EP2170495A1

EP2170495A1 - Method for producing chlorine by multi step adiabatic gas phase oxidation

Info

Publication number: EP2170495A1
Application number: EP08784551A
Authority: EP
Inventors: Ralph Schellen; Stephan Schubert; Leslaw Mleczko; Aurel Wolf; Oliver Felix-Karl SCHLÜTER
Original assignee: Bayer Technology Services GmbH
Current assignee: Bayer Intellectual Property GmbH
Priority date: 2007-07-13
Filing date: 2008-06-26
Publication date: 2010-04-07
Also published as: CN101687160A; JP2010533113A; US20100260660A1; WO2009010168A1

Abstract

The invention relates to the production of chlorine, by catalytic gas phase oxidation of hydrogen chloride with oxygen, wherein the reaction is carried out in 18 to 60 serially arranged catalyst beds under adiabatic conditions and a reactor system for carrying out said method.

Description

Process for the production of chlorine by multi-stage adiabatic gas-phase oxidation

The present invention relates to a process for the production of chlorine by catalytic gas phase oxidation of hydrogen chloride with oxygen, wherein the reaction is carried out on 18 to 60 catalyst beds connected in series under adiabatic conditions, and a reactor system for carrying out the process.

The process of catalytic hydrogen chloride oxidation with oxygen in an exothermic equilibrium reaction, developed by Deacon in 1868, was at the beginning of technical chlorine chemistry:

4 HCl + O ₂ => 2 Cl ₂ + 2 H ₂ O

However, the technical application of the Deacon process was pushed to the background by chloralkali electrolysis. Nearly all of chlorine was produced by electrolysis of aqueous saline solutions. However, the attractiveness of the Deacon process has increased again recently, as the worldwide demand for chlorine is growing faster than the demand for caustic soda, a by-product of NaCl electrolysis. This development is countered by the process for the production of chlorine by oxidation of hydrogen chloride, which is decoupled from the sodium hydroxide production. In addition, the precursor hydrogen chloride is easily accessible; it is produced in large quantities, for example in phosgenation reactions, such as in isocyanate production, as by-product.

The removal and use of the heat of reaction is an important issue in the practice of the Deacon process. An uncontrolled increase in temperature, which could be from 600 ° to 900 ° C. from the beginning to the end of the Deacon reaction, would on the one hand lead to permanent damage to the catalyst, and on the other hand there would be an unfavorable shift in the reaction equilibrium in the direction of the educts at high temperatures a corresponding deterioration of the yield. It is therefore advantageous to maintain the temperature of the catalyst bed in the range of 150 to 600 ⁰ C in the course of the process. In established processes, therefore, the catalyst is used in the form of a fluidized, thermally stabilized bed. According to EP 0 251 731 A2, the catalyst bed is tempered via the outer wall, and according to DE 10 2004 006 610 A1, the fluidized bed is heated by means of a heat exchanger arranged in the bed. The effective heat removal of this process faces problems of non-uniform residence time distribution and catalyst wear, both of which result in a loss of revenue.

A narrow residence time distribution and low catalyst abrasion are possible in reactors with stationary catalyst beds. However, problems arise in such reactors with the thermostating of the catalyst beds. In general, therefore, thermostated tube bundle reactors are used which, especially in the case of large reactors, have a very complicated cooling circuit (WO 2004/052776 A1).

In order to improve the heat removal from the catalyst bed, the use of a fixed bed catalyst made of ruthenium oxide supported on titanium oxide is proposed in the R & D report "Sumitomo Kagaku", vol 2004-100 In addition to the high catalyst activity, the good thermal conductivity of the catalyst system is mentioned as an advantage Since the heat conductivity of the bed remains low even with a high thermal conductivity within the catalyst pellets, the heat dissipation is not significantly improved by this measure.

EP 1 170 250 A1 has proposed the use of catalyst fillings in tube bundle reactors which have different activities in different regions of the cooled contact tubes. As a result, the progress of the reaction is slowed down so much that the resulting heat of reaction can be more easily removed via the wall of the catalyst tubes. A similar result should be achieved by the targeted dilution of the catalyst bed with inert material. A disadvantage of these solutions is that two or more catalyst systems must be developed and used in the catalyst tubes or that by using inert material, the reactor capacity is impaired.

In the published patent applications WO 2004/037718 and WO 2004/014845, the possibility of an adiabatic catalytic hydrogen chloride

Oxidation mentioned in addition to the preferred isothermal process. Concrete execution However, forms of adiabatic hydrogen chloride oxidation are not described. It thus remains completely unclear how the heat of reaction of the exothermic reaction can be dissipated in a completely adiabatic mode of operation of the overall process and damage to the catalyst can be avoided. In fact, however, the hydrogen chloride oxidation according to these documents is carried out isothermally as a fixed-bed process in tube bundle reactors, which, as already mentioned, require cooling which is extremely complicated to control. In principle, all described tube bundle reactors are very complex and cause high investment costs. With the size rapidly increasing problems in terms of mechanical strength and uniform thermostating the catalyst bed make large aggregates of such types uneconomical.

It would therefore be advantageous to provide a simple process that can be carried out in a simple reactor without complex system for heat management in the reactor. Such reactors would be easily scaled up and are inexpensive and robust in all sizes. The reaction enthalpy is reflected quantitatively in this reactor type in the temperature difference between educt and product gas.

For the exothermic gas phase oxidation of hydrogen chloride with an oxygen-containing gas stream, neither the use of such reactors is yet described, nor suitable catalysts and suitable methods are shown.

The catalysts used initially for the Deacon process, for example supported catalysts with the active composition CuCl ₂ , had only a low activity. Although the activity could be increased by increasing the reaction temperature, it was disadvantageous that the volatility of the active components at high temperature led to rapid deactivation of the catalyst. The oxidation of hydrogen chloride to chlorine is also an equilibrium reaction. The position of the equilibrium shifts with increasing temperature to the detriment of the desired end product.

Usually, therefore, catalysts with the highest possible activity are used, which allow the reaction to proceed at low temperature. Known highly active catalysts are based on ruthenium. DE-A 197 48 299 describes supported catalysts with the active material ruthenium oxide or ruthenium mixed oxide. The content of ruthenium oxide is 0.1 wt .-% to 20 wt .-% and the average Particle diameter of ruthenium oxide 1.0 nra to 10.0 nm. The reaction is carried out at a temperature between 90 ° C and 150 ° C. Further supported catalysts based on ruthenium are known from DE-A 197 34 412: ruthenium chloride catalysts containing at least one compound of titanium oxide or zirconium oxide, ruthenium-carbonyl complexes, ruthenium salts of inorganic acids, ruthenium-nitosyl complexes, ruthenium-amine complexes , Ruthenium complexes of organic amines or ruthenium-acetylacetonate complexes. The reaction is carried out at a temperature between 100 ° C and 500 ° C, preferably 200 ° C and 380 ° C. In the two applications DE-A 197 48 299 and DE-A 197 34 412, the catalyst is used in a fixed bed or in a fluidized bed. The oxygen source used is air or pure oxygen. However, the Deacon reaction remains an exothermic reaction and temperature control is also required in the application of such highly active catalysts.

It was therefore an object to provide a process for the catalytic oxidation of hydrogen chloride to chlorine, which can be carried out in a simple reactor without a complex system for heat retention in the reactor.

The inventors of the present invention have surprisingly found that it is possible to achieve the objects described above by carrying out the reaction on 18 to 60 catalyst beds arranged in series under adiabatic conditions.

The process gas may in addition to oxygen and hydrogen chloride still have minor components, eg. As nitrogen, carbon dioxide, carbon monoxide or water. The hydrogen chloride can upstream production process, eg. As for the production of polyisocyanates, originate and other impurities, eg. B. phosgene.

According to the invention, carrying out the process under adiabatic conditions on the catalyst beds means that substantially no heat is supplied to the catalyst from the outside in the respective catalyst beds nor is heat removed (with the exception of the heat which is supplied or removed by the reaction gas entering or leaving). , Technically, this is achieved by insulating the catalyst beds in a conventional manner. The individual catalyst beds are operated adiabatically, so they are in particular no means of heat dissipation in them are provided. Considering the process as a whole, the invention also includes the case in which the heat of reaction is removed, for example, by means of heat exchangers connected between the individual catalyst beds.

The advantages of the adiabatic driving method according to the invention of the 18 to 60 catalytic beds connected one behind the other compared to the conventional isothermal mode of operation are, above all, that no means for heat removal must be provided in the catalyst beds, which entails a considerable simplification of the construction. This results in particular simplifications in the manufacture of the reactor and in the scalability of the process and an increase in reaction conversions.

Under catalyst bed is here an arrangement of the catalyst in all known forms, e.g. Fixed bed, fluidized bed or fluidized bed understood. Preferred is a fixed bed arrangement. This comprises a catalyst bed in the true sense, d. H. loose, supported or unsupported catalyst in any form and in the form of suitable packings:

The term catalyst bed as used herein also encompasses contiguous areas of suitable packages on a support material or structured catalyst supports. These would be e.g. to be coated ceramic honeycomb carrier with comparatively high geometric surfaces or corrugated layers of metal wire mesh on which, for example, catalyst granules is immobilized.

Stationary catalyst beds are preferably used in the new process.

In a preferred embodiment of the method according to the invention, the reaction is carried out at 20 to 40, preferably 22 to 30 consecutive Katal ysatorbetten.

A preferred further embodiment of the method is characterized in that the process gas mixture emerging from at least one catalyst bed is subsequently passed over at least one heat exchanger arranged downstream of the catalyst bed.

In a particularly preferred further embodiment of the method is located after each catalyst bed at least one, preferably a heat exchanger, through which the exiting process gas mixture is passed. In a preferred embodiment, at least one heat exchanger is located behind at least one catalyst bed. Particularly preferably, at least one, more preferably in each case exactly one heat exchanger is located behind each of the catalyst beds, via which the gas mixture emerging from the catalyst bed is passed.

The catalyst beds can either be arranged in a reactor or arranged divided into several reactors. The arrangement of the catalyst beds in a reactor leads to a reduction in the number of apparatuses used.

In addition, individual ones of the series catalyst beds can be independently replaced or supplemented by one or more catalyst beds in parallel. The use of catalyst beds connected in parallel allows in particular their replacement or supplementation during ongoing continuous operation of the process.

However, the process according to the invention preferably has 18 to 60 catalyst beds connected in series. Parallel and successively connected catalyst beds can in particular also be combined with one another. However, the process according to the invention particularly preferably has exclusively catalyst beds connected in series.

If parallel-connected catalyst beds are used, then in particular a maximum of 5, preferably 3, more preferably a maximum of 2 process beds arranged in series are connected in parallel.

The reactors which are preferably used in the process according to the invention can consist of simple containers with one or more thermally insulated catalyst beds, as described, for example, in Ullmann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, VoI B4, pages 95-104, pages 210-216) become. This means that it is possible, for example, to use single-stage or multistage fixed-bed reactors, radial-flow reactors or even shallow-bed reactors., However, tube bundle reactors are preferably not used because of the disadvantages described above, since heat is removed from the catalyst beds according to the invention , Such reactor types for receiving the catalyst beds are also unnecessary. The catalysts or the catalyst beds thereof are applied in a manner known per se to or between gas-permeable walls of the reactor. Especially with thin catalyst beds, technical devices for uniform gas distribution are installed above, below or above and below the catalyst beds. These may be perforated plates, bubble-cap trays, valve trays, or other internals which cause uniform entry of the gas into the catalyst bed by producing a small but uniform pressure drop.

The empty tube velocity of the gas in the catalyst bed is preferably from 0.1 to 10 m / s in the case of the embodiment using a fixed bed.

In a particular embodiment of the process according to the invention, a molar ratio of between 0.25 and 10 equivalents of oxygen per equivalent of hydrogen chloride before entry into the catalyst bed is preferably used. By increasing the ratio of equivalents of oxygen per equivalent of hydrogen chloride, on the one hand, the reaction can be accelerated and thus the space-time yield (produced amount of chlorine per reactor volume) can be increased; on the other hand, the equilibrium of the reaction is shifted positively toward the products.

In a further particularly preferred embodiment of the method, the inlet temperature of the material entering a first catalyst bed gas mixture from 150 to 630 ° C, preferably 200-480 ⁰ C.

The hydrogen chloride and oxygen-containing feed gas stream can also be fed preferably only in front of the first catalyst bed. This has the advantage that the entire feed gas stream can be used for the absorption and removal of the heat of reaction in all catalyst beds. However, it is also possible to meter in hydrogen chloride and / or oxygen into the gas stream before one or more of the catalyst beds following the first catalyst bed as required. In addition, the temperature of the reaction can be controlled via the supply of gas between the catalyst beds used.

In a particularly preferred embodiment of the process according to the invention, the reaction gas is cooled after at least one of the catalyst beds used, more preferably after each of the catalyst beds used. This is what leads you to that

Reaction gas through one or more heat exchangers, which are behind the respective Catalyst beds are located. These may be the heat exchanger known to those skilled in the art, such as, for example, tube bundle, plate ring groove, spiral, finned tube, micro heat exchanger. In a particular embodiment of the method, steam is generated on cooling the product gas at the heat exchangers.

In a preferred embodiment of the process, the catalyst beds connected in series are operated at increasing or decreasing average temperature from catalyst bed to catalyst bed. This means that the temperature of catalyst bed to catalyst bed can be both increased and decreased within a sequence of catalyst beds. Thus, it may be particularly advantageous to first increase the average temperature from catalyst bed to catalyst bed to increase the catalyst activity, and then to lower the average temperature in the following last catalyst beds again to shift the equilibrium. This can be set, for example, via the control of the heat exchangers connected between the catalyst beds. Further options for setting the average temperature are described below.

In a preferred step subsequent to the new process, the chlorine formed is separated off. The separation step usually comprises several stages, namely the separation and, if appropriate, recycling of unreacted hydrogen chloride from the product gas stream of the catalytic hydrogen chloride oxidation, drying of the obtained, essentially chlorine and oxygen-containing stream and the separation of chlorine from the dried stream.

The separation of unreacted hydrogen chloride and water vapor formed can be carried out by condensation of aqueous hydrochloric acid from the product gas stream of hydrogen chloride oxidation by cooling. Hydrogen chloride can also be absorbed in dilute hydrochloric acid or water.

In a preferred embodiment of the method unreacted hydrogen chloride and / or oxygen after removal of chlorine and water from the product stream and after branching off a small amount of gas to keep constant gaseous components, the z. T. be entrained with the starting materials, the reaction again fed. The recirculated hydrogen chloride and / or oxygen are recycled in front of one or more of the catalyst beds and before, if necessary brought back to the inlet temperature by means of a heat exchanger. Advantageously, the cooling of the product gas and the warming-up of the recirculated hydrogen chloride and / or oxygen are carried out by passing the gas streams in counterflow through heat exchangers to one another.

The new process is preferably operated at a pressure of 1 to 30 bar, preferably from 1 to 20 bar, more preferably from 1 to 15 bar.

The temperature of the educt gas mixture is preferably before each of the catalyst beds of 150 to 630 ° C, preferably from 200 to 480 ⁰ C, more preferably from 250 to 470 ° C. The gas mixture is preferably homogenized before entering the individual catalyst bed.

The thickness of the flow-through catalyst beds can be chosen the same or different, and is suitably 1 cm to 8 m, preferably 5 cm to 5 m, particularly preferably 30 cm to 2.5 m.

The catalyst is preferably used immobilized on a support. The catalyst preferably contains at least one of the following elements: copper, potassium, sodium, chromium, cerium, gold, bismuth, uranium, ruthenium, rhodium, platinum, and the elements of VIII. Subgroup of the Periodic Table of the Elements. These are preferably used as oxides, halides, or mixed oxides / halides, in particular chlorides or oxides / chlorides. These elements or compounds thereof can be used alone or in any combination.

Preferred compounds of these elements include copper chloride, copper oxide, potassium chloride, sodium chloride, chromium oxide, bismuth oxide, uranium oxide, ruthenium oxide, ruthenium chloride, ruthenium oxychloride, rhodium oxide.

Particularly preferably, the catalyst portion consists completely or partially of ruthenium and / or uranium or compounds thereof, more preferably the catalyst consists of halide and / or oxygen-containing uranium and / or ruthenium compounds.

Most preferably, all or part of the catalyst portion is uranium oxides such as UO ₃ , UO ₂ , UO or the non-stoichiometric phases resulting from mixtures of these species such as U ₃ O ₅ , U ₂ O ₅ , U ₃ O ₇ , U ₃ O ₈ , U ₄ θ ₉ . The carrier fraction may be wholly or partly composed of: titanium oxide, tin oxide, aluminum oxide, zirconium oxide, uranium oxide, vanadium oxide, ceria, chromium oxide, uranium oxide, silicon oxide, silica, carbon nanotubes or a mixture or compound of said substances, in particular mixed oxides such as silicon-aluminum oxides. Particularly preferred support materials are tin oxide, carbon nanotubes, uranium oxides such as UO ₃ , UO ₂ , UO or the non-stoichiometric phases resulting from mixtures of these species, such as U ₃ O ₅ , U ₂ O ₅ , U ₃ O ₇ , U ₃ O ₈ , U ₄ O ₉ ..

The ruthenium-supported catalysts can be obtained, for example, by impregnation of the support material with aqueous solutions of RuCl ₃ and optionally a promoter for doping. The shaping of the catalyst can take place after or preferably before the impregnation of the support material.

For doping the catalysts are suitable as promoters alkali metals such as lithium, sodium, potassium, rubidium and cesium, preferably lithium, sodium and potassium, more preferably potassium, alkaline earth metals such as magnesium, calcium, strontium and barium, preferably magnesium and calcium, particularly preferably magnesium, Rare earth metals such as scandium, yttrium, lanthanum, cerium, praseodymium and neodymium, preferably scandium, yttrium, lanthanum and cerium, more preferably lanthanum and cerium, or mixtures thereof.

The shaped bodies can then be dried at a temperature of 100 to 400 ° C., preferably 100 to 300 ° C., for example under a nitrogen, argon or air atmosphere, and optionally calcined. Preferably, the moldings are first dried at 100 to 150 ° C and then calcined at 200 to 400 ° C.

The temperature of the catalyst in the catalyst beds is suitably in a range of 150 ° C to 800 ° C, preferably 200 ⁰ C to 450 ° C, more preferably 250 ° C to 400 ° C. The control of the temperature in the catalyst beds is preferably carried out by at least one of the following measures:

Dimensioning of the catalyst beds,

Control of heat removal between the catalyst beds,

Addition of feed gases between the catalyst beds, Molar ratio of the educts,

Concentrations of the educts

Addition of inert gases, in particular nitrogen, carbon dioxide, before and / or between the catalyst beds

In principle, the catalysts or the supported catalysts may have any desired form, for. As balls, rods, Raschig rings or granules or tablets.

The composition of the catalysts in the catalyst beds used according to the invention may be identical or different. In a preferred embodiment, the same catalysts are used in each catalyst bed. However, it is also advantageous to use different catalysts in the individual catalyst beds. Thus, in particular in the first catalyst bed, when the concentration of the reaction products is still high, a less active catalyst can be used and in the further catalyst beds the activity of the catalyst can be increased from catalyst bed to catalyst bed. The control of the catalyst activity can also be carried out by dilution with inert materials or carrier material.

With the process according to the invention, 0.1 g / h to 10 g / h of chlorine, preferably 0.5 g / h to 5 g / h of chlorine, can be prepared per 1 g of catalyst. The inventive method is thus characterized by high space-time yields, combined with a reduction of the apparatus sizes and a simplification of the apparatus or reactors.

The educt for the process according to the invention is hydrogen chloride, which is e.g. is produced and adopted as by-product from the phosgenation of organic amines, especially diamines to isocyanates, in particular diisocyanates or the gas phase phosgenation of phenol to diphenyl carbonate.

Oxygen can be supplied as pure oxygen or preferably in the form of an oxygen-containing gas, in particular air.

The produced chlorine can be used, for example, for the production of phosgene and possibly recycled into connected production processes. In a further embodiment of the present invention, the process is conducted such that a continuous exchange of a fixed bed catalyst takes place.

In a further embodiment of the present invention, unreacted educt gases are recycled back to the process. Unreacted educt gases are in particular hydrogen chloride and oxygen. The process is therefore operated as a cyclic process.

The invention further provides a reactor system for reacting a gas containing hydrogen chloride and oxygen, at least containing feed lines for hydrogen chloride and oxygen or for a mixture of hydrogen chloride and oxygen and 18 to 60 thermally insulated catalyst beds connected in series. The reactor system may also comprise 20 to 40 or 22 to 30 catalyst beds.

The present invention will be explained with reference to FIGS. 1 and 2. Show it:

1 shows a process according to the invention with 18 catalyst beds divided into separate reactors

FIG. 2 shows a process according to the invention with 18 catalyst beds in an integrated reactor

FIG. 1 shows a method according to the invention with 18 catalyst beds divided into separate reactors. Before the first of the reactors, the educt gases (1, 2) are mixed to gas mixture (3) and fed to the reactor. The reactors each comprise a catalyst bed (20). The product gases of the reactors (4) are passed through heat exchangers (30). The heat exchanger (30) comprises feeds (5) and discharges (6) of cooling medium. In Fig. 1 it is symbolized that a repeat unit of reactor with catalyst bed (20) and heat exchanger (30) repeated 16 times in total, so that a total of 18 units are shown. The product gas mixture is finally subjected to a separation of substances (40) and separated into hydrogen chloride (7), oxygen (8), chlorine (9) and water (10). It is also possible to return unreacted hydrogen chloride gas (7) and oxygen gas (8) back to the reactors. This is not shown here.

2 shows a process according to the invention with 18 catalyst beds in an integrated reactor. Before the first of the reactors, the educt gases (1, 2) mixed to gas mixture (3) and fed to the reactor. The reactors each comprise a catalyst bed (20). The product gases of the reactors (4) are passed through heat exchangers (30). The heat exchanger (30) comprises feeds (5) and discharges (6) of cooling medium. FIG. 2 symbolizes that a repeat unit of reactor with catalyst bed (20) and heat exchanger (30) is repeated a total of 16 times, so that a total of 18 units are shown. The product gas mixture is finally subjected to a separation of substances (40) and separated into hydrogen chloride (7), oxygen (8), chlorine (9) and water (10). It is also possible to return unreacted hydrogen chloride gas (7) and oxygen gas (8) back to the reactors. This is not shown here.

The present invention will be further illustrated by the following Examples 1 and 2. These examples relate to the number of catalyst beds and the temperature profile of the process gas mixture when it reacts in the reaction zones according to the inventive method and is cooled again in downstream heat exchangers. Furthermore, the examples relate to the conversion of HCl obtained.

Examples

Example 1 :

In this example, the process gas mixture flowed through a total of 24 catalyst stages, ie through 24 reaction zones. After each catalyst stage there was a heat exchanger which cooled the process gas mixture before entering the next catalyst stage. The process gas used at the outset was a mixture of HCl (38.5 mol%), O ₂ (38.5 mol%) and inert gases (Ar, Cl ₂ , N ₂ , CO ₂ , totaling 23 mol%). The inlet pressure of the process gas mixture was 5 bar. The length of the catalyst stages, ie the reaction zones was uniformly 7.5 cm. The activity of the catalyst was adjusted to be the same in all catalyst stages. The procedure was carried out so that a load of 1, 2 kg of HCl per kg of catalyst and hour was achieved. There was no replenishment of process gas components before the individual catalyst stages. The total residence time in the plant was 2.3 seconds.

The results are shown in FIG. Here, the individual catalyst stages are listed on the x-axis, so that a spatial course of developments in the process is visible. The temperature of the process gas mixture is indicated on the left y-axis. The temperature profile over the individual catalyst stages is shown as a solid line. On the right y-axis the total conversion of HCl is indicated. The course of the conversion over the individual catalyst stages is shown as a dashed line.

It can be seen that the inlet temperature of the process gas mixture before the first catalyst stage is about 340 ° C. Due to the exothermic reaction to chlorine gas under adiabatic conditions, the temperature rises to about 370 ° C, before the process gas mixture is cooled by the downstream heat exchanger again. The inlet temperature before the next catalyst stage is about 344 ° C. By exothermic adiabatic reaction, it rises again to about 370 ⁰ C. The sequence of heating and cooling continues. The inlet temperatures of the process gas mixture upstream of the individual catalyst stages increase with increasing number of stages. This is possible since the amount of reactants capable of reacting in the later stages of the reaction is lower and accordingly the danger of a through exothermic reaction conditional leaving the optimum temperature range of the process decreases. Consequently, the temperature of the process gas mixture can be kept closer to optimal for the respective composition.

The conversion of HCl after the 24th stage totaled 88.1%

Example 2:

In this example, the process gas mixture flowed through a total of 18 catalyst stages, ie through 18 reaction zones. After each catalyst stage there was a heat exchanger which cooled the process gas mixture before entering the next catalyst stage. The process gas used at the outset was a mixture of HCl (38.5 mol%), O ₂ (38.5 mol%) and inert gases (Ar, Cl ₂ , N ₂ , CO ₂ , totaling 23 mol%) The inlet pressure of the process gas mixture was 5 bar. The length of the catalyst stages, ie the reaction zones, was uniformly 15 cm in each case. The activity of the catalyst was adjusted to increase with the number of catalyst stages. The relative catalyst activities were as follows:

Levels 1 and 2 30%

Levels 3 and 4 40%

Levels 5 and 6 50%

Stages 7 and 8 60%

Levels 9 and 10 70%

Levels 11 and 12 80%

Steps 13 and 14 90%

Levels 15 and 16 100%

Levels 17 and 18 100%

The procedure was carried out to achieve a load of 1.12 kg of HCl per kg of catalyst per hour. There was no replenishment of process gas components before the individual catalyst calls. The total residence time in the plant was 3.5 seconds.

The results are shown in FIG. Here are the individual on the x-axis

Catalyst stages listed so that a spatial course of developments in the process is visible. On the left y-axis is the temperature of the Process gas mixture specified. The temperature profile over the individual catalyst stages is shown as a solid line. On the right y-axis the total conversion of HCl is indicated. The course of the conversion over the individual catalyst stages is shown as a dashed line.

It can be seen that the inlet temperature of the process gas mixture before the first catalyst stage is about 350 ° C. Due to the exothermic reaction to chlorine gas under adiabatic conditions, the temperature rises to about 370 ° C, before the process gas mixture is cooled by the downstream heat exchanger again. The inlet temperature before the next catalyst stage is again about 350 ° C. By exothermic adiabatic reaction, it rises again to about 370 ⁰ C. The sequence of heating and cooling continues. The inlet temperatures of the process gas mixture upstream of the individual catalyst stages increase more slowly with increasing number of stages than in the case of Example 1. Overall, the fluctuation range of the process gas temperatures is even lower. The desired lower activity of the catalyst in the early stages makes it possible to introduce the process gas mixture with a higher inlet temperature, without fear of undesired overheating. Consequently, the temperature of the process gas mixture can be kept closer to optimal for the respective composition.

The conversion of HCl after the 18th stage totaled 88.1%.

LIST OF REFERENCE NUMBERS

1 hydrogen chloride (educt)

2 oxygen (educt)

3 mixed educt gas stream

4 product gases of the reactors

5 supply of cooling medium

6 discharge of cooling medium

7 hydrogen chloride (from product gas)

8 oxygen (from product gas)

9 chlorine

10 water

16 does not represent a reference symbol, but symbolizes 16 repeating units Reactor bed Heat exchanger Substance separation

Claims

claims

1. A process for the preparation of chlorine by catalytic gas phase oxidation of hydrogen chloride with oxygen, characterized in that the reaction of hydrogen chloride with oxygen at 18 to 60 connected in series

Catalyst beds is carried out under adiabatic conditions.

2. The method according to claim 1, characterized in that the reaction takes place at 20 to 40, preferably 22 to 30 successively connected catalyst beds.

3. The method according to claim 1 or 2, characterized in that the temperature of the catalyst in the catalyst beds, in particular during the reaction 150th

° C to 800 ⁰ C, preferably 200 to 450 ° C, is.

4. The method according to at least one of claims 1 to 3, characterized in that the emerging from at least one catalyst bed process gas mixture is then passed through at least one catalyst bed downstream of the heat exchanger.

5. The method according to claim 4, characterized in that after each catalyst bed at least one heat exchanger, preferably a single heat exchanger is located, through which the exiting process gas mixture is passed and that the heat exchanger is preferably selected from the group comprising tube bundle, plate ring groove , Spiral, finned tube and / or

Micro heat exchanger.

6. The method according to claim 4 or 5, characterized in that the heat of reaction dissipated in the heat exchangers is used for steam extraction.

7. The method according to at least one of claims 1 to 6, characterized in that the reaction is carried out at a pressure of 1 to 30 bar.

8. The method according to at least one of claims 1 to 7, characterized in that the inlet temperature of the entering into a first catalyst bed gas mixture of 150 to 630 ° C, preferably from 200 to 480 ° C, is.

9. The method of claim 8, wherein the inlet temperature of the material entering each of the catalyst beds gas mixture from 150 to 630 ° C, preferably from 200 to 480 ° C, particularly preferably 250 to 470 ⁰ C,.

10. The method according to any one of claims 1 to 9, characterized in that the successively connected catalyst beds in the catalyst bed to

Catalyst bed rising or falling average temperature can be operated.

11. The method according to any one of claims 1 to 10, characterized in that the molar ratio of oxygen to hydrogen chloride before entering each catalyst bed 0.25 to 10 equivalents of oxygen, preferably 0.5 to 5 equivalents

Oxygen, per equivalent of hydrogen chloride.

12. The method according to any one of claims 1 to 11, characterized in that one or more individual catalyst beds can be independently replaced by two or more parallel catalyst beds.

13. The method according to any one of claims 1 to 12, characterized in that the hydrogen chloride and oxygen-containing feed gas stream is supplied only to the first catalyst beds.

14. The method according to any one of claims 1 to 13, characterized in that before one or more of the first catalyst bed downstream catalyst beds fresh hydrogen chloride and / or oxygen are metered into the process gas stream.

15. The method according to any one of claims 1 to 14, characterized in that the catalyst comprises at least one element selected from the group consisting of: copper, potassium, sodium, chromium, cerium, gold, bismuth, uranium, ruthenium , Rhodium, platinum, as well as an element of the VIII. Subgroup of the

Periodic Table of the Elements.

16. The method according to any one of claims 1 to 15, characterized in that the catalyst is based on ruthenium and / or uranium or a ruthenium and / or uranium compound.

17. The method according to any one of claims 1 to 16, characterized in that the activity of the catalysts in the individual catalyst beds is different and in particular increasing from catalyst bed to catalyst bed.

18. The method according to any one of claims 1 to 17, characterized in that the catalyst of the catalyst beds is applied to an inert carrier.

19. The method according to claim 18, characterized in that the catalyst support consists wholly or partly of titanium oxide, tin oxide, alumina, zirconium oxide, vanadium oxide, ceria, chromium oxide, silica, uranium oxide, silica, carbon nanotubes, or a mixture or compound of said substances.

20. The method according to any one of claims 1 to 19, wherein a continuous exchange of a fixed bed catalyst take place.

21. The method according to any one of claims 1 to 20, wherein unreacted educt gases are recycled back into the process.

22. A reactor system for reacting a hydrogen chloride and oxygen-containing gas, at least containing feed lines for hydrogen chloride and oxygen or for a mixture of hydrogen chloride and oxygen and 18 to 60 in series thermally insulated catalyst beds.