Process for producing propylene

Abstract

The present invention relates to a process for producing propylene from a hydrocarbon mixture having an olefin content of 5-50 wt% and boiling over 90 vol% between 35 DEG C and 280 DEG C, or from a hydrocarbon feed comprising paraffins, naphthenic compounds and / or aromatic compounds and optionally up to 10 wt% olefins, propylene is first separated in the reactor by contacting the feed with a low acid density cracking catalyst, and subsequently contacting the residue with a high acid density cracking catalyst in the reactor at a higher temperature, and the residue is recycled to the first cracking reactor and the second cracking reactor. Aromatics may be added in the first and second cracking steps to improve 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 above 90 vol% between 35 °C and 280 °C, and / or from a hydrocarbon feedstock comprising paraffins, naphthenic compounds, aromatic compounds and optionally up to 10 wt% olefins by contacting the feedstock with a cracking catalyst in a reactor. Background Art

[0002] More than 50% of the propylene produced is by steam cracking processes. Typical feedstocks are straight-run naphthas obtained when refining crude oil sources, which generally include unsaturated compounds such as paraffins and naphthenic compounds optionally mixed with aromatic compounds.

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

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

[0005] GB2345294 describes a method in which olefin-containing C 4The residual oil feedstock contacts a cracking catalyst in a fixed bed reactor. The catalyst consists of ZSM-5 containing silver instead of protons, where 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 disadvantages of the prior art methods are the low propylene yield and the high formation of coke on the catalyst. This results in a short cycle time (i.e., the time between decoking operations), where propylene can be produced in one reactor. The object of the present invention is to provide a method that is capable of producing propylene in high yield while the formation of coke on the catalyst is maintained at a level that allows for an acceptable cycle time.

[0007] The applicant has now found that the following method does not have such disadvantages. A method for producing propylene from a hydrocarbon mixture having an olefin content of 5-50 wt% and boiling above 90 vol% between 35 °C and 280 °C, and / or from a hydrocarbon feedstock comprising paraffins, naphthenic compounds, and / or aromatic compounds and optionally up to 10 wt% olefins, wherein the method comprises the following steps: a) feeding a hydrocarbon mixture at a temperature of 450 °C to 750 °C, optionally mixed with a recycle stream, to a reactor, where the feed contacts a low acid density cracking catalyst at a hydrocarbon partial pressure below 3 bar and a weight hourly space velocity of 0.5 to 100 h -1 ; b) separating propylene and optionally other low-boiling compounds from the effluent of step a), leaving a first high-boiling fraction; c) feeding the first high-boiling fraction at a temperature of 400 °C to 750 °C, wholly or partly optionally mixed with a recycle stream, to a reactor, where the first high-boiling fraction contacts a high acid density cracking catalyst at a hydrocarbon partial pressure below 3 bar and a weight hourly space velocity of 0.5 to 100 h -1 , where the temperature of the hydrocarbon mixture optionally mixed with a recycle stream fed to the reactor in step a) is lower than the temperature of the first high-boiling fraction optionally mixed with a recycle stream fed to the reactor in step c); d) separating propylene and optionally other low-boiling compounds from the effluent of step c), leaving a second high-boiling fraction; and e) recycling wholly or partly the second high-boiling fraction to step a) and / or step c) as an optional recycle stream.

[0008] The Applicant has now found that hydrocarbon mixtures comprising paraffins or paraffins and olefins can be effectively converted into propylene and other lower olefins in two cracking steps. In the first step (a), the 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, coke formation is minimized. The first high-boiling fraction will have a higher paraffin content than the above-mentioned olefin feed. This allows the feed to be cracked under more severe conditions by contacting it with a high acid density cracking catalyst. Accordingly, a method is provided for converting olefins and paraffins, even C 5 paraffins in a hydrocarbon mixture into propylene in 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 fraction contains little or at least a small amount of olefins. Further advantages will be described below.

[0009] The feed used in step (a) is a hydrocarbon mixture. The mixture will comprise paraffins optionally mixed with aromatic compounds and / or cycloaliphatic compounds and olefins having an olefin content of 5-50 wt% and boiling above 90 vol% between 35°C and 280°C, preferably boiling above 90 vol% between 35°C and 240°C. The hydrocarbon mixture will suitably comprise paraffins, cycloaliphatic compounds and / or aromatic compounds in addition to the olefins. Such a mixture can be obtained from any source. Suitably, the hydrocarbon mixture comprises a fraction separated from the effluent of a fluid catalytic cracking process or is a fraction separated from the effluent of a fluid catalytic cracking process, such as light catalytic cracked naphtha, medium catalytic cracked naphtha and heavy catalytic cracked naphtha. Other examples include delayed coker naphtha, pyrolysis naphtha and ebullated bed naphtha. Such a mixture can also comprise aromatic compounds, paraffins and / or cycloaliphatic compounds and suitably comprises aromatic compounds, paraffins and cycloaliphatic compounds. The hydrocarbon mixture can also comprise or be a fraction separated from the effluent of a steam cracking process.

[0010] As an alternative or addition to the above-mentioned olefin feed, the process according to the invention can also convert more paraffin and / or cycloalkane mixtures into propylene. Such a hydrocarbon feed comprises paraffins, cycloaliphatic compounds and / or aromatic compounds and optionally up to 10 wt% olefins. Preferably, such an additional or alternative feed boils above 90 vol% between 35°C and 280°C, preferably boils above 90 vol% between 35°C and 240°C. When used as an additional feed, the feed is preferably fed directly to the reactor in step (c) together with the first high-boiling fraction (the first high-boiling fraction is optionally mixed with a recycle stream). When used as an alternative feed, it is preferably used in step (a).

[0011] The olefin content of such a more suitable alkane and / or cycloalkane mixture is less than 10 wt%, more preferably less than 5 wt%, and even more preferably less than 1 wt%. Examples of such mixtures are, for example, refinery naphtha fractions of straight-run naphtha and light straight-run naphtha, or fractions obtained during the refinery hydrotreating process of a hydrocracker or hydrotreating process, also known as hydrotreated naphtha and hydrocracked naphtha. Other examples are polymerized naphtha and reformed naphtha and natural gas liquids.

[0012] 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 alkanes, cycloalkane compounds, and / or aromatic compounds, to the olefin mixture. The heat capacity per mass of olefin of the feed will then increase. This is advantageously achieved by recycling a portion of the first high-boiling fraction obtained in step (b) to step (a). Thus, the weight fraction of a hydrocarbon mixture having an olefin content between 5 and 50 wt% and boiling above 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.

[0013] The applicant has found that the presence of aromatic compounds in the mixtures fed to the reactor in steps (a) and (c) increases the cycle time, stabilizes the activity, and contributes to maximizing the conversion rate. This effect is most significant when starting with the more alkane and / or cycloalkane mixtures as described above, when the feed to the reactor has a low olefin content. When starting with an olefin feed, it may be more preferred to have aromatic compounds in the feed to step (c). The presence of aromatic compounds essentially does not affect the selectivity. Without wishing to be bound by the following theory, it is believed that the presence of aromatic compounds reduces coke formation on the cracking catalyst either by competitive adsorption on the catalyst surface or by dilution of coke precursors. Secondly, the presence of essentially inert aromatic compounds can increase the conversion rate in the reactor by providing heat to the endothermic cracking. The aromatic compounds are preferably aromatic compounds having a boiling point substantially in the same range as or slightly higher than that of the olefin mixture. Examples of suitable aromatic compounds are benzene, toluene, xylene, ethylbenzene, and other aromatic compounds having 8 or more carbon atoms, preferably up to and including 11 carbon atoms.

[0014] Aromatic compounds may be present in the above-mentioned olefin or alkane / naphthene feedstocks used in process steps (a) and (c), or may be present in the recycle stream. Preferably, aromatic compounds are added to the above process purposefully. For the olefin feedstock entering the reactor in step (a), it is preferred that there is at least 10 wt% of aromatic compounds, 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, where the maximum conversion in a single 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 fed to the reactor in step (a) containing an optional one or more recycle streams, 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%.

[0015] For the mixture fed 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 at most 80 wt%, and even more preferably at most 40 wt%. Such an aromatic compound content is particularly preferred when the olefin conversion in step (a) results in an olefin content in the first high-boiling fraction obtained in step (b) of less than 10 wt%, more preferably less than 5 wt%. When the aromatic compound content is outside the above ranges, such an olefin content is also preferred to carry out step (c) with the desired cycle time and selectivity.

[0016] As described above, it is preferred to add aromatic compounds purposefully to the feedstock entering the reactor in step (a) and / or step (c). These aromatic compounds can come from other parts of, for example, an oil refinery or a steam cracker. Preferably, in a separate process step, part of the first high-boiling fraction and / or all or part of the second high-boiling fraction in step (f) is contacted 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 can also be recycled to step (f) itself. Preferably, part of the aromatic compounds is separated from these recycle streams to avoid the accumulation of substantially inert aromatic compounds. This is not disadvantageous per se, as these aromatic compounds (such as benzene, toluene, and xylene) represent ideal compounds to be used in this way.

[0017] This aromatic conversion step (f) is known per se and is called reforming. Step (f) can be carried out using known reforming processes such as those provided by UOP. Step (f) can be carried out at a temperature between 400 °C and 700 °C, preferably between 400 °C and 650 °C, even more preferably between 400 °C and 550 °C, with 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 , with a hydrocarbon partial pressure below 10 bar and a hydrogen partial pressure below 10 bar for step (f).

[0018] The aromatic conversion catalyst can be any reforming catalyst or a heterogeneous catalyst comprising ZnO, mesoporous zeolite and binder. Suitably, the mesoporous zeolite is ZSM-5 and the binder is alumina. Suitably, the binder comprises some P 2 O 5 . Preferred catalysts comprise 25 - 60 wt% of ZSM-5, 5 - 35 wt% of ZnO and 2.5 - 20 wt% of P 2 O 5 and alumina binder. Such a catalyst can preferably be prepared by adding 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 basis gel alumina (Catapal B, Sasol) and, for example, 10 parts of dry basis zinc nitrate (technical grade, Alpha Aesar), and then kneaded. For example, 5 parts of P 2 O 5 can be added as diluted phosphoric acid to the kneaded mass. The mixture is extruded and, for example, dried at 120 °C for 1 hour and calcined at 600 °C for 1 hour.

[0019] The reactor in which step (f) can be carried out can be a fixed bed reactor, a radial bed reactor, a moving bed reactor, a bubble 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.

[0020] In steps (b) and (d), propylene and optionally other low-boiling compounds are separately separated from the effluents of steps (a) and (c), leaving a high-boiling fraction. The other low-boiling compounds can be, for example, ethane, ethylene, hydrogen, water, propane, and butene. This separation can include distillation and / or flash distillation. Since the selectivity for propylene over the total amount of propylene and propane is increased, less propane is formed. This is advantageous because a less difficult separation of propylene and propane is required to obtain, for example, polymer-grade propylene. In this separation, ethylene can be separated from the low-boiling compounds. The C 4 fraction including butane and butene can be recovered in this way or together as part of the high-boiling compounds.

[0021] The reactor of step (a) can be a fixed-bed reactor, such as a radial-bed reactor, a moving-bed reactor, a bubble-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 bubble-bed reactor, or a fluidized-bed reactor. The reactor of step (a) and / or step (c) may in particular not be a fixed-bed reactor. The above reactors for step (a) and / or step (c) can be provided with an internal 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) by indirect heat exchange, such that the conversion of the reactants in the reactor has a relatively high conversion level.

[0022] The catalyst present in the reactor of step (a) can be any cracking catalyst having a relatively low acid density. The low acid density has a relatively large distance between the acid sites, which avoids the reaction intermediate compounds from forming coke. The catalyst is active in the conversion of olefins, while paraffins hardly react. Possible low-acid-density catalysts are amorphous catalysts, such as catalysts containing amorphous silica-alumina, silica-zirconia, and / or silica-borate as amorphous low-acid-density components. The preferred low-acid-density catalyst in step (a) is a heterogeneous catalyst containing mesoporous or macroporous zeolite with a silica-alumina ratio of 2 to 1000, preferably 10 to 1000, more preferably 10 to 300, even more preferably 20 to 300, and most preferably 20 to 100. For example, it can start with a fresh catalyst having a relatively low silica-alumina ratio. In due course, this ratio can be increased to a higher ratio due to dealumination treatment. The resulting decrease in activity can be compensated by operating at a higher temperature. Examples of suitable mesoporous or macroporous zeolites are ZSM-5, ZSM-11, and beta-zeolite. Examples of suitable low-acid-density catalysts include up to 70 wt% of ZSM-5, 1-20 wt% of P 2 O 5and a binder. Examples of suitable binders are 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.

[0023] The catalyst present in the reactor of step (c) can be any cracking catalyst having a relatively high acid density. When the first high-boiling fraction contains a high olefin content, especially when the olefin content is higher than that in the hydrocarbon mixture fed to step (a), a catalyst with a lower acid density can also be used, such as the above-mentioned catalyst used in step (a). Thus, the high acid density catalyst and the low acid density catalyst preferably comprise macroporous or mesoporous zeolites with a relatively low silica-alumina ratio. Preferably, the silica-alumina ratio of the high acid density catalyst is lower than that of the low acid density catalyst. Examples of suitable mesoporous or macroporous zeolites are ZSM-5, ZSM-11, and beta-zeolite. An example of a suitable high acid density catalyst in step (c) is one containing up to 80 wt%, preferably up to 70 wt% of ZSM-5, 1 - 20 wt% of P 2 O 5 and a binder, the silica-alumina ratio of the ZSM-5 being from 2 to 1000, preferably 10 to 1000, more preferably 10 to 300, even more preferably 25 to 100. Examples of suitable binders are 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.

[0024] The catalyst of step (a) and / or step (c) can be steamed before use. This is to partially deactivate the catalyst initially to limit the activity range. Within a more restricted activity range, the conversion can be controlled by adjusting the temperature. If the activity range is too large, it is impossible to control the conversion by adjusting the temperature. Steaming can be carried out by contacting the catalyst in the reactor with a gas containing 1 - 100 vol%, preferably 5 - 100 vol%, more preferably 5 - 10 vol%, even more preferably 70 - 95 vol% steam. The preferred pressure can 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 can range from 1 hour to 5 days, with the contact time preferably about 1 day.

[0025] The equilibrium catalyst of step (c) that has to be replaced due to its higher silica-alumina ratio can be advantageously used as the catalyst for step (a).

[0026] The zeolite-containing catalyst for 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 containing zeolite and an alumina binder. The particles are calcined, for example, in air at 600 °C for 1 hour. Subsequently, the calcined particles are impregnated with phosphoric acid and calcined again, for example, at 600 °C for 1 hour.

[0027] In the process according to the invention, the weight hourly space velocity is defined based on the total hydrocarbons fed to the reactor. Thus, it also includes an optional recycle stream and / or a stream rich in aromatics added. This also applies to the temperature of the feed entering the reactor. The temperature value refers to the temperature of the total hydrocarbons fed to the reactor.

[0028] Preferably, the weight hourly space velocity of 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 。Even more preferably, it is 1 to 5 h -1 。The weight hourly space velocity of 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 h -1 。Preferably, the weight hourly space velocity of step (a) is greater than that of step (c). The hydrocarbon partial pressure is preferably below 1 bar, more preferably below 0.5 bar, and even more preferably below 0.2 bar. The hydrocarbon partial pressure of hydrocarbons other than aromatics in step (a) and step (c) is preferably below 1 bar, more preferably below 0.5 bar, and most preferably below 0.2 bar.

[0029] The temperature of the hydrocarbon mixture optionally mixed with the recycle stream in step (a) is 450 °C to 750 °C, preferably 450 °C to 650 °C, and more preferably 500 °C to 650 °C.

[0030] The temperature of the first high-boiling fraction optionally mixed with the recycle stream in step (c) is 400 °C to 750 °C, preferably 450 °C to 700 °C, more preferably 450 °C to 650 °C, and most preferably 500 °C to 650 °C.

[0031] The reactor in step (a), step (c) and / or step (f) is preferably configured as more than one reactor. For example, such a configuration can be a set of reactors operating in parallel. These reactors can be the reactors listed above or a combination of these listed reactors.

[0032] As described above, the applicant has found that the presence of aromatic compounds is beneficial for obtaining long cycle times, stable activity and helps to maximize the conversion of the alkane / cycloalkane mixture. For this reason, the present invention also relates to the following method.

[0033] A method for preparing propylene from a hydrocarbon starting feed comprising alkanes, cycloalkane compounds, aromatic compounds and optionally up to 10 wt% of olefins, by adding an aromatic compound to the hydrocarbon starting feed to produce an improved feed containing 10 - 70 wt%, preferably 20 - 50 wt%, more preferably 25 - 40 wt% of aromatic compounds, wherein the aromatic compound content in the starting feed is below the lower limit of these ranges, and wherein the improved feed is catalytically cracked to propylene and other reaction products in the presence of an acidic cracking catalyst.

[0034] The aromatic compound added to the hydrocarbon starting feed can be any aromatic compound containing aromatic compounds 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 products of the present method and reused in the present method. It appears 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 substantially the same as the total amount of aromatic compounds in the reactor effluent.

[0035] Such a cyclic process is described in the following method according to the present invention. A method for preparing propylene from a hydrocarbon feed comprising alkanes, cycloalkane compounds, aromatic compounds and optionally up to 10 wt% of olefins, wherein the method comprises the following steps: (aa) Feeding the feed, which is at a temperature of 450 °C to 700 °C, preferably 550 °C to 700 °C and mixed with a recycle stream, to a continuously operating reactor containing a high acid density cracking catalyst, in the reactor, the mixture is at a hydrocarbon partial pressure 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 except for aromatic compounds and a weight hourly space velocity of 1 to 30 h -1 、preferably 2 to 30 h -1 and contacting with the high acid density cracking catalyst; (bb) separating propylene and optionally other low-boiling compounds from the effluent of step (aa), leaving a high-boiling fraction; (cc) recycling the high-boiling fraction to the reactor of step (aa), wherein optionally by additionally feeding aromatic compounds further comprising a hydrocarbon mixture to the reactor, 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%.

[0036] The high-acid density catalyst in step (aa) can be as described in the earlier step (c). The preferred conditions and catalysts for operating this method are the same as those described in the above steps (c) to (e). Also preferably, an aromatic conversion is added, which uses the high-boiling fraction and hydrogen in step (dd) in the presence of an aromatic conversion catalyst present in the 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 (aa) as a further hydrocarbon mixture. The conditions and catalysts are as described above.

[0037] The feed for this method can boil above 90 vol% between 35 °C and 280 °C, preferably above 90 vol% between 35 °C and 240 °C. Appropriately, the olefin content of the feed is less than 10 wt%, preferably less than 5 wt%, more preferably free of olefins. Examples of the mixture are naphtha fractions obtained in refinery hydrotreating processes (such as hydrocracking processes or hydrotreating processes).

[0038] The present invention also relates to a process configuration suitable for preparing propylene from olefins containing a hydrocarbon mixture, the configuration comprising: (i) one or more first reactors operating in parallel, which contain an amorphous heterogeneous cracking catalyst or a heterogeneous cracking catalyst containing a mesoporous or macroporous zeolite with a silica-alumina ratio of 1 to 1000; (ii) a first distillation and / or flash separation unit fluidly connected to the outlet of one or more first reactors, the first reactor having at least an outlet for propylene containing a fraction and an outlet for high-boiling compounds; (iii) means for recycling the high-boiling compounds from the outlet of the distillation and / or flash separation unit to the inlet of one or more first reactors operating in parallel; (iv) one or more second reactors operating in parallel containing a heterogeneous cracking catalyst, the heterogeneous cracking catalyst comprising up to 80 wt% of ZSM-5 with a silica-alumina ratio of 2 to 1000 (preferably 25 to 50), 1-20 wt% of P 2 O 5and an adhesive, wherein the inlet of the second reactor is fluidly connected to the outlet for the high-boiling compounds of the first distillation and / or flash separation unit; (v) a second distillation and / or flash separation unit fluidly connected to the outlet of one or more second reactors of (iv), said second distillation and / or flash separation unit having at least an outlet for propylene containing fractions and an outlet for high-boiling compounds; (vi) means for recycling the high-boiling compounds from the outlet of the second distillation and / or flash separation unit to the inlet of one or more first reactors operating in parallel and to the inlet of one or more second reactors operating in parallel.

[0039] The first reactor and / or the second reactor can be a fixed-bed reactor, a fluidized-bed reactor, a bubble-bed reactor, a boiling-bed reactor, such as 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 invention, the first reactor and / or the second reactor is not a fixed-bed reactor. The first reactor and the second reactor can be provided with internal pipes allowing superheated steam or other superheated media to flow through. This steam will add energy to the endothermic reactions occurring in step (a) and / or step (c) by indirect heat exchange, so that a higher conversion level of the reactants in the reactor is achieved.

[0040] Preferably, the process configuration further includes an inlet means (vii) for fluidly connecting a further hydrocarbon feed to the inlet of one or more second reactors operating in parallel. The flash separation unit can be suitably combined with a vapor-liquid separation step to recover any C 5 compounds above.

[0041] Preferably, the process configuration further includes (viii) one or more aromatics conversion reactors operating in parallel, fluidly connected to the outlet for the high-boiling compounds of the second distillation and / or flash separation unit, and fluidly connected to means for recycling a portion of the effluent of the aromatics conversion reactor to the inlet of one or more first reactors, the inlet of one or more second reactors, and the inlet of the aromatics conversion reactor. Suitably, the aromatics conversion reactor has an inlet for hydrogen and has a bed of a heterogeneous catalyst comprising ZnO, mesoporous zeolite and an adhesive as described above. The aromatics conversion reactor can be a fixed-bed reactor, a fluidized-bed reactor, a bubble-bed reactor, a boiling-bed reactor, a radial-bed reactor or a moving-bed reactor.

[0042] Step (c) of the one-step or two-step process according to the present invention is an energy-intensive process that 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 higher through a feed / effluent heat exchanger and then a natural gas fired furnace. Then, the temperature of the reactor effluent can be further reduced through a combination of an air-cooled heat exchanger and a cold water heat exchanger with a target temperature appropriately between 25 °C and 30 °C. Then, the low-boiling fraction can be separated by single-stage equilibrium flash, where the overhead vapor as the low-boiling product is further separated in the product recovery unit. The high-boiling and unreacted liquid hydrocarbons can be advantageously recycled. This recycle stream, with a temperature of, for example, 25 °C to 30 °C, is combined with the fresh hydrocarbon feed and reheated to a reactor inlet temperature of 450 °C or higher.

[0043] Another less energy-intensive process modifies the above basic process as follows. The reactor effluent is heat-exchanged with the feed in the feed / effluent heat exchange network, and then the reactor effluent is 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 butene and lighter components of the reactor effluent are taken out from the top of the distillation column as low-boiling compounds. The pentane and heavier components of the reactor effluent are the bottom product of the distillation column, i.e., high-boiling compounds. The bottom product is treated as the recycle stream as described above in the above basic configuration. The operating temperature of the recycle stream leaving the bottom of the debutanizer can be between 250 °C and 350 °C. The preferred recycle feed-fresh feed ratio varies between 2.0 and 4.0 (recycle stream mass flow rate / fresh feed mass flow rate). For example, 67 wt% of the reactor feed consists of the recycle stream. The above alternative process eliminates the need to inject up to 420 °C of energy equivalent into the reactor feed of this recycle stream, thus significantly saving energy. The negative consequence of this alternative process flow is that the pressure drop across the debutanizer can be 3.0 psig. This is higher than the approximately 2.0 psig pressure drop generated by air cooling followed by cooling of the reactor effluent in a water-cooled heat exchanger. When using this alternative process, this additional 1.0 psig pressure drop can be added to the reactor inlet operating pressure.

[0044] Since an increase in the inlet pressure of the reactor can have a negative impact on propylene selectivity, it is preferred that the debutanizer operate under partial vacuum in order to eliminate the need for an increase in the reactor inlet pressure. The overhead accumulator of the debutanizer will operate as a knockout drum for the centrifugal compressor. The centrifugal compressor preferably generates 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 another process flow. The reactor inlet pressure of the other process can operate, for example, at a reactor inlet pressure of 18.5 psia, while for the base process it is 37.1 psia, further enhancing propylene selectivity.

[0045] 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 a high-pressure separator to efficiently remove higher molecular weight compounds. The overhead vapor stream from this separator will be taken to a secondary boost centrifugal compressor to pressurize the product separation. In the product recovery section, due to operating at a higher pressure, the liquid hydrocarbon stream from the high-pressure separator can be distilled with lower energy requirements. BRIEF DESCRIPTION OF THE DRAWINGS

[0046] Figure 1 Schematic diagrams showing the conversion rates over time when the embodiments of the present application contain no aromatic compounds and contain aromatic compounds are shown.

[0047] Figure 2 Schematic diagrams showing the conversion rates over time for the first process and the recycle stream in the embodiments of the present application are shown. DETAILED DESCRIPTION

[0048] Example 1 At a weight hourly space velocity of 30 h -1 −1, fluid catalytic cracking (FCC) naphtha boiling between 20°C and 206°C and having the components listed in Table 1 is fed into a pilot fixed bed reactor containing 1.5 grams of a fixed bed catalyst. The temperature in the reactor is 600°C.

[0049] ZSM-5 crystals with SAR30 (CBV 3024E, Zeolyst) and alumina (Sasol) are crushed into a 55 / 45 (wt / wt) mixture and extruded to form shaped bodies. The extruded shaped bodies are dried overnight at 120°C and calcined at 600°C in flowing air for 3 hours. The calcined extrudates are impregnated with phosphoric acid to incipient wetness, then dried overnight at 120°C and calcined at 600°C in flowing air for 3 hours.

[0050] Table 1

[0051] The composition of the reaction product is listed in Table 2.

[0052] Example 2 Example 1 was repeated except that toluene was used in place of 20 wt% of the feed. This caused the overall conversion (defined as: (the production quality of H 2 , C 1 -C 4 hydrocarbons and δ-aromatic compounds) / raw material quality * 100%) to decrease from 22 wt% to 19 wt%, and the conversion of FCC naphtha itself to increase from 22 wt% to 24 wt%. The addition of toluene demonstrated the beneficial effect of the recycle stream containing aromatic compounds on the cracking reactor. In Example 1, the conversion of the feed was 22%. In Example 2, the conversion was: the absolute conversion on the total feed basis was 19%, or the conversion on the FCC naphtha portion of the feed was 24%. The product selectivity was not affected by the addition of aromatic compounds (as shown by the results reported in Table 2).

[0053] Table 2

[0054] Example 3 Example 1 was repeated for about 3000 minutes (50 hours) except that the feed was hexane and the weight hourly space velocity was 60 h -1 . The real-time feed conversion was represented by Figure 1 black dots (without aromatic compounds).

[0055] Example 4 Example 3 was repeated except that toluene was used in place of 20 wt% of the feed. The real-time feed conversion was represented by Figure 1 hollow dots (with aromatic compounds). Figure 1 It was shown that when toluene was added, the conversion of the crackable portion (hexane) of the feed was low at the beginning. Example 4 with toluene showed significantly improved stability, and the catalyst deactivation over the running time due to catalyst coking was significantly reduced. The addition of toluene did not have a negative impact on the product selectivity of propylene (about 35% in both experiments) and butene (about 20 - 24% in both experiments).

[0056] Example 5 At a weight hourly space velocity of 10 h -1Under these conditions, a non-olefin feed having the components 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 as described in Example 1 for approximately 1200 minutes. The temperature in the reactor was 600 °C. The real-time conversion is as Figure 2 shown by the black dots (first run).

[0057] Table 3

[0058] Example 6 Example 5 was repeated, where a portion of the reactor liquid effluent was recycled to the reactor, thereby replacing a portion of the feed such that the current feed consisted of 80 wt% recycle and 20 wt% fresh feed. The recycle contained 2 wt% olefins. The real-time conversion is as Figure 2 shown by the open circles (recycle stream).

[0059] Figure 2 It is shown that when a portion of the liquid effluent is recycled to the reactor, the overall conversion (of the total mixed feed entering the reactor) is higher. When comparing Examples 5 and 6, there is no significant effect on the selectivity to the desired C 3 and C 4 olefins.

Claims

1. A process for preparing propylene from a hydrocarbon feedstock comprising alkanes, naphthenic compounds and / or aromatic compounds and optionally up to 10 wt% of olefins, wherein the process comprises the following steps: (aa) Feed at a temperature of 450 °C to 700 °C and mixed with a recycle stream is supplied to a continuously operating fixed bed reactor containing a high acid density cracking catalyst, where the mixture contacts the high acid density cracking catalyst at a hydrocarbon partial pressure below 3 bar, except for aromatic compounds, and a weight hourly space velocity of 1 to 30 h -1 -1; (bb) separating propylene and optionally other low-boiling compounds from the effluent of step (aa), leaving a high-boiling fraction; (cc) recycling a portion of the high-boiling fraction to the reactor of step (aa), wherein optionally by additionally supplying an aromatic compound comprising another hydrocarbon mixture to the reactor, the total content of aromatic compounds in the mixture supplied to the reactor in step (aa) is maintained between 5 and 50 wt%.

2. The process according to claim 1, wherein in step (bb), propylene and other low-boiling compounds are separated from the effluent of step (aa) in a distillation column operating under partial vacuum, and the high-boiling fraction is obtained as the bottom product of the distillation column.

3. The process according to claim 1, wherein the high-acid density cracking catalyst in step (aa) is a heterogeneous catalyst comprising mesoporous or macroporous zeolite, selected from ZSM-5, ZSM-11 or β-zeolite.

4. The process according to claim 1, wherein the high-acid density cracking catalyst in step (aa) is a heterogeneous catalyst comprising mesoporous or macroporous zeolite with a silica-alumina ratio of 2 to 1000.

5. The process according to claim 1, wherein the hydrocarbon partial pressure of the hydrocarbons other than aromatic compounds in step (aa) is less than 1 bar.

6. The method according to claim 1, wherein the high acid density cracking catalyst in step (aa) is a heterogeneous catalyst comprising up to 80 wt% of ZSM-5 with a silica-alumina ratio of 2 to 1000, 1-20 wt% of P 2 O 5 and a binder.

7. The process according to claim 1, wherein in step (bb), propylene and other low-boiling compounds are separated from the effluent of step (aa) in a debutanizer distillation column operable under partial vacuum, and a second high-boiling fraction is obtained as the bottom product of the debutanizer distillation column.

8. The process according to claim 1, wherein in step (dd), a portion of the high-boiling fraction is contacted with hydrogen in the presence of an aromatic hydrocarbon conversion catalyst present in the reactor to obtain a fraction rich in aromatic compounds, wherein all or part of the fraction rich in aromatic compounds is recycled as another hydrocarbon mixture to step (aa).

9. The method according to claim 8, wherein the contact in step (dd) occurs at a temperature of 400 °C to 550 °C, a weight hourly space velocity of 0.5 to 5 h -1 , the hydrocarbon partial pressure is less than 10 bar, and the hydrogen partial pressure is less than 10 bar.

10. The process according to claim 8, wherein the aromatic hydrocarbon conversion catalyst is a heterogeneous catalyst comprising ZnO, mesoporous zeolite and binder.

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