed, the result is a low hydrogen content of the product gas stream b. As a consequence, in 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.

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.

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. In general, before carrying out the gas wash the product gas stream c by single or multi-stage compression to a pressure in the range of 5 to 25 bar compressed.

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.

In the case of propane dehydrogenation, 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.

If the autothermal alkane dehydrogenation is carried out with the introduction of air as an 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. First of all, steam is separated off by condensation by cooling the product gas stream b and optionally compacting it, with a product gas stream b depleted in water vapor being obtained. Subsequently, 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 also 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.

For this purpose, in the absorption stage at 5 to 40 bar, preferably 8 to 20 bar, particularly preferably 10 to 15 bar, 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 -KoIi hydrocarbons are absorbed in the inert absorbent and a C ₃ - or C ₄ -Koh hydrocarbons laden absorbent and the other gas components containing exhaust c3 are obtained. Essentially, These are carbon oxides, hydrogen, inert gases as well as 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. In a desorption step, the C ₃ - or C ₄ -Koh hydrocarbons are released from the absorbent again.

Inert absorbent used in the absorption stage are generally high-boiling non-polar solvents in which the separated C3 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-cap, valve, and / or sieve trays, columns with structured packings, gen example Gewebepackun- 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. However, there are also trickle and spray towers, graphite block absorbers, surface absorbers such as thick-film and thin-layer absorbers and bubble columns with and without internals into consideration.

Preferably, the absorption column has an absorption part and a rectification part. To increase the concentration of the C ₃ - or C ₄ -hydrocarbons in the solvent in the manner of a rectification, heat can then be introduced into the bottom of the column. Alternatively, a stripping gas stream can be fed into the bottom of the column, for example from nitrogen, air, steam or propane / propene mixtures. Thus, in the rectification section of the absorption column, the laden absorbent is brought into contact with the stripping gas stream. As a result, C ₂ ^" hydrocarbons are stripped out of the loaded absorbent A portion of the top product can be condensed and returned to the column head as reflux to limit solvent losses.

Suitable absorbents are relatively nonpolar organic solvents, for example C ₄ -C ₈ -alkenes, naphtha or aromatic hydrocarbons, such as the paraffin distillation medium fractions, or bulky group ethers, or mixtures of these solvents, such as polar solvents such as 1, 2 and 3. Dimethyl phthalate may be added. Suitable absorbers are also esters of benzoic acid and phthalic acid with straight-chain C 1 -C 6 -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 C ₉ -hydrocarbons, in particular nonanes.

For the desorption of the C ₃ or C ₄ hydrocarbons, the laden absorbent is heated and / or expanded to a lower pressure. Alternatively, the desorption can also be effected by stripping, usually with steam, or in a combination of relaxation, heating and stripping in one or more process steps. For example, 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. Optionally, a portion of this absorbent stream, which may contain C ₄ ⁺ hydrocarbons, is discharged, worked up and recycled, or discarded.

In a variant of the method, the desorption step is carried out by relaxation and / or heating of the loaded absorbent. In a further process variant 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.

Subsequently, the desorbed C ₃ or C ₄ hydrocarbon stream can be cooled, wherein 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. In addition, an aqueous condensate stream and possibly further amounts of the exhaust gas stream c3 incurred. 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. In general, 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. After each compression stage is followed by a cooling step, in which the gas stream is cooled to a temperature in the range of 15 to 80 ⁰ C, preferably 15 to 60 ⁰ C, cooled. Subsequently, the compressed gas mixture is cooled to a temperature of -10 ° C to 60 ⁰ C, preferably -10 ° 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 above-described absorption step can also be carried out as follows:

The absorbent used in the absorption column is the same alkanoic 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 alkanoic acid, can then, if appropriate after further heating and / or compression, be passed directly into the esterification zone. In this case, the absorption of the absorbent into the absorption column can be carried out using the alkanoic acid separated off from the separation step (G) after the ester cleavage zone. This is then not directly in the esterification zone, but in the absorption stage and from there, charged with the C ₃ - or C ₄ -KoIi hydrogens, recycled to 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. For example, if the alkane dehydrogenation is not carried out with the introduction of an oxygen-containing gas and also not in the presence of water vapor, the resulting dehydrogenation gas mixture b, which in the case of propane dehydrogenation consists essentially of propane, propene, hydrogen and low boilers and the n-butane dehydrogenation consists essentially of butane, 1-butene, 2-butene, hydrogen and low boilers, fed directly and without prior separation of the C ₃ - or C ₄ -KoIi hydrogens in the esterification zone and in contact with the organic acid to be brought. If 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. Also the presence of Carbon oxides and other inert gases (atmospheric nitrogen) do 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)).

In a process stage D), the product gas stream b or - if the separation step in the process step C) is carried out - the alkane (II) and the

Alkene (III) containing phase c2 and optionally the residual gas stream c3 or, if the absorption step is carried out as described above with the organic acid, the C3- or C ₄ -Koh hydrocarbons laden alkanoic acid (IV) in an esterification zone with the alkanoic acid (IV ), where a product mixture d containing the corresponding alkyl ester (V) (isopropyl ester or 2-

Butyl ester) of the alkanoic 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. In general, the alkanoic acid (IV), 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, 2-2 , 5 moles per mole of alkene used. In general, 50 to 90% of the alkene reacts to the corresponding alkyl ester. In addition, unsaturated secondary components contained in the product gas stream b or in the stream c2 can react to alkyl esters. These can readily be removed by distillation in the subsequent work-up steps. The alkanes present in stream b or c2 do not naturally react in the ester formation stage D) and do not react. leave esterification step D) together with alkene (III) unreacted in the esterification in the case of operation of esterification step D) in homogeneous liquid phase in stream d or in case of operation of esterification step D) in gas-liquid mode as additional stream d2 ). In addition to alkane and alkene, this stream can still contain small amounts of the ester formed, as well as water and the low boilers already contained in stream c2.

When operating in gas-liquid mode, 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 alkanoic acid (IV) with the alkene (III). In the backwashing zone, the gas stream d2 is largely freed from acid residues by the addition of water. This can be corrosion problems in other parts of the system in the return of the current d2) avoided.

The ester formation can be carried out in fixed-bed reactors in trickle-bed mode or in fluidized bed mode. In principle, the organic acid and the alkene-containing stream can be conducted in cocurrent or countercurrent, but also in crossflow. Preferably, a cascading of several catalyst beds takes place, optionally with intermediate introduction of the acid and / or the alkene-containing stream. Preferably, 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. Also, 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. Alternatively, the esterification can also be carried out in bubble columns or jet loop reactors. A particular embodiment is the tubular reactor in fluidized bed mode, which is operated selectively at the loosening point of the catalyst (so-called floating bed method), with external heat exchanger and 2-fold cascading.

In the backwash zone, removal of acid residues occurs by the addition of water. Apparently, 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 residue in the gas stream d2 <50 ppm, preferably << 10 ppm, more 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 wash water loaded with acid residues can be used for ester cleavage in process stage F). 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 differentiates 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 kind 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. Particularly preferred are those catalysts which contain a high content of acidic groups such as. Amberlyst 35, 36, 40, 49, 119, 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 or with an alcohol, preferably isopropanol or 2-butanol. A group of catalysts related to acidic ion exchange resins is derived from sulfonated polycondensed aromatics or graphitic carbon, respectively. 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. In addition to the polymeric ion exchange resins, a number of inorganic catalysts, such as acidic metal oxide catalysts or acidic zeolites, are also suitable. Sulfurized zirconium oxide and zirconium tungstate or titanium tungstate systems are to be expected in particular with regard to the acid metal oxide catalysts. Furthermore, this group of catalysts include the water-insoluble acidic salts of heteropolyacids, for example, tungstophosphoric 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 ⁺ . In these salts, usually 10-90 mol%, in particular 40-85 mol%, of the acidic sites 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. In addition to the zeolite structure, the ratio of the Si / Al atoms (the modulus) in the zeolite framework is crucial for the catalytic activity of the zeolite. For the process according to the invention are those zeolites which have a modulus between 2 and 500, in particular between 3 and 200 and very particularly preferably between 5 and 100. The inorganic catalysts described herein are typically thermally activated, ie the materials are defined at temperatures from 50 to 900 ⁰ C, preferably 90-500 ⁰ C calcium.

The esterification reaction can also be carried out homogeneously catalyzed. Mineral acids, in particular sulfuric acid, sulfonic acids or the free heteropoly acids 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 potassium, cesium or ammonium salts of the heteropolyacids, and also beta zeolites and faujasites. Particularly preferred are ion exchange resins.

Preferred alkanoic acids (IV) are unbranched or branched C ₄ -C 10 -alkanoic acids, particular preference is given to unbranched or branched C ₄ -C 6 -alkanoic acids, particular preference to butyric acid (butanoic acid), valeric acid (pentanoic acid) and isovaleric acid (3-methylbutanoic acid). For example, propene reacts with butyric acid to form isopropylpyrrole and valeric acid or isovaleric acid to form isopropylvalerate or isopropylisovalerate. 1-Butene and 2-butene react with butyric acid to 2-butyl butyrate and with valeric acid or isovaleric acid to 2-butylvalerate or 2-butylisovalerate.

In process stage E), a gas stream e1 containing the unconverted alkane (II) (propane or n-butane), which is optionally recycled to the dehydrogenation zone, is separated from the product mixture d of the ester formation. For this purpose, the product mixture d of the ester formation is generally expanded, a gas stream e1 containing propane or n-butane being separated off, and a product mixture e2 containing the alkyl ester (V) being obtained. The gas stream e1 can be combined with the gas stream d2 of the backwashing zone in the process part D and fed together to the dehydrogenation zone B.

In general, 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. With the propane or n-butane, if appropriate, low boilers such as ethane, ethene, methane, carbon oxides and inert gases contained in the product mixture are also removed. 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 not already in step C) the above-described (optional) separation of low boilers and hydrogen (residual gas stream c3) is carried out. From the gas stream e1, a partial stream can be 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. 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.

In order to achieve better material separation between alkene / alkane on the one hand and the ester / acid mixture on the other hand than in the case of pure relaxation, distillation or rectification can additionally be carried out. Typically, for such a separation at a reflux ratio between 0.2 and 1.5, between 10 and 20 theoretical plates are required. Here, both 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. Variants I are distinguished by the fact that the C3 components (propane and propene) are separated "sharply", with very little propane and propene in the bottom discharge (corresponding to stream e2) (for example about 100-1000 ppm by mass) The advantage of these variants is that no or only very little C3-hydrocarbons get into the subsequent stages F) and G). If the C3-components are not separated sharply, as it corresponds to the variant II described below, then this can in stages F) and G) bring a higher procedural effort with it.

According to a variant Ia, the gas stream e1 present after the expansion is fed to a first column K1. The pressure in this column is the same or only weak. less than the pressure of the feed stream. The column K1 is preferably the pure enrichment part of a conventional column, ie it 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 of the column K2 is well below the pressure of the column K1. Typical pressures are, for example, for K1 = 30 bar (all pressures absolute) and for K2 = 1.5 bar. In general, 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 reused or further worked up as described above. In this case, the top outlet of the second column K2 is preferably recycled directly into the dehydrogenation, optionally using a compressor.

According to a further variant Ib is carried out as described under variant Ia, except that the pressure of the column K2 is higher here (typical values, for example, K1 = 30 bar, K2 = 5 bar). In general, 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.

According to a further variant Ic is carried out as described under variant Ia, with the difference that 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. As a result, the top draw stream of the column K2 still contains large amounts of alkyl ester and consequently can not be recycled directly to the dehydrogenation. Therefore, this stream is compressed to the pressure of the column K1 and fed to the column K1. In this variant, only the top draw stream of the columns K1 corresponds to the stream e1 and can be used further as described above or further worked up. In general, 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.

According to a further variant Il, 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 for evaporator and bottom part of the column K2 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. In the same way, optionally, the recycle stream d2 can be treated if it contains even larger amounts of alkyl ester (V). Again, 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 alkanoic acid (IV), for example consisting of Isopro- pylisovalerat and isovaleric acid) can be further worked up before the stream e2 in the process step F) out becomes. In this case, in a distillation column, the alkanoic acid (IV), for example isovaleric acid, from the alkyl ester (V), for example Isopropylisovalerat be separated. The alkanoic acid (IV), for example isovaleric acid, can be obtained as the bottom product and recycled to the esterification stage. The alkyl ester (V), for example isopropyl isovalerate, 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 isovaleric acid and isopropyl isovalerate are a pressure of up to 1, 5 bar (all pressures absolute) and a reflux ratio of 0 to 3. The separation of the alkanoic acid (IV) becomes the reaction equilibrium in the hydrolysis reactor the process step G) positively influenced. This allows a higher hydrolysis conversion, with the additional consequence that the recycle streams (water = stream k1, isopropyl isovalerate = stream h2) are greatly reduced. As a result of the much smaller return flows, the investment and energy requirements of, for example, the apparatus (1), (2), (3) and (4) described below are reduced. As a result of the described separation of the alkanoic acid (IV) before carrying out the hydrolysis stage F), addition stage D) and hydrolysis stage F) are additionally decoupled.

In a process stage F), 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 alkanoic acid (IV). In a process stage G), the alkanol (I) is separated off from the product mixture f and the alkanoic acid (IV) is recovered. The alkanoic acid (IV) is generally recycled to the esterification stage D) or it is fed as an absorbent into the absorption stage 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 alkanoic acid (IV) and the alkyl ester (V) and water and a stream g2 essentially consisting of the alkanol (I) and water. The stream gl is separated by distillation into the alkanoic acid (IV) and into a mixture of alkyl ester (V) and water. The Alkanoic acid (IV) is recycled to the esterification stage D or to the adsorption stage to separate the alkene and alkane from the gas phase c3, while the mixture of alkyl ester (V) and water is recycled to the ester cleavage zone F. The separation of the g2 consisting of alkanol (I) / water can, 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 under Use of an extractant (eg of ionic liquids or acetic acid) as well as by membrane processes (pervaporation or vapor permeation) take place.

The ester cleavage can be carried out both homogeneously and heterogeneously catalyzed. For the ester cleavage by hydrolysis, in principle, 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.

Ester cleavage and distillation can be carried out in separate process steps. In a variant of the process according to the invention, the mixture e2 containing the alkyl ester (V) is reacted with water in an ester cleavage reactor to give a product mixture f containing the alkanol (I), the alkanoic acid (IV), the alkyl ester (V) and water, and then this mixture in at least two distillation columns connected in series, optionally in conjunction with a dividing wall column, separated

In the case of heterogeneously catalyzed ester cleavage, the ester cleavage reactor can be designed as a fixed bed, trickle bed, fluidized bed or slurry reactor. In the case of the fluidized bed reactor, the catalyst can be operated selectively at the loosening point (so-called floating bed mode). In the case of homogeneously catalyzed ester cleavage, 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 off with the alkanoic acid in process step G), generally via the bottom draw stream of the second distillation column. In this case, before recycling the alkanoic acid to the esterification step D), the catalyst may be separated and the separated catalyst recycled to the ester cleavage 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 by a combination of phase separation and thermal separation. Is used for the ester cleavage of the same homogeneous catalyst 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.

For example, in a first specific embodiment of the process according to the invention, isovaleric acid is used as the alkanoic acid (IV) and a product mixture e2 containing isopropyl isovalerate is obtained, where

in step F) the isopropyl isovalerate-containing product mixture e2 is reacted in an ester cleavage reactor with water to give a product mixture f comprising isopropanol, isovaleric acid, isopropyl isovalerate and water, and

in step G), the product mixture f in a first distillation column (1) into a stream gl consisting essentially of isovaleric acid, isopropyl isovalerate and water and a stream g2 consisting essentially of isopropanol and water is separated. The stream gl is separated in a second distillation column (2) into a stream h1 (which is generally obtained as a bottom draw stream) consisting essentially of isovaleric acid and a stream h2 (which is generally obtained as a top draw stream) consisting essentially of isopropyl isovalerate and water , The stream h1 can be recycled to the esterification zone or to the optional adsorption step to separate the alkane / alkene mixture. The stream h2 is returned directly to the hydrolysis reactor. This variant is shown in FIG.

For example, in a second specific embodiment of the process according to the invention, n-valeric acid is used as the alkanoic acid (IV) and a product mixture e2 containing isopropyl valency is obtained, where

in step F) the isopropyl valerate-containing product mixture e2 is reacted in an ester cleavage reactor with water to give a product mixture f comprising isopropanol, valeric acid, isopropyl valerate and water, and

in step G), the product mixture f in a first distillation column (1) is separated into a stream gl consisting essentially of valeric acid, isopropyl valerate and water and a stream g2 essentially consisting of isopropanol and water. In a second distillation column (2), the stream g1 is converted into a stream h1 (which is generally obtained as a bottom draw stream) essentially consisting of valeric acid and a stream h2 (which is generally used as topcoat stream). ström) is essentially separated consisting of isopropyl valerate and water. The stream h1 can be recycled to the esterification zone or to the optional adsorption step to separate the alkane / alkene mixture. The stream h2 is returned directly to the hydrolysis reactor.

For example, in a third specific embodiment of the process according to the invention, the acid used as alkanoic acid (IV) is butyric acid, and a product mixture e2 containing isopropyl butyrate is obtained, where

in step F) the isopropyl butyrate-containing product mixture e2 is reacted in an ester cleavage reactor with water to give a product mixture f containing isopropanol, butyric acid, isopropyl butyrate and water, and

in step G), the product mixture f in a first distillation column (1) is separated into a stream gl consisting essentially of butyric acid, isopropyl butyrate and water and a stream g2 consisting essentially of isopropanol and water. The stream gl is converted in a second distillation column (2) into a stream h1 (which is generally obtained as a bottom draw stream) consisting essentially of butyric acid and a stream h2 (which is generally obtained as a top draw stream) consisting essentially of isopropyl butyrate and Water separated. The stream h1 can be recycled to the esterification zone or to the optional adsorption step to separate the alkane / alkene mixture. The stream h2 is returned directly to the hydrolysis reactor.

The water contained in stream g2 can be separated off by azeotropic distillation using an entraining agent such as benzene, cyclohexane or diisopropyl ether using an azeotrope distillation column (3), a phase separator (5) and the water separation column (4) and recycled as stream k1 into the ester cleavage reactor become. If diisopropyl ether is obtained as a by-product in the ester cleavage, preference is given to using diisopropyl ether as the entraining agent. The bottom draw stream i1 of the azeotropic distillation column (3) is pure isopropanol.

Compared with the use of formic and acetic acids as carboxylic acids in the esterification step D), the inventive use of higher alkanoic acids having 3 or more carbon atoms, for example the use of isovaleric acid, offers significant advantages. Thus, no alkyl ester / isopropanol azeotrope and no alkyl ester / isopropanol / water azeotrope is formed. The alkyl ester can be separated from isopropanol by a simple and "sharp" method, and most of the water contained in the product mixture f from the ester cleavage zone can be mixed with the alkyl ester by means of the first and the second distillation column (1). and (2) separated and recycled to the ester cleavage zone. As a result, the water load of the downstream azeotropic distillation for separating the alkanol / water mixture is significantly reduced.

The ester cleavage is generally carried out at a pressure of 1 to 20 bar, preferably 2 to 5 bar, and a temperature of generally 50 to 150 ⁰ C, preferably 80 to 120 ⁰ C, performed. Preference is given to working in the presence of an acidic ion exchanger as a catalyst. Water can be used in stoichiometric excess or in excess, generally in stoichiometric excess based on alkyl ester (V). For example, it is used in amounts of from 0.5 to 0.9 mol per mole of isopropyl isovalerate. The ester cleavage reaction can be carried out in a fluidized bed reactor in suspended bed mode or in a fixed bed reactor.

The product mixture f obtained in the ester cleavage reaction contains, for example, 5 to 15% by weight of isopropanol, 10 to 40% by weight of isovaleric acid (or valeric acid or butyric acid), 45 to 80% by weight of isopropyl isovalerate (or isopropyl valerate or isopropyl butyrate) and 1 to 10 wt .-% water. The stream gl, which is generally obtained as the bottom draw stream of the first column (1), contains, for example, 32% by weight of isovaleric acid (or valeric acid or butyric acid) and 64% by weight of isopropyl isovalerate (or isopropyl valerate or isopropyl butyrate) and 4% by weight of water. The stream g2, which is generally obtained as the top draw stream of the first column (1), contains, for example, 85% by weight of isopropanol and 15% by weight of water. The stream h1, which is generally obtained as the bottom draw stream of the second column (2), contains 99% by weight of isovaleric acid (or valeric acid or butyric acid) and 1% by weight of isopropyl isovalerate (or isopropyl valerate or isopropyl butyrate) and the stream h2 which is generally obtained as a top draw stream of the second column (2) contains, for example, 93% by weight of isopropyl isovalerate (or isopropyl valerate or isopropyl butyrate), 2% by weight of isovaleric acid (or valeric acid or butyric acid) and 5% by weight Water.

The stream i1, which is generally obtained as the bottom draw stream of the third column (3), preferably contains at least 98% by weight of isopropanol. The stream i2, which is generally obtained as the top draw stream of the third column (3), contains, for example, 25% by weight of isopropanol, 7% by weight of water and 68% by weight of the entraining agent, for example cyclohexane.

In the phase separator (5), the current i2 is separated by adding an entrainer, for example cyclohexane, into an aqueous and an organic phase. The aqueous phase is added as reflux to the fourth column (4) and contains preferably at least 60% by weight of water, up to 40% by weight of isopropanol and small amounts of the entraining agent. The organic phase is added as reflux to the third column (3) and preferably contains 20 to 50% by weight of isopropanol, less than 10% by weight of water and 50 to 80% by weight of the entraining agent. The bottom draw stream k1 of the fourth column (4) preferably contains more than 98% by weight of water and is recycled to the ester cleavage zone.

The first column (1) generally has 20 to 40 theoretical plates and is operated at a pressure of 0.5 to 2 bar. The second column (2) generally has 15 to 35 theoretical plates and is operated at a pressure of 0.5 to 1.5 bar. The third column (3) generally has 10 to 30 theoretical plates and is operated at a pressure of 0.5 to 4 bar. The fourth column (4) generally has 1 to 15 theoretical plates and is operated at a pressure of 0.5 to 2 bar.

However, ester cleavage and distillation can also be carried out-at least in part-simultaneously in one and the same process step, wherein the mixture formed by ester cleavage is at least partially separated by distillation and the alkanoic acid (IV), optionally together with the homogeneous catalyst, is recovered. The reactive distillation column may be preceded by a pre-reactor.

This reactive distillation can be carried out homogeneously or heterogeneously catalyzed. The reactive distillation column may contain conventional internals such as packings, packing and trays. Heterogeneous catalysts can be present in the form of catalytic units, for example as catalyst packings or catalyst packings, or in the form of a catalyst suspension.

The invention is further illustrated by the following examples.

Examples

example 1

Using the commercial simulation program AspenPlus, an integrated overall process for the production of isopropanol from propane was calculated by way of example. The method in this example comprises steps (1) to (5) as described above. These are shown in Figure 1 with the associated streams. Isovaleric acid is used as the alkanoic acid. The calculated compositions of the individual streams can be found in Table 1. The simulation calculation based on a thermodynamic model based on measurements of vapor-liquid and liquid-liquid phase equilibria.

Example 2

The procedure according to Example 1 was simulated, but instead of isovaleric acid, butyric acid was assumed to be alkanoic acid (IV). The results are summarized in Table 2.

Table 1

O

Table 2

10

Claims

claims

1 . Process for the preparation of 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, comprising the steps of:

A) a feed gas stream a containing the alkane (II) is provided;

B) the feed gas stream a containing the alkane (II) is fed into a dehydrogenation zone and the alkane (II) is subjected to dehydrogenation to the alkene (III), a product gas stream b comprising the alkene (III), unreacted alkane (II ), optionally high boilers, water vapor, hydrogen and low boilers is obtained;

C) the product gas stream b is at least compressed, optionally the product gas stream b into an aqueous phase c1, one alkene (III), the alkane

(II) and optionally high boilers containing phase c2 and a gas phase c3 containing hydrogen and low boilers separated;

D) the product gas stream b or the alkene (III) and the alkane (II) containing phase c2 is reacted in an esterification zone with an alkanoic acid (IV) having 3 or more carbon atoms, wherein a product mixture d comprising the corresponding alkyl ester (V ) of the alkanoic acid (IV) and the unreacted alkane (II) is obtained;

E) 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 (V) is obtained;

F) 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) and the alkanoic acid (IV);

G) from the product mixture f, the alkanol (I) and the alkanoic acid (IV) are separated, and the alkanoic acid (IV) is optionally recycled to the esterification zone.

2. The method according to claim 1, characterized in that the dehydrogenation in step B) is carried out in the presence of oxygen.

3. The method according to claim 1, characterized in that the dehydrogenation is carried out in the presence of water vapor.

4. The method according to claim 1, characterized in that step F) is carried out in a reactive distillation column.

5. The method according to claim 1, characterized in that step F) is carried out at least partially in an ester cleavage pre-reactor and at least partially in a reactive distillation column.

6. The method according to claim 1, characterized in that as alkane (II) propane is used, wherein in step E) an isopropyl ester-containing product mixture e2 is obtained, and

in step F) the isopropyl ester-containing product mixture e2 is reacted in an ester cleavage reactor with water to give a product mixture f containing isopropanol, alkanoic acid (IV), isopropyl ester and water, and

in step G) the product mixture f in a first distillation column into a stream gl consisting essentially of alkanoic acid (IV), isopropyl ester and

Water and a stream g2 essentially consisting of isopropanol and water is separated, the stream gl is separated in a second distillation column into a stream h1 consisting essentially of alkanoic acid (IV) and a stream h2 essentially consisting of isopropyl ester and water, the in Stream g2 contained water is separated by azeotropic distillation using an entraining agent by means of an azeotropic distillation, a phase separator and a water separation column, and substantially pure isopropanol is recovered as the bottom draw stream i1 of the azeotropic distillation.

7. The method according to any one of claims 1 to 6, characterized in that as alkanoic acid (IV) butyric acid, valeric acid or isovaleric acid is used.

8. The method according to claim 7, characterized in that as alkanoic acid (IV) isovaleric acid is used, wherein a isopropyl isovalerate containing product mixture e2 is obtained,

in step F) the isopropyl isovalerate-containing product mixture e2 is reacted in an ester-cleavage reactor with water to give a product mixture f comprising isopropanol, isovaleric acid, isopropyl isovalerate and water, and

in step G) the product mixture f in a first distillation column into a

Stream gl consisting essentially of isovaleric acid, isopropyl isovalerate and water and a stream g2 essentially consisting of isopropanol and water is separated, the stream gl is separated in a second distillation column into a stream h1 consisting essentially of isovaleric acid and a stream h2 consisting essentially of isopropyl isovalerate and water, the water contained in stream g2 by azeotropic distillation using an entraining agent by means of an azeotropic distillation , a phase separator and a water separation column is separated, and substantially pure isopropanol is recovered as the bottom draw stream i1 of the azeotropic distillation.