Method for preparing propylene

Abstract

The present invention relates to a process for preparing propylene from a hydrocarbon mixture having an olefin content of 5-50 wt% and boiling more than 90 vol% between 35°C and 280°C, or from a hydrocarbon feed containing paraffins, cycloparaffins and / or aromatic compounds and optionally up to 10 wt% of olefins, first in a reactor by contacting the feed with a low acid density cracking catalyst, separating the propylene, and then contacting the residue with a high acid density cracking catalyst at a higher temperature in the reactor, separating the propylene, and recycling the residue to a first cracking reactor and a second cracking reactor. Aromatic compounds may be added in the first and second cracking steps to improve the cycle time.

Description

Technical Field

[0001] The present invention relates to a process for preparing propylene from a hydrocarbon mixture having an olefin content of 5-50 wt% and boiling more than 90 vol% between 35°C and 280°C, and / or from a hydrocarbon feed comprising paraffins, cycloparaffins, aromatics and optionally up to 10 wt% of olefins by contacting the feed with a cracking catalyst in a reactor. Background Art

[0002] Propylene produced by the steam cracking process accounts for more than 50%.A typical feedstock is straight run naphtha obtained when refining a crude oil source, which generally comprises unsaturated compounds such as paraffins and cycloparaffins optionally mixed with aromatic compounds.

[0003] Propylene is also produced in a refinery environment as a byproduct of the fluid catalytic cracking (FCC) process. Since the late 1990s, some FCC units have been operating at a relatively severe level to achieve a propylene yield of 10-12 wt% of fresh FCC feed. In order to further increase the propylene yield, different methods have been developed around the FCC configuration of the refinery, and it is reported that propylene yields have reached 20 wt% of fresh FCC feed. One way to increase the propylene yield is to add a medium pore zeolite to the FCC catalyst, such as described in DE4114874. A variety of variants have been developed in which the medium pore catalyst and the FCC catalyst contact the hydrocarbon fraction in the FCC riser reactor. The disadvantage of these methods is that the medium pore zeolite catalyst will undergo a regeneration step together with the FCC catalyst, causing the medium pore zeolite catalyst to degrade.

[0004] The naphtha fraction obtained in the FCC process can also be contacted in a separate process, where the feed is contacted with a cracking catalyst in a fixed bed reactor. One such process is described in WO99 / 29804, which discloses a fixed bed reactor process in which an olefin-rich feedstock is contacted with a crystalline silicate catalyst. In an embodiment, light cracked naphtha (LCN) is cracked using a crystalline silicate catalyst. The propylene yield is about 18 wt% based on the feed. Experiments using ZSM-5 and using 1-hexene feed show that the highest propylene yield using ZSM-5 with a Si / Al atomic ratio of 350 (SAR=750) is 28.8 wt%, while experiments using ZSM-5 with Si / Al atomic ratios of 40 and 25 (SAR=80, SAR=50) show lower propylene yields and more coke formation.

[0005] GB2345294 describes a process in which a C4 residual oil liquid feed containing olefins is contacted with a cracking catalyst in a fixed bed reactor. The catalyst consists of ZSM-5 containing silver instead of protons, wherein the SAR of ZSM-5 is 300. The reaction temperature is 600 ° C and the weight hourly space velocity is 47 h -1 . Summary of the invention

[0006] The disadvantage of the prior art processes is that the propylene yield is low, while the formation of coke on the catalyst is high. This leads to short cycle times (i.e. the time between decoking operations), in which propylene can be produced in one reactor. The object of the present invention is to provide a process which enables the production of propylene in high yields while the formation of coke on the catalyst is maintained at a level that results in an acceptable cycle time.

[0007] The applicant has now found that the following process does not have such disadvantages. A process for producing propylene from a hydrocarbon mixture having an olefin content of 5-50 wt% and boiling more than 90 vol% between 35°C and 280°C, and / or from a hydrocarbon feed comprising paraffins, cycloparaffins and / or aromatics and optionally up to 10 wt% of olefins, wherein the process comprises the following steps:

[0008] (a) feeding a hydrocarbon mixture at a temperature of 450°C to 750°C, optionally mixed with a recycle stream, to a reactor in which the feed is heated at a hydrocarbon partial pressure below 3 bar and a temperature of 0.5 to 100 h -1 contacting with a low acid density cracking catalyst at a weight hourly space velocity of 1.5;

[0009] (b) separating propylene and optionally other low boiling point compounds from the effluent of step (a), wherein a first high boiling point fraction remains;

[0010] (c) feeding a first high boiling point fraction having a temperature of 400° C. to 750° C., all or part of which is optionally mixed with a circulating stream, to a reactor, wherein the first high boiling point fraction is heated to 100 % by pressure at a hydrocarbon partial pressure below 3 bar and at a temperature of 0.5 to 100 h. -1 contacting with a high acid density cracking catalyst at a weight hourly space velocity of 1000 ℃, wherein the temperature of the hydrocarbon mixture fed to the reactor, optionally mixed with the recycle stream, in step (a) is lower than the temperature of the first high boiling point fraction fed to the reactor, optionally mixed with the recycle stream, in step (c);

[0011] (d) separating propylene and optionally other low boiling point compounds from the effluent of step (c) leaving a second higher boiling point fraction; and

[0012] (e) recycling all or part of the second higher boiling point fraction to step (a) and / or step (c) as an optional recycle stream.

[0013] The applicant has now found that a hydrocarbon mixture containing paraffins or paraffins and olefins can be effectively converted into propylene and other lower olefins in two cracking steps. In the first step (a), olefins and cycloalkanes (if any) are mainly converted into propylene, other lower olefins and paraffins. This is achieved under relatively mild reaction conditions in the presence of a low acid density cracking catalyst. Under these conditions, the formation of coke is minimized. The first high boiling point fraction will have a higher paraffin content than the above-mentioned olefin feed. This makes it possible to crack the feed under more severe conditions by contacting with a high acid density cracking catalyst. Therefore, a method is provided that can convert olefins and paraffins in a hydrocarbon mixture, even C5 paraffins, into propylene with high yield. In addition, the applicant has found that less coke formation can be maintained. It is believed that this is the reason why the first high boiling point fraction is almost free of olefins or at least contains a small amount of olefins. Further advantages will be described below.

[0014] The feed used in step (a) is a hydrocarbon mixture. The mixture will include paraffins optionally mixed with aromatic compounds and / or cycloalkanes and olefins having an olefin content of 5-50wt% and boiling more than 90vol% between 35°C and 280°C, preferably boiling more than 90vol% between 35°C and 240°C. The hydrocarbon mixture will appropriately include paraffins, cycloalkanes and / or aromatics next to olefins. This mixture can be obtained from any source. Suitably, the hydrocarbon mixture includes a fraction separated from the effluent of a fluid catalytic cracking process, or a fraction separated from the effluent of a fluid catalytic cracking process, such as light catalytic cracking naphtha, medium catalytic cracking naphtha and heavy catalytic cracking naphtha. Other examples include delayed coking naphtha, pyrolysis naphtha and ebullient bed naphtha. This mixture may also include aromatic compounds, paraffins and / or cycloalkanes, and appropriately include aromatic compounds, paraffins and cycloalkanes. The hydrocarbon mixture may also comprise or be a fraction separated from the effluent of a steam cracking process.

[0015] As an alternative or addition to the above-mentioned olefin feed, the method according to the present invention can also convert more paraffin and / or cycloparaffin mixtures into propylene. This hydrocarbon feed includes paraffins, cycloparaffins and / or aromatic compounds and optionally up to 10wt% of olefins. Preferably, this additional or alternative feed boils more than 90vol% between 35°C and 280°C, preferably boils more than 90vol% between 35°C and 240°C. When used as an additional feed, the feed is preferably fed directly to the reactor of step (c) together with the first high boiling point fraction (the first high boiling point fraction is optionally mixed with the circulating stream). When used as an alternative feed, it is preferably used in step (a).

[0016] The olefin content of this more suitable paraffinic and / or cycloparaffinic mixture is less than 10wt%, more preferably less than 5wt%, even more preferably less than 1wt%. Examples of such mixtures are, for example, refinery naphtha fractions of straight run naphtha and light straight run naphtha, or fractions obtained during refinery hydrotreating such as hydrocrackers or hydrotreating processes, also referred to as hydrotreated naphtha and hydrocracked naphtha. Other examples are polymerized naphtha and reformed naphtha and natural gas liquids.

[0017] The conversion of olefins in step (a) is an endothermic reaction. The required energy can be added to the reactor in various ways. A preferred method is to add inert hydrocarbons, such as paraffins, cycloparaffins and / or aromatic compounds to the olefin mixture. The heat capacity per olefin mass of the feed will then increase. This is advantageously achieved by recycling part of the first high-boiling fraction obtained in step (b) to step (a). Therefore, the weight fraction of the hydrocarbon mixture with an olefin content of between 5 and 50 wt% and boiling more than 90 vol% between 35°C and 280°C, preferably between 35°C and 240°C, can be 25-75 wt% in the total feed to the reactor.

[0018] The applicant has found that the presence of aromatic compounds in the mixture fed to the reactor in step (a) and step (c) increases the cycle time, stabilizes the activity and helps to maximize the conversion rate. When the feed entering the reactor has a low olefin content, this effect is most significant as the above-mentioned more paraffinic and / or cycloparaffinic mixture is used as the starting feed. When starting from an olefin feed, it may be more preferred to have aromatic compounds in the feed of step (c). The presence of aromatic compounds does not substantially affect the selectivity. It is not desirable to be bound by the following theory that the presence of aromatic compounds is to reduce the coke formation on the cracking catalyst by competitive adsorption on the catalyst surface or by dilution of coke precursors. Secondly, the presence of substantially inert aromatic compounds can increase the conversion rate in the reactor by providing heat to endothermic cracking. The aromatic compound is preferably an aromatic compound that is substantially in the same range as the boiling point of the olefin mixture or slightly higher than its boiling point. Examples of suitable aromatic compounds are benzene, toluene, xylene, ethylbenzene and other aromatic compounds with 8 or more carbon atoms, preferably up to and including 11 carbon atoms.

[0019] Aromatic compounds may be present in the above-mentioned olefin or paraffin / cycloparaffin feed used in steps (a) and (c) of the present process, or may be present in the recycle stream. Preferably, aromatic compounds are purposefully added to the above-mentioned process. For the olefin feed entering the reactor in step (a), preferably at least 10 wt% of aromatic compounds are present, more preferably at least 20 wt%, even more preferably at least 30 wt%, further preferably at least 40 wt%, and even more preferably 50 wt% of aromatic compounds. The optimum value of the aromatic compound content can be determined by those skilled in the art, wherein the maximum conversion rate per pass can be a determining factor. The upper limit of the aromatic compound content can be 80 wt%, and more preferably, in the hydrocarbon mixture containing one or more optional recycle streams fed to the reactor in step (a), the content of aromatic compounds is between 10 and 80 wt%, more preferably between 20 and 70 wt%, even more preferably between 30 and 60 wt% and most preferably between 40 and 50 wt%.

[0020] For the mixture supplied to the reactor in step (c), it is preferred that it contains at least 5 wt%, more preferably at least 10 wt% of aromatic compounds, even more preferably at least 20 wt%, preferably up to 80 wt%, and even more preferably up to 40 wt%. When the olefin conversion in step (a) is such that the olefin content in the first high boiling point fraction obtained in step (b) is less than 10 wt%, more preferably less than 5 wt%, such an aromatic content is particularly preferred. When the aromatic content is outside the above range, such an olefin content is also preferred to carry out step (c) at the desired cycle time and selectivity.

[0021] As mentioned above, aromatic compounds are preferably added purposefully to the feed entering the reactor in step (a) and / or step (c). These aromatic compounds may come from, for example, other parts of an oil refinery or steam cracker. Preferably, in a separate process step, aromatic compounds are prepared by contacting part of the first high boiling point fraction and / or all or part of the second high boiling point fraction in step (f) with hydrogen in the presence of an aromatic conversion catalyst present in the reactor to obtain a fraction rich in aromatic compounds. By recycling all or part of the fraction rich in aromatic compounds to step (a) and / or step (c), the desired aromatic compound content can be obtained. Part of the fraction rich in aromatic compounds may also be recycled to step (f) itself. Preferably, part of the aromatic compounds are separated from these circulating streams to avoid the accumulation of substantially inert aromatic compounds. This is not disadvantageous in itself because these aromatic compounds (such as benzene, toluene and xylene) represent ideal compounds to be used in this way.

[0022] This aromatic conversion step (f) is known per se and is referred to as reforming. Step (f) may be carried out using known reforming processes such as those provided by UOP. The temperature may be between 400°C and 700°C, preferably between 400°C and 650°C, even more preferably between 400°C and 550°C, and at a weight hourly space velocity (WHSV) of 0.1 to 50 h -1 , preferably 0.5 to 25 h -1 , even more preferably 0.5 to 5 h -1 , step (f) is carried out at a hydrocarbon partial pressure below 10 bar and a hydrogen partial pressure below 10 bar.

[0023] The aromatic conversion catalyst can be any reforming catalyst or a heterogeneous catalyst comprising ZnO, a medium-pore zeolite and a binder. Suitably, the medium-pore zeolite is ZSM-5, and the binder is alumina. Suitably, the binder includes some P2O5. Preferred catalysts include 25-60wt% ZSM-5, 5-35wt% ZnO and 2.5-20wt% P2O5 and an alumina binder. Such a catalyst can preferably be prepared by adding a ZSM-5 zeolite (e.g., 50 parts of ZSM-5 crystals, SAR30, Zeolyst) to a certain amount of water. This aqueous mixture can be added to, for example, 35 parts of dry gel alumina (Catapal B, Sasol) and, for example, 10 parts of dry zinc nitrate (industrial grade, Alpha Aesar), and then kneaded. For example, 5 parts of P2O5 can be added to the kneaded material as diluted phosphoric acid. The mixture is extruded and dried, for example, at 120° C. for 1 hour and calcined at 600° C. for 1 hour.

[0024] The reactor in which step (f) may be performed may be a fixed bed reactor, a radial bed reactor, a moving bed reactor, a bubbling bed reactor or a fluidized bed reactor. The preferred reactor is a fixed bed reactor. In some embodiments of the present invention, the reactor is not a fixed bed reactor.

[0025] In step (b) and step (d), propylene and optional other low boiling point compounds are separated respectively from the effluent of step (a) and step (c), and the high boiling point fraction is left. The other low boiling point compounds can be, for example, ethane, ethylene, hydrogen, water, propane and butenes. This separation can include distillation and / or flash distillation. Since the selectivity of propylene to the total amount of propylene and propane is improved, the propane formed is reduced. This is advantageous because less difficult propylene and propane separation is required to obtain, for example, polymer-grade propylene. In this separation, ethylene can be separated from the low boiling point compounds. The C4 fraction comprising butanes and butenes can be recovered like this or recovered together as a part of the high boiling point compounds.

[0026] The reactor of step (a) can be a fixed bed reactor, such as a radial bed reactor, a moving bed reactor, a bubbling bed reactor or a fluidized bed reactor. The reactor of step (c) can be a fixed bed reactor, a radial bed reactor, a moving bed reactor, a bubbling bed reactor or a fluidized bed reactor. The reactor of step (a) and / or step (c) may not be a fixed bed reactor in particular. The above-mentioned reactor used in step (a) and / or step (c) may be provided with an inner pipe allowing superheated steam or other superheated medium to flow through. This steam will add energy to the endothermic reaction occurring in step (a) and / or step (c) through indirect heat exchange, so that the conversion of the reactants in the reactor has a higher conversion level.

[0027] The catalyst present in the reactor of step (a) can be any cracking catalyst with relatively low acid density. The low acid density has a relatively large distance between acid sites, which avoids that the reaction intermediate compound can form coke. The catalyst is active in the conversion of olefins, while paraffins are hardly reactive. Possible low acid density catalysts are amorphous catalysts, such as catalysts containing amorphous silica-alumina, silica-zirconia and / silica-borate as amorphous low acid density components. The preferred low acid density catalyst in step (a) is a heterogeneous catalyst containing a medium or large pore zeolite, the silica-alumina ratio of the medium or large pore zeolite is 2 to 1000, preferably 10 to 1000, more preferably 10 to 300, even more preferably 20 to 300, most preferably 20 to 100. For example, it is possible to start with a fresh catalyst with a relatively low silica-alumina ratio. In due course, the ratio can be increased to a higher ratio due to dealumination treatment. The resulting reduced activity can be compensated by operating at a higher temperature. Examples of suitable medium or large pore zeolites are ZSM-5, ZSM-11 and beta zeolite. Examples of suitable low acid density catalysts include up to 70 wt% ZSM-5, 1-20 wt% P2O5 and a binder. Examples of suitable binders are alumina, such as boehmite, optionally mixed with clay to enhance strength. The catalyst preferably includes 25-80 wt%, more preferably 25-70 wt%, even more preferably 35-50 wt% ZSM-5.

[0028] The catalyst present in the reactor of step (c) may be any cracking catalyst having a relatively high acid density. When the first high boiling point fraction comprises a high olefin content, in particular when the olefin content is higher than the olefin content in the hydrocarbon mixture fed to step (a), a catalyst with a lower acid density, such as the above-mentioned catalyst for step (a), may also be used. Therefore, the high acid density catalyst and the low acid density catalyst may preferably comprise a large or medium pore zeolite having a relatively low silica-alumina ratio. Preferably, the silica-alumina ratio of the high acid density catalyst is lower than the silica-alumina ratio of the low acid density catalyst. Examples of suitable medium or large pore zeolites are ZSM-5, ZSM-11 and β-zeolite. An example of a suitable high acid density catalyst in step (c) is a catalyst comprising up to 80 wt%, preferably up to 70 wt%, ZSM-5 having a silica-alumina ratio of 2 to 1000, preferably 10 to 1000, more preferably 10 to 300, even more preferably 25 to 100, 1-20 wt% of PO5 and a binder. An example of a suitable binder is alumina, such as boehmite, optionally mixed with clay to enhance strength. The catalyst preferably comprises 25-80 wt%, more preferably 25-70 wt%, even more preferably 35-50 wt% of ZSM-5.

[0029] The catalyst of step (a) and / or step (c) may be steamed before use. This is to initially deactivate part of the catalyst in order to limit the activity range. Within a more limited activity range, the conversion can be controlled by adjusting the temperature. If the activity range is too large, it is not possible to control the conversion by adjusting the temperature. The steaming may be carried out by contacting the catalyst in the reactor with a gas comprising 1-100 vol%, preferably 5-100 vol%, more preferably 5-10 vol%, even more preferably 70-95 vol% steam. The preferred pressure may be between about atmospheric pressure and a maximum pressure of 10 bar. The preferred temperature is between 300°C and 800°C, more preferably 400°C to 750°C, and most preferably 450°C to 600°C. The contact time may be from 1 hour to 5 days, wherein the contact time is preferably about 1 day.

[0030] The equilibrated catalyst of step (c) which has to be replaced due to its higher silica-alumina ratio can advantageously be used as the catalyst for step (a).

[0031] The catalyst comprising zeolite used in step (a) and step (c) can be prepared starting from a zeolite having the desired silica-alumina ratio. The zeolite is suitably slurried in distilled water and mixed with an alumina gel. The gel is prepared, for example, by using nitric acid and Catapal B from Sasol. The mixture is extruded to obtain particles comprising the zeolite and an alumina binder. The particles are calcined, for example, at 600°C in air for 1 hour. The calcined particles are subsequently impregnated with phosphoric acid and calcined again, for example at 600°C for 1 hour.

[0032] According to the process of the invention, the weight hourly space velocity is defined based on the total hydrocarbons fed to the reactor. Therefore, the optional recycle stream and / or the added aromatics-rich stream are also included. The same applies to the temperature of the feed entering the reactor. The temperature values ​​refer to the temperature of the total hydrocarbons fed to the reactor.

[0033] Preferably, the weight hourly space velocity in step (a) is 0.5 to 100 h -1 , preferably 0.5 to 50 h -1 , more preferably 1 to 25 h -1 , and most preferably 1 to 10 h -1 More preferably, the time is 1 to 5 h. -1 The weight hourly space velocity in step (c) is 0.5 to 100 h -1 , preferably 1 to 100 h -1 , more preferably 1 to 50 h -1 , and most preferably 2 to 30 h -1 or 2 to 20 hours -1 . Preferably, the weight hourly space velocity of step (a) is greater than the weight hourly space velocity of step (c). The hydrocarbon partial pressure is preferably below 1 bar, more preferably below 0.5 bar, even more preferably below 0.2 bar. The hydrocarbon partial pressure excluding aromatic compounds in step (a) and step (c) is preferably below 1 bar, more preferably below 0.5 bar, most preferably below 0.2 bar.

[0034] The temperature of the hydrocarbon mixture of step (a) optionally mixed with the recycle stream is from 450°C to 750°C, preferably from 450°C to 650°C, more preferably from 500°C to 650°C.

[0035] The temperature of the first high boiling fraction optionally mixed with the recycle stream in step (c) is 400 to 750°C, preferably 450 to 700°C, more preferably 450 to 650°C, most preferably 500 to 650°C.

[0036] The reactor in step (a), step (c) and / or step (f) is preferably a configuration of more than one reactor. For example, this configuration can be a group of reactors operating in parallel. These reactors can be the reactors listed above or a combination of these listed reactors.

[0037] As mentioned above, the applicant has found that the presence of aromatic compounds is beneficial for obtaining long cycle times, stabilizing activity and helping to maximize conversion of paraffin / cycloparaffin mixtures. For this reason, the present invention also relates to the following process.

[0038] Process for preparing propylene from a hydrocarbon starting feed comprising paraffins, cycloparaffins, aromatics and optionally up to 10 wt% of olefins, by adding aromatics to the hydrocarbon starting feed to produce an improved feed containing 10-70 wt%, preferably 20-50 wt%, more preferably 25-40 wt% of aromatics, wherein the aromatic content of the starting feed is below the lower limits of these ranges, and wherein the improved feed is catalytically cracked into propylene and other reaction products in the presence of an acidic cracking catalyst.

[0039] The aromatic compound added to the hydrocarbon starting feed can be any aromatic compound that is included in the aromatic compound boiling in the diesel boiling range. Suitably, the above aromatic compound is added to the hydrocarbon starting feed. The acidic cracking catalyst can be as described in the present application. The reactor and conditions can be the reactor and conditions as described in the present invention. Alternatively, the reactor can be a fluidized bed. The aromatic compound added to the hydrocarbon starting feed can be an aromatic compound separated from the reaction product of the present method, which is reused in the present method. It seems that the aromatic compound is not cracked into other products at a significant level, and the total amount of aromatic compounds in the improved feed is roughly consistent with the total amount of aromatic compounds in the reactor effluent.

[0040] This cyclic process is described in the following process according to the present invention: A process for producing propylene from a hydrocarbon feed comprising paraffins, cycloparaffins, aromatics and optionally up to 10 wt% of olefins, wherein the process comprises the following steps:

[0041] (aa) feeding said feed at a temperature of 450°C to 700°C, preferably 550°C to 700°C, mixed with a recycle stream, to a continuously operated reactor containing a high acid density cracking catalyst, wherein the mixture is heated in the reactor at a partial pressure of hydrocarbons other than aromatic compounds of less than 3 bar, preferably less than 1 bar, more preferably less than 0.5 bar and most preferably less than 0.2 bar and at a temperature of 1 to 30 h. -1 , preferably 2 to 30 h -1 contacting with a high acid density cracking catalyst at a weight hourly space velocity of 1.5;

[0042] (bb) separating propylene and optionally other low boiling point compounds from the effluent of step (aa), wherein a high boiling point fraction remains;

[0043] (cc) recycling the high boiling point fraction to the reactor of step (aa), wherein the total content of aromatic compounds in the mixture fed to the reactor in step (aa) is maintained between 5 and 50 wt%, preferably between 10 and 40 wt%, more preferably between 20 and 30 wt%, optionally by additionally feeding to the reactor further aromatic compounds containing a hydrocarbon mixture.

[0044] The high acid density catalyst in step (aa) may be as described earlier in step (c). Preferred conditions and catalysts for operating this process are the same as described above for steps (c) to (e). It is also preferred to add an aromatic conversion, which uses the high boiling point fraction with hydrogen in step (dd) in the presence of an aromatic conversion catalyst present in the reactor to obtain an aromatics-rich fraction, and wherein all or part of the aromatics-rich fraction is recycled to step (aa) as a further hydrocarbon mixture. The conditions and catalysts are as described above.

[0045] The feed for the process may boil more than 90 vol% between 35°C and 280°C, preferably more than 90 vol% between 35°C and 240°C. Suitably, the feed has an olefin content of less than 10 wt%, preferably less than 5 wt%, more preferably no olefins. An example of the mixture is a naphtha fraction obtained in a refinery hydroprocessing process, such as a hydrocracking process or a hydrotreating process.

[0046] The present invention also relates to a process configuration suitable for preparing propylene from olefins comprising a hydrocarbon mixture, the configuration comprising:

[0047] (i) one or more first reactors operated in parallel containing an amorphous heterogeneous cracking catalyst or a heterogeneous cracking catalyst containing a medium or large pore zeolite having a silica-alumina ratio of 1 to 1000;

[0048] (ii) a first distillation and / or flash separation unit fluidly connected to the outlet of one or more first reactors, said first reactors having at least an outlet for the propylene containing fraction and an outlet for high boilers;

[0049] (iii) means for recycling high boiling point compounds from the outlet of the distillation and / or flash separation unit to the inlet of one or more first reactors operated in parallel;

[0050] (iv) one or more second reactors operated in parallel containing a heterogeneous cracking catalyst comprising up to 80 wt% of ZSM-5 having a silica-alumina ratio of 2 to 1000 (preferably 25 to 50), 1-20 wt% of P2O5 and a binder, wherein the inlet of the second reactor is fluidly connected to the outlet for high boilers of the first distillation and / or flash separation unit;

[0051] (v) a second distillation and / or flash separation unit fluidly connected to the outlet of the one or more second reactors of (iv), said second distillation and / or flash separation unit having at least an outlet for the propylene containing fraction and an outlet for high boilers;

[0052] (vi) means for recycling high boiling point compounds from the outlet of the second distillation and / or flash separation unit to the inlet of one or more first reactors operated in parallel and to the inlet of one or more second reactors operated in parallel.

[0053] The first reactor and / or the second reactor may be a fixed bed reactor, a fluidized bed reactor, a bubbling bed reactor, an ebullating bed reactor, for example a radial bed reactor or a moving bed reactor or a combination of these reactors. Preferably, the first reactor and the second reactor are fixed bed reactors, and in one embodiment of the present invention, the first reactor and / or the second reactor are not fixed bed reactors. The first reactor and the second reactor may be provided with an inner pipe that allows superheated steam or other superheated medium to flow through. This steam will add energy to the endothermic reaction occurring in step (a) and / or step (c) through indirect heat exchange, so that the conversion of the reactants in the reactor has a higher conversion level.

[0054] Preferably, the process arrangement further comprises inlet means (vii) for connecting a further hydrocarbon feed stream to the inlet of one or more parallel second reactors. The flash separation unit may be suitably combined with a vapour-liquid separation step to recover any C5+ compounds that may be present in the gaseous effluent.

[0055] Preferably, the process configuration further comprises (viii) one or more aromatic conversion reactors operated in parallel, which are fluidically connected to the outlet for high boiling point compounds of the second distillation and / or flash separation unit, and are fluidically connected to a device for recycling part of the effluent of the aromatic conversion reactor to the inlet of one or more first reactors, the inlet of one or more second reactors and the inlet of the aromatic conversion reactor. Suitably, the aromatic conversion reactor has an inlet for hydrogen and a bed of a heterogeneous catalyst comprising ZnO, a medium pore zeolite and a binder as described above. The aromatic conversion reactor can be a fixed bed reactor, a fluidized bed reactor, a bubbling bed reactor, an ebullating bed reactor, a radial bed reactor or a moving bed reactor.

[0056] Step (c) of the one-step process or two-step process according to the present invention is an energy-intensive process, which can obtain a large amount of light olefins, such as ethylene, propylene and butene. The hydrocarbon feed can be heated to a reactor inlet temperature of 450°C or above by a feed / effluent heat exchanger and a subsequent natural gas gas burner. The temperature of the reactor effluent can then be further reduced by a combination of an air-cooled heat exchanger and a cold water heat exchanger with a target temperature of appropriately 25°C to 30°C. The low boiling point fraction can then be separated by a single-stage equilibrium flash, wherein the overhead steam as a low boiling point product is further separated in a product recovery unit. High boiling point and non-reactive liquid hydrocarbons can be advantageously recycled. This recycled material, which has a temperature of, for example, 25°C to 30°C, is combined with fresh hydrocarbon feed and reheated to a reactor inlet temperature of 450°C or above.

[0057] Another less energy intensive process is as follows to modify the basic process described above. The reactor effluent is heat exchanged with the feed in the feed / effluent heat exchange network and then fed to a distillation column (e.g., a debutanizer). The operating temperature of the reactor effluent leaving the feed / effluent heat exchange can be between 200°C and 300°C. The butenes and lighter components of the reactor effluent are taken from the top of the distillation column as low boiling point compounds. The pentane and heavier components of the reactor effluent are the bottom product of the distillation column, i.e., high boiling point compounds. The bottom product is treated as a recycle feed as described above in the basic configuration described above. The operating temperature of the recycle stream leaving the debutanizer bottom stream can be between 250°C and 350°C. The preferred recycle feed-fresh feed ratio varies between 2.0 and 4.0 (recycle feed mass flow rate / fresh feed mass flow rate). For example, 67 wt% of the reactor feed consists of recycle feed. The alternative process described above eliminates the need to inject up to 420°C energy equivalents into the reactor feed of the recycle stream, thereby saving significant energy. A negative consequence of this alternative process flow is that the pressure drop across the debutanizer column may be 3.0 psig. This is higher than the approximately 2.0 psig pressure drop produced by cooling the reactor effluent in an air-cooled followed by a water-cooled heat exchanger. This additional 1.0 psig pressure drop can be added to the reactor inlet operating pressure when this alternative process is used.

[0058] Since an increase in reactor inlet pressure can negatively impact propylene selectivity, it is preferred that the debutanizer be operated at a partial vacuum to eliminate the required increase in reactor inlet pressure. The overhead accumulator tank of the debutanizer will operate as a branch tank for the centrifugal compressor. The centrifugal compressor may preferably produce a total vacuum of about 10 to 15 psia. This will eliminate the need to increase the reactor inlet pressure while obtaining the benefits of the alternative process flowsheet. The reactor inlet pressure of the alternative process may, for example, be operated at a reactor inlet pressure of 18.5 psia, compared to 37.1 psia for the base process, resulting in a further increase in propylene selectivity.

[0059] The centrifugal compressor discharge can be passed through a cold water heat exchanger to reduce the operating temperature to about 30°C. The cooled hydrocarbon stream will be taken to the high-pressure separator for efficient removal of higher molecular weight compounds. The overhead vapor stream from the separator will be taken to the secondary booster centrifugal compressor to pressurize the product separation. In the product recovery section, the liquid hydrocarbon stream of the high-pressure separator can be distilled with lower energy requirements due to operation at a higher pressure. BRIEF DESCRIPTION OF THE DRAWINGS

[0060] Figure 1 A schematic diagram showing the conversion rate changing with time when there is no aromatic compound and when there is an aromatic compound in the examples of the present application is shown.

[0061] Figure 2 A schematic diagram showing the conversion rate of the first process and the circulation flow changing with time in the embodiment of the present application is shown. DETAILED DESCRIPTION

[0062] Example 1

[0063] At a weight hourly space velocity of 30 h -1 Fluid catalytic cracking (FCC) naphtha boiling between 20°C and 206°C and having the composition listed in Table 1 was fed into a pilot fixed bed reactor containing 1.5 g of fixed bed catalyst. The temperature in the reactor was 600°C.

[0064] ZSM-5 crystals with SAR30 (CBV 3024E, Zeolyst) and alumina (Sasol) were crushed into a 55 / 45 (wt / wt) mixture and extruded to form a shaped body. The extruded shaped body was dried at 120°C overnight and calcined at 600°C for 3 hours in flowing air. The calcined extrudate was impregnated with phosphoric acid to incipient wetness, then dried at 120°C overnight and calcined at 600°C for 3 hours in flowing air.

[0065] Table 1

[0066]

[0067] The composition of the reaction product is listed in Table 2.

[0068] Example 2

[0069] Example 1 was repeated except that toluene was used to replace 20 wt% of the feed. This resulted in a decrease in the total conversion (defined as: (mass of H2, C1-C4 hydrocarbons and delta-aromatics produced) / mass of feed * 100%) from 22 wt% to 19 wt%, and an increase in the conversion of the FCC naphtha itself from 22 wt% to 24 wt%. The addition of toluene demonstrates the beneficial effect of a recycle feed containing aromatics on the cracking reactor. In Example 1, the conversion of the feed was 22%. In Example 2, the conversion was: 19% absolute conversion on the total feed basis, or 24% conversion on the FCC naphtha portion of the feed. Product selectivity was not affected by the addition of aromatics (as shown in the results reported in Table 2).

[0070] Table 2

[0071]

[0072] Example 3

[0073] Except that the feed is hexane and the weight hourly space velocity is 60 h -1 In addition, Example 1 was repeated for about 3000 minutes (50 hours). The real-time feed conversion rate is Figure 1 The black dots (without aromatic compounds) are shown.

[0074] Example 4

[0075] Example 3 was repeated except that toluene was used to replace 20 wt% of the feed. The real-time feed conversion was given by Figure 1 The hollow dots (containing aromatic compounds) are indicated. Figure 1 It is shown that when toluene is added, the conversion of the crackable portion of the feed (hexane) is initially lower. Example 4 with toluene shows significantly improved stability, and catalyst deactivation due to catalyst coking over time is significantly reduced. The addition of toluene does not negatively affect the product selectivity of propylene (about 35% for both experiments) and butenes (about 20-24% for both experiments).

[0076] Example 5

[0077] At a weight hourly space velocity of 10 h -1Under the conditions, a non-olefin feed having the composition listed in Table 3 and boiling between 20°C and 220°C was fed to a pilot fixed bed reactor containing 3 grams of the fixed bed catalyst described in Example 1 for about 1200 minutes. The temperature in the reactor was 600°C. The real-time conversion was as follows: Figure 2 The black dots (first process) are shown.

[0078] Table 3

[0079]

[0080] Example 6

[0081] Example 5 was repeated, wherein part of the reactor liquid effluent was recycled to the reactor, thereby replacing part of the feed so that the feed now consisted of 80 wt % recycled feed and 20 wt % fresh feed. The recycled feed contained 2 wt % olefins. The real-time conversion was as follows: Figure 2 The hollow circles (circulating flow) are shown.

[0082] Figure 2 It is shown that the overall conversion (of the total mixed feed entering the reactor) is higher when part of the liquid effluent is recycled to the reactor. When comparing Examples 5 and 6, the selectivity to the desired C3 and C4 olefins is not affected in any significant way.

Claims

1. A method for preparing propylene from a hydrocarbon mixture having an olefin content of 5-50 wt% and boiling more than 90 vol% between 35° C. and 280° C., wherein the method comprises the following steps: (a) feeding a hydrocarbon mixture having a temperature of 450°C to 750°C mixed with a recycle stream to a reactor in which the feed is heated at a hydrocarbon partial pressure below 3 bar and at a temperature of 0.5 to 100 h. -1 contacting with a low acid density cracking catalyst at a weight hourly space velocity of 1.5; wherein the low acid density cracking catalyst in step (a) is a heterogeneous catalyst comprising a medium or large pore zeolite having a silica-alumina ratio of 2 to 1000; (b) separating propylene and other low boiling point compounds from the effluent of step (a) thereby retaining a first high boiling point fraction; (c) feeding the first high boiling point fraction having a temperature of 500° C. to 750° C., which is mixed in whole or in part with the circulating stream, to a reactor, wherein the first high boiling point fraction is heated to 100 % by pressure at a hydrocarbon partial pressure of less than 3 bar and at a temperature of 0.5 to 100 h. -1 contacting with a high acid density cracking catalyst at a weight hourly space velocity of 1000 rpm, wherein the temperature of the hydrocarbon mixture mixed with the recycle stream fed to the reactor in step (a) is lower than the temperature of the first high boiling point fraction mixed with the recycle stream fed to the reactor in step (c); wherein the high acid density cracking catalyst in step (c) comprises a medium-pore or large-pore zeolite, wherein the silica-alumina ratio of the high acid density cracking catalyst is lower than the silica-alumina ratio of the low acid density cracking catalyst; (d) separating propylene and other low boiling point compounds from the effluent of step (c) thereby leaving a second highest boiling point fraction; and (e) recycling all or part of the second highest boiling point fraction to step (a) and step (c) as a recycle stream.

2. The method of claim 1, wherein the weight hourly space velocity in step (a) is greater than the weight hourly space velocity in step (c).

3. The process according to claim 1, wherein part of the first high boiling point fraction and / or all or part of the second high boiling point fraction is contacted with hydrogen in step (f) in the presence of an aromatic conversion catalyst present in a reactor to obtain a fraction rich in aromatic compounds, and wherein all or part of the fraction rich in aromatic compounds is recycled to step (c) as a recycle stream.

4. The method according to claim 3, wherein the contacting in step (f) is carried out at a temperature of 400°C to 700°C for 0.1 to 50 h. -1 The reaction takes place at a weight hourly space velocity of 1000 rpm, a hydrocarbon partial pressure of less than 10 bar and a hydrogen partial pressure of less than 10 bar. 5 . The method according to claim 3 , wherein the aromatic conversion catalyst is a heterogeneous catalyst comprising ZnO, a medium-pore zeolite, and a binder.

6. The process of claim 1, wherein the heterogeneous catalyst in step (a) comprises up to 70 wt% ZSM-5, 1-20 wt% P2O5 and a binder.

7. The process according to claim 6, wherein the heterogeneous catalyst in step (a) comprises 25-70 wt% of ZSM-5.

8. The method according to claim 1, wherein the weight hourly space velocity in step (a) is 0.5 to 50 h -1 . 9 . The method according to claim 1 , wherein the medium-pore or large-pore zeolite of the heterogeneous catalyst in step (c) is ZSM-5, ZSM-11 or β-zeolite.

10. The process according to claim 9, wherein the high acid density cracking catalyst in step (c) comprises up to 80 wt% of ZSM-5 having a silica-alumina ratio of 2 to 1000, 1-20 wt% of P2O5 and a binder.

11. The process of claim 1, wherein the low acid density heterogeneous catalyst comprises 35-50 wt% ZSM-5.

12. The method according to claim 1, wherein the weight hourly space velocity in step (c) is 1 to 50 h -1 .

13. The process of claim 1, wherein the partial pressure of hydrocarbons other than aromatic compounds in step (a) and step (c) is less than 1 bar.

14. The process according to claim 1, wherein the hydrocarbon mixture comprises fractions separated from the effluent of a fluid catalytic cracking process and / or separated from a steam cracking process.

15. The process according to claim 1, wherein the content of aromatic compounds in the hydrocarbon mixture including the recycle stream fed to the reactor in step (a) is 10-80 wt%.

16. The process according to claim 1, wherein part of the second high boiling point fraction obtained in step (d) is recycled to step (c) as a recycle stream.

17. The process according to claim 16, wherein in step (d), propylene and other low boiling compounds are separated from the effluent of step (c) in a debutanizer operated under partial vacuum, and wherein the second higher boiling fraction is obtained as the bottom product of the debutanizer.

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