WO2014128191A1

WO2014128191A1 - Process for removing oxygen from c4 hydrocarbon streams

Info

Publication number: WO2014128191A1
Application number: PCT/EP2014/053275
Authority: WO
Inventors: Gauthier Luc Maurice Averlant; Alireza REZAI; Sonja Giesa; Martin Dieterle
Original assignee: Basf Se
Priority date: 2013-02-21
Filing date: 2014-02-20
Publication date: 2014-08-28

Abstract

The process for removing oxygen from a C₄ hydrocarbon stream comprising free oxygen by catalytic combustion, in which the hydrocarbon stream comprising free oxygen is converted over a catalyst bed by catalytic combustion in the presence or absence of free hydrogen to obtain an oxygen-depleted hydrocarbon stream is characterized in that the catalytic combustion is performed continuously, the inlet temperature in the catalyst bed is at least 300°C and the maximum temperature in the catalyst bed is not more than 700°C.

Description

The invention relates to a process for the removal of oxygen from free-oxygen-containing C4 hydrocarbon streams.

In various chemical processes, free oxygen-containing hydrocarbon streams may be obtained, from which the free oxygen should or must be removed.

For example, free oxygen contained in a gas stream containing ethylenically unsaturated hydrocarbon streams can result in the formation of peroxides that are difficult to handle for safety reasons. Free oxygen containing butadiene may, for. B. by oxidative dehydrogenation of n-butenes (1-butene and / or 2-butene) are obtained. As the starting gas mixture for the oxidative dehydrogenation of n-butenes to butadiene, any n-butenes containing mixture 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 starting gas. Furthermore, n-butenes containing gas mixtures can be used as the starting gas, which were obtained by catalytic fluid catalytic cracking (FCC). Gas mixtures containing n-butenes, which are used as the starting gas in the oxidative dehydrogenation of n-butenes to butadiene, can also be prepared by non-oxidative dehydrogenation of n-butane-containing gas mixtures. WO2005 / 063658 discloses a process for producing butadiene from n-butane by the steps

A) providing a n-butane containing feed gas stream a;

B) feeding the n-butane-containing feed gas stream a into at least one first hydrogenation zone and non-oxidative catalytic dehydrogenation of n-butane, wherein a product gas stream b containing n-butane, 1-butene, 2-butene, butadiene, hydrogen , low-boiling secondary constituents and optionally water vapor is obtained;

C) feeding the product gas stream b of the non-oxidative catalytic dehydrogenation and an oxygen-containing gas into at least one second dehydrogenation zone and oxidative dehydrogenation of 1-butene and 2-butene, wherein a product gas stream c containing n-butane,

2-butene, butadiene, hydrogen, low-boiling secondary components and water vapor is obtained, which has a higher content of butadiene than the product gas stream b; D) separating hydrogen, the low-boiling minor components and water vapor to obtain a C4 product gas stream d consisting essentially of n-butane, 2-butene and butadiene;

E) separating the C4 product gas stream d into a recycle stream e1 consisting essentially of n-butane and 2-butene and a stream e2 consisting essentially of butadiene by extractive distillation and recycling the stream e1 into the first de-hydrogenation zone.

This process is characterized by a particularly effective utilization of raw materials. Thus, losses of the raw material n-butane are minimized by recycling unreacted n-butane into the dehydrogenation. The coupling of non-oxidative catalytic dehydrogenation and oxidative dehydrogenation achieves a high butadiene yield. The process is characterized by high selectivity as compared to the production of butadiene by cracking. There are no by-products. It eliminates the costly separation of butadiene from the product gas mixture of the cracking process.

WO 2006/075025 describes a process for the preparation of butadiene from n-butane by non-oxidative, catalytic dehydrogenation of n-butane, subsequent oxidative dehydrogenation and workup of the product mixture. After oxidative dehydrogenation, the oxygen remaining in the product gas stream can be removed, for example by catalytically reacting it with hydrogen. A corresponding C4 product gas stream can be 20 to 80% by volume butadiene, 20 to 80% by volume n-butane, 5 to 50% by volume 2-butene and 0 to 20% by volume 1-butene and also low Contain quantities of oxygen.

The residual oxygen can have a disturbing effect insofar as it can act as an initiator for polymerization reactions in downstream process steps. This danger is particularly present in the distillative removal of butadiene and may lead to deposits of polymers (formation of so-called "popcorn") in the extractive distillation column, therefore an oxygen removal is carried out immediately after the oxidative dehydrogenation, generally by a catalytic combustion stage, in which oxygen is reacted with the hydrogen contained in the gas stream in the presence of a catalyst .This results in a reduction of the oxygen content down to a few traces.A suitable alumina catalyst is described, which contains 0.01 to 0.1 wt % Of platinum and 0.01 to 0.1% by weight of tin Alternatively, catalysts containing reduced-form copper are also indicated.

WO 2010/130610 describes a process for the preparation of propylene oxide by reacting propene with hydrogen peroxide, separating the propylene oxide to obtain a gas mixture containing propene and oxygen. Hydrogen is added to this gas mixture and the oxygen contained is at least partially converted by reaction with the Hydrogen reacted in the presence of a copper-containing catalyst. The catalyst contains 30 to 80 wt .-% copper, calculated as CuO.

WO 2006/050969 describes a process for the preparation of butadiene from n-butane, in which first butane is catalytically dehydrogenated to butene, followed by an oxidative dehydrogenation (ODH) to form butadiene. It is stated that the product gas stream may still contain small amounts of oxygen and in the presence of larger amounts of oxygen, a catalytic combustion stage is connected, in which the oxygen is reacted with the hydrogen contained in the gas stream in the presence of a catalyst. This is to achieve a reduction of the oxygen content down to a few traces. In a simulation example, the flow discharged from the ODH contains 4.5 vol.% Oxygen.

Similar methods are described in DE-A-10 2004 059 356 and DE-A-10 2004 061 514. It is stated in each case that the remaining oxygen in the product gas of the oxidative dehydrogenation can be removed by catalytically reacting with hydrogen.

Besides "popcorn" formation, in oxygenated gas mixtures, especially butadiene and oxygen-containing gas mixtures, the oxygen content can contribute to the deactivation of catalysts, soot deposits, peroxide formation and to the deterioration of the adsorption properties of solvents in the work-up process.

In particular, in the production of butadiene from n-butane selective oxygen removal is a prerequisite for the economic feasibility of the process, since any loss of the target product butadiene is associated with higher costs.

It is an object of the present invention to provide a process for the removal of oxygen from a free-oxygen C4 hydrocarbon stream by catalytic combustion, wherein a residual oxygen content of less than 100 ppm or less than 80 ppm or less than 50 ppm or less than 30 ppm or less than 10 ppm or less than 1 ppm or Oppm is accessible and preferably as little as possible C ₄ - hydrocarbon is consumed as desired product.

Particularly preferably, the residual oxygen content should be less than 50 ppm. It is determined by electrochemical oxygen sensors such as KE 25 from Figgres or A-3, B-1 or B-3 from Teledyne. After calibration, a measurement accuracy of about 10 ppm O2 is achieved.

The object is achieved by a method for the removal of oxygen from a free oxygen-containing C ₄ hydrocarbon stream by catalytic combustion, in which the free oxygen-containing hydrocarbon stream is reacted on a catalyst bed by catalytic combustion in the presence or absence of free hydrogen Obtaining an oxygen-depleted hydrocarbon stream, characterized in that the catalytic combustion is carried out continuously, the inlet temperature in the catalyst bed is at least 300 ° C and the maximum temperature in the catalyst bed is not more than 700 ° C.

According to the invention, the term "C 4 -hydrocarbon stream" means a hydrocarbon stream in which at least 60% by volume, preferably at least 80% by volume, in particular at least 95% by volume, of the hydrocarbons are C 4 -hydrocarbons.

Preferably, the C4 hydrocarbon stream is derived from butane dehydrogenation or butene dehydrogenation, especially from the oxydehydrogenation of butene to butadiene. It preferably contains 0.5 to 8.0% by volume of free oxygen, preferably 1.0 to 8.0% by volume, particularly preferably 2.0 to 7.0% by volume, in particular 3.0 to 6.5% by volume.

According to one embodiment of the invention, the free oxygen-containing hydrocarbon stream contains a sufficient amount of free hydrogen to react with the free oxygen and / or it is added thereto, and the free oxygen is reacted with the free hydrogen.

Alternatively, the free oxygen-containing hydrocarbon stream does not contain free hydrogen and no free hydrogen is added to it.

In this case, it is preferable to react the free oxygen with hydrocarbon contained in the free oxygen or with added methanol, natural gas and / or synthesis gas as the reducing agent. According to one embodiment of the invention, the C4 hydrocarbon stream used according to the invention is obtained according to the following steps:

- Providing an n-butane-containing feed gas stream a;

Feed of the n-butane-containing feed gas stream a into at least one first hydrogenation zone and non-oxidative, catalytic dehydrogenation of n-butane, wherein a

Gas stream b containing n-butane, 1-butene, 2-butenes, butadiene, hydrogen, optionally water vapor, optionally carbon oxides and optionally inert gases is obtained;

Feed of a stream of butane, butenes, butadiene obtained from the gas stream b and of an oxygen-containing gas into at least one second dehydrogenation zone and oxidative dehydrogenation of 1-butene and 2-butenes, a gas stream g comprising Butane, unreacted (s) 1-butene and 2-butenes, butadiene, water vapor, optionally carbon oxides, optionally hydrogen and optionally inert gases is obtained, and

- Removal of residual oxygen contained in the gas stream g by means of catalytic combustion to obtain an oxygen-depleted stream h. For a description of the butane dehydrogenation and oxydehydrogenation, reference may be made to the documents mentioned at the outset, in particular to DE-A-10 2004 059 356 (WO 2006/061202), DE-A-10 2004 061 514 (WO 2006/066848), WO 2010/130610, WO 2006/0 50969, DE-A-10 2005 002 127 (WO 2006/075025).

The product gas stream leaving the oxidative dehydrogenation contains, in addition to butadiene and non-separated n-butane, hydrogen, carbon oxides, oxygen and water vapor. As minor constituents, it may also contain inert gas, such as nitrogen, methane, ethane, ethene, propane and propene, as well as oxygen-containing hydrocarbons, so-called oxygenates.

In general, the product gas stream leaving the oxidative dehydrogenation has from 2 to 40% by volume of butadiene, from 5 to 80% by volume of n-butane, from 0 to 15% by volume of 2-butenes, from 0 to 5% by volume of 1-butene. From 5 to 70% by volume of steam, from 0 to 10% by volume of low-boiling hydrocarbons (methane, ethane, ethene, propane and propene), from 0.1 to 15% by volume of hydrogen, from 0 to 70% by volume of inert gas, 0 to 10% by volume of carbon oxides, 0 to 10% by volume of oxygen and 0 to 10% by volume of oxygenates, the total amount of the ingredients being 100% by volume. Oxygenates may be, for example, furan, acetic acid, methacrolein, maleic anhydride, maleic acid, phthalic anhydride, propionic acid, acetaldehyde, acrolein, formaldehyde, formic acid, benzaldehyde, benzoic acid and butyraldehyde. In addition, acetylene, propyne and 1, 2-butadiene may be present in traces.

Other sources of the free oxygen-containing C4 hydrocarbon are, for. Raffinate II and products of ethylene dimerization. If the product gas stream contains more than just traces of oxygen, the process stage according to the invention for removing residual oxygen from the product gas stream is carried out. The residual oxygen may be disturbing insofar as he, z. B. in the case of butadiene, in downstream process steps can cause butadiene peroxide formation and can act as an initiator for polymerization reactions.

Preferably, the oxygen removal in the case of butadiene production is carried out immediately after the oxidative dehydrogenation.

The free-oxygen C4 hydrocarbon stream may contain sufficient free hydrogen to react with the free oxygen. Missing amounts or the total amount of free hydrogen needed can be added to the hydrocarbon stream. In this reaction, the free oxygen can be reacted with the free hydrogen, so that only a very small proportion of the hydrocarbon is reacted with the oxygen. Despite the presence of hydrogen hardly hydrogenation of the hydrocarbon occurs according to Invention.

In an alternative embodiment, the hydrocarbon stream containing free oxygen contains no free hydrogen and no free hydrogen is also added to it. In In this case, the free oxygen can be reacted with the hydrocarbon stream containing the free oxygen-containing hydrocarbon or with added methanol, natural gas and / or synthesis gas as a reducing agent. The process can be isothermal or adiabatic. Advantage of the implementation of the hydrogen is the formation of water as a reaction product. The water formed can easily be separated by condensation.

In addition, a low reaction pressure may be advantageous since this results in a separate compression step z. B. can be avoided after the oxidative dehydrogenation. A lower reaction pressure allows less expensive reactor production and is advantageous for safety reasons.

The process according to the invention is therefore preferably carried out at an absolute pressure of 0.5 to 20 bar, preferably 0.9 to 10 bar, particularly preferably 0.9 to 5 bar, more preferably 0.9 to 3 bar, in particular 0.9 up to 2 bar.

The reaction preferably takes place at an inlet temperature in the catalyst bed of from 300 to 450.degree. C., more preferably from 320 to 400.degree. This temperature is especially true at an oxygen content of 1 to 3.5 vol .-%. At higher oxygen levels (eg, 8% by volume), intercooling may be required to meet the temperature requirements.

The maximum temperature in the catalyst bed is preferably not more than 650 ° C. Preferably, it is in the range of 500 to 650 ° C, especially 580 to 650 ° C.

If the reaction temperature is too low, then z. B. butadiene can be hydrogenated. If it is too high, cracking processes can occur.

The reactor type is not restricted according to the invention. For example, the reaction can be carried out in a fluidized bed, in a tray oven, in a solid-tube or shell-and-tube reactor or in a plate-type heat exchanger. The fixed bed catalyst can be operated adiabatically on a large scale. The flow through the bed can be oriented both axially and radially. A radial reactor could be advantageous for large volume flows. Due to the high expected starting temperature (up to 600 ° C), an infeed from the outside to the inside can be advantageous. A flow from the inside out nevertheless results in smaller pressure losses. Monolith reactors can be advantageously used as an adiabatic reactor for reactions requiring little catalyst. Since the residual oxygen concentration should be low, a backmixed system should be avoided. Therefore, the reactor concept of the tubular reactor (fixed bed reactor) has proven itself. A cascading of backmixed fluidized bed reactors or the use of Horden reactors would also be conceivable.

By structuring a fixed bed of the catalyst with inert materials, the temperature control can be adjusted and the maximum temperature can be kept within the optimum range. If hydrogen is used in deficit in the process according to the invention, the reaction with hydrogen can serve to achieve a sufficiently high temperature for the necessary reaction between hydrocarbons and oxygen.

If no hydrogen or a hydrogen deficiency is used, the oxygen reacts predominantly with the most reactive molecule, for example butadiene. This leads to the formation of carbon oxides and water. Since the reaction of oxygen with the hydrocarbons at low temperature proceeds more slowly than with hydrogen, the hydrogen is first completely consumed.

A further embodiment of the invention is the implementation of this catalytic reaction together with an oxidative dehydrogenation in a reactor with 2 catalysts and, optionally, intermediate feed of the combustion gas after the Dehydrierschüttung.

According to the invention, the term "catalyst bed" means the area of a reactor in which the catalyst is present as a fixed-bed catalyst.This can be a catalyst bed, one or more catalyst monoliths or other structured packings.If necessary, the reactor used for the catalytic combustion contains The inlet temperature in the catalyst bed refers to the area of the catalyst bed into which the gas mixture to be reacted enters.

The catalytic combustion can be carried out on any suitable catalysts, as described, for example, in the above-mentioned prior art, in particular in WO 2006/061202. According to the invention, preference is given to using a catalyst which comprises at least one noble metal and / or at least one transition metal on a support.

Suitable noble metals are in particular Pt, Pd, Ir, Rh, Ru, Au and Ag and mixtures thereof.

Suitable transition metals are preferably those of groups 7 to 14 of the Periodic Table of the Elements, more preferably Mn, Fe, Ni, Co, Cu, Zn, Sn. Sn is particularly preferably used.

Platinum or a platinum-containing alloy is preferably used as the noble metal. Particularly preferred is the combination of platinum and tin as active metals. The proportion of platinum, based on the total catalyst, preferably 0.01 to 1 wt .-%, more preferably 0.02 to 0.5 wt .-%, in particular 0.05 to 0.2 wt .-% , Tin is likewise preferably present in an amount, based on the total catalyst, of from 0.01 to 1% by weight, more preferably from 0.02 to 0.5% by weight, in particular from 0.05 to 0.2% by weight. % used.

The weight ratio of tin to platinum is preferably 1: 4 to 4: 1, more preferably 1: 2 to 2: 1, in particular about 1: 1.

In addition to platinum and tin, alkali metal and / or alkaline earth metal compounds may optionally be used in amounts of <2% by weight, in particular <0.5% by weight, based on the total catalyst. Particularly preferably, the catalyst contains only platinum and tin as active metals.

As catalyst support, any suitable solid support materials can be used. Preferably, the carrier is α-alumina or zeolite A, especially α-alumina. It preferably has a BET surface area of 0.5 to 15 m ² / g, preferably 2 to 14 m ² / g, in particular 7 to 1 1 m ² / g. The carrier used is preferably a shaped body. Preferred geometries are, for example, tablets, ring tablets, spheres, cylinders, star strands or gear-shaped strands with diameters of 1 to 10 mm, preferably 2 to 6 mm. Particularly preferred are balls or cylinders, in particular balls. When using this catalyst, starting from a butadiene-containing C ₄ -

Hydrocarbon stream butadiene loss is suppressed and at the same time the residual oxygen is safely removed. By using the catalyst in the temperature range according to the invention, few secondary reactions occur and the reaction can be carried out with an excess of hydrogen, hydrogen deficiency or in the absence of hydrogen.

The reaction is alternatively carried out on a catalyst containing on zeolite A as a carrier 0.01 to 0.5 wt .-% of platinum, based on the catalyst, and optionally tin, wherein the weight ratio Sn: Pt is 0-10.

This catalyst preferably contains zeolite A as a carrier. In this case, based on the carrier, preferably at least 80 wt .-%, more preferably at least 90 wt .-%, in particular at least 95 wt .-% zeolite A in the carrier before. In particular, the support is constructed entirely of zeolite A.

Zeolite A is a synthetic, crystalline aluminosilicate and in its hydrated sodium form has the empirical formula Nai2 ((AlO2) i2 (SiO2) i2) x 27 H2O. The term "zeolite A" encompasses various variants of this compound, all of which have the same aluminosilicate lattice. However, instead of sodium ions, they may contain other ions such as potassium or calcium. Also low-water or anhydrous forms are counted according to the invention to zeolite A. Other designations are molecular sieve A, molecular sieve A, LTA (Linde type A), MS 5 A (with Ca), MS 4 A (with Na), NF3 A (K), Sasil ^®.

Zeolite A has a framework structure of AIC and SiC tetrahedra. They form a covalent lattice with cavities, which usually contain water. AIC and SiC tetrahedra are present in a ratio of 1: 1. In this case, aluminum and silicon atoms are alternately connected to one another via oxygen atoms. Overall, there is a negative charge, which is compensated by ionic compounds with cations such as sodium ions. As a spatial structure, zeolite A has a so-called sodalite cage.

This catalyst preferably contains 0.01 to 0.5 wt .-%, preferably 0.05 to 0.4 wt .-%, in particular 0.1 to 0.3 wt .-% of platinum, based on the catalyst. It may also contain tin, wherein the weight ratio Sn: Pt 0 to 10, preferably 0 to 7, particularly preferably 0 to 3. When using tin, the weight ratio Sn: zPt is preferably 0.5 to 10, particularly preferably 0.7 to 4, in particular 0.9 to 1.1. Especially preferred is a weight ratio Sn: Pt of 1: 1. In addition to platinum and tin, this catalyst may preferably also contain further metals, for example alkali metal and / or alkaline earth metal compounds, preferably in amounts of <2% by weight, in particular <0, 5 wt .-%, based on the catalyst. Particularly preferably, the catalyst contains only platinum and optionally tin as active metals. In the finished catalyst, the BET surface area is preferably 10 to 80 m ² / g, particularly preferably 15 to 50 m ² / g, in particular 20 to 40 m ² / g.

The catalyst can be used in any suitable form. Preferably, it is used as a shaped body having an average diameter in the range of 1 to 10 mm, more preferably 2 to 8 mm, in particular 2.5 to 5 mm. The molding may have any suitable shape, it may be in the form of a strand, tablet, granules, split or preferably in spherical form with the specified average diameter. Other possible shaped bodies are ring tablets, cylinders, star strands or gear-shaped strands. Alternatively, said catalysts may be present as a monolith, wherein the monolith may comprise the catalyst as a washcoat on a support structure. This support structure can dictate the three-dimensional structure of the monolith. For example, the support structure may be constructed of cordierite. The proportion of washcoat in the entire monolith is preferably 0.5 to 5 g / in ³ .

The catalyst can be prepared by any suitable method. Preferably, it is obtained by impregnating the carrier with a solution of a platinum compound and optionally a tin compound and then dried and calcined. For example, platinum nitrate can be used as an aqueous solution for impregnating the carrier. The impregnation may be followed by drying, preferably at 80 to 150 ° C. and calcination, preferably at 200 to 500 ° C. It is preferably dried for a period in the range of 1 to 100 hours, more preferably 5 to 20 hours. The calcining is preferably carried out for a period of 1 to 20 hours, more preferably 2 to 10 hours.

To the actual catalyst production, a silylation can connect, for example by use of an aqueous colloidal dispersion of very small silica particles such as are available for example under the name Ludo ^® from Helm AG. This silylation can also be effected by impregnation with subsequent drying and calcination, as described above. The catalyst used according to the invention is distinguished in particular by a long-term stability, especially in the butane or butene dehydrogenation for the preparation of butadiene, wherein free oxygen is to be separated from the butadiene-containing product stream.

The catalyst used with preference has the advantage that it catalyzes, in particular, the reaction of hydrogen with oxygen, without there being any appreciable conversion of hydrocarbon with the free oxygen. In the case of the production of butadiene from butene or n-butane preferably does not result in a reaction of the butadiene with the free oxygen. Another advantage of using the catalyst of the invention is its stability to water in the feed, especially at 5 to 30% water in the feed.

For example, the production of butadiene from n-butane by feeding a feed stream containing n-butane in at least a first dehydrogenation zone and non-oxidative catalytic dehydrogenation of n-butane, wherein a product gas stream containing n-butane, 1-butene, 2 Butene, butadiene, hydrogen, low-boiling minor components, optionally carbon oxides and optionally water vapor is obtained. This product gas stream, together with an oxygen-containing gas, is passed into at least one second dehydrogenation zone for oxidative dehydrogenation to give a product gas stream containing n-butane, 2-butenes, butadienes, low-boiling minor components, carbon oxides and water vapor.

The invention is explained in more detail by the following examples. Examples

Catalytic oxygen removal was studied in an adiabatic reactor. Fig. 1 shows schematically the structure of the reactor whose dimensions are listed as follows: Length: 200 cm

Outer diameter: 2.5 cm

Wall thickness: 0.2 cm

Inner diameter: 2.1 cm

Outer diameter of the thermowell: 3.1 mm

Material: Steel (1 .4841)

In this case, in FIG. 1:

In: inert material

Co: copper block

Ca: catalyst

He: heating element in zone 3

Th: thermocouple

The reactor consists essentially of 2 zones. In the first heating zone, an inert bed is preheated to the desired temperature by an accompanying heating. Between the accompanying heating and the reactor is a copper block to allow a homogeneous distribution of heat in the first zone. The catalyst is introduced into the second zone. The reactor is framed in this section with insulation material to keep heat loss low. Two companion heaters have also been placed in the insulation to allow for temperature compensation between the interior of the reactor and the exterior.

In the middle of the reactor is a thermowell, in which thermocouples are placed. These thermocouples enable the recording of an axial temperature profile in the catalytic bed. A pneumatically operated, multiple thermocouple with four measuring points was used to determine the temperature profiles with a resolution of 2 cm in the catalyst bed. The catalyst bed was packed between an inert material (steatite) which served as a guard bed.

The reactor is flowed through from top to bottom with the gas.

The reactor was operated under the following typical conditions

Catalyst volume: 75 ml_

Catalyst mass: 54.6 g

Inlet temperature: 150 - 450 ° C

Outlet pressure: 1 .5 - 2.5 bara

GHSV: 10000 - 12 000 Nlgas IKat-1 h-1

Oxygen entry concentration: 3 vol.%

Ratio hydrogen / oxygen: 2.1 - 2.5% Hydrocarbon concentration: approx. 20% by volume

Water concentration: approx. 13 Vol .-%

Rest: nitrogen

Preparation of the catalyst

The catalyst contains 99.7% by weight of zeolite A, molecular sieve 3A (of Roth GmbH), 0.3 mm of type 562 C, bead form, spheres with a diameter in the range of 2.5 to 5 mm, and 0, 3% by weight of platinum.

For the preparation, 1000 g of molecular sieve and 5.2 g of platinum nitrate are used. Platinum nitrate is dissolved in water and the solution is made up to 460 ml of total solution. The carrier is then soaked to 100% of its water absorption. For this purpose, the molecular sieve was distributed in two porcelain dishes, separated the impregnation solution, and it was mixed well.

It was followed by drying for 16 hours at 120 ° C in a circulating air T rockenschrank and calcination for four hours at 400 ° C in a muffle furnace. For silylation, the catalyst thus obtained was placed in a beaker and a solution of Ludox and water in the ratio 1:10 (final concentration 4 wt .-%) was prepared. The amount was calculated so that the catalyst in the beaker could be well covered. The mixture was stirred at regular intervals and filtered off after 40 minutes through a pleated filter. It was then followed by drying for 16 hours at 120 ° C in a convection oven and subsequent calcination for 4 hours at 400 ° C in a muffle furnace.

The elemental analysis showed a Pt content in the catalyst of 0.27 wt .-%.

The experiments with this catalyst show that the oxygen removal is a fast reaction and thus can be operated at high loads (up to 1 1000 Nlgas / IKat / h-1). The specification of 100 ppm can be fulfilled for inlet temperatures in the catalyst bed of equal size 290 ° C for GHSV of about 1 1000 h-1. If the temperature is too low (<290 ° C), the hydrocarbons present in the educt stream are significantly converted. For example, at 290 ° C, the conversion of total hydrocarbons is about 4%. The main products are butene (selectivity over 75%) and COx. If the temperature at the reactor inlet is increased, the specification is still around 02, but the conversion of the hydrocarbons decreases significantly (2% at about 350 ° C).

Temperature control:

Fig. 2 shows the determined axial temperature profiles for different inlet temperatures. The temperature is plotted in ° C over the length of the (catalyst) bed in cm. From zero length, catalyst is present. The oxygen content in the input stream was 3 vol.%, The molar Ratio of hydrogen to oxygen 2.6, the GHSV about 1 1000 r ¹ , the temperature of the heating sleeves about 496 ° C. In each experiment, the 02 specification was met and reflected by the temperature increase, which is almost the same for all experiments. It also corresponds to the adiabatic temperature increase. It is apparent that after half the bed length, the temperature increase has already reached 90% of the maximum temperature increase. Should the residual oxygen specification be less than 100 ppm, the GHSV could be further increased.

Without hydrogen:

As an alternative to the process with hydrogen, the oxygen content in the exhaust gas stream can be reduced by catalytic reaction with the hydrocarbons present in the gas. The oxygen will react predominantly with the most reactive molecule, i. in this case butadiene, and leads to the formation of CO2 and H2O. The reaction of O2 with the hydrocarbons is slower at low temperature than with hydrogen. Thus, this reaction should preferably be carried out at lower catalyst loadings and / or at a higher temperature (compared to the hydrogen mode). Nevertheless, a fast reaction of the O2 seems to inhibit soot formation, so that higher temperatures would be preferable. At higher temperatures, this reaction is comparably fast as the oxyhydrogen reaction and can similarly be carried out at high loads (experiment with GHSV = 10500 h-1, inlet temperature 400 ° C).

Hybrid mode: Should a gas flow contain H2 (e.g., from the BDH stage), it could be introduced to the 02 removal stage for 02 removal. Oxygen will react with hydrogen, preferably at lower temperatures. If hydrogen is in excess, the remainder of O2 reacts with the hydrocarbons. The reaction with H2 can also serve to achieve a sufficiently high temperature for the reaction between the hydrocarbons and oxygen (ignition). For example, depending on the catalyst used, the ODH stage is operated between 320 and 420 ° C. Should this stage be operated at the low temperature, a heat exchanger between the ODH and O2 removal stages would be advantageous to bring the gas mixture to the desired temperature. However, experience shows that a mixture of butadiene and oxygen tends to form polymer-like deposits for temperatures higher than 250 ° C. For this reason, a rapid increase in temperature with rapid degradation of O2 is desirable. For this purpose, hydrogen can be metered in so that a sufficiently high temperature for the catalytic combustion of butadiene with O2 is achieved. The operation of an additional heat exchanger is spared the process, and the risk of clogging of the system is thereby reduced.

Example 2

Oxygen removal with hydrogen: The catalytic oxygen removal was investigated in a wall-cooled reactor. Fig. 3 shows schematically the structure of the reactor whose dimensions are listed as follows:

Length: 200 cm

Outer diameter: 2.5 cm

Wall thickness: 0.2 cm

Inner diameter: 2.1 cm

Outer diameter of the thermowell 3.1 mm

Material: Steel (1 .4841)

In this case, in FIG. 3:

HA: main stream

Ku: copper blocks

In: inert packing

Ka: catalyst

Re: reactor wall

Th: thermowell

Be: heating

Wä: thermal insulation

From: Exhaust gas flow The reactor consists of 3 heating zones and is provided with copper blocks so that a uniform temperature field can be set on the reactor wall. In the first heating zone, an inert bed is preheated to the desired temperature. In the second heating zone, the wall temperature of the catalytic bed is adjusted. In the middle of the reactor is a thermowell, in which thermocouples are placed. These thermocouples enable the recording of an axial temperature profile in the catalytic bed. A pneumatically operated, multiple thermocouple with four measuring points was used to determine the temperature profiles with a resolution of 2 cm in the catalyst bed. The catalyst bed was packed between an inert material (steatite) which served as a guard bed.

The reactor was operated under the following typical conditions:

Catalyst volume: 0.05-0.1 I

Catalyst mass: 0.010 - 0.1 kg

Inlet temperature: 150 - 450 ° C

Outlet pressure: 1 .5 - 2.5 bara

GHSV: 2000 - 12 000 Nlgas IKat-1 h-1 Oxygen entry concentration: 3 vol.%

Ratio hydrogen / oxygen: 2.1 - 2.5%

Hydrocarbon concentration: approx. 20 vol.

Water concentration: approx. 13 vol.

Rest: nitrogen

Preparation of the catalyst:

The catalyst contains 99.7% by weight of zeolite A, molecular sieve 3A (of Roth GmbH), 0.3 mm of type 562 C, bead form, spheres with a diameter in the range of 2.5 to 5 mm, and 0, 3 wt .-% platinum and was prepared as described in Example 1.

The experiments with this catalyst showed that the oxygen removal is a fast reaction and thus can be operated at high loads (up to 1 1000 Nlgas / IKat / h-1). The specification of 100 ppm can be fulfilled for inlet temperatures in the catalyst bed equal to or greater than 320 ° C. If the temperature is too low (<300 ° C), up to 12% of the hydrocarbons are converted. These conditions are predominantly a hydrogenation of butadiene to butene. For this reason, an isothermal driving at low temperature level is less desirable. If the inlet temperature is raised above 380 ° C, H2 reacts more than 90% with O2, so that the butadiene conversion is kept low. This low conversion is further promoted by choosing the excess of hydrogen low and keeping the residence time on the catalyst short.

Temperature control:

Suitable wall-cooled reactors are common in the chemical industry. In the case of fixed bed reactors, preference may be given to tube bundle reactors in which the tubes are filled with the catalyst and the heat resulting from the reaction is removed by means of a cooling medium in the outer space. For temperatures greater than 300 ° C Salzbadreaktoren are particularly suitable. However, the salt is subject to a gradual

Decomposition usually for temperatures above 460 ° C. This defines a temperature window in which the method is preferably operated.

4 shows the determined axial temperature profiles for different wall temperatures. The temperature is recorded in ° C over the length of the (catalyst) bed in cm. From length 0 catalyst is present. The oxygen content in the input stream was 3.1% by volume, the molar ratio of hydrogen to oxygen was 2.1, the GHSV about 10500 r ¹ . Furthermore, the respective wall temperature is entered, as well as the maximum temperature difference between the inlet temperature in the catalyst bed and maximum temperature. In each experiment, the 02 specification was met and the hydrocarbon conversion was less than 2%.

Regardless of the wall temperature, the temperature increase is defined as the difference between the maximum temperature at the hot spot and the inlet temperature in the catalytic converter. see bed, almost equal and corresponds in this case more than half of the adiabatic temperature increase (3% O2 correspond to about 250 K adiabatic temperature increase). The location of the hot spot remains unchanged for all experiments, suggesting that higher temperatures do not significantly accelerate the reaction between H2 and O2. This means that higher temperatures (> 380 ° C) do not significantly affect the reaction process, and a strict control of the height of the hot spot is not urgently required. Nevertheless, the maximum temperature required is preferably set at 600 ° C in order not to thermally stress the catalyst. The height of the hot spot can be influenced by various parameters, such as the flow rate, a dilution of the catalyst and the tube diameter. It is also known that the heat transfer is subject to a resistance between the bed and the wall, so that temperature gradients of more than 30 ° C are common. Thus, the temperature at the wall is much lower than at the hot spot, which is an advantage for the salt of a brine bath heat carrier. Thus, the use of a salt bath reactor is possible.

Depending on the catalyst used, the ODH stage is preferably operated at temperatures between 320 and 420 ° C. in a salt bath reactor. Previous experiments have shown that the hydrogen is only partially reacted on the ODH catalyst. It is thus possible to couple the oxygen removal with the ODH stage by metering in hydrogen at the reactor inlet. Since the 02 removal with H2 is a fast reaction, the extension of the tubes is less than 1 m. To reach the minimum temperature of 380 ° C, a two-zone salt bath reactor (with two different salt bath temperatures) can be used. Without hydrogen:

As an alternative to the hydrogen process, the oxygen content in the exhaust gas stream can be reduced by catalytic reaction with the hydrocarbons present in the gas. The oxygen will react predominantly with the most reactive molecule, i. in this case butadiene, and leads to the formation of CO2 and H2O. The reaction of O2 with the hydrocarbons is slower at low temperature than with hydrogen. This reaction should be carried out at lower catalyst loadings and / or at a higher temperature (in comparison to the H2 mode of operation). Nevertheless, a fast reaction of the O2 seems to inhibit soot formation, so that higher temperatures would be preferable. At higher temperatures, this reaction is as fast as the oxyhydrogen gas reaction and, similarly, can be carried out in a wall-cooled, high-load reactor.

(Experiment with GHSV = 10500 h-1, inlet temperature 400 ° C). The sequence of the ODH stage and the O 2 removal in a single salt bath reactor, in which the minimum required temperature prevails in the second zone, is possible. Hybrid driving style:

Should a gas stream contain H2 (e.g., from the BDH stage), it could be introduced to the 02 removal stage for O2 removal. Oxygen will react with hydrogen, preferably at lower temperatures. If hydrogen is in excess, the remainder of O2 reacts with the hydrocarbons. The reaction with H2 can also serve to achieve a sufficiently high temperature for the reaction between the hydrocarbons and oxygen (ignition).

Example 3: Influence of the temperature on the selectivities

In the reactor as described in Example 1, a gas stream consisting of 3% O 2, 14% butadiene, 5.5% butane, 12% steam with nitrogen as the remainder is introduced into the reactor. The catalyst bed consists of 75 ml of undiluted catalyst (DA301 with Pd). The total volume flow is 800 Nl / h. The inlet temperature in the catalyst bed was varied between 170 and 350 ° C. The form was kept constant at 1 .5 bara. The temperature profile along the catalyst bed is recorded for each setting. The hot-spot temperature is between 400 ° C and 650 ° C depending on the test. Hydrogen is presented in excess with a molar ratio H2: 02 between 2.1 and 2.8.

The total butane-butene conversion (C4 conversion) is determined as follows: _v _ Butctn + ™ buta-diene

Where ri _| the molar flow at the exit and the molar flow at the entrance of the reactor from the components i are.

The yields of butene, CO and CO 2 are based on the educts butane and butene and are calculated as follows.

Butyl-diene

Where i is CO, CO 2 or butene, and μ _{έ is} the stoichiometric coefficient. μ _ί = 1 for butene and μ _ί = 4 for CO and C02.

The selectivities are then calculated as follows:

The results were as follows (Figure 5): The C ₄ conversion was 7.2 mol% at 270 ° C., of which 6.6 mol% were hydrogenation products and 0.6 mol%.

At 330 C, the C ₄ conversion was 3.2 mol%, of which 2.3 mol% were hydrogenation products and 0.9 mol%.

At 350 ° C., the C ₄ conversion was 2.0 mol%, of which 0.7 mol% were hydrogenation products and 0.3 mol%.

If the inlet temperature in the catalyst bed is increased from 270 ° C to 350 ° C, the C4 conversion decreases from about 7% to 2%. At low temperature (<300 ° C), the selectivity to butene is> 75%, indicating hydrogenation of butadiene. This hydrogenation decreases significantly in favor of the combustion products with increasing temperature.

Example 4:

Preparation of the Catalyst: The catalyst contains 99.8% by weight of Al 2 O 3 (Axes), SPH-512, bead form, spheres having a diameter in the range from 2.5 to 5 mm, and 0.1% by weight. Platinum and 0.1% by weight Sn.

For preparation, 100 g of carrier, 0.25 g of hexachloroplatinic acid (IV) hydrate and 0.19 g of SnCl2-2H20 g are used. The hexachloroplatinic acid is dissolved in 10 ml of water. The

SnCl ₂ -2H ₂ 0 is dissolved in a mixture of 17.5 ml of 65.0% by weight of HN0 ₃ and 17.5 ml H ₂ 0th The two solutions are mixed together with stirring and supplemented with H2O to 100 ml.

The support is then impregnated with the solution, dried at 120 ° C for 30 min and then calcined at 500 ° C for 3 hrs.

Example 4.1 with the above-mentioned catalyst:

In the reactor as described in Example 1, a gas stream consisting of 3% by volume O 2, 8.0% by volume of butadiene, 2.0% butane, 20.0% steam with nitrogen as the remainder is introduced into the reactor. The catalyst bed consists of 75 mL undiluted catalyst. The total volume flow is 800 Nl / h. The form was kept constant at 1 .5 bara. Hydrogen is metered into the line before the reactor into the gas stream (with a H2: B2 volume ratio of 2.5: 1). Under the test conditions no oxygen could be detected during the whole operating time of 200 hours. The inlet temperature of the catalyst bed is 390 ° C and the hot-spot temperature is 580 ° C. The butadiene loss is 1, 16 mol .-%. 0.56 mol% thereof refers to COx (CO and CO 2), and the remainder (0.60 mol%) refers to hydrogenation products (butenes) (see Fig. 6).

Example 4.2 with the above catalyst:

In the reactor as described in Example 1, a gas stream consisting of 3 vol.% O 2, 8.0 vol.% Butadiene, 2.0% butane, 20.0% steam with nitrogen remainder is introduced into the reactor. The catalyst bed consists of 75 mL undiluted catalyst. The total volume flow is 800 Nl / h. The form was kept constant at 1 .5 bara. Hydrogen is metered into the line before the reactor into the gas stream (with a H2: 02 volume ratio of 2.0: 1).

Under the test conditions no oxygen could be detected during the whole operating time of 200 hours. The inlet temperature of the catalyst bed is 395 C and the hot-spot temperature is 600 ° C. The loss of butadiene is 1, 32 mol .-%. 0.65 mol% refer to COx (CO and CO 2), and the remainder (0.67 mol%) refers to hydrogenation products (butenes) (see FIG. 6).

Example 4.3 with the above-mentioned catalyst: In the reactor as described in Example 1 is a gas stream consisting of 3 vol.% O 2, 8.0 vol.% Butadiene, 2.0% butane, 20.0% water vapor with nitrogen in introduced the reactor. The catalyst bed consists of 75 mL undiluted catalyst. The total volume flow is 800 Nl / h. The form was kept constant at 1 .5 bara. Hydrogen is metered into the line before the reactor into the gas stream (with an H2: B2 volume ratio of 1.5: 1).

Under the test conditions no oxygen could be detected during the whole operating time of 200 hours. The inlet temperature of the catalyst bed is 401 ° C and the hot-spot temperature is 596 ° C. The butadiene loss is 1, 76 mol .-%. 1.16 mol% refer to CO _x (CO and CO ₂ ), and the remainder (0.60 mol%) refers to hydrogenation products (butenes) (see FIG. 6).

Claims

Patent claims

Process for removing oxygen from a C ₄ hydrocarbon stream containing free oxygen by catalytic combustion, in which the hydrocarbon stream containing free oxygen is reacted on a catalyst bed by catalytic combustion in the presence or absence of free hydrogen to obtain an oxygen-depleted hydrocarbon stream, characterized in that the catalytic combustion is carried out continuously, the inlet temperature in the catalyst bed is at least 300 ° C and the maximum temperature in the catalyst bed is not more than 700 ° C.

Method according to claim 1, characterized in that the hydrocarbon stream used contains 0.5 to 8% by volume of free oxygen.

Method according to claim 1 or 2, characterized in that the hydrocarbon stream containing free oxygen contains a sufficient amount of free hydrogen for reaction with the free oxygen and / or this is added to it and the free oxygen is reacted with the free hydrogen.

Method according to claim 1 or 2, characterized in that the hydrocarbon stream containing free oxygen does not contain any free hydrogen and no free hydrogen is added to it.

Method according to claim 4, characterized in that the free oxygen is reacted with hydrocarbon contained in the free oxygen-containing hydrocarbon stream or with added methanol, natural gas and / or synthesis gas as a reducing agent.

Process according to one of claims 1 to 5, characterized in that in the C ₄ hydrocarbon stream at least 80% by volume of the hydrocarbons are C ₄ hydrocarbons.

Method according to one of claims 1 to 6, characterized in that the inlet temperature in the catalyst bed is 300 to 450°C.

Method according to one of claims 1 to 7, characterized in that the maximum temperature in the catalyst bed is not more than 700 °C.

Method according to one of claims 1 to 8, characterized in that the catalytic

cal combustion is carried out at a pressure in the range of 0.5 to 20 bar absolute.

10. The method according to any one of claims 1 to 9, characterized in that a catalyst containing at least one noble metal and / or at least one transition metal on a support is used.

1 1 . Method according to claim 10, characterized in that the noble metal in the catalyst is platinum or a platinum-containing alloy.

12. The method according to claim 10 or 1 1, characterized in that the transition metal in the catalyst is Mn, Fe, Ni, Co, Cu, Zn, Sn or a mixture thereof.

13. The method according to any one of claims 10 to 12, characterized in that the carrier is α-aluminum oxide or zeolite A.

14. The method according to any one of claims 10 to 13, characterized in that the catalyst is present as a shaped body with an average diameter in the range of 1 to 10 mm or as a monolith, wherein the monolith can have the catalyst as a washcoat on a support structure.