DE102007033106A1 - Production of chlorine by catalytic gas-phase oxidation of hydrogen chloride with oxygen, involves using a multi-zone reactor with special heat exchangers between successive zones to cool the process gas - Google Patents

Abstract

A method for the production of chlorine by the adiabatic gas-phase oxidation of hydrogen chloride with oxygen on a catalyst bed in a reactor with at least two separate zones from at least one of which the process gas is then passed through a heat exchanger, involves using a heat exchanger with a stack of connected plates, each with at least two separate flow channels in a predetermined pattern, arranged so that the process gas flows in one direction and the heat exchange medium in another. An independent claim is included for a reactor system for this process as described above.

Description

The The present invention relates to a process for the preparation of Chlorine by catalytic gas-phase oxidation of hydrogen chloride with oxygen, wherein in a reactor, the process gas mixture in at least two separate reaction zones under adiabatic Reacts to catalyst beds and wherein the at least Subsequently, a process gas mixture exiting a reaction zone passed through one of the respective reaction zone downstream heat exchanger becomes. It further relates to a reactor system for the production of Chlorine by catalytic gas-phase oxidation of hydrogen chloride with oxygen by means of the method according to the invention.

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: 4HCl + O ₂ ⇒ 2Cl ₂ + 2H ₂ O

By Chloralkali electrolysis became the technical application of the Deacon process however, strongly pushed into the background. Almost the whole Production of chlorine was carried out by electrolysis of aqueous Saline solutions. The attractiveness of the Deacon process However, it has increased again recently as the world Chlorine demand grows faster than the demand for Sodium hydroxide, a by-product of NaCl electrolysis. This development is coming the process for the production of chlorine by oxidation of hydrogen chloride contrary, which is decoupled from the sodium hydroxide production. About that In addition, the precursor hydrogen chloride is easily accessible; it falls in large quantities, for example in phosgenation reactions, as in the isocyanate production, as by-product.

The Exhaustion and use of the heat of reaction is an important point in the implementation of the Deacon process. An uncontrolled rise in temperature from the beginning to the end of the Deacon reaction could be 600 ° C to 900 ° C, on the one hand would lead to permanent damage lead the catalyst, on the other hand it comes at high Temperatures to an unfavorable shift of the reaction equilibrium in the direction of the educts with a corresponding deterioration the yield. It is therefore desirable to have the temperature of Catalyst bed in the process in one area from 150 to 600 ° C.

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. Similar problems with the volatility of the catalyst components also occur with the use of more active ruthenium chloride / oxide. 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.

In established processes, therefore, the catalyst is used in the form of a fluidized, thermally stabilized bed. After EP 0 251 731 A2 the catalyst bed is tempered over the outer wall, according to the DE 10 2004 006 610 A1 the fluidized bed is tempered via 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.

In the publications WO 2004/037718 and WO 2004/014845 while mention is made in general terms of the possibility of adiabatic catalytic hydrogen chloride oxidation in addition to the preferred isothermal processes. However, concrete embodiments of an adiabatically conducted 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 require complex cooling to be controlled. 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.

Plate-shaped ducted heat exchangers as components of a chemical reactor are disclosed in US Pat WO 2001/54806 disclosed. However, this application does not relate to the Deacon process.

Thus, there remains a need for a process for the production of chlorine by adiabatic catalytic gas phase oxidation of hydrogen chloride with oxygen, in which the temperature of the reaction mixture and also of the catalyst can be better controlled. In particular, the maximum temperature should be limited to to avoid damage to the catalyst and the minimum temperature should not be too low to obtain a sufficiently high space-time yield.

The The object of the present invention is to provide such To provide method. In particular, she has the task provided a method of producing chlorine by catalytic To provide gas phase oxidation of hydrogen chloride with oxygen, wherein in a reactor, the process gas mixture in at least two separated reaction zones under adiabatic conditions reacts on catalyst beds and wherein the at least one Reaction zone exiting process gas mixture then by one of the respective reaction zone downstream heat exchanger is directed.

Solved according to the invention The task is characterized in that the heat exchanger to each other layered and interconnected plates, wherein the individual plates according to a predetermined Pattern at least two separate fluid flow channels and provided with fluid flow channels Plates are arranged so that the process gas mixture in a first flow path direction and in the heat exchanger used heat exchange medium in a second flow path direction flow through the heat exchanger.

in the In accordance with the present invention, a reactor is the overall plant to understand in which the educts hydrogen chloride and oxygen be initiated, react with each other and the reaction products be discharged. The educt of hydrogen chloride can, for example from the reaction of amines with phosgene for the synthesis of isocyanates. The reactor comprises reaction zones which are spatially separated represent separated areas in which the desired Reaction takes place. Due to the corrosive reaction gases is the reactor is preferably made of stainless steel, such as 1.4571 or 1.4828 or nickel 2.4068 or nickel base alloys such as 2.4610, 2.4856 or 2.4617, Inconel or Hastelloy.

In the reaction zones are catalyst beds. Under catalyst bed Here is an arrangement of the catalyst in all known forms of appearance, For example, fixed bed, fluidized bed or fluidized bed understood. Preferred is a fixed bed arrangement. This comprises a catalyst bed in the proper sense, that is, loose, supported or unsupported Catalyst in any form and in the form of suitable packings.

Of the Concept of catalyst bed as used herein also includes related areas more appropriate Packages on a carrier material or structured catalyst carrier. This were for example. 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.

Of the Heat exchanger is designed to work as a sequence of stacked and interconnected plates can be described. The plates can be form-fitting or cohesively connected to each other. An example for a cohesive connection is the welding or diffusion bonding.

In the plates are incorporated fluid flow channels, by a fluid from one side of a plate to the other side, For example, to the opposite side, flow can. The channels can be linear, so the train the shortest possible way. You can but also train a longer way by doing according to one wavy, meandering or zigzag patterns are created. The cross-sectional profile the channels may be semicircular, for example, elliptical, square, rectangular, trapezoidal or be triangular. That at least two separate plates per plate Fluid flow channels are present means that these channels run across the plate and the fluid flowing therein is not between the channels can change.

The Flow path direction can be determined by the vector between the Plane in which the starting points of the fluid flow channels lie and the plane in which the end points of the fluid flow channels a disk or a disk stack are defined. So it gives the general direction of the flow of Fluids through the heat exchanger. This is the name of a first Flow path direction the direction in which the process gas mixture through the heat exchanger or, in continuation, flows through the reaction zone. A second flow path direction denotes the path of the heat exchange medium. This can for example, in cocurrent, countercurrent or cross flow to the process gas mixture stream.

All in all the heat exchanger works so effectively that the temperature of the process gas mixture entering the catalyst bed of next reaction zone even when the reaction starts This causes a local overheating of the catalyst entry.

By means of the method according to the invention, a flow rate, expressed in terms of annual tons of chlorine gas produced, from ≥ 100 to ≤ 400,000, from ≥ 1000 to ≤ 300000 or from ≥ 10000 to ≤ 200,000 can be achieved.

through the process of the invention leaves a conversion of HCl from ≥ 10% to ≤ 99%, of ≥ 50% reach ≤ 95% or from ≥ 80% to ≤ 90%.

By the inventive method will be an effective Temperature control of the Deacon process is achieved so that the Formation of uncontrolled zones with elevated temperature, the so-called hot spots, especially in the entrance area of the catalyst bed can be avoided. This enables service lives of the catalyst which, expressed in years, from ≥ 1 to ≤ 10, from ≥ 2 to ≤ 6 or from ≥ 3 to ≤ 4 can be.

In an embodiment of the present invention is the Catalyst bed designed as a structured packing. In a another embodiment of the present invention the catalyst in the catalyst bed as a monolithic catalyst in front. The use of structured catalysts such as monoliths, structured packings, but also coated catalysts has mainly a reduction of the pressure loss to the advantage. Besides the advantages for the overall process may be at a lower specific pressure loss the volume to be introduced into the design of the reactor the catalyst and the heat exchanger surface a smaller flow cross-section for longer Reaction and heat exchanger stages can be realized. One Another advantage of using structured catalysts is that that in the thinner catalyst layers shorter Diffusion pathways of the reactants are necessary, which with a Increase the catalyst selectivity can go along.

in the structured catalyst bed can fluid flow channels be incorporated, wherein the hydraulic diameter of the fluid flow channels ≥ 0.1 mm to ≤ 10 mm, preferably ≥ 0.3 mm to ≤ 5 mm, more preferably ≥ 0.5 mm to ≤ 2 mm. The specific surface of the catalyst grows when the hydraulic diameter drops. If the diameter is too small, enter too much pressure loss. Furthermore, at a Impregnation with a catalyst suspension also a channel clog.

In another embodiment of the present invention is the hydraulic diameter of the fluid flow channels in the heat exchanger ≥ 10 μm to ≤ 10 mm, preferably ≥ 100 μm to ≤ 5 mm, more preferably ≥ 1 mm to ≤ 2 mm. With these diameters is an effective heat exchange especially guaranteed.

In a further embodiment of the present invention, the method comprises ≥ 6 to ≦ 50, preferably ≥ 10 to ≦ 40, more preferably ≥ 20 to ≦ 30 reaction zones. With such a number of reaction zones, the use of materials can be optimized with regard to the conversion of HCl gas. A smaller number of reaction zones would result in an unfavorable temperature control. The inlet temperature would have to be set lower, which would make the catalyst less active. Furthermore, then also decreases the average temperature of the reaction. A higher number would not justify the cost and material costs due to the low increase in sales. Especially the handling of the highly corrosive gases HCl, O ₂ and Cl ₂ requires resistant and correspondingly expensive materials for the reactor.

In another embodiment of the present invention Hydrogen chloride and oxygen are simultaneously in the reactor fed. This can be a mixing in an antechamber without a catalyst bed mean or the simultaneous introduction of the gases into the first Reaction zone. This has the advantage that the entire feed gas stream for the uptake and removal of the heat of reaction in all catalyst beds can be used. Furthermore, it is possible the gases in an upstream heat exchanger to guide them heat. With the method according to the invention an apparatus simplification of the reactor is also possible. The absence of additional piping allows a better temperature control. Generally it is also possible that the waste heat of the previous reaction stages to heat up of the process gas mixture before the next reaction zone is used.

In another embodiment of the present invention the length of at least one reaction zone is ≥ 0.01 m to ≤ 5 m, preferably ≥ 0.03 m to ≤ 1 m, more preferably ≥ 0.05 m to ≤ 0.5 m. As a length Here, the length of the reaction zones in the flow direction to understand the process gas mixture. The reaction zones can all the same length or different lengths be. For example, the early reaction zones be short, as there are enough starting materials available and excessive heating the reaction zone should be avoided. The late reaction zones can then be long to the total sales of the process to increase, being excessive Heating the reaction zone less to fear is. The stated lengths themselves have proven to be advantageous proved, because at shorter lengths the reaction can not run with the desired sales and at greater lengths of flow resistance increases too much compared to the process gas mixture. Farther For larger lengths, the catalyst exchange is heavier perform.

In a further embodiment of the before In the present invention, the catalyst in the reaction zones independently comprises substances selected from the group consisting of copper, potassium, sodium, chromium, cerium, gold, bismuth, iron, ruthenium, osmium, uranium, cobalt, rhodium, iridium, nickel, palladium and / or platinum and oxides, chlorides and / or oxychlorides of the aforementioned elements. Particularly preferred compounds include: copper (I) chloride, copper (II) chloride, copper (I) oxide, copper (II) oxide, potassium chloride, sodium chloride, chromium (III) oxide, chromium (IV) oxide, chromium (VI) oxide, bismuth oxide, ruthenium oxide, ruthenium chloride, ruthenium oxychloride and / or rhodium oxide.

Of the Catalyst can be applied to a support. Of the Carrier component may comprise: titanium oxide, tin oxide, aluminum oxide, Zirconium oxide, vanadium oxide, 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. Furthermore, particularly preferred carrier materials are Tin oxide and carbon nanotubes.

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.

to Doping of the catalysts are suitable as promoters alkali metals such as lithium, sodium, rubidium, cesium and especially potassium, Alkaline earth metals such as calcium, strontium, barium and especially magnesium, Rare earth metals such as scandium, yttrium, praseodymium, neodymium and especially Lanthanum and cerium, continue to cobalt and manganese and mixtures of the aforementioned promoters.

The Shaped bodies can then be used in a Temperature from ≥ 100 ° C to ≤ 400 ° C dried under a nitrogen, argon or air atmosphere and optionally calcined. The shaped bodies are preferred initially at ≥ 100 ° C to ≤ 150 ° C dried and then at ≥ 200 ° C. Calcined to ≤ 400 ° C.

In another embodiment of the present invention is the particle size of the catalyst independent of each other ≥ 1 mm to ≤ 10 mm, preferably ≥ 1.5 mm to ≤ 8 mm, more preferably ≥ 2 mm to ≤ 5 mm. The particle size can be at approximately spherical catalyst particles the Diameter correspond or at approximately cylindrical Catalyst particles of expansion in the longitudinal direction. The particle size ranges mentioned have changed proved to be advantageous because smaller particle sizes a high pressure loss occurs and larger Particles the usable particle surface in proportion to the particle volume decreases and thus the achievable space-time yield less becomes.

In another embodiment of the present invention In different reaction zones, the catalyst has a different one Activity, preferably the activity of the catalyst in the reaction zones, along the flow direction the process gases seen increases. If the concentration of the educts in the early stages of reaction is high, as a result of which their reaction also the temperature of the process gas mixture strong increase. To avoid unwanted temperature increase in Therefore, it can be a catalyst be selected with a lower activity. An effect of this is also that less expensive catalysts can be used. To the highest possible Reaction of the remaining starting materials in late reaction zones to achieve more active catalysts can be used there become. Overall, it is therefore due to the different activity the catalysts in the individual reaction zones possible, the temperature of the reaction in a narrower and therefore too to maintain a more favorable temperature range.

One Example of a change in the catalyst activity would be if the activity in the first reaction zone 30% of the maximum activity and per reaction zone in increments of 5%, 10%, 15% or 20% until the activity increases 100% in the last reaction zone.

The Activity of the catalyst can be, for example adjust it so that with the same basic material of the carrier, same promoter and same catalytically active compound the quantitative proportions of the catalytically active Connection are different. Furthermore, in the sense a macroscopic dilution also particles without activity be mixed.

In another embodiment of the present invention a continuous exchange of a fixed bed catalyst is carried out.

In a further embodiment of the present invention, the absolute inlet pressure of the process gases before the first reaction zone is ≥ 1 bar to ≦ 60 bar, preferably ≥ 2 bar to ≦ 20 bar, more preferably ≥ 3 bar to ≦ 8 bar. The absolute inlet pressure determines the amount of starting material and the reaction kinetics in the process gas mixture. The ranges given have proven to be favorable, since lower pressures cause economically low, non-attractive conversions of the educts and at higher pressures the required Verdichterleis big, which entails cost disadvantages.

In another embodiment of the present invention the inlet temperature of the process gases is before one Reaction zone ≥ 250 ° C to ≤ 400 ° C, preferably ≥ 310 ° C to ≤ 390 ° C, more preferably ≥ 330 ° C to ≤ 370 ° C. The inlet temperature can be the same for all zones or individually different. She is responsible for how fast and how high the temperature in the process gas mixture increases. The selected inlet temperatures allow as much as possible high conversion in the reaction zone without the temperature within the Zone increases to unwanted levels.

In another embodiment of the present invention the maximum temperature in a reaction zone is ≥ 340 ° C to ≤ 420 ° C, preferably ≥ 350 ° C. to ≤ 400 ° C, more preferably ≥ 365 ° C up to ≤ 385 ° C. The maximum in a reaction zone The prevailing temperature can be the same for all zones or individually different. It can by process parameters as pressure or composition of the process gas mixture, activity of the Catalyst and length of the reaction zone can be adjusted. The maximum temperature determines both the reaction conversion as also the extent of the discharge or deactivation of the Catalyst. The selected temperatures leave one as possible high conversion in the reaction zone, without the catalyst significantly discharged or deactivated.

• – Dimensionierung der Katalysatorbetten,
• – Steuerung der Wärmeabfuhr zwischen den Katalysatorbetten,
• – Zusatz von Einsatzgasen zwischen den Katalysatorbetten,
• – Molares Verhältnis der Edukte,
• – Konzentrationen der Edukte
• – Zusatz von Inertgasen, insbesondere Stickstoff, Kohlendioxid, vor und/oder zwischen den Katalysatorbetten

In general, the control of the temperature in the catalyst beds can preferably be 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 can the catalysts or the supported Catalysts have any shape, z. B. balls, chopsticks, Raschig rings or granules or tablets.

In another embodiment of the present invention become the successively switched reaction zones at a changing average temperature operated. This can, for example, via the controller the heat exchanger connected between the catalyst beds be set. It means that you are within a sequence of catalyst beds, the temperature of catalyst bed to catalyst bed can both rise and fall. So it can be particularly beneficial be the average temperature first of catalyst bed to catalyst bed to increase the catalyst activity to rise and then to the postponement of Equilibrium the average temperature in the following last Let catalyst beds sink again. On the other hand it can be advantageous, the successively connected reaction zones at to operate at a rising average temperature. Thus, can the reaction of the reactants initially with a larger one Safety distance to the desired upper temperature limit be performed. Later in the implementation, if less starting materials are present, the conversion can be achieved by increasing the average temperature continues to be driven.

In another embodiment of the present invention is the residence time of the process gases in the reactor total ≥ 0.5 s to ≤ 60 s, preferably ≥ 1 s to ≤ 30 s, more preferably ≥ 2 s to ≤ 10 s. Lower residence times and the associated low space-time yield are not economically attractive. At higher residence times There is no significant additional increase in space-time yield a, so that such a procedure also not is economically attractive. Continues rising at a higher Residence time the outlet temperature above the maximum desired temperature.

In another embodiment of the present invention Unreacted educt gases are returned to the beginning of the reactor introduced. Consequently, it is a circular process. Unreacted educt gases are in particular hydrogen chloride and Oxygen.

In a further embodiment of the present invention, the heat exchange medium which flows through a heat exchanger selected from the group comprising liquids, boiling liquids, gases, organic heat carriers, molten salts and / or ionic liquids, wherein preferably water, partially evaporating water and / or steam selected become. Under partially evaporating water is to be understood that in the individual fluid flow channels of the heat exchanger liquid water and water vapor are present side by side. This offers the advantages of a high heat transfer coefficient on the side of the heat exchange medium, a high specific heat absorption by the enthalpy of evaporation of the heat exchange medium and a constant temperature over the channel of the heat exchange medium. Especially in cross-flow to the reactant flow guided heat exchange medium, the constant evaporation temperature is advantageous because it allows a uniform heat removal across all reaction channels. The regulation of the Reaktandentemperatur can be done via the adjustment of the pressure level and thus the temperature for the evaporation of the heat exchange medium.

In another embodiment of the present invention is the mean logarithmic temperature difference between heat exchange medium and the product stream ≥ 5 K to ≤ 300 K, preferably ≥ 10 K to ≤ 250 K, more preferably ≥ 50 K to ≤ 150 K. At lower logarithmic temperature differences will be the needed Heat exchanger surface too large, which cost disadvantages with brings. At higher logarithmic temperature differences the coolant must be very cold. For example, though low-energy steam also in the overall economic balance a plant less valuable.

In a further embodiment of the present invention, the method is carried out so that the space-time yield, expressed in kg of Cl ₂ per kg of catalyst, ≥ 0.1 to ≤ 10, preferably ≥ 0.3 to ≤ 3, more preferably ≥ 0.5 to ≤ 2.

In another embodiment of the present invention becomes the heat of reaction dissipated in the heat exchangers used for steam extraction. This makes the whole procedure more economical and allows, for example, the process beneficial in a compound plant or a composite site operate.

In another embodiment of the present invention is the molar ratio of oxygen to Hydrogen chloride before entering the first reaction zone ≥ 0.25 to ≤ 10, preferably ≥ 0.5 to ≤ 5, more preferably ≥ 0.5 to ≤ 2. By an increase the ratio of equivalents of oxygen On the one hand, the reaction can be per equivalent of hydrogen chloride accelerates and thus the space-time yield (produced chlorine quantity per reactor volume), on the other hand, the equilibrium the reaction shifted positively towards the products.

In another embodiment of the present invention The process gases comprise an inert gas, preferably nitrogen and / or Carbon dioxide. Furthermore, the inert gas has a share of the process gases from ≥ 15 mol% to ≤ 30 mol%, preferably ≥ 18 mol% to ≦ 28 mol%, more preferably ≥ 20 mol% to ≤ 25 mol%. By means of inert gases can be favorable influence on the reaction temperature and the reaction kinetics to take. The stated proportions have proven to be favorable proved because with too small inert gas streams requires more reaction stages and at too large inert gas streams, in particular when running the process in a circulating current, the operating costs rise too much.

object The present invention furthermore relates to a reactor system for the production of chlorine by catalytic gas phase oxidation of hydrogen chloride with oxygen by the method according to present invention. In particular, the present invention relates a reactor system, wherein the heat exchanger to each other layered and interconnected plates, wherein the individual plates according to a predetermined Pattern at least two separate fluid flow channels and provided with fluid flow channels Plates are arranged so that the process gas mixture in a first flow path direction and in the heat exchanger used heat exchange medium in a second flow path direction flow through the heat exchanger. Furthermore is it is advantageous if the reactor system 6 to ≦ 50, preferably ≥ 10 to ≤ 40, more preferably ≥ 20 to ≤ 30 reaction zones.

The The present invention will become apparent from the following Examples 1 and 2 further explained. These examples relate to the temperature profile the process gas mixture, if this according to the invention Process reacts in reaction zones and in downstream heat exchangers is cooled again. Furthermore, the examples relate to the achieved sales of HCl.

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 per 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 in 1 shown. Here, the individual catalyst stages are listed on the x-axis, so that a spatial course of Ent turns in the process becomes 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.

you recognizes that the inlet temperature of the process gas mixture before the first catalyst stage is about 340 ° C. By the exothermic reaction to chlorine gas under adiabatic conditions the temperature rises to about 370 ° C before the process gas mixture cooled again by the downstream heat exchanger becomes. The inlet temperature before the next catalyst stage is about 344 ° C. By exothermic adiabatic Reaction rises again to about 370 ° C. The sequence from heating and cooling continues. The inlet temperatures of the process gas mixture in front of the individual Catalyst levels increase with increasing number of stages. This is possible, as in the course of the reaction later stages the amount of reactants capable of reacting is lower and accordingly the risk of an exothermic reaction Leaving the optimum temperature range of the process decreases. Consequently, the temperature of the process gas mixture can be closer kept optimal for the respective composition become.

Of the Total HCI after the 24th step was 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% Steps 3 and 4 40% Levels 5 and 6 50% Steps 7 and 8 60% Steps 9 and 10 70% Steps 11 and 12 80% Steps 13 and 14 90% Stages 15 and 16 100% Steps 17 and 18 100%

The Procedure was carried out so that a load of 1.12 kg of HCl per kg of catalyst and hour was reached. It happened no replenishment of process gas components before the individual Catalyst stages. The residence time in the plant totaled 3.5 seconds.

The results are in 2 shown. 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.

Of the Temperature profile over the individual catalyst stages away is shown as a solid line. On the right y-axis is the total sales of HCl indicated. The course of sales over the individual catalyst stages is shown as a dashed line shown.

you recognizes that the inlet temperature of the process gas mixture before the first catalyst stage is about 350 ° C. By the exothermic reaction to chlorine gas under adiabatic conditions the temperature rises to about 370 ° C before the process gas mixture cooled again by the downstream heat exchanger becomes. The inlet temperature before the next catalyst stage is again about 350 ° C. By exothermic adiabatic reaction it rises again to about 370 ° C. The sequence of warming and cooling continues. The inlet temperatures of the process gas mixture before the individual catalyst stages increase with increasing number of stages slower 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 allows the process gas mixture to introduce with a higher inlet temperature, without fear of an unwanted overheating must become. Consequently, the temperature of the process gas mixture closer to optimal for the respective composition being held.

Of the Total HCl conversion after the 18th stage was 88.1%.

QUOTES INCLUDE IN THE DESCRIPTION

This list The documents listed by the applicant have been automated generated and is solely for better information recorded by the reader. The list is not part of the German Patent or utility model application. The DPMA takes over no liability for any errors or omissions.

Cited patent literature

• EP 0251731 A2 [0006]
• - DE 102004006610 A1 [0006]
• WO 2004/037718 [0007]
• WO 2004/014845 [0007]
• WO 2001/54806 [0008]

Claims (24)

A process for the preparation of chlorine by catalytic gas phase oxidation of hydrogen chloride with oxygen, wherein in a reactor, the process gas mixture in at least two separate reaction zones under adiabatic conditions of catalyst beds and wherein the emerging from at least one reaction zone process gas mixture is then passed through a respective reaction zone downstream heat exchanger characterized in that the heat exchanger comprises plates stacked and interconnected with each other, the individual plates having at least two separate fluid flow channels according to a predetermined pattern, and the fluid flow channeled plates being arranged such that the process gas mixture is in a first flow path direction and that in the heat exchanger used heat exchange medium in a second Strömungswegrichtung flow through the heat exchanger. The method of claim 1, wherein the catalyst bed is designed as a structured package. The method of claim 1, wherein the catalyst present in the catalyst bed as a monolithic catalyst. Method according to one of the preceding claims, wherein the hydraulic diameter of the fluid flow channels in the heat exchanger ≥ 10 μm to ≤ 10 mm, preferably ≥ 100 μm to ≤ 5 mm, more preferably ≥ 1 mm to ≤ 2 mm. Method according to one of the preceding claims, comprising ≥ 6 to ≤ 50, preferably ≥ 10 to ≤ 40, more preferably ≥ 20 to ≤ 30 Reaction zones. Method according to one of the preceding claims, hydrogen chloride and oxygen simultaneously in the reactor be fed. Method according to one of the preceding claims, wherein the length of at least one reaction zone ≥ 0.01 m to ≤ 5 m, preferably ≥ 0.03 m to ≤ 1 m, more preferably ≥ 0.05 m to ≤ 0.5 m. Method according to one of the preceding claims, wherein the catalyst is independent in the reaction zones comprises substances selected from one another the group comprising copper, potassium, sodium, chromium, cerium, gold, Bismuth, iron, ruthenium, osmium, uranium, cobalt, rhodium, iridium, Nickel, palladium and / or platinum and oxides, chlorides and / or Oxychlorides of the aforementioned elements. Method according to one of the preceding claims, the particle size of the catalyst, independently from each other ≥ 1 mm to ≤ 10 mm, preferably ≥ 1.5 mm to ≤ 8 mm, more preferably ≥ 2 mm to ≤ 5 mm. Method according to one of the preceding claims, wherein in different reaction zones, the catalyst is a different Activity, preferably the activity of the catalyst in the reaction zones, along the flow direction the process gases seen increases. Method according to one of the preceding claims, wherein a continuous exchange of a fixed bed catalyst is carried out. Method according to one of the preceding claims, where the absolute inlet pressure of the process gases before the first Reaction zone ≥ 1 bar to ≤ 60 bar, preferably ≥ 2 bar up to ≤ 20 bar, more preferably ≥ 3 bar to ≤ 8 bar is. Method according to one of the preceding claims, wherein the inlet temperature of the process gases upstream of a reaction zone ≥ 250 ° C. to ≤ 400 ° C, preferably ≥ 310 ° C to ≤ 390 ° C, more preferably ≥ 330 ° C to ≤ 370 ° C. Method according to one of the preceding claims, the maximum temperature in a reaction zone ≥ 340 ° C to ≤ 420 ° C, preferably ≥ 350 ° C. to ≤ 400 ° C, more preferably ≥ 365 ° C to ≤ 385 ° C. Method according to one of the preceding claims, wherein the successively switched reaction zones are changing Average temperature can be operated. Method according to one of the preceding claims, wherein the residence time of the process gases in the reactor total ≥ 0.5 s to ≤ 60 s, preferably ≥ 1 s to ≤ 30 s, more preferably ≥ 2 s to ≤ 10 s. Method according to one of the preceding claims, wherein unreacted educt gases back into the beginning of the reactor be introduced. Method according to one of the preceding claims, wherein the heat exchange medium, which flows through a heat exchanger, is selected from the group comprising liquids, boiling liquids, gases, organic heat carriers, molten salts and / or ionic liquid in which preferably water, partially evaporating water and / or water vapor are selected. Method according to one of the preceding claims, wherein the mean logarithmic temperature difference between heat exchange medium and the product stream ≥ 5 K to ≤ 300 K, preferably ≥ 10 K to ≤ 250 K, more preferably ≥ 50 K to ≤ 150 K is. Method according to one of the preceding claims, wherein the method is conducted so that the space-time yield, expressed in kg of Cl ₂ per kg of catalyst, ≥ 0.1 to ≤ 10, preferably ≥ 0.3 to ≤ 3, more preferably ≥ 0.5 to ≤ 2. Method according to one of the preceding claims, wherein the heat of reaction dissipated in the heat exchangers is used for steam extraction. Method according to one of the preceding claims, the molar ratio of oxygen to hydrogen chloride before entering the first reaction zone ≥ 0.25 to ≤ 10, preferably ≥ 0.5 to ≤ 5, more preferably ≥ 0.5 to ≤ 2. Method according to one of the preceding claims, wherein the process gases an inert gas, preferably nitrogen and / or Carbon dioxide, and the inert gas continue to make up a share the process gases has from ≥ 15 mol% to ≤ 30 mol%, preferably ≥ 18 mol% to ≤ 28 mol%, more preferably ≥ 20 mol% to ≤ 25 mol%. Reactor system for the production of chlorine by catalytic Gas phase oxidation of hydrogen chloride with oxygen by means of the Method according to one of the preceding claims.