Patent Publication Number: US-2011060149-A1

Title: Process for preparing ethylene oxide

Description:
The present invention relates to a process for preparing ethylene oxide by catalytic gas-phase oxidation of ethylene by means of oxygen, wherein the reaction is carried out in from 5 to 50 reaction zones connected in series under adiabatic conditions, and also a reactor system for carrying out the process. 
     Ethylene oxide is generally prepared from gaseous ethylene and oxygen in the presence of e.g. silver catalysts, in an exothermic, catalytic reaction according to formula (I): 
       H 2 C=CH 2  (g)+0.5·O 2  (g)→C 2 H 4 O (g); ΔH=−107 kJ/mol  (I)
 
     The ethylene oxide prepared by means of the reaction according to formula (I) forms an essential starting material for many other syntheses in the chemical industry. Ethylene oxide is preferably used in the preparation of polymers, such as polyethylene glycols, but also as a material for the chemical sterilization of materials which are noxious and not amenable to a heat treatment. 
     The removal and use of the heat of reaction is an important issue in carrying out the ethylene oxide synthesis. An uncontrolled temperature rise can lead to permanent damage to the catalyst. In addition, the possibility of secondary reactions to form more or less large amounts of carbon dioxide and water, which arise through total oxidation of ethylene and also ethylene oxide, exists at high temperatures. It is therefore advantageous to control the temperature of the catalysts during the process so as to keep them at a level which allows a rapid reaction with minimization of the secondary reactions and/or catalyst deactivation. 
     EP 0 821 678 B1 discloses a process for preparing ethylene oxide from ethylene and gases comprising oxygen, in which the conversion is carried out over a silver-based heterogeneous catalyst in a single reaction zone which is carried out multiply in parallel. It is also disclosed that, in particular, controlled cooling of the reaction zone presents a technical problem which is said to be solved in this disclosure by means of a particular flow and temperature regime of the cooling liquid. The catalyst temperatures disclosed are in the range from 150° C. to 350° C. Adiabatic operation is not disclosed. The reactor inlet pressures disclosed are in the range from 1 to 40 bar. 
     The process disclosed is disadvantageous because a high technical outlay has to be employed in the single reaction zone in order to keep the exothermic reaction within temperature ranges which are advantageous for the reaction of ethylene with oxygen to form ethylene oxide. Furthermore, it can be assumed that, owing to the high throughputs which are to be achieved and the associated large spatial dimensions of the reaction zone, the temperature profile established therein is, at least in subregions, either below the optimum temperature for achieving a high selectivity of the conversion of ethylene into ethylene oxide or above this. Furthermore, this process does not allow the reaction enthalpy to be utilized for increasing the reaction rate. 
     U.S. Pat. No. 6,172,244 B1 discloses an apparatus and a process for the heterogeneous catalytic reaction of ethylene with oxygen to form ethylene oxide, with a plurality of parallel reaction zones being surrounded by surrounding cooling walls. 
     It is also disclosed that further gases apart from ethylene and oxygen can be fed into the process/apparatus to prevent hot spots. Such gases can be, for instance, nitrogen or methane. It is pointed out that, especially the secondary reaction to form carbon dioxide and water, is particularly exothermic and should therefore be avoided. Neither customary operating temperatures nor entry temperatures of the process gases are disclosed. 
     The apparatus disclosed and the process carried out therein are disadvantageous because, here too, as in the disclosure of EP 0 821 678 B1, a high outlay in terms of apparatus is required in order to bring about direct cooling of the reaction zone. This in turn automatically leads to the reaction enthalpy of the main reaction not being utilized for increasing the reaction rate. At the same time, countermeasures which do not lead to a reduced throughput of the process or its total cessation when the reaction reaches a critical temperature level cannot readily be undertaken with constant cooling of the total single-stage, parallel process since the overall reactor system and not individual reaction zones is always affected. 
     A two-stage process is disclosed in U.S. Pat. No. 6,717,001 B2, in which a fresh catalyst is used in a first reaction zone and an aged catalyst is used in a further reaction zone. To compensate for the decrease in activity of the aged catalyst, an increased proportion of ethylene in the process gas of from 1.1 times to 4 times the proportion in the first reaction zone is used in the further reaction zone. Furthermore, it is disclosed that at reaction temperatures of from 180° C. to 325° C. the oxygen concentrations in the reaction zones should be kept in a range such that ignition of the process gases cannot take place. The possible inlet pressures of the process are in the range from 10 to 35 bar. It is also disclosed that the proportion of ethylene in the process gases fed to the process should be below 50 mol %. Cooling between the first reaction zone and the further reaction zone is not disclosed. 
     The process disclosed is disadvantageous because no means of cooling the process gases between the reaction zones are disclosed. The process is therefore disadvantageous safetywise since in the event of the temperature in the reaction zones increasing to above the planned level there are no means available for preventing ignition of the process gases in the process. The possibility of controlling the temperatures in the process is not disclosed in U.S. Pat. No. 6,717,001 B2. 
     EP 1 251 951 (B1) discloses an apparatus and the opportunity of carrying out chemical reactions in the apparatus, where the apparatus is characterized by a cascade of reaction zones and heat exchange apparatuses which are in contact with one another and are integrated with one another in terms of material. The process to be carried out therein is thus characterized by contact of the various reaction zones with a respective heat exchange apparatus in the form of a cascade. A disclosure in respect of the usability of the apparatus and of the process for the synthesis of ethylene oxide from gaseous oxygen and ethylene is not to be found. It therefore remains unclear how, proceeding from the disclosure of EP 1 251 951 (B1), such a reaction can be carried out by means of the apparatus and the process carried out therein. Furthermore, for reasons of unity, it has to be assumed that the process disclosed in EP 1 251 951 (B1) is carried out in an apparatus identical or similar to the disclosure in respect of the apparatus. As a result, due to the large-area contact of the heat exchange zones with the reaction zones as per the disclosure, a significant amount of heat is transferred by thermal conduction between the reaction zones and the adjacent heat exchange zones. The disclosure in respect of the oscillating temperature profile can thus only be interpreted as meaning that the temperature peaks found here would be larger if this contact did not exist. A further indication of this is the exponential rise in the disclosed temperature profiles between the individual temperature peaks. These indicate that some heat sink which has an appreciable but limited capacity and can reduce the temperature rise is present in each reaction zone. It can never be ruled out that some removal of heat (e.g. by radiation) takes place, but a reduction in the possible heat removal from the reaction zone would be indicated by a linear temperature profile or a temperature profile having a degressive gradient, since no further introduction of starting materials is provided and after an exothermic reaction, the reaction would proceed ever more slowly and thus with a reduced evolution of heat. Thus, EP 1 251 951 (B1) discloses multistage processes in cascades of reaction zones from which heat is removed in an undefined amount by thermal conduction. Accordingly, the process disclosed has the disadvantage that precise temperature control of the process gases of the reaction is not possible. 
     Proceeding from the prior art, it would therefore be advantageous to provide a process which can be carried out in simple reaction apparatuses and allows precise, simple temperature control so that it allows high conversions at very high product purities. Such simple reaction apparatuses would be simple to scale up to an industrial scale and are inexpensive and robust in all sizes. 
     As just indicated, neither suitable reactors nor suitable processes which allow these objectives to be achieved have hitherto been described for the catalytic gas-phase oxidation of ethylene by means of oxygen to form ethylene oxide. 
     It is therefore an object of the invention to provide a process for the catalytic gas-phase oxidation of ethylene by means of oxygen to form ethylene oxide, which process can be carried out with precise temperature control in simple reaction apparatuses and thus allows high conversions at high product purities, with the heat of reaction being able to be utilized to the benefit of the reaction or in another way. 
     It has surprisingly been found that a process for preparing ethylene oxide from ethylene and oxygen in the presence of heterogeneous catalysts, characterized in that it comprises from 5 to 50 reaction zones which are connected in series and have adiabatic conditions, is able to achieve this object. 
     The term ethylene refers, in the context of the present invention, to a process gas which is introduced into the process of the invention and comprises ethylene. The proportion of ethylene in the process gases fed to the process is usually in the range from 15 to 50 mol %, preferably from 20 to 40 mol %. 
     The term oxygen refers, in the context of the present invention, to a process gas which is introduced into the process of the invention and comprises oxygen. The proportion of oxygen in the process gases fed to the process is usually in the range from 5 to 30 mol %, preferably from 15 to 25 mol %. 
     Apart from the essential components of the process gases ethylene and oxygen, these gases can also comprise secondary components. Nonexhaustive examples of secondary components which can be present in the process gases are, for instance, argon, nitrogen, carbon dioxide, methane and/or ethane. 
     In general, process gases are, in the context of the present invention, gas mixtures which comprise oxygen and/or ethylene and/or ethylene oxide and/or secondary components. 
     For the purposes of the invention, carrying out the process under adiabatic conditions means that essentially no heat is either actively introduced or actively removed from the reaction zone from or to the outside. It is generally known that complete insulation against introduction or removal of heat can be achieved only by complete evacuation and ruling out heat transfer by radiation. Therefore, in the context of the present invention, adiabatic means that no measures for introducing or removing heat are taken. 
     In an alternative embodiment of the process of the invention, heat transfer can be reduced by, for example, insulation by means of generally known insulation materials, e.g. polystyrene insulation materials, or by means of sufficiently large distances to heat sinks or heat sources, with the insulation material being air. 
     An advantage of the adiabatic mode of operation according to the invention of the 5 to 50 reaction zones connected in series over a nonadiabatic mode of operation is that no means of removing heat have to be provided in the reaction zones, which results in a considerable simplification of the construction. Simplifications in the manufacture of the reactor and also in the scalability of the process and an increase in the reaction conversions are, in particular, obtained in this way. In addition, the heat generated during the course of the exothermic reaction is utilized in a controlled manner in the individual reaction zone to increase the conversion. 
     A further advantage of the process of the invention is the possibility of very precise temperature control by means of the close spacing of adiabatic reaction zones. It is thus possible for a temperature advantageous to the progress of the reaction to be set and controlled in each reaction zone. 
     The catalysts used in the process of the invention are usually catalysts comprising a material which not only have catalytic activity for the reaction according to formula (I) but are also characterized by sufficient chemical resistance under the conditions of the process and also by a high specific surface area. Catalyst materials which are characterized by such a chemical resistance under the conditions of the process are, for example, catalysts comprising silver which is supported on aluminum oxide. 
     The term specific surface area refers, in the context of the present invention, to the area of the catalyst material which can be reached by the process gas, based on the mass of catalyst material used. 
     A high specific surface area is a specific surface area of at least 10 m 2 /g, preferably at least 20 m 2 /g. 
     The catalysts used according to the invention are in each case located in the reaction zones and can be present in all forms known per se, e.g. fixed bed, moving bed, fluidized bed. 
     Preference is given to fixed beds and moving beds. 
     The fixed-bed arrangement comprises a catalyst bed in the actual sense, i.e. loose, supported or unsupported catalyst of any shape, and also in the form of suitable packings. The term catalyst bed as used here also encompasses contiguous regions of suitable packings on a support material or structured catalyst support. These would be, for example, ceramic honeycomb bodies having comparatively high geometric surface areas to be coated or corrugated layers of woven metal wire mesh on which, for example, catalyst granules are immobilized. In the context of the present invention, a special form of packing is the presence of the catalyst in monolithic form. 
     A particularly preferred embodiment of a fixed-bed arrangement is a monolithic catalyst comprising silver supported on aluminum oxide. 
     If a catalyst in monolithic form is used in the reaction zones, the catalyst present in monolithic form is, in a preferred embodiment of the invention, provided with channels through which the process gases flow. The channels usually have a diameter of from 0.1 to 3 mm, preferably a diameter of from 0.2 to 2 mm, particularly preferably from 0.3 to 1 mm. 
     A monolithic catalyst having channels of the diameter indicated is particularly advantageous since protection against explosion can be ensured thereby. This is achieved by uptake of the enthalpy by the wall of the monolith, and for this reason the further spread of flames is suppressed. 
     If a moving-bed arrangement of the catalyst is used, the catalyst is preferably present in loose beds of particles. 
     In a preferred embodiment of the process of the invention, the reaction is carried out in from 7 to 40, particularly preferably from 10 to 30, reaction zones connected in series. 
     A preferred further embodiment of the process is characterized in that the process gas leaving at least one reaction zone is subsequently passed through at least one heat exchange zone located downstream of this reaction zone. 
     In a particularly preferred further embodiment of the process, each reaction zone is followed by at least one, preferably precisely one, heat exchange zone through which the process gas leaving the reaction zone is passed. 
     The reaction zones can either be arranged in one reactor or be divided between a plurality of reactors. The arrangement of the reaction zones in one 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 one reactor or in any combinations of reaction zones with heat exchange zones in a plurality of reactors. 
     If reaction zones and heat exchange zones are present in one reactor, a thermal insulation zone is, in an alternative embodiment of the invention, present between these in order to be able to maintain adiabatic operation of the reaction zone. 
     In addition, individual reaction zones among the reaction zones connected in series can also, independently of one another, be replaced or supplemented by one or more reaction zones connected in parallel. The use of reaction zones connected in parallel allows, in particular, replacement or supplementation of these during ongoing continuous overall operation of the process. 
     Parallel reaction zones and reaction zones connected in series can, in particular, also be combined with one another. However, the process of the invention particularly preferably has exclusively reaction zones connected in series. 
     The reactors which are preferably used in the process of the invention can comprise simple vessels having one or more reaction zones, as are described, for example, in Ullmanns Encyclopedia of Industrial Chemistry (Fifth, Completely Revised Edition, Vol B4, pages 95-104, pages 210-216), with thermal insulation zones being able to be additionally provided in each case between the individual reaction zones and/or heat exchange zones. 
     In an alternative embodiment of the process, at least one thermal insulation zone is thus located between a reaction zone and a heat exchange zone. Preference is given to a thermal insulation zone being present around each reaction zone. 
     The catalysts or the fixed beds of catalysts are applied in a manner known per se to or between gas-permeable walls comprising the reaction zone of the reactor. Particularly in the case of thin fixed beds, technical devices for obtaining uniform distribution of gas can be installed upstream of the catalyst beds. These can be perforated plates, bubble cap trays, valve trays or other internals which, by producing a small but uniform pressure drop, bring about uniform entry of the process gas into the fixed bed. 
     In a particular embodiment of the process of the invention, preference is given to using an excess of from 0 to 50% of oxygen based on the molar flow of ethylene before entry into the reaction zone. An increase in the ratio of oxygen to ethylene enables the reaction to be accelerated and thus the space-time yield (amount of ethylene oxide produced per mass of catalyst material) to be increased. 
     In a further particularly preferred embodiment of the process, the entry temperature of the process gas entering the first reaction zone is from 10 to 290° C., preferably from 50 to 270° C., particularly preferably from 100 to 250° C. 
     In another particularly preferred embodiment of the process, the absolute pressure at the entrance into the first reaction zone is in the range from 3 to 30 bar, preferably from 5 to 25 bar, particularly preferably from 7 to 20 bar. 
     In a further particularly preferred embodiment of the process, the residence time of the process gas in a reaction zone is in the range from 1 to 60 s, preferably from 2 to 30 s, particularly preferably from 5 to 20 s. 
     The ethylene and the oxygen are preferably fed in only upstream of the first reaction zone. This has the advantage that the entire process gas can be utilized for taking up and removing the heat of reaction in all reaction zones. In addition, such a mode of operation enables the space-time yield to be increased or the mass of catalyst necessary to be reduced. However, it is also possible to introduce ethylene and/or oxygen into the process gas as required before one or more of the reaction zones following the first reaction zone. The introduction of gas between the reaction zones additionally allows the temperature of the reaction to be controlled. 
     In a preferred embodiment of the process of the invention, the process gas is cooled after at least one of the reaction zones used, particularly preferably after each of the catalyst beds used. For this purpose, the process gas leaving a reaction zone is passed through one or more of the abovementioned heat exchange zones which are located downstream of the respective reaction zones. These can be configured as heat exchange zones in the form of the heat exchangers known to those skilled in the art, e.g. shell-and-tube, plate, annular groove, spiral, finned tube, micro heat exchangers. The heat exchangers are preferably microstructured heat exchangers. 
     The term microstructured means, in the context of the present invention, that the heat exchanger has, for the purposes of heat transfer, fluid-conducting channels which are characterized in that they have a hydraulic diameter in the range from 50 μm to 5 mm. The hydraulic diameter is given by four times the cross-sectional area of the fluid-conducting channel through which flow occurs divided by the circumference of the channel. 
     In a further embodiment of the process, steam is generated by the heat exchanger during cooling of the process gas in the heat exchange zones. 
     Within this further embodiment, preference is given to carrying out a vaporization, preferably partial vaporization, on the side of the cooling medium in the heat exchangers comprising the heat exchange zones. 
     In the context of the present invention, partial vaporization is vaporization in which a gas/liquid mixture of a substance is used as cooling medium and in which a gas/liquid mixture of a substance is still present after heat transfer in the heat exchanger. 
     Carrying out a vaporization is particularly advantageous because the achievable heat transfer coefficient from/to process gases to/from cooling/heating medium becomes particularly high as a result and efficient cooling can therefore be achieved. 
     The carrying out of a partial vaporization is particularly advantageous because the uptake/release of heat by the cooling medium then no longer results in a temperature change in the cooling medium but only produces a shift in the gas/liquid equilibrium. As a result, the process gas is cooled against a constant temperature over the entire heat exchange zone. This in turn reliably prevents occurrence of temperature profiles in the flow of the process gases, as a result of which control over the reaction temperatures in the reaction zones is improved and, in particular, the formation of local hot spots due to temperature profiles is prevented. 
     In an alternative embodiment, a mixing zone can be provided instead of a vaporization/partial vaporization before the entrance to a reaction zone in order to even out any temperature profiles in the flow of the process gases arising during cooling by mixing transverse 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 which increases or decreases from reaction zone to reaction zone. This means that, within a sequence of reaction zones, the temperature can both increase and decrease from reaction zone to reaction zone. This can be achieved, for example, by control of the heat exchange zones located between the reaction zones. Further possibilities for setting the average temperature are described below. 
     The thickness of the reaction zones through which flow occurs can be made identical or different and is derived according to laws generally known to those skilled in the art from the above-described residence time and the amounts of process gas put through the process in each case. The mass flows of product gas (ethylene oxide) which can be put through the process according to the invention, from which the amounts of process gas to be used are also derived, are usually in the range from 0.01 to 45 t/h, preferably from 0.1 to 40 t/h, particularly preferably from 1 to 35 t/h. 
     The maximum exit temperature of the process gas from the reaction zones is usually in the range from 260° C. to 320° C., preferably from 270° C. to 310° C., particularly preferably from 280° C. to 300° C. The control of the temperature in the reaction zones is preferably effected by means of at least one of the following measures: dimensioning of the adiabatic reaction zone, control of the heat removal between the reaction zones, addition of gas between the reaction zones, molar ratio of the starting material/excess of oxygen used, addition of inert gases, in particular nitrogen or methane and/or argon, before and/or between the reaction zones. 
     The composition of the catalysts in the reaction zones according to the invention can be identical or different. In a preferred embodiment, the same catalysts are used in each reaction zone. However, different catalysts can also advantageously be used in the individual reaction zones. Thus, it is possible, in particular, to use a less active catalyst in the first reaction zone where the concentration of the reactants is still high and to increase the activity of the catalyst from reaction zone to reaction zone in the further reaction zones. The catalyst activity can also be controlled by dilution with inert materials or support material. The use of a catalyst which is particularly stable toward deactivation at the temperatures of the process in the first and/or second reaction zones in these reactions zones is likewise advantageous. 
     The process of the invention makes it possible to produce, per 1 kg of catalyst, from 0.1 kg/h to 2 kg/h, preferably from 0.2 kg/h to 1 kg/h, particularly preferably from 0.3 kg/h to 0.5 kg/h, of ethylene oxide. 
     The process of the invention is thus characterized by high space-time yields, combined with a reduction in the sizes of the apparatuses and a simplification of the apparatuses or reactors. This surprisingly high space-time yield is made possible by interaction of the inventive and preferred embodiments of the novel process. In particular, the interaction of gradated, adiabatic reaction zones with heat exchange zones located between them and the defined residence times makes possible precise control of the process and the resulting high space-time yields and also a reduction in the by-products formed, e.g. CO 2  and water, is achieved. 
     The invention further provides a reactor system for reacting ethylene and oxygen to form ethylene oxide, characterized in that it comprises feed lines (Z) for a process gas comprising ethylene and oxygen or for at least two process gases of which at least one comprises ethylene and at least one comprises oxygen and comprises from 5 to 50 reaction zones (R) which are connected in series and are in the form of fixed beds of a heterogeneous catalyst, where thermal insulation zones (I) in the form of insulation material are located between the reaction zones and heat exchange zones (W) in the form of plate heat exchangers which are connected to the reaction zones via feed lines and discharge lines for the process gases and comprise feed lines and discharge lines for a cooling medium are located between these thermal insulation zones. 
     The reactor system can also comprise from 7 to 40, preferably from 10 to 30, reaction zones in the form of fixed beds. 
     The insulation material of the thermal insulation zones is preferably a material having a coefficient of thermal conductivity λ less than or equal to 0.08. 
     
       
         
           
             
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     Particular preference is given to, for instance, polystyrene, polyurethanes, glass wool or air. 
    
    
     
       The present invention will be illustrated with the aid of the drawings, but is not restricted thereto. 
         FIG. 1  schematically shows an embodiment of the reactor system of the invention, where the following reference numerals are used in the drawing: 
       Z: feed line(s) 
       R: reaction zone(s) 
       I: thermal insulation zone(s) 
       W: heat exchange zone(s) 
         FIG. 2  shows reactor temperature (T), ethylene conversion (U) and ethylene oxide selectivity (Y) over a number of 18 reaction zones (S) followed by heat exchange zones (as per example 1). 
         FIG. 3  shows reactor temperature (T), ethylene conversion (U) and ethylene oxide selectivity (Y) over a number of 12 reaction zones (S) followed by heat exchange zones (as per example 2). 
     
    
    
     The present invention will also be illustrated by examples 1 and 2 below, without being restricted thereto. 
     EXAMPLES 
     Example 1: 
     In this example, the process gas flows through a total of 18 fixed catalyst beds in the form of monoliths which have channel diameters of 0 5 mm and are coated with a catalyst comprising silver on aluminum oxide, i.e. through 18 reaction zones. After each reaction zone, there is a heat exchange zone in which the process gas was cooled before entering the next reaction zone. The process gas used at the entry to the first reaction zone contains 31.9 mol % of ethylene, 21.1 mol % of oxygen, 32.3 mol % of methane, 2.2 mol % of CO 2 , 10.1 mol % of argon, 1.1 mol % of ethane and 1.3 mol % of nitrogen. The absolute entry pressure of the process gas directly before the first reaction zone is 10 bar. The length of the fixed catalyst beds, i.e. the reaction zones, is in each case 1 m. The amount of catalyst coated on the monoliths is 20% by weight. No further gas is introduced before the individual catalyst stages. The total residence time in the plant is 14 seconds. 
     The results are shown in  FIG. 2 . Here, the individual reaction zones are shown on the x axis, so that a spatial course of the developments in the process can be seen. The temperature of the process gas is indicated on the left-hand y axis. The course of the temperature over the individual reaction zones is shown as a bold, solid line. The total conversion of ethylene and the selectivity to ethylene oxide is indicated on the right-hand y axis. The course of the conversion over the individual reaction zones is shown as a bold broken line. The course of the selectivity is shown as a thin solid line. 
     It can be seen that the entry temperature of the process gas before the first reaction zone is about 289° C. As a result of the exothermic reaction to form ethylene oxide under adiabatic conditions, the temperature rises to about 300° C. in the first reaction zone before the process gas is cooled again in the following heat exchange zone. The entry temperature before the next reaction zone is again about 289° C. As a result of the exothermic adiabatic reaction, it increases again to about 300° C. The sequence of heating and cooling continues. The entry temperatures of the process gas before the individual reaction zones changes only slightly to about 291° C. over the course of the process. 
     A further feature of the operation of the reaction zones under adiabatic conditions is shown in  FIG. 2 . When the shape of the temperature profile within the reaction zones and the shape of the temperature profile thereof are examined, it can be seen that the gradient of the temperature rise never increases over the reaction zone. This shows the important property of the process that no significant heat sink is present in the reaction zones. 
     A conversion of ethylene of 11.3 mol % is obtained. The selectivity obtained is 88.3 mol %. The space-time yield achieved, based on the mass of catalyst used, is 0.43 kg ethylene oxide /kg cat h. 
     Example 2: 
     In this example, the process gas flows through a total of 12 reaction zones, i.e. through 12 fixed catalyst beds in the form of monoliths, these now having channel diameters of 0.8 mm but otherwise being the same as those in example 1. After each reaction zone, there is a heat exchange zone in which the process gas is cooled before entering the next reaction zone. The process gas used at the beginning and also the entry pressure before the first reaction zone are identical to those in example 1. The length of the reaction zones is always 1 m. The amount of catalyst coated on the monoliths is 35% by weight. Thus, oscillation in a temperature window from 280° C. to 300° C. is achieved after the first reaction zone, into which temperature window the process oscillates. No further gas is introduced before the individual catalyst stages. The total residence time in the plant is 10 seconds. 
     The results are shown in  FIG. 3 . Here, the individual reaction zones are shown on the x axis, so that a spatial course of the developments in the process can be seen. The temperature of the process gas is indicated on the left-hand y axis. The course of the temperature over the individual reaction zones is shown as a bold, solid line. The total conversion of ethylene and the selectivity to ethylene oxide is indicated on the right-hand y axis. The course of the conversion over the individual reaction zones is shown as a bold broken line. The course of the selectivity is shown as a thin solid line. 
     It can be seen that the entry temperature of the process gas before the first reaction zone is about 282° C. As a result of the exothermic reaction to form ethylene oxide under adiabatic conditions, the temperature rises to about 300° C. in the first reaction zone before the process gas is cooled again in the following heat exchange zone. The entry temperature before the next reaction zone is again about 283° C. As a result of the exothermic adiabatic reaction, it increases again to about 300° C. The sequence of heating and cooling continues. The entry temperatures of the process gas before the individual reaction zones changes only slightly to about 286° C. over the course of the process. A level of performance of the process analogous to example 1 can therefore be achieved by using a higher proportion of catalyst or by using a more active catalyst, if appropriate in fewer stages, without there being the risk of the process overheating. 
     A conversion of 11.6 mol % of the ethylene used initially in the first reaction zone, calculated from the remaining mass at the exit from the last reaction zone, is obtained. The selectivity to ethylene oxide is about 87.7 mol %. The space-time yield achieved, based on the mass of catalyst used, is 0.37 kg ethylene oxide /kg cat h.