EP3497073A1

EP3497073A1 - Method of starting up a reactor for the oxidative dehydrogenation of n-butenes

Info

Publication number: EP3497073A1
Application number: EP17754659.5A
Authority: EP
Inventors: Jan Ungelenk; Oliver HAMMEN; Ulrich Hammon; Rainer ECKRICH; Signe Unverricht; Christian Walsdorff; Heinz Boelt; Hendrik Reyneke; Christine TOEGEL; Anton Wellenhofer; Ulrike Wenning
Original assignee: BASF SE; Linde GmbH
Current assignee: BASF SE; Linde GmbH
Priority date: 2016-08-09
Filing date: 2017-08-08
Publication date: 2019-06-19
Also published as: KR20190046848A; WO2018029215A1; JP2019528264A; CN109843836A; US20190337870A1

Abstract

Process for preparing butadiene from n-butenes, which has a start-up phase and an operating phase and the operating phase of the process comprises the steps: A) provision of a feed gas stream (a1) comprising n-butenes; B) introduction of the feed gas stream (a1) comprising n-butenes, an oxygen-comprising gas stream (a2) and an oxygen-comprising recycle gas stream (d2) into at least one oxidative dehydrogenation zone and oxidative dehydrogenation of n-butenes to butadiene, giving a product gas stream (b) comprising butadiene; C) cooling and compression of the product gas stream (b), giving at least one aqueous condensate stream (c1) and a gas stream (c2) comprising butadiene; D) introduction of the gas stream (c2) into an absorption zone and separation of incondensable and low-boiling gas constituents as gas stream (d) from the gas stream (c2) by absorption of the C₄ hydrocarbons in an absorption medium, giving an absorption medium stream loaded with C₄ hydrocarbons and the gas stream (d), and recirculation of the gas stream (d) as recycle gas stream (d2) to the oxidative dehydrogenation zone, where the start-up phase comprises the steps, in the order i) to iv): i) introduction of a gas stream (d2') having a composition corresponding to the recycle gas stream (d2) in the operating phase into the dehydrogenation zone and setting of the recycle gas stream (d2) to at least 70% of the total volume flow in the operating phase; ii) optionally additional introduction of a steam stream (a3) into the dehydrogenation zone; iii) additional introduction of the feed gas stream (a1) comprising butenes at a lower volume flow than in the operating phase and raising of this volume flow until at least 50% of the volume flow of the feed gas stream (a1) in the operating phase has been attained, with the total gas flow through the dehydrogenation zone corresponding to not more than 120% of the total gas flow during the operating phase; iv) additional introduction, when at least 50% of the volume flow of the feed gas stream (a1) comprising butenes in the operating phase has been attained, of an oxygen-comprising stream (a2) at a lower volume flow than in the operating phase and raising of the volume flows of the feed gas streams a1 and a2 until the volume flows in the operating phase have been attained, with the total gas flow through the dehydrogenation zone corresponding to not more than 120% of the total gas flow during the operating phase.

Description

Process for starting up a reactor for the oxidative dehydrogenation of n-butenes

description

The invention relates to a process for starting up a reactor for producing 1,3-butadiene from n-butenes by oxidative dehydrogenation (ODH). Butadiene is an important basic chemical and is used for example for the production of synthetic rubbers (butadiene homopolymers, styrene-butadiene rubber or nitrile rubber) or for the production of thermoplastic terpolymers (acrylonitrile-butadiene-styrene copolymers). Butadiene is further converted to sulfolane, chloroprene and 1, 4-hexamethylenediamine (over 1, 4-dichlorobutene and adiponitrile). By dimerization of butadiene, vinylcyclohexene can also be produced, which can be dehydrogenated to styrene.

Butadiene can be prepared by thermal cracking (steam cracking) of saturated hydrocarbons, usually starting from naphtha as the raw material. Steam cracking of naphtha produces a hydrocarbon mixture of methane, ethane, ethene, acetylene, propane, propene, propyne, allenes, butanes, butenes, butadiene, butynes, methylalls, Cs and higher hydrocarbons.

Butadiene can also be obtained by oxidative dehydrogenation of n-butenes (1-butene and / or 2-butene). As input gas for the oxidative dehydrogenation (oxydehydrogenation, ODH) of n-butenes to butadiene, any mixture containing n-butenes can be used. For example, a fraction containing n-butenes (1-butene and / or 2-butene) as a main component and obtained from the C ₄ fraction of a naphtha cracker by separating butadiene and isobutene can be used. Furthermore, gas mixtures which comprise 1-butene, cis-2-butene, trans-2-butene or mixtures thereof and which have been obtained by dimerization of ethylene can also be used as input gas. Further, n-butenes containing gas mixtures obtained by catalytic fluid cracking (FCC) can be used as the input gas. The reaction of the gas streams containing butenes is generally carried out industrially in tube bundle reactors which are operated in a salt bath as a heat carrier. The product gas stream is cooled behind the reactor by direct contact with a coolant in a quenching stage and then compressed. Then, the C ₄ components are absorbed in an absorption column in an organic solvent. Inert gases, low boilers, CO, CO2 and others leave the column overhead. This top stream is partly called

Circulating gas fed to the ODH reactor. Hydrocarbons and oxygen can create an explosive atmosphere. The concentration of combustible gas constituents (mainly hydrocarbons and CO) may be reduced to avoid ignitable mixtures. below the lower explosion limit (LEL) or above the upper explosion limit (LEL). Below the lower explosion limit, the oxygen concentration can be freely selected without the formation of an explosive gas mixture. However, then the concentration of input gas is low, which is economically unfavorable. Therefore, reaction with a reaction gas mixture above the upper explosion limit is preferred. Here it depends on the oxygen concentration, whether it can come to an explosion. Below a certain oxygen concentration, the LOC (limiting oxygen concen- tration), the concentration of flammable gas constituents can be freely selected without the formation of an explosive gas mixture. Both LEL, OEG and LOC are temperature and pressure dependent.

On the other hand, coke precursors can be formed depending on the oxygen concentration in the oxidative dehydrogenation of n-butenes to butadiene, which can eventually lead to coking, deactivation and irreversible destruction of the multimetal oxide catalyst. This is still possible if the oxygen concentration in the reaction gas mixture of the oxydehydrogenation at the entrance of the reactor is above the LOC.

The need for an excess of oxygen for such catalyst systems is well known and manifests itself in the process conditions when using such catalysts. Representing the recent work of Jung et al. (Catalan Surv. Asia 2009, 13, 78-93, DOI 10.1007 / s10563-009-9069-5 and Applied Catalysis A: General 2007, 317, 244-249, DOI 10.1016 / j.apcata.2006.10.021) ,

However, the presence of high oxygen concentrations in addition to hydrocarbons such as butane, butene and butadiene or the organic absorbents used in the work-up part is fraught with risks. So can form explosive gas mixtures. If you only work with a small distance to the potentially explosive area, it is not always technically possible to prevent this area from entering due to fluctuations in the process parameters. Particularly critical in terms of risk of explosion and coking of the catalyst is the period in which the reactor is started up and flowed through with reaction gas mixture.

Processes for the oxidative dehydrogenation of butenes to butadiene are known in principle. For example, US 2012 / 0130137A1 describes such a process using catalysts containing oxides of molybdenum, bismuth and, as a rule, other metals. For the sustained activity of such catalysts for the oxidative dehydrogenation, a critical minimum oxygen partial pressure in the gas atmosphere is required in order to avoid too extensive a reduction and thus a loss of performance of the catalysts. For this reason, it is generally also not possible to work with a stoichiometric oxygen feed or complete oxygen conversion in the oxydehydrogenation reactor (ODH reactor). For example, US 2012/0130137 A1 describes an oxygen content of 2.5 to 8% by volume in the product gas. The problem of the formation of any explosive mixtures after the reaction step is discussed in paragraph [0017]. In particular, it is pointed out that in a so-called "rich" mode of operation above the upper explosive limit in the reaction part, the problem is that after absorption of a large part of the organic constituents in the workup, the gas composition crosses the explosive region in a transition from rich to lean gas mixture it is stated in Sections 0061-0062 that according to the invention it is necessary for the concentration of combustible gas constituents in the gas mixture introduced into the oxidative dehydrogenation reactor to be above the upper explosive limit, and for starting the oxidative dehydrogenation reaction first to equal the oxygen concentration in the mixed gas at the reactor inlet Value is set below the oxygen limit concentration (LOC) by first adjusting the amount of oxygen-containing gas and water vapor passed into the reactor, and then introducing flammable gas (essentially outlet sgas) is started. Subsequently, the introduction amount of oxygen-containing gas, for example, air, and combustible gas can be increased, so that the concentration of combustible gas components in the mixed gas becomes larger than the upper explosion limit. On the other hand, as the introduction amount of combustible gas components and oxygen-containing gas increases, the introduction amount of nitrogen and / or water vapor is lowered to make the introduction amount of mixed gas stable.

On the other hand, it is also pointed out that there is a risk of catalyst deactivation by coking in the case of a consistently lean mode of operation in the reaction part. However, US 2012 / 0130137A1 does not provide a solution to this problem.

Incidentally, paragraph [0106] mentions how the occurrence of explosive atmospheres in the absorption step can be avoided, for example, by diluting the gas stream with nitrogen before the absorption step. In the detailed description of the absorption step in paragraphs [0132] ff, the problem of the formation of explosive gas mixtures will not be discussed further.

The document does not describe which conditions must be met in order to prevent coking of the catalyst. Furthermore, the document does not relate to a process with circulating gas driving. Furthermore, the currents are set in succession, which means a high cost of operation.

JP2016-69352 also discusses a process for the oxidative dehydrogenation of butenes to butadiene and describes the problem that there is a risk of catalyst deactivation by coking in a low-oxygen mode of operation in the reaction part. At the same time, attention is drawn to the problem that the concentrations for fuel gas and oxygen can not be chosen arbitrarily, given the formation of any explosive mixtures. Paragraphs [0045-0046], in particular, describe that the LOC increases with increasing proportions of carbon dioxide when a mixture of nitrogen and carbon is used as the inert gas. is selected. However, this approach is economically unfavorable, since the provision of inert gases is expensive and must be provided with carbon dioxide, in addition to nitrogen, another inert gas. JP2010-280653 describes starting an ODH reactor. The reactor should be started without catalyst deactivation or an increase in pressure loss. This is to be achieved by driving the reactor to more than 80% of full load within 100 hours. In section 0026 it is stated that, according to the invention, at the start of the reaction less than 100 hours after starting supply of the raw material gas reactor, the supply quantity per unit time of the raw material gas reactor is set to more than 80% of the maximum allowable supply quantity, and during this time the supply quantity of the raw material is controlled together with the raw material gas into the reactor nitrogen gas, the elemental oxygen-containing gas and water vapor so that the composition of the mixed gas of raw material gas, nitrogen gas, elemental oxygen-containing gas and water vapor does not get into the explosion area. The document does not describe which conditions must be met in order to prevent coking of the catalyst. Furthermore, the document does not relate to a process with circulating gas driving. Furthermore, the document does not consider the problem of explosion in the processing part of the process.

EP 1 180 508 describes the start-up of a reactor for the catalytic gas-phase oxidation. Specifically described is the oxidation of propylene to acrolein. A process is described in which, when starting up the reactor, an area is run through in which the oxygen content in the reaction gas mixture is greater than the LOC and the concentration of combustible gas constituents is less than the LEL. In the stationary operating state, the O 2 concentration is then lower than the LOC, and the concentration of combustible gas components is greater than the OEG.

DE 1 0232 482 describes a method for the safe operation of an oxidation reactor for the gas-phase partial oxidation of propylene to acrolein and / or acrylic acid with a computer-aided shutdown mechanism. This is based on the deposition of an explosion diagram and the concentration determination of C ₄ and O 2 by measuring the O 2 and C ₃ hydrocarbon concentration in the recycle gas and the volume flows of recycle gas, C 3 hydrocarbon stream and oxygen-containing gas. Sections 0076-0079 describe the starting of the reactor. It is stated in section 0079 that the release for opening the supply of first air and then propene is issued only when the inflow of the diluent gas (water vapor and / or recycle gas) has increased to a minimum value, for example, 70% of the maximum possible air supply , The concentration of O2 in the recycle gas during the start-up process is already identical to stationary operation (3.3% by volume). WO2015 / 104397 discloses a process for producing butadiene from n-butenes having a start-up phase and an operating phase, the process comprising, in the operating phase, the steps of: A) providing a feed gas stream a1 containing n-butenes;

B) feed of the n-butenes containing feed gas stream a1, an oxygen-containing gas stream d2 and an oxygen-containing cycle gas stream d2 in at least one oxidative dehydrogenation and oxidative dehydrogenation of n-butenes to butadiene, wherein a product gas stream b containing butadiene, unreacted n-butenes, Water vapor, oxygen, low-boiling hydrocarbons, high-boiling secondary components, optionally carbon oxides and optionally inert gases is obtained;

C) cooling and compression of the product gas stream b and condensation of at least a portion of the high-boiling secondary components, at least one aqueous condensate stream c1 and a gas stream c2 containing butadiene, n-butenes, water vapor, oxygen, low-boiling hydrocarbons, optionally carbon oxides and optionally inert gases is obtained; D) feeding the gas stream c2 in an absorption zone and separation of non-condensable and low-boiling gas components comprising oxygen, low-boiling hydrocarbons, optionally carbon oxides and optionally inert gases as gas stream d from the gas stream c2 by extensive absorption of C4 hydrocarbons comprising butadiene and n-butenes in an absorbent to obtain an absorbent stream laden with C4 hydrocarbons and the gas stream d, and recycle, optionally after separating a purge gas stream p, the gas stream d as a recycle gas stream d2 into the oxidative dehydrogenation zone; wherein the start-up phase in the order i) to iv) comprises the steps of: i) feeding an oxygen-containing gas stream and an inert gas stream into the dehydrogenation zone in such a ratio that the oxygen content of the circulating gas stream d2 is 30 to 80% of the oxygen content of the circulating gas stream d2 in the operating phase corresponds; ii) adjusting the circulating gas flow d2 to at least 70% of the volume flow of the circulating gas d2 in the operating phase; iii) optionally feeding, at an initial oxygen content of the cycle gas stream d2, from 30 to 80% of the oxygen content of the cycle gas stream d2 in the operating phase, a water vapor stream a3 into the dehydrogenation zone; iv) feeding, at an initial oxygen content of the cycle gas stream d2, from 30 to 80% of the oxygen content of the cycle gas stream d2 in the operating phase, an oxygen-containing gas stream a2 'and a butene-containing feed gas stream a1' with lower volume flows than in the operating phase in a ratio k = a2 '/ a1', and increase the volume flows of the gas streams a1 'and a2' until reaching the volume flows of the gas streams a1 and a2 in the operating phase, wherein the circulating gas flow d2 is at least 70% and at most 120% of the volume flow in the operating phase.

According to WO2015 / 104397, the high oxygen content of the cycle gas is thus initially mixed with an inert gas, e.g. Nitrogen, so far diluted that the oxygen content of the cycle gas preferably corresponds to 50 to 60% of the oxygen content in the operating phase. Based on this, the oxygen content and the content of butenes are increased so that a sufficient distance from the explosion limit is always maintained. The oxygen content is thereby increased by adding an oxygen-containing gas, preferably air. However, this approach is economically unfavorable, since the oxygen content of the available cycle gas is first lowered by dilution with an additional inert gas, to be subsequently raised again by adding an additional oxygen-containing gas. It should be noted that the provision of inert gas is expensive and also the addition of oxygen-containing gas, e.g. Air, costs caused because the air must first be compressed to the required process pressure. The object of the invention is to provide a safe and economical process for starting a reactor for the oxidative dehydrogenation of n-butenes to butadiene and of downstream units for working up the product gas mixture.

The object is achieved by a process for the preparation of butadiene from n-butenes with a start-up phase and an operating phase, the process comprising the steps in the operating phase:

A) providing a n-butenes containing feed gas stream a1; B) feed of the n-butenes containing feed gas stream a1, an oxygen-containing gas stream a2 and an oxygen-containing cycle gas stream d2 in at least one oxidative dehydrogenation and oxidative dehydrogenation of n-butenes to butadiene, wherein a product gas stream b containing butadiene, unreacted n-butenes, water vapor, Oxygen, low-boiling hydrocarbons, high-boiling secondary components, optionally carbon oxides and optionally inert gases is obtained;

C) cooling and compression of the product gas stream b and condensation of at least a portion of the high-boiling secondary components, at least one aqueous condensate stream c1 and a gas stream c2 containing butadiene, n-butenes, water vapor, oxygen, easily boiling hydrocarbons, optionally carbon oxides and optionally inert gases becomes; D) feeding the gas stream c2 into an absorption zone and separating off non-condensable and low-boiling gas constituents comprising oxygen, low-boiling hydrocarbons, optionally carbon oxides and optionally inert gases as gas stream d from the gas stream c2 by extensive absorption of the C4 hydrocarbons comprising butadiene and n Butenes in an absorbent to give an absorbent stream laden with C4 hydrocarbons and the gas stream d, and recycling, optionally after separating a purge gas stream p, the gas stream d as a circulating gas stream d2 into the oxidative dehydrogenation zone; wherein the start-up phase in the order i) to iv) comprises the steps of: i) feeding a gas stream d2 'having a composition corresponding to the cycle gas stream d2 in the operating phase into the dehydrogenation zone and setting the cycle gas stream d2 to at least 70% of the total volume flow in the operating phase; ii) optionally additionally feeding a steam stream a3 into the dehydrogenation zone; iii) additionally feeding the butene-containing feed gas stream a1 at a lower volume flow than in the operating phase and increasing this volume flow until reaching at least 50% of the volume flow of the feed gas stream a1 in the operating phase, the total gas flow through the dehydrogenation zone being at most 120% of the total gas flow during corresponds to the operating phase; iv) additional feeding, upon reaching at least 50% of the volume flow of the Buene-containing feed gas stream a1 in the operating phase, an oxygen-containing stream a2 with a lower volume flow than in the operating phase, and increasing the volume flows of the feed gas flows a1 and a2 until reaching the volume flows in the operating phase, wherein the total gas flow through the dehydrogenation zone corresponds to a maximum of 120% of the total gas flow during the operating phase.

The starting method according to the invention has the advantage over the procedure described in WO 2015/104397 that oxygen-rich conditions are also present at the start of the starting phase, since the circulating gas is not diluted with an inert gas. This counteracts coking of the catalyst. The conditions during the start-up phase in terms of load, gas velocity, residence time and composition of the circulating gas flow correspond more to the conditions during the operating phase. Since the gas velocity is essentially constant, a developing hot spot does not migrate within the reactor. The stationary operating state and the optimal performance of the catalyst in terms of space-time yield and selectivity are achieved faster overall.

Preferably, in step i) the circulating gas stream d2 is adjusted to 90 to 1 10% of the total volume flow in the operating phase. The total volume flow is the sum of the volume flows a1, a2, d2 and optionally a3. In a particularly preferred embodiment the circulating gas flow d2 is set to 95-105% of the total volume flow in the operating phase, more preferably the circulating gas flow d2 is set to 100% of the total volume flow in the operating phase. The set circulating gas flow d2 is reduced in the subsequent steps iii) and iv) so that the total gas flow through the Dehnernerzone, that is, the sum of the streams a1, a2, d2 and optionally a3, during the further start-up phase at least 70% and at most 120 %, preferably at least 90% and at most 1 15% of the total gas flow during the operating phase. Preferably, the total gas flow during the start-up phase remains essentially constant and varies by a maximum of +/- 10% by volume, in particular +/- 5% by volume, ie during the start-up phase preferably amounts to 90 to 110% by volume, in particular 95% to 105% by volume of the total gas flow during the operating phase.

In step i), a gas stream d2 'with a composition corresponding to the circulating gas stream d2 in the operating phase is fed into the dehydrogenation zone. Preferably, as the gas stream d2 ', part of the circulating gas stream is taken from one or more reactors operated in parallel to produce butadiene from n-butenes which are in the operating phase. In this case, the dehydrogenation zone of step B) thus comprises a plurality of reactors, at least one of the reactors, preferably at least two of the reactors, being in the operating phase and generating a total cycle gas stream from which the gas stream d2 ^' or d2 is withdrawn. In general, the gas stream d2 'and correspondingly the circulating gas stream d2 in the operating phase contain 5.5 to 8.5% by volume of O 2 and 89.4 to 94.5% by volume of inert gases selected from nitrogen, noble gases (in particular argon) and carbon oxides (CO, CO2). In addition, the gas stream d2 'and correspondingly the circulating gas stream d2 may contain 0 to 0.5% by volume of water vapor. In addition, the circulating gas stream d2 and corresponding to the

Gas stream d2 'still 0 to 1, 5 vol .-% oxygenates such as acrolein and 0 to 0.1 wt .-% hydrocarbons.

A variant with several reactors is shown in FIG. Mean in it

R ¹ , R ² , R ⁿ n reactors operated in parallel, of which R ¹ are in the start-up phase and R ² , R ⁿ are in the operating phase,

a1 ¹ , a2 ¹ , a3 ¹ , a1 ² , a2 ² , a3 ² , a1 ⁿ , a2 ⁿ , a3 ⁿ the material flows a1, a2, a3 assigned to the individual reactors,

d2 _to tai the total cycle gas flow,

d2 ¹ , d2 ² , d2 ⁿ the partial cycle gas streams assigned to the individual reactors,

b the butadiene-containing C4 product gas stream,

Q is a quenching step,

K is a compression level,

c1 an aqueous condensate stream,

c2 the butadiene-containing C4 product gas stream,

A is the absorption level,

d1 the butadiene-containing C4 product gas stream,

d the total cycle gas stream before separation of a purge stream, and

p the purge stream. The gas stream d2 'has a composition corresponding to the circulating gas stream d2 in the operating phase when its oxygen content deviates at most +/- 2% by volume from the oxygen content of the circulating gas stream d2 in the stationary operating phase.

In step ii), in addition, a steam stream a3 can be fed into the dehydrogenation zone. In general, the amount of water vapor in the dehydrogenation zone during steps ii) to iv) is 0.5 to 10% by volume, preferably 1 to 7% by volume. This can also be humidity.

In step iii), the butene-containing feed gas stream a1 is additionally fed to the dehydrogenation zone until at least 50% of the volume flow in the operating phase has been reached. The volume flow is generally increased in stages, for example, starting with 10% of the volume flow in the operating state, in steps of 10% until reaching at least 50% of the volume flow in the operating state. The volume flow can also be increased in the form of a ramp. If appropriate, the circulating gas stream d2 is correspondingly reduced so far that the total gas flow through the dehydrogenation zone corresponds to a maximum of 120% of the total gas flow during the operating phase. The content of C4 hydrocarbons (butenes and butanes) in the total gas flow through the dehydrogenation zone is generally from 7 to 9% by volume at the end of step (iii).

The volume flow of butene-containing feed gas stream a1 can be raised in step iii) until reaching at least 60% of the volume flow in the operating phase, but at most until reaching a maximum of 75% of the volume flow in the operating phase.

In step iv), when at least 50% and at most 75% of the volume flow of the butenes containing feed gas stream a1 are reached in the operating phase, in addition to the butenes containing feed gas stream a1 an oxygen-containing stream a2 with lower flow than in the operating phase fed to the dehydrogenation zone and the volume flows of the feed gas flows a1 and a2 raised to reach the volume flows in the operating phase. In general, in a first step, the volume flow of the oxygen-containing gas stream a2 is increased in one or more stages until a ratio of oxygen to hydrocarbons corresponding to the ratio of oxygen to hydrocarbons in the operating phase is reached, and then both volume flows a1 and a2 become in stages in the operating phase, the ratio of oxygen to hydrocarbons remains substantially constant and corresponds to the ratio of oxygen to hydrocarbons in the operating phase. In this case, the volume flow is generally increased in steps, for example beginning with 50% of the volume flow of the gas flow a1 in the operating state, in steps of, for example, 10%, wherein the stages for the steps for increasing the volume flow of the gas flow a2 are chosen such that the ratio Oxygen to hydrocarbon During the start-up phase remains substantially constant until reaching 100% of the volume flows in the operating state.

The ratio of oxygen to hydrocarbons in the start-up phase corresponds to the ratio of oxygen to hydrocarbons in the operating phase if it does not deviate more than 10% from this ratio. The ratio of oxygen to hydrocarbons in the operating phase is generally from 0.65: 1 to 1.5: 1, preferably 0.65: 1 to 1 at an n-butenes content of 50 to 100% by volume in the feed gas stream a1 , 3: 1. The ratio may change during the operating phase.

The content of C4 hydrocarbons (butenes and butanes) in the total gas flow through the dehydrogenation zone at the end of step (iv) is generally 7 to 9 vol.%, The oxygen content generally 12 to 13 vol.%. In general, the pressure in the dehydrogenation zone during the start-up phase is 1 to 5 bar absolute, preferably 1.05 to 2.5 bar absolute.

In general, the pressure in the absorption zone during the start-up phase is 2 to 20 bar, preferably 5 to 15 bar.

In general, the temperature of the heat exchange medium during the start-up phase is between 220 to 490 ° C and preferably between 300 to 450 ° C and more preferably between 330 and 420 ° C. In general, the duration of the start-up phase is from 15 to 2000 minutes, preferably from 15 to 500 minutes and more preferably from 20 to 120 minutes. Afterwards the operating phase begins.

In general, step C) comprises steps Ca) and Cb):

Ca) cooling the product gas stream b in at least one cooling stage, wherein the cooling takes place in at least one cooling stage by contacting with a coolant, and condensation of at least a portion of the high-boiling secondary components; Cb) compression of the remaining product gas stream b in at least one compression stage, wherein at least one aqueous condensate stream c1 and a gas stream c2 containing butadiene, n-butenes, water vapor, oxygen, low-boiling hydrocarbons, optionally carbon oxides and optionally inert gases is obtained. In general, step D) comprises steps Da) and Db):

Da) separation of non-condensable and low-boiling gas constituents comprising oxygen, low-boiling hydrocarbons, optionally carbon oxides and, where appropriate, optionally inert gases as gas stream d from the gas stream c2 by absorption of the C ₄ -Kohlen- hydrogens containing butadiene and n-butenes in an absorbent, wherein a charged with C ₄ - hydrocarbons absorbent stream and the gas stream d are obtained, and

Db) subsequent desorption of the C ₄ hydrocarbons from the loaded absorbent stream, whereby a C ₄ product gas stream d1 is obtained.

Preferably, the steps E) and F) are subsequently carried out:

E) separation of the C ₄ product stream d1 by extractive distillation with a solvent which is selective for butadiene into a butadiene and the stream e1 containing the selective solvent and a stream e2 containing n-butenes; F) Distillation of the butadiene and the selective solvent-containing material stream f2 in a substantially consisting of the selective solvent stream gl and a butadiene-containing stream g2.

In general, the gas stream d contained in step Da) is recirculated to at least 10%, preferably at least 30%, as recycle gas stream d2 in step B).

In general, in the cooling step Ca), aqueous coolants or organic solvents or mixtures thereof are used. Preferably, in the cooling stage Ca), an organic solvent is used. These generally have a much higher solvent power for the high-boiling by-products, which can lead to deposits and blockages in the downstream of the ODH reactor plant components, as water or alkaline aqueous solutions. Preferred organic solvents used as coolants are aromatic hydrocarbons, for example toluene, o-xylene, m-xylene, p-xylene, diethylbenzenes, triethylbenzenes, diisopropylbenzenes, triisopropylbenzenes and mesitylene or mixtures thereof. Especially preferred is mesylene.

The following embodiments are preferred or particularly preferred variants of the method according to the invention:

The stage Ca) is carried out in several stages in stages Ca1) to Can), preferably in two stages in two stages Ca1) and Ca2). In this case, particularly preferably at least part of the solvent is fed after passing through the second stage Ca2) as cooling agent of the first stage Ca1).

The stage Cb) generally comprises at least one compression stage Cba) and at least one cooling stage Cbb). In the at least one cooling stage Cbb), the compressed gas in the compression stage Cba) is brought into contact with a cooling agent. More preferably, the cooling agent of the cooling step Cbb) contains the same organic solvent used in step Ca) as a cooling agent. In a particularly preferred variant, at least part of this cooling agent is fed after passing through the at least one cooling stage Cbb) as cooling agent of the stage Ca).

Preferably, the stage Cb) comprises a plurality of compression stages Cba1) to Cban) and cooling stages Cbb1) to Cbbn), for example four compression stages Cba1) to Cba4) and four cooling stages Cbb1) to Cbb4).

Preferably, step D) comprises the steps Da1), Da2) and Db):

Da1) absorption of the C4 hydrocarbons comprising butadiene and n-butenes in a high-boiling absorbent, wherein a C4 hydrocarbons laden absorbent flow and the gas stream d are obtained

Da2) removal of oxygen from the loaded with C4 hydrocarbons absorbent stream from step Da) by stripping with a non-condensable gas stream, and

Db) desorption of the C4 hydrocarbons from the loaded absorbent stream to obtain a C4 product gas stream d1 consisting essentially of C4 hydrocarbons and comprising less than 100 ppm oxygen. Preferably, the high-boiling absorbent used in step Da) is an aromatic hydrocarbon solvent, particularly preferably the aromatic hydrocarbon solvent used in step Ca), in particular mesitylene. It is also possible to use, for example, diethylbenzenes, triethylbenzenes, diisopropylbenzenes and triisopropylbenzenes or mixtures containing these substances.

Embodiments of the method according to the invention are shown in FIG. 1 and will be described in detail below.

As feed gas stream pure n-butenes (1-butene and / or cis- / trans-2-butene), but also containing butene gas mixtures can be used. It is also possible to use a fraction containing n-butenes (1-butene and cis- / trans-2-butene) as the main component and obtained from the C4 fraction of naphtha cracking by separating butadiene and isobutene. Furthermore, gas mixtures which comprise pure 1-butene, cis-2-butene, trans-2-butene or mixtures thereof and which have been obtained by dimerization of ethylene can also be used as input gas. Furthermore, gas mixtures containing n-butenes which have been obtained by fluid catalytic cracking (FCC) can be used as the input gas. In one embodiment of the process according to the invention, the input gas containing n-butenes is obtained by non-oxidative dehydrogenation of n-butane. By coupling a non-oxidative catalytic dehydrogenation with the oxidative dehydrogenation of the n-butenes formed, a high yield of butadiene, based on n-butane used, can be obtained. In the non-oxidative catalytic n-butane dehydrogenation, a gas mixture is obtained which, in addition to butadiene 1-butene, 2-butene and unreacted n-butane, contains minor constituents. Common secondary constituents are hydrogen, water vapor, nitrogen, CO and CO2, methane, ethane, ethene, propane and propene. The composition of the gas mixture leaving the first dehydration zone can vary greatly depending on the mode of operation of the dehydrogenation. Thus, when carrying out the dehydrogenation with the introduction of oxygen and additional hydrogen, the product gas mixture has a comparatively high content of water vapor and carbon oxides. When operating without oxygen feed, the product gas mixture of the non-oxidative dehydrogenation has a comparatively high content of hydrogen.

In step B), the feed gas stream containing n-butenes and an oxygen-containing gas are fed into at least one dehydrogenation zone (one or more ODH reactors R operated in parallel) and the butenes contained in the gas mixture are oxidatively dehydrogenated to form butadiene in the presence of an oxydehydrogenation catalyst.

The molecular oxygen-containing gas generally contains more than 10% by volume, preferably more than 15% by volume, and more preferably more than 20% by volume of molecular oxygen. It is preferably air. The upper limit of the content of molecular oxygen is generally 50% by volume or less, preferably 30% by volume or less, and more preferably 25% by volume or less. In addition, any inert gases may be contained in the molecular oxygen-containing gas. Possible inert gases include nitrogen, argon, neon, helium, CO, CO2 and water. The amount of inert gases for nitrogen is generally 90% by volume or less, preferably 85% by volume or less, and more preferably 80% by volume or less. In the case of components other than nitrogen, it is generally 10% by volume or less, preferably 1% by volume or less.

For carrying out the oxidative dehydrogenation at full conversion of n-butenes, preference is given to a gas mixture which has a molar oxygen: n-butenes ratio of at least 0.5. Preference is given to working at an oxygen: n-butenes ratio of 1.25 to 1.6. To set this value, the input gas stream with oxygen or at least one oxygen-containing gas, such as air, and optionally additional inert gas or steam can be mixed. The resulting oxygen-containing gas mixture is then fed to the oxydehydrogenation. Furthermore, inert gases such as nitrogen and also water (as water vapor) may be contained together in the reaction gas mixture. Nitrogen can be used to adjust the oxygen concentration and to prevent the formation of an explosive gas mixture. The same applies to water vapor. Steam also serves to control the coking of the catalyst and to dissipate the heat of reaction.

Catalysts suitable for oxydehydrogenation are generally based on a Mo-Bi-O-containing multimetal oxide system, which generally additionally contains iron. In general, the catalyst contains other additional components such as potassium, cesium, magnesium, zirconium, chromium, nickel, cobalt, cadmium, tin, lead, germanium, lanthanum, manganese, tungsten, phosphorus, cerium, aluminum or silicon. Iron-containing ferrites have also been proposed as catalysts.

In a preferred embodiment, the multimetal oxide contains cobalt and / or nickel. In a further preferred embodiment, the multimetal oxide contains chromium. In a further preferred embodiment, the multimetal oxide contains manganese. Examples of Mo-Bi-Fe-O-containing multimetal oxides are Mo-Bi-Fe-Cr-O or Mo-Bi-Fe-Zr-O-containing multimetal oxides. Preferred catalysts are described, for example, in

US 4,547,615 (Moi2BiFeo, iNi ₈ ZrCr ₃ Ko, 20x and Moi2BiFeo, iNi ₈ AICr ₃ Ko, 20x), US 4,424,141 (Moi2BiFe ₃ Co4,5Ni2, ₅ Po, 5Ko, iOx + Si0 ₂ ), DE-A 25 30 959 ₍₃ Moi2BiFe Co4,5Ni2,5Cro, 5KO, iO _x, Moi _3, 7 ₅ ₃ BiFe Co4,5Ni2,5Geo, 5KO, 80 _x, Moi2BiFe ₃ Co4,5Ni _2, 5Mno, 5KO, _x and OK

Moi2BiFe Co4 _3, ₅ Ni _2, 5Lao, 5KO, lox), U.S. 3,91 1, 039 ₍₃ Moi2BiFe Co4, ₅ Ni2,5Sno, ₅ Ko, lox), DE-A 25 30 959 and DE-A 24 47 825 (Moi2BiFe ₃ Co4, ₅ Ni2.5Wo, ₅ Ko, iOx).

Suitable multimetal oxides and their preparation are further described in US 4,423,281 (Moi2BiNi ₈ Pbo, ₅ Cr ₃ Ko, 20x and Moi2BibNi ₇ Al ₃ Cro, 5Ko, 50x), US 4,336,409 (Moi2BiNi ₆ Cd2Cr ₃ Po, ₅ Ox), DE-A 26 00 128 (Moi2BiNi ₀ , 5Cr ₃ Po, ₅ Mg7, ₅ Ko, iOx + Si0 ₂ ) and DE-A 24 40 329

(Moi2BiCo4.5Ni ₂ , 5Cr ₃ Po, ₅ Ko, iOx).

Particularly preferred catalytically active, molybdenum and at least one further metal-containing multimetal oxides have the general formula (Ia):

Moi2BiaFebCOcNidCr _e X ¹ fX ² gOy (la), with

X ¹ = Si, Mn and / or Al,

X ² = Li, Na, K, Cs and / or Rb,

0.2 <a <1,

0.5 <b <10,

0 <c <10,

0 <d <10,

2 <c + d <10

0 <e <2,

0 <f <10

0 <g <0.5 y = a number determined on the assumption of charge neutrality by the valence and frequency of the elements other than oxygen in (1a).

Preference is given to catalysts whose catalytically active oxide composition of the two metals Co and Ni has only Co (d = 0). X ¹ is preferably Si and / or Mn and X ² is preferably K, Na and / or Cs, particularly preferably X ² = K. Particular preference is given to a largely Cr (VI) -free catalyst.

The reaction temperature of the oxydehydrogenation is generally controlled by a heat exchange medium located around the reaction tubes. As such liquid heat exchange agents come z. B. melting of salts or salt mixtures such as potassium nitrate, potassium nitrite, sodium nitrite and / or sodium nitrate and melting of metals such as sodium, mercury and alloys of various metals into consideration. But ionic liquids or heat transfer oils are used. The temperature of the heat exchange medium is between 220 to 490 ° C and preferably between 300 to 450 ° C and more preferably between 330 and 420 ° C.

Due to the exothermic nature of the reactions taking place, the temperature in certain sections of the interior of the reactor during the reaction may be higher than that of the heat exchange medium and a so-called hotspot is formed. The location and height of the hotspot is determined by the reaction conditions, but it may also be regulated by the dilution ratio of the catalyst layer or the flow rate of mixed gas. The difference between hotspot temperature and the temperature of the heat exchange medium is generally between 1 and 150 ° C, preferably between 10 and 100 ° C and more preferably between 20 and 80 ° C. The temperature at the end of the catalyst bed is generally between 0 and 100 ° C, preferably between 0.1 and 50 ° C, more preferably between 1 and 25 ° C above the temperature of the heat exchange medium.

The oxydehydrogenation can be carried out in all fixed-bed reactors known from the prior art, such as, for example, in a hearth furnace, in a fixed-bed or shell-and-tube reactor or in a plate heat exchanger reactor. A tube bundle reactor is preferred.

Preferably, the oxidative dehydrogenation is carried out in fixed bed tubular reactors or fixed bed bundle bundle reactors. The reaction tubes are (as well as the other elements of the tube bundle reactor) usually made of steel. The wall thickness of the reaction tubes is typically 1 to 3 mm. Their inner diameter is usually (uniformly) at 10 to 50 mm or 15 to 40 mm, often 20 to 30 mm. The number of reaction tubes accommodated in the tube bundle reactor is generally at least 1000, or 3000, or 5000, preferably at least 10,000. Frequently, the number of reactor tubes accommodated in the tube bundle reactor is 15,000 to 30,000 or up to 40,000 or up to 50 000. The length of the reaction tubes normally extends to a few meters, typical is a reaction tube length in the range of 1 to 8 m, often 2 to 7 m, often 2.5 to 6 m. Furthermore, the catalyst layer configured in the ODH reactor (s) R may consist of a single layer or of two or more layers. These layers may be pure catalyst or diluted with a material that does not react with the input gas or components of the product gas of the reaction. Furthermore, the catalyst layers may consist of solid material and / or supported shell catalysts.

In addition to butadiene, the product gas stream leaving the oxidative dehydrogenation generally contains unreacted 1-butene and 2-butene, oxygen and water vapor. As secondary components it furthermore generally contains carbon monoxide, carbon dioxide, inert gases (mainly nitrogen), low-boiling hydrocarbons such as methane, ethane, ethene, propane and propene, butane and isobutane, optionally hydrogen and optionally oxygen-containing hydrocarbons, so-called oxygenates. Oxygenates may be, for example, formaldehyde, furan, acetic acid, maleic anhydride, formic acid, methacrolein, methacrylic acid, crotonaldehyde, crotonic acid, propionic acid, acrylic acid, methyl vinyl ketone, styrene, benzaldehyde, benzoic acid, phthalic anhydride, fluorenone, anthraquinone and butyraldehyde.

The product gas stream at the reactor exit is characterized by a temperature near the temperature at the end of the catalyst bed. The product gas stream is then brought to a temperature of 150 to 400 ° C, preferably 160 to 300 ° C, more preferably 170 to 250 ° C. It is possible to isolate the line through which the product gas stream flows to maintain the temperature in the desired range, or to use a heat exchanger. This heat exchanger system is arbitrary as long as the temperature of the product gas can be maintained at the desired level with this system. As an example of a heat exchanger, spiral heat exchangers, plate heat exchangers, double tube heat exchangers, multi-tube heat exchangers, boiler spiral heat exchangers, shell-shell heat exchangers, liquid-liquid contact heat exchangers, air heat exchangers, direct contact heat exchangers and finned tube heat exchangers can be cited. Since, while the temperature of the product gas is adjusted to the desired temperature, part of the high-boiling by-products contained in the product gas may precipitate, the heat exchanger system should preferably have two or more heat exchangers. If two or more intended heat exchangers are arranged in parallel, and so a distributed cooling of the product gas obtained in the heat exchangers is made possible, the amount of high-boiling by-products which are deposited in the heat exchangers, and thus their operating time can be extended. As an alternative to the above-mentioned method, the two or more intended heat exchangers may be arranged in parallel. The product gas is supplied to one or more, but not all, heat exchangers, which are replaced after a certain period of operation of other heat exchangers. In this method, the cooling can be continued, a portion of the heat of reaction recovered and in parallel can be removed in one of the heat exchanger deposited high-boiling by-products. As a coolant mentioned above, a solvent can be used as long as it is capable of dissolving the high-boiling by-products. Examples are aromatic hydrocarbon solvents, such as. As toluene and xylenes, diethylbenzenes, triethylbenzenes, diisopropylbenzenes, Triisopropylbenzole. Particularly preferred is mesitylene. It is also possible to use aqueous solvents. These can be both acidic and alkaline, such as an aqueous solution of sodium hydroxide. Subsequently, a large part of the high-boiling secondary components and the water is separated from the product gas stream by cooling and compression. Cooling is by contacting with a coolant. This step is also referred to below as quench Q. This quench can consist of only one stage or of several stages. The product gas stream is thus brought directly into contact with a preferably organic cooling medium and thereby cooled. Suitable cooling media are aqueous coolants or organic solvents, preference is given to using aromatic hydrocarbons, particularly preferably toluene, o-xylene, m-xylene, p-xylene or mesitylene, or mixtures thereof. It is also possible to use all possible isomers of diethylbenzene, triethylbenzene, diisopropylbenzene and triisopropylbenzene and mixtures thereof.

Preferred is a two-step quench, i. H. the stage Ca) comprises two cooling stages Ca1) and Ca2), in which the product gas stream b is brought into contact with the organic solvent. In a preferred embodiment of the invention, therefore, the cooling stage Ca) is carried out in two stages, wherein the solvent of the second stage Ca2) loaded with secondary components is passed into the first stage Ca1). The solvent removed from the second stage Ca2) contains less secondary components than the solvent removed from the first stage Ca1).

A gas stream is obtained which comprises n-butane, 1-butene, 2-butenes, butadiene, optionally oxygen, hydrogen, water vapor, in small quantities methane, ethane, ethene, propane and propene, isobutane, carbon oxides, inert gases and parts of the solvent used in the quench. Furthermore, traces of high-boiling components can remain in this gas stream, which were not quantitatively separated in the quench.

The product gas stream from the solvent quench is compressed in at least one compression stage K and subsequently cooled further in the cooling apparatus, whereby at least one condensate stream is formed. There remains a gas stream containing butadiene, 1-butene, 2-butene, oxygen, water vapor, optionally low-boiling hydrocarbons such as methane, ethane, ethene, propane and propene, butane and isobutane, optionally carbon oxides and optionally inert gases. Furthermore, this product gas stream may still contain traces of high-boiling components. The compression and cooling of the gas stream can take place in one or more stages (n-stage). Generally, a total pressure is compressed in the range of 1.0 to 4.0 bar (absolute) to a pressure in the range of 3.5 to 20 bar (absolute). After each stage of compression, a cooling stage follows, in which the gas flow is reduced to a temperature in the range of 15 to 60 ° C. is cooled. The condensate stream can therefore also comprise a plurality of streams in the case of multistage compression. The condensate stream consists to a large extent of water and optionally the organic solvent used in the quench. Both streams (aqueous and organic phase) may also contain minor components such as low boilers, C ₄ hydrocarbons, oxygenates and carbon oxides.

The butadiene, n-butenes, oxygen, low-boiling hydrocarbons (methane, ethane, ethene, propane, propene, n-butane, isobutane), optionally water vapor, optionally carbon oxides and optionally inert gases and optionally traces of minor components containing gas stream is used as the output stream the further processing supplied.

In a step D) are non-condensable and low-boiling gas components comprising oxygen, low-boiling hydrocarbons (methane, ethane, ethene, propane, propene), carbon oxides and inert gases in an absorption column A as a gas stream from the process gas stream by absorbing the C ₄ Hydrocarbons separated in a high-boiling absorbent and subsequent desorption of C ₄ hydrocarbons. Preferably, step D) comprises the steps Da1), Da2) and Db):

Da1) absorption of the C ₄ hydrocarbons comprising butadiene and n-butenes in a high-boiling absorbent, wherein a stream of absorbent laden with C ₄ hydrocarbons and the gas stream are obtained,

Da2) removal of oxygen from the C ₄ hydrocarbon laden absorbent stream from step Da) by stripping with a non-condensable gas stream to obtain an absorbent stream laden with C ₄ hydrocarbons, and

Db) desorption of the C ₄ hydrocarbons from the loaded absorbent stream to obtain a C ₄ product gas stream consisting essentially of C ₄ hydrocarbons.

For this purpose, in the absorption stage Da1), the gas stream is contacted with an inert absorbent and the C ₄ hydrocarbons are absorbed in the inert absorbent to obtain an absorbent laden with C ₄ hydrocarbons and an offgas containing the remaining gas constituents. In a desorption step, the C ₄ hydrocarbons are released from the high-boiling absorbent again.

The absorption stage can be carried out in any suitable absorption column known to the person skilled in the art. The absorption can be carried out by simply passing the product gas stream through the absorbent. But it can also be done in columns or in rotational absorbers. It can be used in cocurrent, countercurrent or cross flow. Preferably, the absorption is carried out in countercurrent. Suitable absorption columns are z. B. tray columns with bell, centrifugal and / or sieve bottom, colon NEN with structured packs, z. B. Sheet metal packings with a specific surface area of 100 to 1000 m ² / m ³ as Mellapak® 250 Y, and packed columns. However, trickle and spray towers, graphite block absorbers, surface absorbers such as thick-layer and thin-layer absorbers, as well as rotary columns, rags, cross-flow scrubbers and rotary scrubbers are also suitable.

In one embodiment, an absorption column in the lower region of the butadiene, n-butenes and the low-boiling and non-condensable gas components containing gas stream is supplied. In the upper region of the absorption column, the high-boiling absorber is applied.

Inert absorbent used in the absorption stage are generally high-boiling non-polar solvents in which the C4-hydrocarbon mixture to be separated has a significantly higher solubility than the other gas constituents to be separated off. Suitable absorbents are comparatively nonpolar organic solvents, for example aliphatic Ce to Cis alkanes, or aromatic hydrocarbons, such as the paraffin distillation from the paraffin distillation, toluene or ethers with bulky groups, or mixtures of these solvents, these being a polar solvent such as 1, 2-dimethylphthalate may be added. Suitable absorbers are also esters of benzoic acid and phthalic acid with straight-chain d-Cs-alkanols, as well as 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 ^®. Often, this solvent mixture contains di-methyl phthalate in an amount of 0.1 to 25 wt .-%.

In a preferred embodiment in the absorption stage Da1) the same solvent as in the cooling stage Ca) is used.

Preferred absorbents are solvents which have a solubility for organic peroxides of at least 1000 ppm (mg of active oxygen / kg of solvent). Preference is given to aromatic hydrocarbons, particularly preferably toluene, o-xylene, p-xylene and mesitylene, or mixtures thereof. It is also possible to use all possible isomers of diethylbenzene, triethylbenzene, diisopropylbenzene and triisopropylbenzene and mixtures thereof. At the top of the absorption column, a gas stream d is withdrawn, which is essentially oxygen, low-boiling hydrocarbons (methane, ethane, ethene, propane, propene), the hydrocarbon solvent, optionally C 4 -hydrocarbons (butane, butenes, butadiene), optionally inert gases, optionally carbon oxides and optionally still contains water vapor. This stream is at least partially supplied as circulating gas stream d2 the ODH reactor. Thus, for example, the inlet flow of the ODH reactor can be adjusted to the desired C4 hydrocarbon content. In general, if appropriate after separation of a purge gas stream, at least 10% by volume, preferably at least 30% by volume, of the gas stream d is recirculated as circulating gas stream d2 into the oxidative dehydrogenation zone. In general, the recycle stream is 10 to 70% by volume, preferably 30 to 60% by volume, based on the sum of all material streams fed into the oxidative dehydrogenation B). The purge gas stream may be subjected to thermal or catalytic afterburning. In particular, it can be thermally utilized in a power plant.

At the bottom of the absorption column residues in the absorbent dissolved oxygen are discharged in a further column by flushing with a gas. The remaining oxygen content should be so small that the stream leaving the desorption column and containing butane, butene and butadiene only contains a maximum of 100 ppm of oxygen.

The stripping of the oxygen in step Da2) can be carried out in any suitable column known to the person skilled in the art. The stripping can be carried out by simply passing non-condensable gases, preferably gases which are not or only weakly absorbable in the absorbent stream, such as methane, through the loaded absorption solution. With stripped C4 hydrocarbons are washed in the upper part of the column back into the absorption solution by the gas stream is passed back into this absorption column. This can be done both by a piping of the stripping column and a direct assembly of the stripping column below the absorber column. Since the pressure in the stripping column part and the absorption column part is the same, this direct coupling can take place. Suitable Stippkolonnen are z. B. tray columns with bell, centrifugal and / or sieve tray, columns with structured packings, eg. B. Sheet metal packings with a specific surface area of 100 to 1000 m ² / m ³ as Mellapak® 250 Y, and packed columns. But there are also trickle and spray towers and rotary columns, dishwashers, cross-flow scrubbers and rotary scrubbers into consideration. Suitable gases are for example nitrogen or methane.

In one embodiment of the method, stripping is carried out in step Da2) with a methane-containing gas stream. In particular, this gas stream (stripping gas) contains> 90% by volume of methane.

The loaded with C4 hydrocarbons absorbent stream can be heated in a heat exchanger and then passed into a desorption. In a variant of the method, the desorption step Db) is carried out by relaxing and stripping the loaded absorbent by a steam stream.

The absorbent regenerated in the desorption stage can be cooled in a heat exchanger. The cooled stream still contains water in addition to the absorbent, which is separated in the phase separator. The C4 product gas stream consisting essentially of n-butane, n-butenes and butadiene generally contains from 20 to 80% by volume of butadiene, from 0 to 80% by volume of n-butane, from 0 to 10% by volume of 1-butene , 0 to 50% by volume of 2-butenes and 0 to 10% by volume of methane, the total amount being 100% by volume. Furthermore, small amounts of iso-butane may be included. A portion of the condensed head effluent of the desorption column, which contains mainly C 4 hydrocarbons, can be returned to the top of the column to increase the separation efficiency of the column.

The liquid or gaseous C ₄ product streams leaving the condenser can subsequently be separated by extractive distillation in step E) with a butadiene-selective solvent into a material stream containing the selective solvent and a stream comprising butanes and n-butenes.

In a preferred embodiment of the method according to the invention is additionally prevented by means of a shutdown that the oxidative dehydrogenation reactor is charged with a reactive gas mixture whose composition is an explosive, the shutdown mechanism is designed as follows: a) in a computer is a characteristic of the reaction mixture explosion diagram deposited, in which, depending on the composition of the reaction gas mixture explosive and non-explosive compositions are delimited from each other; b) by determining the amount and optionally composition of the reactor for generating the reaction gas mixture supplied gas streams, a record is determined, which is forwarded to the computer; c) from the data set obtained under b), the computer calculates a current operating point of the reaction gas mixture in the exploded diagram; d) If the distance between the operating point and the nearest explosive limit falls below a specified minimum value, the supply of gas streams to the reactor is automatically interrupted.

The minimum value is preferably calculated from a statistical error analysis of the measured variables necessary for calculating the operating point. The process makes it possible to carry out heterogeneously catalyzed gas-phase partial oxidations and oxidative dehydrogenations of at least one organic compound with increased safety at oxygen contents of the reaction gas mixture which is> = 0.5, or> = 0.75, or> = 1, or> = 2, or> = 3, or> = 5, or> = 10 volume percentage points above the oxygen limit concentration. As already described, the oxygen limit concentration (LOC) is understood as meaning the percentage volume fraction of molecular oxygen of the reaction gas mixture which falls below the volume proportions of the other constituents of the reaction gas mixture, namely in particular the organic compound to be oxidized and the inert diluent gas, regardless of the quantity , one by one Local ignition (such as local overheating or sparking in the reactor) initiated combustion (explosion) at a given pressure and temperature of the reaction gas mixture in the same is no longer able to spread from the ignition source ago. For reasons of safety, it may be expedient not to store the course of the experimentally determined explosion limit as an explosion diagram, but rather a so-called switching curve shifted by a safety distance to this in the computer. The safety distance is suitably chosen so that all incoming error in the determination of the operating point of the reaction gas mixture sources of error and measurement inaccuracies are taken into account. The safety distance can be determined both by an absolute error analysis and by a statistical error analysis. As a rule, a safety margin of 0.1 to 0.4 vol.% O2 points is sufficient.

Since the explosive behavior of butane and n-butenes is comparable and water vapor and nitrogen have a barely distinguishable effect on the exploded diagram of butane and / or butene, as a characteristic explosion diagram to be deposited in the computer according to the invention, for example: a) the Chart Butene / O 2 - N ₂ ;

b) the diagram butane / O 2 - N ₂ ;

c) the diagram butenes / O 2 - H ₂ O;

d) the diagram butane / O 2 - H ₂ O;

e) the diagram butene / O 2 - (N 2 / H 2 O);

f) the diagram butane / O 2 - (N 2 / H 2 O);

According to the invention, the exploded diagram of butene / 02-N2 will preferably be stored in the computer.

The temperature used in the experimental determination of the explosion diagram will be a temperature not too far from the temperature range covered by the partial oxidation.

To calculate a meaningful current operating point of the reaction gas mixture in the exploded diagram, for example, the experimental determination of the following measured quantities is sufficient: a) the amount of air supplied to the reactor per unit time in Nm ³ ;

b) the amount of butene-containing input gas fed to the reactor per unit time in Nm ³ ;

c) the amount of water vapor and / or circulating gas fed to the reactor per unit time in Nm ³ ;

d) the 02 content of the circulating gas. The oxygen and nitrogen content of the air is known, the amount of input gas containing butenes and the optionally used amount of water vapor result as an immediate measurement result and the cycle gas is assumed, apart from its oxygen content, consisting exclusively of nitrogen. Should the circulating gas still contain flammable components, this will not adversely affect the safety issue since its presence in the explosion diagram would only mean a rightward shift of the actual operating point with respect to the calculated operating point. Steam contained in circulating gas in small quantities or contained carbon oxides can be considered as nitrogen in terms of safety relevance.

The quantity measurement of the gas streams supplied to the reactor can be carried out with any suitable measuring device. Such measuring instruments include, for example, all flowmeters such as throttle devices (eg diaphragms or Venturi tubes), displacement flowmeters, variable area apparatus, turbine, ultrasonic, vortex and mass flow devices. For reasons of low pressure loss Venturi tubes are preferred according to the invention. By taking pressure and temperature into account, the measured volume flows can be converted into Nm ³ .

The determination of the oxygen content in the recycle gas may e.g. inline as described in DE-A 101 17678. In principle, however, it can also be carried out online by taking a sample of the product gas mixture coming from the oxidative dehydrogenation before it enters the product separation (workup) and analyzing it online in such a way that the analysis takes place over a period of time shorter than the residence time of the product Product gas mixture in the workup. That The gas volume to the analyzer must be correspondingly large via an analysis gas bypass and the piping system to the analyzer selected accordingly small. It goes without saying that an O 2 determination in the reaction gas could also be carried out instead of the recycle gas analysis. Of course, both can be done. In application-technically expedient manner, the determination of the operating point for the application of the safety-related, programmable logic controller (SSPS) according to the invention is multichannel, preferably of at least three channels.

That Each quantity measurement is preferably carried out by means of three fluid flow indicators (FFI) connected in series or in parallel. The same applies to the 02 analysis. If one of the three operating points of the reaction gas mixture calculated from the three data sets falls below the predetermined minimum distance in the exploded diagram, the gas flow is reduced, for example. closed in the order of air, with time delay butene-containing input gas and finally, optionally water vapor and / or recycle gas automatically closed.

From the point of view of later recommissioning, it may be expedient to continue circulating water vapor and / or circulating gas. Alternatively, from the three individual measurements, a mean operating point can also be calculated in the exploded diagram. If its distance to the explosion limit falls below a minimum value, an automatic shutdown takes place as described above. In principle, the method according to the invention can be used not only for stationary operation but also for starting and stopping the partial oxidation.

Examples: The tube reactor (R) consists of stainless steel 1, 4571, has an internal diameter of 29.7 mm and 5 m in length and is filled with a mixed oxide catalyst (2500 ml). In the center of the pipe is a thermo-sleeve (outer diameter 6 mm) with internal thermocouples installed to detect the temperature profile in the bed. The tube is lapped with a molten salt to keep the outer wall temperature constant. A stream of butenes and butanes (a1), water vapor, air and oxygen-containing circulating gas are fed to the reactor. Furthermore, nitrogen can be fed to the reactor.

The exhaust gas (b) is cooled in a quenching device (Q) to 45 ° C, wherein the high-boiling by-products are separated. The stream is compressed in a compressor stage (K) to 10 bar and cooled again to 45 ° C. In the cooler, a condensate stream c1 is discharged. The gas stream c2 is fed to an absorption column (A). The absorption column is operated with mesitylene. From the absorption column, a liquid stream rich in organic products and a gaseous stream d at the top of the absorption column are obtained. The entire workup is designed so that water and the organic components are completely separated. A portion of the stream d is recycled as cycle gas d2 back into the reactor.

Example 1 Figure 1 shows schematically the experimental setup. The circulating gas flow d2 is set to 5025 Nl / h and kept constant. The oxygen concentration in the circulating gas is 7.5% by volume, ie it is approximately as high as in the later stationary operating state. 285 Nl / h steam are fed to the reactor. Thereafter, the reactor is a current a1, consisting of 51 vol .-% butenes and 49 vol .-% butanes supplied. Starting with a flow of 250 NL / h butene / butane, the flow is gradually increased within 25 minutes. After 25 minutes, the butene / butane stream a1 is 420 NL / h, at the same time the circulating gas stream d2 is reduced and amounts to 4820 NL / h after 25 min. Then the volume flow a1 is increased to 630 NL / h and at the same time air at 1338 NL / h is supplied, and at the same time the circulating gas flow d2 is reduced to 3270 NL / h. After 35 minutes, in a further step, the volume flow a1 is increased to 845 NL / h and at the same time the volume flow of air is increased to 2695 NL / h, and at the same time the circulating gas flow d2 is reduced to 1700 NL / h. At the reactor inlet results in a final value of 12.5 vol .-% oxygen, wherein the oxygen comes from both the circulating gas stream d2 and from the supplied air. The total gas flow is during the entire start-up Ganges kept approximately constant by reducing the circulating gas flow d2. The height of the hotspot can be easily controlled by the addition of the butene / butane stream a1 and controlled specifically. FIG. 2 shows the volume flows of C4-hydrocarbons a1 and oxygen-containing gas a2 and the resulting residual oxygen content at the outlet of the quench during the starting process according to the invention. The 02 concentration in the feed initially amounts to 7.5% by volume in order to simulate the O 2 concentration in the circulating gas stream d 2. Additional inert gas (N2) for dilution is not needed.

FIG. 3 shows the concentration profile of butanes / butenes (fuel gas), oxygen and the remaining gas components (100% C Crenngas - C02) before the reactor ("reactor"), and between quench and compression stage ("absorption") together with the explosion diagrams for the reactor ("Ex. Reactor") and the absorption column ("Ex. Absorption"). All concentrations are in volume percent. On the ordinate the concentration of the fuel gas is plotted, on the abscissa the concentration of oxygen.

Immediately before the addition of the butene / butane stream a1, the oxygen concentration in front of and behind the reactor, due to the dilution with water vapor, is 7.1% by volume, between quench and absorption column, and in the cycle gas 7.5% by volume. , Until immediately before the addition of air and while the fuel gas flow is raised, the oxygen concentration between quench and absorption column drops to about 4 vol .-%. After addition of air and while the fuel gas flow is raised to the final value, the oxygen concentration upstream of the reactor rises to 12.5% by volume, behind the reactor to 6% by volume and between quench and absorption column to approximately 7.8 Vol .-%, but without crossing the explosion area. Before air is fed in, the oxygen concentration in front of and behind the reactor and between the quench and the absorption column can not exceed the value of the circulating gas. Thus, a safe start can be guaranteed. Comparative Example 1

The reactor and the workup part are first purged with a stream of 1000 Nl / h of nitrogen. After one hour, the measured oxygen content behind the reactor and in the recycle gas is less than 0.5% by volume. Then, 240 Nl / h of air and 1000 Nl / h of nitrogen are introduced into the reactor. The circulating gas flow is set to 2190 Nl / h. The circulating gas flow is kept constant. After 20 minutes, the oxygen concentration in the circulating gas stream is 4.1% by volume. The air and nitrogen supply to the reactor are stopped simultaneously and fed to the reactor 225 Nl / h of water vapor. Thereafter, air and a stream consisting of 80 vol .-% butenes and 20 vol .-% butanes are fed to the reactor, wherein the ratio of air flow to butenes / butane stream is controlled so that this ratio is constant about 3.75. Starting with a flow of 44 Nl / h of butenes / butanes and 165 Nl / h of air, the streams are raised within one hour at a constant ramp and become low one hour 440 Nl / h butene / butane and 1650 Nl / h air. The circulating gas flow is kept constant during the entire starting process and is 2190 Nl / h.

The system is operated for 4 days, with a steady state, in which the concentrations of the gas components do not change by more than 5% / h. The concentration course of butanes / butenes (fuel gas), oxygen and the remaining gas components (100% - Cßrenngas - C02) before the reactor ("reactor"), as well as between quench and compression stage ("absorption") and in the recycle gas ("recycle gas"). ) is shown together with the explosion diagrams for the reactor ("Ex. reactor") and the absorption column ("Ex. absorption") in FIG. All concentrations are in volume percent. On the ordinate the concentration of the fuel gas is plotted, on the abscissa the concentration of oxygen. Immediately before the start of the addition of fuel gas (butenes and butanes), the oxygen concentration in front of and behind the reactor, between quench and absorption column and in the recycle gas is 4.1% by volume. While the fuel gas flow is raised to the final value, the oxygen concentration in the recycle gas increases to a final value of about 7.5% by volume. Even before the reactor and between the quench and the absorption column, the oxygen concentration increases, but without crossing the explosion area. Thus, a safe start can be guaranteed. Comparative Example 1 corresponds to Example 1 of WO2015 / 104397. The explosion area is avoided in the reactor and in the workup. However, the procedure is so far less favorable than the starting process according to the invention, since the oxygen content of the available cycle gas is first reduced by dilution with an inert gas and then raised again by adding an additional oxygen-containing gas (air). However, the injection of inert gas causes additional costs, as well as the supply of an oxygen-containing gas (air), since this must be compressed to the required pressure.

Comparative Example 2 The reactor is first purged with a stream of 1000 Nl / h of nitrogen as in Comparative Example 1. After one hour, the measured oxygen content behind the reactor and in the recycle gas is less than 0.5% by volume. Then 620 Nl / h of air and 1000 Nl / h of nitrogen are introduced into the reactor. The circulating gas flow is set to 2190 Nl / h and kept constant. After 20 minutes, the oxygen concentration in the recycle gas is 7.9 vol .-%. The oxygen concentration in the circulating gas stream is therefore approximately as high as in the later stationary operating state, cf. Table 1. Air and nitrogen supply to the reactor are stopped simultaneously. 225 Nl / h steam are fed to the reactor. Thereafter, air and a stream consisting of 80 vol .-% butenes and 20 vol .-% butanes are fed to the reactor, wherein the ratio of air flow to butenes / butane stream is controlled so that it is constant about 3.75. Starting with a flow of 44 Nl / h of butenes / butanes and 165 Nl / h of air, the streams are raised within one hour at a constant ramp. After one hour, the butene / butane stream is 440 Nl / h and the air flow 1650 Nl / h. The circulating gas flow is kept constant during the entire starting process and amounts to 2190 NL / h. The concentration course of butanes / butenes (fuel gas), oxygen and the remaining gas components (100% - Cßrenngas - C02) before the reactor ("reactor"), as well as between quench and compression stage ("absorption") and in the recycle gas ("recycle gas"). ) is shown together with explosion diagrams for the reactor ("Ex. reactor") and the absorption column ("Ex. absorption") in FIG. All concentrations are in volume percent. On the ordinate the concentration of the fuel gas is plotted, on the abscissa the concentration of oxygen. Immediately before the start of the addition of fuel gas (butenes and butanes), the oxygen concentration in front of and behind the reactor, between the quench and the absorption column and in the recycle gas is 7.9% by volume. While the fuel gas flow is raised to the final value, the oxygen concentration in the recycle gas changes only slightly to a final value of about 7.6 vol .-%. Before the reactor and between quench and absorption column, the oxygen concentration increases, it can be seen that the distance to the explosion region in the reactor is very low during reactor commissioning. Safe process operation is difficult to realize here.

Claims

claims

Process for the preparation of butadiene from n-butenes having a start-up phase and an operating phase, the process comprising the steps in the operating phase:

A) providing a n-butenes containing feed gas stream a1;

B) feed of the n-butenes containing feed gas stream a1, an oxygen-containing gas stream a2 and an oxygen-containing cycle gas stream d2 in at least one oxidative dehydrogenation and oxidative dehydrogenation of n-butenes to butadiene, wherein a product gas stream b containing butadiene, unreacted n-butenes, water vapor, Oxygen, low-boiling hydrocarbons, high-boiling secondary components, optionally carbon oxides and optionally inert gases is obtained;

C) cooling and compression of the product gas stream b and condensation of at least a portion of the high-boiling secondary components, at least one aqueous condensate stream c1 and a gas stream c2 containing butadiene, n-butenes, water vapor, oxygen, low-boiling hydrocarbons, optionally carbon oxides and optionally inert gases is obtained;

D) feeding the gas stream c2 in an absorption zone and separation of non-condensable and low-boiling gas components comprising oxygen, low-boiling hydrocarbons, optionally carbon oxides and optionally inert gases as gas stream d from the gas stream c2 by absorption of C4 hydrocarbons comprising butadiene and n-butenes in an absorbent, wherein an absorbent stream loaded with C ₄ -hydrocarbons and the gas stream d are obtained, and recycling, optionally after separation of a purge gas stream p, of the gas stream d as circulating gas stream d 2 into the oxidative dehydrogenation zone, wherein the start-up phase in the order i) to iv) comprises the steps of: i) feeding a gas stream d2 'having a composition corresponding to the circulating gas stream d2 in the operating phase into the dehydrogenating zone and adjusting the circulating gas stream d2 to at least 70% of the total volume flow in the operating phase; ii) optionally additionally feeding a steam stream a3 into the dehydrogenation zone; iii) additional feeding of the butene-containing feed gas stream a1 with a lower volume flow than in the operating phase and raising this volume flow until to achieve at least 50% of the volume flow of the feed gas stream a1 in the operating phase, wherein the total gas flow through the dehydrogenation zone corresponds to a maximum of 120% of the total gas flow during the operating phase; iv) additional feeding, upon reaching at least 50% of the volume flow of the butenes containing feed gas stream a1 in the operating phase, an oxygen-containing stream a2 with a lower volume flow than in the operating phase, and increasing the volume flows of the feed gas streams a1 and a2 until reaching the volume flows in the Operating phase, wherein the total gas flow through the dehydrogenation corresponds to a maximum of 120% of the total gas flow during the operating phase.

A method according to claim 1, characterized in that the circulating gas stream d2 is set to 95 to 105% of the total volume flow in the operating phase.

A method according to claim 1 or 2, characterized in that in step iii) the volume flow of the butene-containing feed gas stream is raised to a maximum of 75% of the volume flow in the operating phase.

Method according to one of claims 1 to 3, characterized in that in step iv) first only the volume flow of the oxygen-containing gas stream a2 is raised until reaching a ratio of oxygen to hydrocarbons, which corresponds to the ratio of oxygen to hydrocarbons in the operating phase, and then both volume flows a1 and a2 are raised until reaching in each case 100% of the volume flow of the gas flows a1 and a2 in the operating phase.

Method according to one of claims 1 to 3, characterized in that the ratio of oxygen to hydrocarbons in the operating phase at an n-butenes content of 50 to 100 vol .-% in the feed gas stream a1 of 0.65: 1 to 1, 5: 1 amounts to.

Method according to one of claims 1 to 5, characterized in that the total gas flow, comprising the streams a1, a2, d2 and optionally a3, by the dehydrogenation zone during steps (ii), (iii) and (iv) remains substantially constant, and 90 to 1 10 vol .-% of the total gas flow through the dehydrogenation during the operating phase corresponds.

Method according to one of claims 1 to 6, characterized in that the amount of water vapor in the dehydrogenation zone during steps ii), iii) and iv) is 0.5 to 10 vol .-%.

Method according to one of claims 1 to 7, characterized in that as the gas stream d2 'a part of the circulating gas stream of one or more parallel operated reactors for the production of butadiene from n-butenes by oxidative dehydrogenation, which are in the operating phase, is removed.

9. The method according to any one of claims 1 to 8, characterized in that the pressure in the dehydrogenation zone during the start-up phase is 1 to 5 bar.

10. The method according to any one of claims 1 to 9, characterized in that the pressure in the absorption zone during the start-up phase is 2 to 20 bar.

1 1. Method according to one of claims 1 to 10, characterized in that step D) the steps Da) and Db):

Da) separation of non-condensable and low-boiling gas constituents comprising oxygen, low-boiling hydrocarbons, optionally carbon oxides and optionally inert gases as gas stream d from the gas stream c2 by absorption of the C4 hydrocarbons containing butadiene and n-butenes in an absorbent, wherein one with C4- Hydrocarbons laden absorbent stream and the gas stream d are obtained, and

Db) subsequent desorption of the C4 hydrocarbons from the loaded Absorpti- onsmittelstrom, wherein a C4 product gas stream d1 is obtained comprises.

12. The method according to any one of claims 1 to 1 1 with the additional steps:

E) separation of the C4 product stream d1 by extractive distillation with a solvent which is selective for butadiene into a butadiene and the stream e1 containing the selective solvent and a stream e2 containing n-butenes; F) Distillation of the butadiene and the selective solvent-containing material stream f2 in a substantially consisting of the selective solvent stream gl and a butadiene-containing stream g2.

13. The method according to any one of claims 1 to 12, characterized in that the absorbent used in step D) is an aromatic hydrocarbon solvent.

14. The method according to any one of claims 1 to 13, characterized in that it is prevented by means of a shutdown mechanism that the oxidative dehydrogenation reactor is charged with a feed gas mixture whose composition is explosive, wherein the shutdown mechanism is designed as follows: a) in a computer, a characteristic for the feed gas mixture explosion diagram is deposited, in which, depending on the composition of the feed gas mixture explosive and non-explosive compositions are delimited against each other; b) by determining the amount and optionally composition of the dehydrogenation zone for generating the feed gas mixture supplied gas streams, a record is determined, which is forwarded to the computer; c) from the data set obtained under b) the computer calculates a current operating point of the feed gas mixture in the exploded diagram; d) If the distance between the operating point and the nearest explosive limit falls below a predetermined minimum value, the supply of gas streams to the dehydrogenation zone is automatically interrupted.