fig1 shows the basic process diagram . an educt stream o containing at least one oxygenate and an educt stream c containing at least one c 4 + olefin are introduced into a reactor 20 and simultaneously converted there on a solid catalyst , such as an alumosilicate of the pentasil type zsm - 5 . the resulting product stream p subsequently is introduced into a first separating device 30 . in this separating device 30 , the stream p is separated into a mixture h rich in c 3 − olefins and a mixture b rich in c 4 + gasoline hydrocarbons . the stream b subsequently is introduced into a second separating device 40 and is separated there into a stream f substantially containing c 4 fractions , a stream a rich in c 5 + gasoline hydrocarbons , and a recycling stream r . the recycling stream r , which can chiefly contains c 4 to c 6 olefins in gaseous form , subsequently is at least partly recirculated into the reactor 20 . fig2 shows the configuration of a plant according to the invention . for this purpose , a stream containing water , such as water evaporated already , is admixed via conduit 1 to a stream containing at least one c 4 + olefin in conduit 2 . via conduit 3 and conduit 4 , the mixture of these streams is fed into a heat exchanger 10 and preheated there . the quantity of c 4 + olefin is controlled via the dosing device 6 . a likewise present dosing means for the steam is not shown in the figure . in addition , an oxygenate stream supplied via conduit 5 and dosing means 9 is evaporated and possibly superheated in this heat exchanger 10 . via conduit 11 , the olefin / steam stream is delivered into an undergrate firing 12 and from there via conduits 13 and 14 passed on into the reactor 20 . the quantity of the stream is adjusted via dosing devices in the conduits 1 and 2 as well as 43 . via conduit 15 , the pre - evaporated and optimally superheated oxygenate stream enters into a dosing device 8 , from where it is fed in directly before the fixed catalyst beds 20 b and 20 c via conduit 16 and conduit 18 . via conduit 19 , it is also possible to at least partly introduce the oxygenate stream into the olefin / steam mixture preheated to the reactor inlet temperature and along with the same guide it into the reactor 20 upstream of the catalyst bed 20 a . the quantity of the oxygenate stream is to be put into the relation to the olefin quantity contained in the olefin stream as defined according to the invention , which is why the oxygenate stream must be dosed in by control devices such as 7 or 8 . this dosing or allocation to the individual fixed catalyst beds is effected such that the temperature of a reaction gas stream exiting from a bed is adjusted to the desired value by dosing the oxygenates before the bed . the last catalyst bed 20 c is designed with a low load of oxygenates , so that the oxygenate conversion there becomes almost quantitative . corresponding to the c 4 + olefin content and the fraction of the supplied oxygenate , a decrease or increase in temperature by few degrees celsius each is obtained across a bed , whereby across the entire reactor 20 a temperature difference of about − 5 to + 13 ° c . is obtained . in principle , it would also be conceivable to at least partly directly feed the stream containing water into the reactor 20 via conduit 1 , which would require a further dosing device for the supplied quantity of water . at least partly admixing the stream of water into the oxygenate stream also is conceivable in principle . it is likewise conceivable to guide a part of the c 4 + olefin educt stream from conduit 2 or a part of the olefin - containing recycling stream from conduit 43 to conduit 5 or directly to one or more of the conduits 18 , 16 , 16 ′ and hence supply it to the reactor 20 . this would also require at least one further dosing means . via conduits 13 ′ and 19 ′ parts of the educt streams also can be introduced into a parallel reactor , and via conduit 19 ″ parts of the stream from conduit 19 can be fed into conduit 14 . here as well , a dosing means is possible . parallel reactors on the one hand serve to buffer the aging behavior of the catalyst by regenerating reactors at different times . a parallel reactor either can be operated just like the reactor 20 or also have another configuration , for example include another quantity of fixed catalyst beds . such parallel connection of reactors provides for a greater flexibility of the plant with regard to the converted mass flows . via conduit 21 ′, a product stream obtained in a possibly existing parallel reactor can be admixed to the product stream from reactor 20 in conduit 21 . the sum of the product streams enters into the heat exchanger 10 via conduit 21 and hence at the same time is utilized to heat the educt streams . in the heat recovery device 10 — a likewise existing integrated final cooler is not shown in the figure — a separation of the aqueous phase from the organic phase also is effected already . via conduit 22 , the gas phase is supplied to a compressor 23 , from where it is delivered into a first separating device 30 via conduit 24 . in the separating device 30 , the product mixture compressed to about 25 bar initially is separated into the c 3 − olefins and a fraction of the longer - chain olefins . while the c 3 − olefins are discharged via conduit 32 , the bottom product with the longer - chain olefins is delivered via conduit 31 into the second separating device 40 , which just like the separating device 30 can be is designed as distillation column . in this second separating device 40 , the c 4 hydrocarbons are separated as top product and the hydrocarbons containing at least five carbon atoms ( c 5 + hydrocarbons ) are separated as bottom product . from the separating device 40 , a return stream with a rather high residual content of c 4 to c 6 olefins is recirculated into conduit 3 via conduit 43 . the quantity of this stream is determined via the control means 44 . in an embodiment , the return stream is recirculated to the reactor in gaseous form . for this purpose , the separating means 40 is operated at a pressure of about 4 bar above the reactor pressure , and via conduit 43 a partial gas stream is withdrawn from the separating means at a suitable position . when the separating unit 40 is designed as integrated distillation column as shown in fig6 , a c 4 fraction is withdrawn via conduit 42 , wherein the quantity is determined by the reboiler at the upper part of the column . via conduit 46 , the residual quantity of c 4 is delivered into the lower part of the column along with c 5 hydrocarbons and via conduit 43 is guided to the reactor with an advantageous quantity of c 5 - 6 hydrocarbons . via the control of the reboiler at the lower part of the column it is determined which part of the c 5 + hydrocarbons is guided to the top of the column or is recovered as bottom product . the fraction c 6 + rich in aromatics chiefly is discharged from the process as product via conduit 31 . *** in the unit for heat recovery 10 , steam is condensed and separated from the reaction gases . in a non - illustrated stripper , a fraction containing oxygenates can be recovered from the aqueous condensate and be recirculated into the reactor 20 . the minimum configuration shown in fig2 can be employed above all in cases with a comparatively low use of oxygenates . a residual content of highly volatile reaction products , e . g . dimethyl ether ( dme ), compromising the propylene product can be separated with the bottom product propane at the latest in the non - illustrated c 3 splitter without a particular expenditure of apparatus . if this is not possible , because particular requirements are to be satisfied by the propane product or because a high propylene yield should be produced with a very high oxygenate feed , a more expensive configuration can be employed , which is shown in fig3 . the construction of the reactor unit is similar to the one described already in fig2 . however , the reactor is provided as four - stage reactor , which is why a further partial oxygenate stream is fed into the reactor 20 via conduit 16 ′ and dosing device 8 ′. after the compression to about 25 bar in the compressor 23 , the product streams initially are supplied via the conduits 24 and 25 to a separating device consisting of the columns 30 and 50 . this separating means is described in de 10 2008 058 931 a1 and by means of a washing agent — here methanol — separates the product mixture into a c 3 − fraction free from oxygenates ( conduit 51 ), a c 4 + stream ( conduit 31 ), and the washing agent chiefly loaded with dme / c 4 - 5 hydrocarbons ( conduit 54 ). via conduit 24 , the liquid product stream is guided to the separating device 30 and divided into a short - chain fraction ( c 3 − olefins ) and a long - chain fraction ( c 4 + olefins ). via conduit 32 , the short - chain fraction is admixed to conduit 25 and the combined stream — containing the entire c 3 − olefin yield — is supplied , to the separating means 50 . there , residual quantities of c 4 + hydrocarbons are separated from the c 3 − top product and at the same time highly volatile oxygenates , above all dme , are quantitatively removed from the c 3 − product by means of an oxygenate washing agent , such as methanol . the c 3 − olefins then are discharged via conduit 51 . in the further course — going beyond de 10 2008 058 931 a1 — the methanol / oxygenate / olefin stream is admixed either directly via conduit 61 to the oxygenate feed in conduit 5 and via conduit 5 ′ fed into the apparatus for heat recovery 10 . or an additional separating device 60 is connected , by which the contained c 4 - 5 olefins can be separated and can be withdrawn via conduit 62 . conduit 62 opens into conduit 43 , whereby the united olefin stream can be recirculated into conduit 3 via conduit 44 . the methanol bottom product of the column 60 is admixed to the oxygenate feed via conduit 61 . the separation of c 4 - 5 hydrocarbons from methanol has the effect that the hydrocarbons can be supplied to the reactor at a high temperature and accordingly improve the energy balance at very low oxygenate / olefin ratios . on the other hand , a further separating device 60 can be omitted with a relatively high oxygenate feed , because due to the “ dilution ” of the oxygenate by olefins a stronger cooling effect is achieved before charging the same to the catalyst bed , which is desirable with high oxygenate feeds . via conduit 31 , the longer - chain hydrocarbons from the separating means 30 are introduced into a separating device 40 , in which the c 4 hydrocarbons are separated from the c 5 + hydrocarbons . from the top of the distillation column 40 c 4 hydrocarbons are discharged via conduit 42 , while the c 5 + hydrocarbons leave the column via conduit 41 from the bottom of the column 40 . via conduit 43 , a recycling stream furthermore can be recirculated from the column 40 via conduit 44 into conduit 3 . the position of the conduit 43 in the column 40 is determined such that corresponding to the composition of the feed in conduit 31 a maximum c 4 to c 6 fraction of the recycling stream can be adjusted in conduit 43 . in an embodiment , the column is operated at a pressure of & gt ; 4 bar above the reactor pressure , so that the recycling stream can be recirculated to the reactor in gaseous form without compression . in fig4 , essential ones of the above - mentioned relationships for a typical c 4 to c 6 olefin cut from a cracker plant are illustrated . at a reactor pressure of 1 . 8 bar ( a ), a temperature profile of − 5 to + 13 ° c . is obtained across the entire reactor with a methanol feed of 15 - 30 % as defined according to the quantity ratio v of the invention . the temperature holding range is adjusted by different preheating temperatures of the methanol . the propylene : methanol ratio is 3 . 6 - 1 . 5 . with a total olefin conversion of 80 - 75 %, the achievable selectivity based on the total quantity of ethylene and propylene is about 59 . 5 - 57 %. like in the mtp ® process , a rather low pressure — i . e . about 1 . 7 - 1 . 9 bar ( a ) at the reactor inlet — promotes the propylene and ethylene yields . however , because c 4 + olefins in part represent undesired and lower - rated products , the inlet pressure in a reactor according to the invention can be raised to 3 - 5 bar , without thereby ruining the economy of the process . the high pressure provides for a compact construction of the plant and lower investment costs and costs for resources . under these boundary conditions , a maximum propylene yield is renounced . the higher the pressure , the lower the propylene yield and the less methanol ( oxygenate ) is required for adjusting the temperature . fig5 shows how at a reactor pressure of 4 . 5 bar ( a ) with a methanol feed (= v * 100 ) of 5 - 23 % a temperature profile of − 5 to + 13 ° c . is obtained . the propylene : methanol ratio is 10 : 2 . with a total olefin conversion of 85 - 80 %, the achievable selectivity based on the total quantity of ethylene and propylene is about 49 - 48 %. in example 1 it is assumed that from a c 4 + olefin cut of 62 t / h with 67 mass -% of olefins , of which 53 % are c 4 olefins , which is typical for a cracker plant , a rather large quantity of propylene and ethylene should be produced in an economically expedient way . for this purpose , an olefin - methanol conversion plant , as it is shown in fig3 , is configured , which should be operated at a reactor pressure of 1 . 8 bar and a relatively high recirculation rate ( 1 . 38 ) and methanol feed ( 24 %). the reactor target temperature is about 497 ° c . in the case of only three separating devices 30 , 40 and 50 a temperature increase of 7 . 5 ° c . is to be expected . the separating unit 40 is designed as integrated column as shown in fig6 . in table 1a mass balance is listed , and in table 2 the feed / product quantities and important characteristics of such process are listed . for the reactor a total height of about 9 m and a diameter of 5 . 5 m is assumed . in example 2 , a reconstruction of an existing cracker plant into a plant according to the invention ( revamp situation ) is assumed . from the same c 4 + olefin cut as in example 1 additional quantities of propylene and ethylene should be produced with a minimum of investment costs . at the same time , the mass flow and olefin content of the c 4 + streams should be reduced distinctly , in order to save the capacity of the downstream processing plants . for this purpose , an olefin / methanol conversion plant , as shown in fig2 , is configured , which should be operated at a reactor pressure of 4 . 5 bar and a mean recirculation rate ( 1 . 0 ) and only 9 % methanol feed . the reactor target temperature is about 497 ° c . here as well , the separating unit 40 is designed as integrated column according to fig6 . in table 2 , the feed / product quantities and important characteristics are listed . for the reactor a total height of about 7 m and a diameter of 3 . 4 m is assumed . as compared to example 1 , the dimensions of the downstream plants up to the compressor are reduced correspondingly . in this example , too , a revamp situation is assumed in a cracker plant . from a c 4 + olefin cut ( 40 t / h , 50 % olefin content ), additional quantities of propylene and ethylene should be produced with a minimum of investment costs . in this case , 3 t / h of alcohol mixture with c 2 to c 5 alcohols , which are obtained as by - product in a bioethanol plant , are available beside methanol . for this purpose , an olefin / methanol conversion plant , as it is shown in fig2 , is configured , which should be operated at a reactor pressure of 4 . 5 bar and a mean recirculation rate ( 1 . 0 ). the oxygenate fraction is 29 %. the reactor target temperature is about 484 ° c . ; the temperature profile across the reactor is − 4 . 5 ° c . the slight decrease in temperature in this case above all is due to the fact that the alcohol mixture also contains ethanol , which in a highly endothermal reaction chiefly is converted to ethylene . in table 2 , the feed / product quantities and important characteristics are listed . for the reactor a total height of about 7 m and a diameter of 2 . 8 m is assumed . in table 3 , characteristic data for various olefin conversion plants according to the invention are listed and compared with the propylur process . in the example , the same c 4 + cut as in examples 1 and 2 is used as feed . with respect to the yields of propylene and ethylene and the consumption of resources , distinct , in part drastic advantages of the process of the invention as compared to the propylur process are to be noted . the same also would apply for the ocp process . in these examples , the separating unit 40 is designed as simple column with gaseous side draw . in a fischer - tropsch plant ( f - t plant ) olefins and alcohols are obtained as by - products . the same are hydrogenated or used for undergrate firing purposes . in an olefin conversion plant according to the invention , which is integrated with the f - t plant , the olefin - and alcohol - containing streams can be used for propylene production . the yields are shown in table 4 . 1 - 5 conduits 6 - 9 dosing devices 10 heat recovery device 11 conduit 12 undergrate firing 13 - 19 conduits 20 reactor 20 a - d fixed catalyst beds 21 - 22 conduits 23 compressor 24 - 25 conduits 30 separating means 31 - 32 conduit 40 separating means 41 - 43 conduits 44 dosing device 45 conduit 46 conduit 50 separating means 51 , 52 , 54 conduits 60 separating means 61 conduit o oxygenate - containing educt stream c stream containing c 4 + olefins p product stream h c 3 − olefin stream b stream containing c 4 + olefins f stream containing c 4 olefins a stream containing c 5 + olefins r recycling stream containing c 4 - 6 olefins