WO2009143971A1

WO2009143971A1 - Process for preparing phosgene

Info

Publication number: WO2009143971A1
Application number: PCT/EP2009/003499
Authority: WO
Inventors: Ralph Schellen; Evin Hizaler Hoffmann; Leslaw Mleczko; Stephan Schubert
Original assignee: Bayer Technology Services Gmbh
Priority date: 2008-05-29
Filing date: 2009-05-16
Publication date: 2009-12-03
Also published as: DE102008025850A1

Abstract

The present invention relates to a process for preparing phosgene by catalytic gas-phase oxidation of carbon monoxide by means of chlorine, wherein the reaction is carried out over from 6 to 40 catalyst beds connected in series under adiabatic conditions, and also a reactor system for carrying out the process.

Description

Process for the preparation of phosgene

The present invention relates to a process for the preparation of phosgene by catalytic gas phase oxidation of carbon monoxide with chlorine, wherein the reaction is carried out on 6 to 40 catalyst beds connected in series under adiabatic conditions, and a reactor system for carrying out the process.

Phosgene is generally prepared under the catalytic effect of activated carbon from gaseous carbon monoxide and gaseous chlorine in an exothermic, catalytic equilibrium reaction according to formula (I):

CO (g) + Cl ₂ (g) <→ COCl ₂ (g); ΔH = -108 kJ / mol (I)

The phosgene produced by the reaction of formula (I) forms an essential starting material for the synthesis of polymers such as polyurethanes (via isocyanates) and polycarbonates.

The removal and use of the heat of reaction is an important issue in the practice of phosgene synthesis. An uncontrolled rise in temperature, which may be up to 500 ° C during the reaction, can lead to permanent damage to the catalyst. On the other hand, at high temperatures there is an unfavorable shift in the reaction equilibrium in the direction of the educts with a corresponding deterioration of the yield. In addition, at high temperatures, there is the potential for side reactions to produce more or less large amounts of, inter alia, carbon tetrachloride (CCl ₄ ) as an impurity. It is therefore advantageous to maintain the temperature of the catalyst bed controlled in the course of the process at a level which allows a rapid conversion while minimizing side reactions and / or catalyst inactivation.

EP 1 135 329 (B1) discloses that it is advantageous to carry out the reaction according to formula (I) at low temperatures of not more than 30 to 80 ° C. and pressures of from 1.2 to 4 bar. Possible devices in which such a process can be carried out are disclosed in EP 1 135 329 (B1) as tubular reactors to which the educts are attached

Room temperature to be supplied. It is further disclosed that the apparatus in which the reaction is carried out is cooled. It is not disclosed how high the conversions and selectivities are with respect to the educts according to the disclosed process. By the One-step operation at low temperatures, however, suggests that sales are lower than those at higher temperatures. A temperature control takes place only in that educts with low inlet temperatures are introduced into the process. The method is disadvantageous because within the process, the heat of reaction at a temperature level of up to 80 ° C is dissipated, which further complicates the use of heat, and thus the expected sales is low.

EP 1 640 341 (A2) discloses a method and an apparatus which permit reaction according to formula (I) wherein a reaction zone is cooled during the reaction. By cooling the reaction zone, the temperature of the product stream is kept below 100 ° C. Here, the cooling medium is guided in a closed circuit, so that the energy disadvantages, as set out in the disclosure of EP 1 135 329 (B1), are no longer so serious. Nevertheless, the disadvantages, in particular with regard to the expected turnover of such a procedure, are the same as before.

Other established methods include more than a fixed bed of catalyst beds after which heat exchangers are installed to recool the product gas mixture after the exothermic reaction and optionally to add further conversion in another reaction zone. In this way, the heat of reaction can be obtained at higher temperature levels, as well as increased sales can be achieved.

EP 0 134 506 (B1) discloses a procedure with two reaction zones in tube bundle reactors, wherein at temperatures above 250 ° C. a first reaction takes place in a first reaction zone, then the resulting gas mixture is cooled down to a lower temperature and in a second reaction zone at temperatures from 50 ° C to 100 ° C, a further conversion is carried out. The heat dissipated between the two reaction zones is used to generate steam. It is disclosed that the cooling is necessary for the conversion in the second reaction zone in order to shift the unfavorable position of the reaction equilibrium, as it is present at the temperatures of the first reaction zone, back into an advantageous position for the formation of the desired product (phosgene). The process is disadvantageous in that overheating of the first reaction zone is tolerated, thereby necessitating heavy cooling before entering the second reaction zone. Overheating results in a detrimental shift in reaction equilibrium and increased catalyst inactivation as well as byproduct formation in the first reaction zone. The adverse shift of the reaction equilibrium and the formation of the by-products can be at least partially compensated by the second reaction zone; the catalyst inactivation in the first reaction zone, however, is irreversible and thus leads to the need to replace the catalyst at regular intervals, which is economically unfavorable. Furthermore, the tube bundle reactor used is structurally quite complex, which has an economically disadvantageous effect on the investment costs.

EP 1 251 951 (B1) discloses a device and the possibility of carrying out chemical reactions in the device, wherein the device is characterized by a cascade of reaction zones in contact with one another and heat exchanger devices which are arranged in a composite with one another. The method to be carried out here is thus characterized by the contact of the various reaction zones with a respective heat exchanger device in the form of a cascade. There is no disclosure as to the usability of the apparatus and method for the synthesis of phosgene from gaseous chlorine and carbon monoxide. It thus remains unclear how, starting from the disclosure of EP 1 251 951 (B1), such a reaction is to be carried out by means of the device and the method carried out therein. Further, for the sake of uniformity, it is to be understood that the method disclosed in EP 1 251 951 (B1) is carried out in a device the same as or similar to the disclosure regarding the device. As a result, by the large-area contact of the heat exchange zones according to the disclosure with the reaction zones, a significant amount of heat takes place by heat conduction between the reaction zones and the adjacent heat exchange zones. The disclosure with regard to the oscillating temperature profile can therefore only be understood as meaning that the temperature peaks ascertained here would be stronger if this contact did not exist. Another indication of this is the exponential increase in the disclosed temperature profiles between the individual temperature peaks. These indicate that there is some heat sink of appreciable but limited capacity in each reaction zone which reduces its temperature rise can. It can never be ruled out that a certain dissipation of heat (eg by radiation) takes place, but a reduction in the possible heat removal from the reaction zone would indicate a linear or degressive temperature gradient, since at higher temperatures the equilibrium of exothermic reactions shifted to the side of the starting materials and thus reduce the reaction rate and with it the generated heat of reaction. Thus, EP 1 251 951 (B1) discloses multi-stage processes in cascades of reaction zones from which heat in an undefined amount is removed by heat conduction. Accordingly, the disclosed method is disadvantageous in that accurate temperature control of the process gases of the reaction is not possible.

It would therefore be advantageous to provide a process which can be carried out in simple reaction devices and which allows accurate, simple temperature control so that it allows high conversions with the highest possible purity of the product. Such simple reaction devices would be easily scaled up and are inexpensive and robust in all sizes.

For the catalytic gas-phase oxidation of carbon monoxide with chlorine to phosgene, as described above, neither suitable reactors have yet been described, nor are suitable processes disclosed which permit this.

It is therefore an object to provide a process for the catalytic gas phase oxidation of carbon monoxide and chlorine to phosgene, which can be carried out under precise temperature control in a simple reaction apparatus and thereby high conversions at high purities of the product allowed, the heat of reaction, for example, to set an advantageous temperature level can be used for the reaction or steam generation.

It has surprisingly been found that a process for the preparation of phosgene, which is characterized in that it comprises a conversion of carbon monoxide with chlorine to phosgene according to formula (T)

CO (g) + Cl ₂ (g) ^ COCl ₂ (g); ΔH = -108 kJ / mol (I)

in 6 to 40 connected in series reaction zones under adiabatic conditions in the presence of heterogeneous catalysts, this task is able to solve. Carbon monoxide, in the context of the present invention, refers to a process gas which is introduced into the process according to the invention and which essentially comprises carbon monoxide. Essentially, in the context of the present invention, a proportion of more than 90% by weight.

Chlorine in the context of the present invention refers to a process gas which is introduced into the process according to the invention and which essentially comprises chlorine.

In addition to the essential components of the process gases chlorine and carbon monoxide, these can also include secondary components. Non-exhaustive examples of minor components which may be included in the process gases include nitrogen, carbon dioxide, water and carbon tetrachloride.

In general, in connection with the present invention, process gases are understood as gas mixtures which comprise chlorine and / or carbon monoxide and / or phosgene and / or secondary components. Essentially, however, process gases include chlorine and / or carbon monoxide and / or phosgene.

According to the invention, carrying out the process under adiabatic conditions essentially means that the reaction zone is neither actively supplied with heat nor heat withdrawn from the outside. It is well known that complete isolation against heat input or discharge is possible only by complete evacuation, excluding the possibility of heat transfer by radiation. Therefore, in the context of the present invention, adiabat means that no heat supply or removal measures are taken.

However, in an alternative embodiment of the method according to the invention, by isolation by means of well-known isolation means, e.g. Polystyrene insulating materials, or by sufficiently large distances to heat sinks or heat sources, wherein the insulating agent is air, a heat transfer can be reduced.

An advantage of the adiabatic driving method according to the invention of the 6 to 40 reaction zones connected in series with respect to a non-adiabatic mode of operation is that no means for heat removal must be provided in the reaction zones, which entails a considerable simplification of the construction. This results in particular simplifications in the production of the reactor and in scalability of the process and an increase in reaction revenues. In addition, the heat generated in the course of the exothermic reaction progress can be utilized in the individual reaction zone to increase the conversion in a controlled manner by allowing the process gases and the reaction zone to increase in temperature to near equilibrium limitation during the passage of the reaction zone.

Another advantage of the method according to the invention is the possibility of very accurate temperature control by the close staggering of adiabatic reaction zones. It can thus be adjusted in that reaction zone advantageous in the reaction progress temperature.

The catalysts used in the process according to the invention are usually catalysts which consist of a material which, in addition to its catalytic activity for the reaction according to formula (I), is characterized by sufficient chemical resistance to chlorine under the conditions of the process and by a high specific surface area. Catalyst materials characterized by such chemical resistance to chlorine under the conditions of the process are, for example, silica gels, silicon carbides and / or activated carbon. Activated carbon is preferred because it also has a particularly high specific surface area.

Specific surface area in the context of the present invention refers to the area of the catalyst material that can be reached by the process gases, based on the mass of catalyst material used.

A high specific surface area is a specific surface area of at least 100 m ² / g, preferably of at least 200 m / g.

The catalysts of the invention are each in the reaction zones and can be used in all known forms, e.g. Fixed bed, fluidized bed, fluidized bed, are present.

Preferred are the forms of fixed bed and fluidized bed.

The fixed bed arrangement comprises a catalyst bed in the strict sense, ie loose, supported or unsupported catalyst in any form and in the form of suitable packings. The term catalyst bed as used herein also includes contiguous areas of suitable packages on a carrier. material or structured catalyst support. These would be, 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. As a special form of packing, in the context of the present invention, the presence of the catalyst in monolithic form is considered.

If a fixed-bed arrangement of the catalyst is used, the catalyst is preferably present in beds of particles having average particle sizes of 1 to 10 mm, preferably 2 to 7 mm, particularly preferably 3 to 5 mm.

Also preferably, the catalyst is in a fixed bed arrangement in monolithic form. In the case of a fixed-bed arrangement, a monolithic catalyst which consists of carbon is particularly preferred.

If a fluidized bed arrangement of the catalyst is used, the catalyst is preferably present in loose beds of particles, as have also previously been described for the fixed bed arrangement.

Beds of such particles are advantageous because the particles have a high outer surface of the catalyst material compared to the process gases chlorine and carbon monoxide due to their size and thus a high conversion rate can be achieved. Thus, the mass transport limitation of the reaction by diffusion can be kept low. At the same time, however, the particles are not yet so small that excessively increased pressure losses occur when the fixed bed flows through. The ranges of the particle sizes given in the preferred embodiment of the process, comprising a reaction in a fixed bed, are thus an optimum between the achievable conversion from the reaction according to formula (I) and the pressure drop produced when carrying out the process. Pressure loss is coupled in a direct manner with the necessary energy in the form of compressor performance, so that a disproportionate increase in the same would result in an inefficient operation of the method.

In a preferred embodiment of the process according to the invention, the conversion takes place from 8 to 30, particularly preferably 10 to 25 reaction zones connected in series. A preferred further embodiment of the method is characterized in that the process gas emerging from at least one reaction zone is subsequently passed through at least one heat exchange zone downstream of said reaction zone.

In a particularly preferred further embodiment of the process, after each reaction zone is at least one, preferably exactly one heat exchange zone, through which the process gas leaving the reaction zone is passed.

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

The individual reaction zones and heat exchange zones can also be arranged together in a reactor or in any combination of reaction zones with heat exchange zones in several reactors.

If reaction zones and heat exchange zones are present in a reactor, then in an alternative embodiment of the invention there is a heat insulation zone between them in order to be able to support the adiabatic operation of the reaction zone.

In addition, each of the series-connected reaction zones can be replaced or supplemented independently of one another by one or more reaction zones connected in parallel. In particular, the use of reaction zones connected in parallel allows their replacement or supplementation during ongoing continuous operation of the process.

Parallel and successive reaction zones may in particular also be combined with one another. However, the process according to the invention particularly preferably has exclusively reaction zones connected in series.

The reactors preferably used in the process according to the invention can consist of simple containers with one or more reaction zones, as described, for example, in Ullmann's Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, VoI B4, pages 95-104, pages 210-216), wherein in each case between the individual reaction zones and / or heat exchange zones heat insulation zones can be additionally provided. In an alternative embodiment of the method, there is thus at least one heat-insulating zone between a reaction zone and a heat exchange zone. Preferably, there is a heat isolation zone around each reaction zone.

The catalysts or the fixed beds thereof are mounted in a manner known per se on or between gas-permeable walls comprising the reaction zone of the reactor. Especially with thin fixed beds can in the flow direction before the

Catalyst beds technical devices for uniform gas distribution can be attached. These can be perforated plates or other internals that cause a uniform entry of the process gas into the fixed bed by generating a small but uniform pressure loss.

In a particular embodiment of the process according to the invention, preference is given to using a molar excess of between 0 and 20% carbon monoxide, based on the molar flow of chlorine, before it enters the reaction zone. By increasing the ratio of carbon monoxide per chlorine on the one hand, the reaction can be accelerated and thus the space-time yield (produced amount of phosgene per mass of catalyst material) can be increased, on the other hand, the equilibrium of the reaction is positively shifted in the direction of the product (phosgene). As a result, the conversion of the educt chlorine can be increased.

In a further particularly preferred embodiment of the method, the inlet temperature of the process gas entering the first reaction zone is from 10 to 350.degree. C., preferably from 40 to 250.degree. C., particularly preferably from 80 to 210.degree.

In yet another particularly preferred embodiment of the method, the absolute pressure at the inlet of the first reaction zone is between 1.1 and 40 bar, preferably between 1.5 and 10 bar, more preferably between 2 and 6 bar.

In a further particularly preferred embodiment of the process, the residence time of the process gas in a reaction zone is between 0.1 and 100 s, preferably between 0.5 and 50 s, particularly preferably between 1 and 10 s.

The carbon monoxide and the chlorine are preferably fed only before the first reaction zone. This has the advantage that the entire process gas can be used for the absorption and removal of the heat of reaction in all reaction zones. In addition, through Such a procedure, the space-time yield can be increased, or the necessary catalyst mass can be reduced. However, it is also possible to meter in carbon monoxide and / or chlorine into the process gas as required before one or more of the reaction zones following the first reaction zone. In addition, the temperature profile in the reaction zone can be controlled via the supply of gas between the reaction zones.

In a preferred further development of the process, the process gas leaving the last reaction zone is at least partially reused by being introduced into one of the reaction zones. Particularly preferably, only the proportion of the carbon monoxide is reused by being introduced into the first reaction zone.

This further development is advantageous because the process gas at the outset of the last reaction zone essentially comprises only carbon monoxide and phosgene and thus a further conversion of the carbon monoxide with chlorine to phosgene can be achieved via reuse.

In a particularly preferred embodiment of the method according to the invention, the process gas is cooled after at least one of the reaction zones used, more preferably after each of the catalyst beds used. For this purpose, the process gas is passed after exiting a reaction zone through one or more of the above-mentioned heat exchange zones, which are located behind the respective reaction zones. These may be used as heat exchange zones in the form of heat exchangers known to those skilled in the art, e.g. Tube bundle, plate, Ringnut-, spiral, finned tube, microstructured heat exchanger be executed. Preference is given to microstructured heat exchangers.

In the context of the present invention, microstructured means that the heat exchanger for the purpose of heat transfer comprises fluid-carrying channels, which are characterized in that they have a hydraulic diameter between 50 μm and 5 mm. The hydraulic diameter is calculated as four times the cross-sectional area of the fluid-carrying channel, divided by the circumference of the channel.

In a particular embodiment of the method, steam is generated during cooling of the process gas in the heat exchange zones by the heat exchanger. Within this further embodiment, it is preferable in the heat exchangers, which include the heat exchange zones, to carry out evaporation on the side of the cooling medium, preferably partial evaporation.

Partial evaporation referred to in the context of the present invention, an evaporation in which a gas / liquid mixture of a substance is used as a cooling medium and in which there is still a GasTFlüssigkeitsgemisch a substance even after heat transfer in the heat exchanger.

The carrying out of evaporation is particularly advantageous because in this way the achievable heat transfer coefficient from / to process gases on / from the cooling / heating medium becomes particularly high and thus efficient cooling can be achieved.

Performing a partial evaporation is particularly advantageous because the absorption / release of heat by the cooling medium thereby no longer results in a temperature change of the cooling medium, but only the gas / liquid equilibrium is shifted. This has the consequence that over the entire heat exchange zone, the process gas is cooled to a constant temperature. This in turn safely prevents the occurrence of temperature profiles in the flow of process gases, thereby improving control over the reaction temperatures in the reaction zones and, in particular, preventing the formation of local overheating by temperature profiles.

In an alternative embodiment, instead of evaporation / partial evaporation, a mixing zone can also be provided upstream of the inlet of a reaction zone in order to standardize the temperature profiles in the flow of process gases which may arise during cooling by mixing transversely to the main flow direction.

In a further preferred embodiment of the process, the reaction zones connected in series are operated at an average temperature increasing or decreasing from reaction zone to reaction zone. This means that within a sequence of reaction zones, the temperature can be both increased and decreased from reaction zone to reaction zone. Thus, it may be particularly advantageous to first increase the average temperature from reaction zone to reaction zone to increase the conversion and then to shift the equilibrium the Average temperature in the following last reaction zone to drop again. This can be adjusted, for example, via the control of the heat exchange zones connected between the reaction zone. Further options for setting the average temperature are described below.

The thickness of the flow-through reaction zones can be chosen to be the same or different and results according to laws generally known in the art from the residence time described above and the process gas quantities enforced in the process. The mass flows of product gas (phosgene) which can be carried out according to the invention by the process, and from which the amounts of process gas to be used, are usually between 0.01 and 60 t / h, preferably between 1 and 50 t / h, particularly preferably between 10 and 40 t /H.

The maximum exit temperature of the process gas from the reaction zones is usually in a range from 150 ° C to 650 ° C, preferably from 180 ° C to 360 ° C, more preferably from 260 ° C to 320 ° C. The control of the temperature in the reaction zones is preferably carried out by at least one of the following measures: dimensioning of the adiabatic reaction zone, control of heat dissipation between the reaction zones, addition of gas between the reaction zones, molar ratio of the reactants / excess of carbon monoxide used, addition of inert gases, in particular Nitrogen, carbon dioxide, before and / or between the reaction zones.

The composition of the catalysts in the reaction zones according to the invention may be identical or different. In a preferred embodiment, the same catalysts are used in each reaction zone. However, it is also advantageous to use different catalysts in the individual reaction zones. Thus, especially in the first reaction zone, when the concentration of the reaction educts is still high, a less active catalyst can be used and in the further reaction zones the activity of the catalyst can be increased from reaction zone to reaction zone. The control of the catalyst activity can also be carried out by dilution with inert materials or carrier material. Also advantageous is the use of a catalyst in the first and / or second reaction zone, which is particularly stable against deactivation at the temperatures of the process in these reaction zones. 0.1 kg / h to 50 kg / h, preferably 1 kg / h to 20 kg / h, more preferably 3 kg / h to 10 kg / h of phosgene can be prepared by the process according to the invention per 1 kg 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. This surprisingly high space-time yield is made possible by the interaction of the inventive and preferred embodiments of the new method. In particular, the interaction of staggered adiabatic reaction zones with intervening heat exchange zones and the defined dwell allows precise control of the process and the resulting high space-time yields, as well as a reduction of the by-products formed such as CCl. ₄

Another object of the invention is a reactor system for the reaction of carbon monoxide and chlorine to phosgene, characterized in that it comprises leads (Z) for a process gas comprising carbon monoxide and chlorine or for at least two process gases, of which at least one carbon monoxide and at least one chlorine , and 6 to 40 successive reaction zones (R) in the form of fixed beds of a heterogeneous catalyst, which are between the reaction zones heat insulation zones (I) in the form of insulating material and between these heat exchange zones (W) in the form of plate heat exchangers the reaction zones are connected via supply and discharge lines for the process gases and include the supply and discharge lines for a cooling medium.

The reactor system may also comprise 8 to 30 or 10 to 25 reaction zones in the form of fixed beds.

The insulating material of the heat insulating zones is preferably a material having a

[W

Wärmeleitkoeffizient λ less than or equal to 0. Particularly preferred are about m - R K polystyrene, polyurethanes, glass wool or air.

The plate heat exchangers are preferably microstructured plate heat exchangers. The present invention will be elucidated with reference to the drawings without, however, being limited thereto.

1 shows a schematic representation of an embodiment of the reactor system according to the invention, the following reference symbols being used in the figures:

Z: supply line (s)

R: reaction zone (s)

I: thermal insulation zone (s)

W: Heat exchange zone (s) FIG. 2 shows reactor temperature (T), optimum temperature (T *) and conversion of chlorine (U) over a number of 11 reaction zones (S) with downstream heat exchange zones (according to Example 1).

3 shows reactor temperature (T) and conversion of chlorine (U) over a number of 18 reaction zones (S) with downstream heat exchange zones (according to Example T).

4 shows reactor temperature (T), optimum temperature (T *) and conversion of chlorine (U) over a number of 17 reaction zones (S) with downstream heat exchange zones (according to Example 3).

5 shows reactor temperature (T), optimum temperature (T *) and conversion of chlorine (U) over a number of 17 reaction zones (S) with downstream heat exchange zones (according to Example 4).

The present invention will be further illustrated by the following examples 1 to 4, without being limited thereto. Examples

Example 1 :

In this example, the process gas flows through a total of 11 fixed catalyst beds of activated carbon, ie through 11 reaction zones. Each after a reaction zone is a heat exchange zone in which the process gas is cooled before it enters the next reaction zone. The process gases used at the beginning of the first reaction zone are pure carbon monoxide and pure chlorine, the volume flow of the carbon monoxide being adjusted so that there is an excess of 3 mol% of the carbon monoxide relative to chlorine at the beginning of the first reaction zone. The absolute inlet pressure of the process gases directly in front of the first reaction zone is 5 bar. The length of the fixed catalyst beds, ie the reaction zones, is between 0.1 and 1 m. The exact lengths of the reaction zones is shown in Table 1. The activity of the catalyst used is adjusted to be the same in all catalyst stages (for the other examples, this activity is 100% normalized). There is no replenishment of gas before the individual catalyst stages. The residence time in the system is 2 seconds.

Table 1: Length of the reaction zones according to Example 1

The results are shown in FIG. Here, the individual reaction zones are listed on the x-axis, so that a spatial course of developments in the process is visible. The temperature of the process gas is indicated on the left y-axis. The temperature profile over the individual reaction zones is shown as a thick, solid line; the temperature curve to be obtained in the best case at infinitesimal small reaction zones is shown as a thin dashed line. On the right y-axis the total conversion of chlorine is indicated. The course of the conversion over the individual reaction zones is shown as a thick dashed line. It can be seen that the inlet temperature of the process gas before the first reaction zone is about 100 ° C. Due to the exothermic reaction to phosgene under adiabatic conditions, the temperature in the first reaction zone rises to about 600 ° C before the process gas is cooled in the downstream heat exchange zone again. The inlet temperature before the next reaction zone is about 200 ° C. By exothermic adiabatic reaction, it rises again to about 550 ° C. The sequence of heating and cooling continues. The inlet temperatures of the process gas upstream of the individual reaction zones decrease with increasing reaction zone number. This is advantageous since, in the course of the reaction later reaction zones, the amount of reactants capable of the reaction is lower and, accordingly, a more advantageous position of the reaction equilibrium is sought after the conversion. Consequently, the temperature of the process gas can be kept closer to the optimum for the particular composition. Under certain circumstances, a process according to this specific embodiment requires an acceptance of lower catalyst service lives, which is advantageous in the case of inexpensive catalysts, such as activated carbon, if the correspondingly increased conversions and the reduced reaction zone number and thus lower investment costs more than compensate for this.

Another feature of the operation of the reaction zones under adiabatic conditions is shown in FIG. Considering in particular the shape of the temperature profile within the first reaction zone and in particular the shape of the temperature profile within the ninth to eleventh reaction zone, it can be seen that the slope of the temperature increase over the reaction zone decreases. This shows the essential property of the process that no significant heat sink is present in the reaction zones. For all temperature profiles in FIGS. 4 to 5, the slope of the individual temperature curves in the reaction zones never increases, but always have a constant or falling value greater than zero, which is an indication of the adiabatic operation of the reaction zones, since the temperature in the reaction zones can be driven in the vicinity of the equilibrium limitation.

It is a conversion of 99.2% of the input of the first reaction zone used chlorine, calculated from the remaining mass output of the last reaction zone obtained. The space-time yield achieved based on the mass of catalyst used is 13.7 kg _CO ci ₂ / kgK _at h. Example 2:

In this example, the process gas flows through a total of 18 reaction zones, ie over 18 fixed catalyst beds of activated carbon. Each after a reaction zone is a heat exchange zone in which the process gas is cooled before it enters the next reaction zone. The process gas used at the beginning, as well as the inlet pressure before the first reaction zone, identical to that of Example 1. The length of the catalyst stages, ie the reaction zones, is shown in Table 2. The activity of the catalyst is adjusted by dilution with catalytically inactive material so that the activity increases as indicated in Table 2. This achieves stable oscillation in a temperature window between 100 ° C and 200 ° C so that the temperature in the reaction zones never reaches a critical level at which catalyst deactivation would be feared. There is no replenishment of gas before the individual catalyst stages. The residence time in the system is 5 seconds in total.

Table 2: Length and activity of the catalyst in the reaction zones according to Example 2

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

It can be seen that the inlet temperature of the process gas before the first reaction zone is about 100 ° C. Due to the exothermic reaction to phosgene under adiabatic conditions on the less active catalyst, the temperature rises to about 200 ° C, before the process gas is cooled in the downstream heat exchange zone. The inlet temperature before the next reaction zone is again about 100 ° C. By exothermic adiabatic reaction, it rises again to about 200 ° C. The sequence of heating and cooling continues. The inlet temperatures of the process gas upstream of the individual reaction zones no longer increase with an increasing number of reaction zones until, starting from reaction zone 15, a smaller increase in the upstream heat exchange zone allows a certain increase to achieve a residual turnover. The reaction temperature is maintained here in the range of 100 ° C to 200 ° C to achieve a minimum catalyst deactivation. Under certain circumstances, a process according to this specific embodiment requires acceptance of lower space-time yields, which is advantageous in the case of expensive catalysts or high standstill costs if the correspondingly increased catalyst service lives justify the increased investment costs due to the increased reaction zone number.

It is a conversion of 99.3% of the input of the first reaction zone used chlorine, calculated from the remaining mass output of the last reaction zone obtained. The space-time yield achieved based on the mass of catalyst used is 6.6 kg _C oci2 / kgKath.

Example 3:

In this example, the process gas flows through a total of 17 reaction zones, ie over 17 fixed catalyst beds, of activated carbon. Each after a reaction zone is a

Heat exchange zone in which the process gas is cooled before moving to the next Reaction zone occurs. The process gas used in the beginning, as well as the inlet pressure before the first reaction zone, identical to that of Example 1. In contrast to the previous examples, the process gas used at the outset is not preheated to 100 ° C, but fed to the first reaction zone at 50 ° C. , The length of the catalyst stages, ie the reaction zones, is given in Table 3. The activity of the catalyst is adjusted by dilution with catalytically inactive material according to Table 3. This achieves a stable oscillation in a temperature window between 200 ° C and 300 ° C, so that the dissipated at this temperature in the heat exchange zones heat is well suited for the production of water vapor. There is no replenishment of gas before the individual reaction zones. The total residence time in the plant is 5.4 seconds.

Table 3: Length and activity of the catalyst in the reaction zones according to Example 3

The results are shown in FIG. Here, the individual reaction zones are listed on the x-axis, so that a spatial course of developments in the process is visible. The temperature of the process gas is indicated on the left y-axis. The temperature profile over the individual reaction zones is shown as a thick, solid line; the temperature curve to be obtained in the best case at infinitesimal small reaction zones is shown as a thin dashed line. Is on the right y-axis the total sales of chlorine indicated. The course of the conversion over the individual reaction zones is shown as a thick dashed line.

Due to the exothermic reaction to phosgene under adiabatic conditions, the temperature in the first reaction zone rises from 50 ° C to about 300 ° C, before the process gas is cooled by the downstream heat exchange zone again. By using the process gas at only 50 ° C at the beginning of the first reaction zone, energy can be saved. In addition, as compared to other procedures, reaction zones can be saved. The inlet temperature before the next reaction zone is about 200 ° C. By exothermic adiabatic reaction, it rises again to about 300 ° C. The sequence of heating and cooling continues. The inlet temperatures of the process gas before the individual reaction zones no longer change with increasing reaction zone number up to reaction zone 15. The reaction temperature can be kept within the stages 1 to 15 in the vicinity of an optimum temperature value for generating steam.

hi the adiabatic reaction zones 14 to 17 can be achieved by the low residual content of the reactants no increase in temperature of 100 ° C more. After reaction zones 14-16, the reaction gases in the heat exchange zones are cooled to individual inlet temperatures for the subsequent reaction zone. The adjusted temperatures allow further shifting of the thermodynamic equilibrium to the side of the phosgene.

It is a conversion of nearly 100% of the input of the first reaction zone used chlorine, calculated from the remaining mass output of the last reaction zone obtained. The residual content of chlorine in the process gas at the end of the last reaction stage is only 1.9 ppm. The space-time yield achieved, based on the mass of catalyst used, is 5.7 kg coc / kgKath.

Example 4:

In example 4, a procedure substantially similar to that in example 3 is carried out. Different are the relative catalyst activities and lengths of the reaction zones (see Table 4). Next, the process gas is reheated to 100 ° C used. The results are shown in FIG. Here, the individual reaction zones are listed on the x-axis, so that a spatial course of developments in the process is visible. The temperature of the process gas is indicated on the left y-axis. The temperature profile over the individual reaction zones is shown as a thick, solid line; the temperature curve to be obtained in the best case at infinitesimal small reaction zones is shown as a thin dashed line. On the right y-axis the total conversion of chlorine is indicated. The course of the conversion over the individual reaction zones is shown as a thick dashed line.

Table 4: Length and activity of the catalyst in the reaction zones according to Example 4

Due to the exothermic reaction to phosgene under adiabatic conditions, the temperature in the first reaction zone rises from 100 ° C to about 350 ° C, before the process gas is cooled by the downstream heat exchange zone again. The inlet temperature before the next reaction zone is about 250 ⁰ C. By exothermic adiabatic reaction, it rises again to about 350 ° C. The sequence of heating and cooling continues. The inlet temperatures of the process gas before the individual reaction zones no longer change with increasing number of reaction zones to reaction zone 13. In the reaction zones 14 to 17 can be achieved by the low residual content of the reactants no increase in temperature of 100 ⁰ C more. After the reaction zones 14 to 16, the reaction gases are cooled in the heat exchange zones to temperatures below 250 ° C. The low temperature allows further shifting of the thermodynamic equilibrium to the side of the phosgene.

The reaction temperature can be kept even closer to an optimal temperature value for generating steam within stages 1 to 13, whereby a steam thus produced is obtained at temperatures of about 200 ° C to 250 ° C and thus present at an overpressure of about 30 bar.

It is a conversion of nearly 100% of the input of the first reaction zone used chlorine, calculated from the remaining mass output of the last reaction zone obtained. The remaining residual content of chlorine in the process gas at the end of the last reaction stage is only 0.4 ppm. The space-time yield achieved, based on the mass of catalyst used, is 5.7 kg coc / kgKath.

Claims

claims

1. A process for the preparation of phosgene, characterized in that it comprises a conversion of carbon monoxide with chlorine to phosgene according to formula (I)

CO (g) + Cl ₂ (g) <→ COCl ₂ (g); ΔH = -108 kJ / mol (I)

in 6 to 40 consecutive reaction zones under adiabatic

Conditions in the presence of heterogeneous catalysts.

2. The method according to claim 1, characterized in that the conversion takes place in 8 to 30, preferably 10 to 25 successive reaction zones.

3. The method according to claim 1 or claim 2, characterized in that the inlet temperature of the entering into the first reaction zone process gas from

10 to 350 ° C, preferably from 40 to 250 ⁰ C, particularly preferably from 80 to 210 ⁰ C.

4. The method according to claim 1 to 3, characterized in that the absolute pressure at the entrance of the first reaction zone is between 1.1 and 40 bar, preferably between 1, 5 and 10 bar, more preferably between 2 and 6 bar.

5. The method according to any one of the preceding claims, characterized in that the residence time of the process gas in all reaction zones together between 0.1 and 100 s, preferably between 0.5 and 50 s, more preferably between 1 and 10 s.

6. The method according to any one of the preceding claims, characterized in that the

Catalysts consist of a material which is characterized in addition to its catalytic activity for the reaction according to formula (I) by sufficient chemical resistance to chlorine under the conditions of the process, as well as by a high specific surface area.

7. The method according to claim 6, characterized in that the catalysts

Silica gels, silicon carbides and / or activated carbon include.

8. The method according to claims 6 or 7, characterized in that the catalysts are present in a fixed bed arrangement.

9. The method according to claim 8, characterized in that the catalysts are present as monoliths.

10. The method according to any one of claims 6 or 7, characterized in that the catalysts are in fluidized bed arrangement.

11. The method according to claims 8 or 10, characterized in that the

Catalysts in beds of particles having average particle sizes of 1 to 10 mm, preferably 2 to 7 mm, particularly preferably 3 to 5 mm are present.

12. The method according to any one of the preceding claims, characterized in that after at least one reaction zone is at least one heat exchange zone through which the process gas is passed.

13. The method according to claim 12, characterized in that after each reaction zone at least one, preferably a heat exchange zone is located, through which the process gas is passed.

14. The method according to any one of the preceding claims, characterized in that between a reaction zone and a heat exchange zone at least one

Heat insulation zone is located.

15. The method according to claim 14, characterized in that there is a heat insulation zone around each reaction zone.

16. Reactor system for carrying out a method according to one of the preceding claims, characterized in that it comprises feed lines (Z) for a process gas comprising carbon monoxide and chlorine or for at least two process gases, of which at least one carbon monoxide and at least one chlorine and 6 to 40 in a row connected reaction zones (R) in the form of fixed beds of a heterogeneous catalyst, wherein between the reaction zones heat insulation zones (I) in the form of insulating material and between them

Heat exchange zones (W) are in the form of plate heat exchangers, which are connected to the reaction zones via supply and discharge lines for the process gases and the supply and discharge lines for a cooling medium.