PROCESS FOR PREPARING PROPYLENE

Abstract

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

Description

PROCESS FOR PREPARING PROPYLENE Field of Invention The invention relates to a process for preparing propylene from a mixture of hydrocarbons having an olefin content of between 5 and 50% by weight and a boiling point greater than 90% by volume between 35 and 280 °C and / or from a hydrocarbon feed comprising paraffins, naphthenic compounds, aromatic compounds and optionally up to 10% by weight of olefins by contacting the feed with a cracking catalyst in a reactor. Background of the Invention Propylene is produced in more than 50% of cases through steam cracking processes. The typical feedstock is straight-run naphtha, which is obtained by refining a crude oil source that is normally composed of unsaturated compounds, such as paraffinic and naphthenic compounds, optionally mixed with aromatic compounds. Propylene is also produced in a refinery setting as a byproduct of the fluid catalytic cracking (FCC) process. Since the late 1990s, some FCC units have been operating under more severe conditions to achieve a propylene yield of 10 to 12 percent. Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ Ref. 338701 percent by weight of the fresh FCC feed. To further increase propylene yield, different processes have been developed around the FCC setup in a refinery, and propylene yields of up to 20 percent by weight of the fresh FCC feed have been reported. One way to increase propylene yield is to add a medium-pore zeolite to the FCC catalyst, as described, for example, in DE4114874. Several variations have been developed in which the medium-pore catalyst and the FCC catalyst make contact with the hydrocarbon fractions in upstream FCC reactors. A disadvantage of these processes is that the medium-pore zeolite catalyst undergoes a regeneration step along with the FCC catalyst, which causes the medium-pore zeolite catalyst to degenerate. The naphtha fraction obtained from an 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 WO 99 / 29804, which describes a fixed-bed reactor process where 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 Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ was approximately 18 wt% based on the feed. Experiments using a ZSM-5 and using a 1-hexene feed showed the highest propylene yield of 28.8 wt% using a ZSM-5 with a Si / Al atomic ratio of 350 (SAR=750), while experiments using ZSM-5 with Si / Al atomic ratios of 40 and 25 (SAR=80, SAR=50) showed a lower propylene yield and more coke formation. Document GB2345294 describes a process in which a C4 raffinate feed containing phosphate is contacted with a cracking catalyst in a fixed-bed reactor. The catalyst consists of ZSM-5 containing silver in place of a proton and wherein ZSM-5 has a SAR of 300. The reaction temperature is 600 °C at a weight space rate of 47 h⁻¹. A disadvantage of prior art processes is the low propylene yield and high catalyst coke formation. This results in short cycle lengths, i.e., short times between coke removal operations, during which propylene can be prepared in a reactor. The object of this invention is to provide a process that can prepare propylene with a high yield while maintaining catalyst coke formation at a rate that results in an acceptable cycle length. Q7C7 ίΠ / ΖΖηΖ / Ε / ΥΙΛΙ Brief Description of the Figures Figure 1 shows the conversion of the power supply over time as open circuits (w / ring). Figure 2 shows the conversion over time as black circles (first stage). Detailed Description of the Invention The applicants have now found that the following process does not have such a disadvantage. The process for preparing propylene from a mixture of hydrocarbons having a defin content of between 5 and 50% by weight and boiling to more than 90% by volume between 35 and 280 °C and / or from a hydrocarbon feed comprising paraffins, naphthonic compounds and / or aromatic compounds and optionally up to 10% by weight of olefins, wherein the process comprises the following steps: (a) feeding the hydrocarbon mixture, optionally mixed with a recycle stream and having a temperature between 450 and 750 °C, to a reactor where the feed is contacted with a low-density acid cracking catalyst at a hydrocarbon partial pressure of less than 3 bar and a weight space velocity of between 0.5 and 100 h-1, (b) isolating propylene and optionally other low-boiling-point compounds from the effluent of stage (a) in which some high-boiling-point first fractions remain Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ boiling, (c) feeding all or part of the first high boiling point fraction optionally mixed with a recycle stream and having a temperature between 400 and 750 °C to a reactor where the first high boiling point fraction is contacted with a high density acid cracking catalyst at a hydrocarbon partial pressure of less than 3 bar and a weight hourly space velocity of between 0.5 and 100 h-1y wherein the temperature of the mixture of hydrocarbons optionally mixed with a recycle stream as fed to the reactor in stage (a) is lower than the temperature of the first high-boiling fraction, optionally mixed with a recycle stream as fed to the reactor in stage (c), (d) isolating propylene and optionally other low-boiling compounds from the effluent of stage (c) in which a second high-boiling fraction remains, and (e) recycling all or part of the second high-boiling fraction to stage (a) and / or to stage (c) as the optional recycle stream. The applicants have now discovered that a mixture of hydrocarbons comprising paraffins or paraffins and propylene can be efficiently converted into propylene and also into other hydrocarbons. Q7C7 iη / ZZΖΠZ / E / YILI lower olefins in two cracking stages. In the first stage (a), olefins and naphthenes, if present, are converted mainly into propylene, other lower olefins, and paraffins. This is achieved under relatively moderate reaction conditions in the presence of a low-density acid cracking catalyst. Under these conditions, coke formation is minimized. The first high-boiling fraction will have a higher paraffin content than the olefinic hydrocarbon feed described above. This makes it possible to crack this feed under more severe conditions by contact with a high-density acid cracking catalyst. Thus, a process is provided that can convert both olefins and paraffins, and even C5 paraffins, into the hydrocarbon mixture with a high yield of propylene. In addition, the applicants found that coke formation can be kept low.This is believed to be the result of the fact that the first high-boiling fraction contains almost no olefins or at least a small amount. Other advantages will be described below. The feed used in stage (a) is a hydrocarbon mixture. The mixture shall comprise paraffins optionally blended with aromatic and / or naphthenic compounds and olefins having an olefin content of between 5 and 50% by weight and boiling at over 90%. Q7C7 iη / ZZΖΠZ / E / YILI % by volume between 35 and 280 °C and preferably for more than 90% by volume between 35 and 240 °C. The hydrocarbon mixture shall suitably comprise paraffins, naphthenic compounds and / or aromatic compounds along with the paraffins. Such mixtures may be obtained from any source. Suitably, the hydrocarbon mixture comprises or is an isolated fraction of the effluent from a fluid catalytic cracking process such as naphtha cracked with a light catalyst, naphtha cracked with a medium catalyst, naphtha cracked with a heavy catalyst. Other examples are delayed coking naphtha, pyrolysis naphtha, and boiling bed naphtha. Such a mixture may also comprise aromatic compounds, paraffins, and / or naphthenic compounds, and appropriately aromatic compounds, paraffins, and naphthenic compounds. The hydrocarbon mixture may also comprise, or is an isolated fraction of, the effluent from a steam cracking process. Instead of or in addition to the olefinic feed described above, the process according to this invention can also convert a more paraffinic and / or naphthenic mixture of hydrocarbons into propylene. Such a hydrocarbon feed comprises paraffins, naphthenic compounds, and / or aromatic compounds, and optionally up to 10% by weight of defins. Preferably, such an additional or alternative feed boils at more than 90% by volume between 35 and 280 °C, and preferably between 35 and 240 °C. When used as Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ Additional feed, the feed is preferably fed directly to the reactor of stage (c) together with the first high-boiling fraction, optionally mixed with a recycle stream. When used as an alternative feed, it is preferred to use this feed in stage (a). Such a suitable, more paraffinic and / or naphthenic hydrocarbon mixture has a diffin content of less than 10% by weight, more preferably less than 5% by weight, and even more preferably less than 1% by weight. Examples of such mixtures are refinery naphtha fractions, such as first distillation naphtha and light first distillation naphtha, or those obtained from a refinery hydroprocessing process, such as a hydrocracker or hydrotreating process, also referred to as hydrotreated naphthas and hydrocracked naphthas. Other examples include polymerization naphthas, reforming naphthas, and natural gas liquids. The conversion of olefins in step (a) is an endothermic reaction. The required energy can be added to the reactor in several ways. A preferred method is to add inert hydrocarbons, such as paraffins, naphthenic compounds, and / or aromatic compounds, to the olefin mixture. The heat capacity of the feed will then increase per mass of olefin. This is advantageously achieved by recycling Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ part of the first high-boiling fraction as obtained in step (b) to step (a). The weight fraction of the hydrocarbon mixture having an olefin content of between 5 and 50% by weight and boiling more than 90% by volume between 35 and 280 °C, preferably between 35 and 240 °C, in the total feed to the reactor, can therefore be between 25 and 75% by weight. The applicants have found that the presence of aromatic compounds in the feed mixture to the reactor in stages (a) and (c) increases cycle time, stabilizes activity, and helps maximize conversion. This effect is more pronounced when the reactor feed has a low olefin content, such as when starting with the more paraffinic and / or naphthenic hydrocarbon mixture described above. When starting with the olefin feed, then having aromatic compounds in the feed for stage (c) may be preferable. The presence of aromatic compounds does not substantially affect selectivities. Without intending to limit ourselves to the following theory, it is believed that the presence of aromatic compounds reduces coke formation on the cracking catalyst by competitive adsorption on the catalyst surface or by diluting the coke precursors.Secondly, the presence of substantially inert aromatic compounds can increase conversion in reactors when supplied. Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ heat to endothermic cracking. Aromatic compounds are preferably aromatic compounds that boil substantially in the same range as, or just above, the olefinic mixture of hydrocarbons. 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. Aromatic compounds may be present in the olefinic or paraffinic / naffenic feeds described above as used in steps (a) and (c) of this process, or they may be present in the recycle streams described. Preferably, the aromatic compounds are intentionally added to the process mentioned above. For the feed of definites to the reactor in step (a), preferably at least 10 wt% aromatic compounds are present, more preferably at least 20 wt%, even more preferably at least 30 wt%, even more preferably at least 40 wt%, and even more preferably 50 wt% of aromatic compounds. The optimum content of aromatic compounds may be determined by a person skilled in the art, where a determining factor may be the maximum conversion in a single step.The upper limit for the content of aromatic compounds can be 80% by weight, although it is more preferable than the content of. Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ aromatic compounds in the hydrocarbon mixture, including the optional recycle stream or streams fed to the reactor in stage (a), shall be between 10 and 80% by weight, more preferably between 20 and 70% by weight, even more preferably between 30 and 60% by weight and even more preferably between 40 and 50% by weight. For the mixture fed to the reactor in step (c), it is preferred that it contain at least 5 wt%, more preferably at least 10 wt%, and even more preferably at least 20 wt%, and preferably at most 80 wt%, more preferably at most 40 wt%. Such aromatic compound contents are especially preferred when the conversion of olefins in step (a) is such that the olefin content in the first high-boiling fraction obtained in step (b) is below 10 wt%, and more preferably below 5 wt%. Such olefin contents are also preferred for operating step (c) at the desired cycle lengths and selectivity when the aromatic compound content is outside the above ranges. As described above, it is preferable to intentionally add aromatic compounds to the reactor feed in stages (a) and / or (c). These aromatic compounds can be sourced from other parts, for example, from Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ a refinery or a steam cracker. Preferably, the aromatic compounds are prepared in a separate processing stage by contacting a portion of the first high-boiling fraction and / or all or part of the second high-boiling fraction in a stage (f) with hydrogen in the presence of an aromatic conversion catalyst such as that present in a reactor to obtain an aromatic-rich fraction. The desired aromatic content can be achieved by recycling all or part of the aromatic-rich fraction to stage (a) and / or to stage (c). A portion of the aromatic-rich fraction can also be recycled to stage (f) itself. Preferably, a portion of the aromatic compounds is isolated from these recycling streams to avoid an accumulation of substantially inert aromatic compounds.This in itself is not disadvantageous because these aromatic compounds, such as benzene, toluene, and xylene, represent desirable compounds for use as such. This aromatic compound conversion step (f) is itself well-known and is also called reforming. Step (f) can be carried out using the well-known reforming processes provided by UOP. Step (f) can take place at a temperature between 400 and 700 °C, preferably between 400 and 650 °C, and even more preferably between 400 and 550 °C. Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ °C, at a weight space hour velocity (WHSV) of between 0.1 and 50 h-1, preferably between 0.5 and 25 h-1 and even more preferably between 0.5 and 5 h-1, a partial pressure of hydrocarbons of less than 10 bar and a partial pressure of hydrogen of less than 10 bar. The aromatic conversion catalyst can be any reforming catalyst or a heterogeneous catalyst comprising ZnO, a medium-pore zeolite, and a binder. The medium-pore zeolite is preferably ZSM-5, and the binder is preferably alumina. The binder may also contain some P₂O₅. A preferred catalyst comprises 25 to 60 wt% ZSM-5, 5 to 35 wt% ZnO, and 2.5 to 20 wt% P₂O₅, along with an alumina binder. Such a catalyst can be prepared by adding a ZSM-5 zeolite, for example, 50 parts of ZSM-5 crystals (SAR 30, formerly Zeolyst), to a quantity of water. This aqueous mixture can be added to a gelled alumina, for example, 35 parts of the dry base, Catapal B, formerly Sasol, and zinc nitrate, for example, 10 parts of the dry base, formerly Alpha Aesar technical grade, and kneaded. P₂O₅, for example, 5 parts, as dilute phosphoric acid, can then be added to this kneaded material.The mixture is extruded and dried, for example, at 120 °C for 1 hour and calcined, for example, for 1 hour at 600 °C. The reactor in which step (f) can be performed Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ can be a fixed-bed reactor, a radial-bed reactor, a moving-bed reactor, a bubbling-bed reactor, or a fluidized-bed reactor. A preferred reactor is a fixed-bed reactor. In some embodiments of this invention, the reactor is not a fixed-bed reactor. In steps (b) and (d), propylene and optionally other low-boiling compounds are isolated from the effluent of steps (a) and (c), respectively, leaving high-boiling fractions. The other low-boiling compounds may include, for example, ethane, ethylene, hydrogen, water, propane, and butylenes. Such separation may involve distillation and / or flash separation. Because the selectivity of propylene over the total propylene and propane is improved, less propane is formed. This is advantageous because it makes obtaining, for example, polymer-grade propylene, less difficult. In this separation, ethylene can be isolated from the low-boiling compounds. The C4 fraction, which includes butane and butylene, can be recovered as such or recycled together as part of the high-boiling compounds as described. The reactor in stage (a) can be a fixed-bed reactor, for example, a radial bed reactor, a moving-bed reactor, a bubbling-bed reactor, or a reactor Q7C7 ίη / ZZΖΠZ / E / YILI of fluidized bed. The reactor in stage (c) may be a fixed-bed reactor, a radial-bed reactor, a moving-bed reactor, a bubbling-bed reactor, or a fluidized-bed reactor. The reactor in stage (a) and / or stage (c) may specifically not be a fixed-bed reactor. The reactors described above for use in stages (a) and / or (c) may be equipped with internal tubes to allow the flow of superheated steam or other superheating medium. This steam will add energy by indirect heat exchange to the endothermic reaction taking place in stages (a) and / or (c), allowing the conversion of the reactant to higher conversion levels within the reactors. The catalyst present in the reactor in step (a) can be any cracking catalyst with a relatively low acid density. Low acid density results in a relatively large distance between acid sites, preventing intermediate reaction compounds from forming coke. The catalyst is active in the conversion of flakes, while paraffins hardly react. A possible low acid density catalyst is an amorphous catalyst, such as one comprising amorphous silica alumina, zirconia silica, and / or silicon borate as the amorphous low acid density component.Preferably, the low-density acid catalyst in step (a) is a heterogeneous catalyst comprising a medium- or large-pore zeolite having a silica-to-alumina ratio of between 2 and 1000, more preferably between 10 and 1000, even more preferably between 10 and 300, even more preferably between 20 and 300, and still more preferably between 20 and 100. For example, one can start with a new catalyst having a relatively low silica-to-alumina ratio. Over time, this ratio can be increased to a higher ratio due to dealumination. The resulting decrease in activity can be compensated for by operating at a higher temperature. Examples of suitable medium- or large-pore zeolites are ZSM-5, ZSM-11, and beta zeolite. An example of a suitable low acid density catalyst comprises up to 70 wt% ZSM-5, between 1-20 wt% P2O5 and a binder.Examples of suitable binders include alumina, such as boehmite, optionally mixed with a clay to increase strength. The catalyst preferably comprises between 25 and 80% by weight, more preferably between 25 and 70% by weight, and even more preferably between 35 and 50% by weight of ZSM-5. The catalyst present in the reactor in stage (c) can be any cracking catalyst that has a relatively high acid density. When the first high-boiling fraction contains a high content of definites, especially in the case that the content of If the Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ defines is greater than the olefin content in the hydrocarbon mixture fed to stage (a), a catalyst of lower acid density may also be used, for example, such as the catalyst described above for stage (a). Therefore, the high- and low-acid-density catalysts may preferably comprise a large-pore or medium-pore zeolite having a relatively low silica-to-alumina ratio. Preferably, the silica-to-alumina ratio of the high-acid-density catalyst is lower than the silica-to-alumina ratio of the low-acid-density catalyst. Examples of suitable medium- or large-pore zeolites are ZSM-5, ZSM-11, and beta zeolite.An example of a suitable high-density acid catalyst in step (c) is a heterogeneous catalyst comprising up to 80 wt%, preferably up to 70 wt%, of ZSM-5 having a silica-to-alumina ratio of 2 to 1000, more preferably 10 to 1000, even more preferably 10 to 300, and even more preferably 25 to 100, 1–20 wt% of P₂O₅, and a binder. Examples of suitable binders include alumina, such as boehmite, optionally blended with a clay to increase strength. The catalyst preferably comprises 25–80 wt%, more preferably 25–70 wt%, and even more preferably 35–50 wt% of ZSM-5. Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ The catalysts in step (a) and / or step (c) may be vaporized prior to use. This is to cause some initial deactivation of the catalyst to limit the activity range. With a narrower activity range, the conversion can be controlled by adjusting the temperature. This is not possible if the activity range is too wide. Vaporization can be carried out by contacting the catalyst in the reactor with a gas comprising 1–100% by volume of vapor, preferably 5–100% by volume of vapor, more preferably 5–10% by volume of vapor, and even more preferably 70–95% by volume. The preferred pressure may be approximately atmospheric pressure with a maximum pressure of 10 bar. The preferred temperature is between 300 and 800 °C, more preferably between 400 and 750 °C, and most preferably between 450 and 600 °C.Contact time can range from 1 hour to 5 days, with contact times of approximately 1 day preferred. The balanced catalysts of step (c) that have to be replaced due to their higher proportion of silica and alumina can be advantageously used as catalysts in step (a). The zeolite catalysts used in steps (a) and (c) can be prepared from a zeolite having the desired silica-to-alumina ratio. The zeolite is suitably suspended in distilled water and mixed with a Q7C7 iη / ZZΖΠZ / E / YILI alumina gel. The gel is prepared, for example, using nitric acid and Sasol's Catapal B. The mixture is extruded, producing a particle comprising a zeolite and an alumina binder. The particle is calcined, for example, in air for 1 hour at 600 °C. The calcined particle is then impregnated with phosphoric acid and calcined again, for example, by exposure to air for 1 hour at 600 °C. In the process according to this invention, the space velocity per hour by weight is defined based on the total hydrocarbons fed to a reactor. Therefore, optional recycle streams and / or streams rich in added aromatic compounds are also included. This also applies to the temperature of the feed to the reactors. The temperature values ​​refer to the temperature of the total hydrocarbons fed to the reactor. Preferably, the hourly space velocity by weight in stage (a) is between 0.5 and 100 h⁻¹, more preferably between 0.5 and 50 h⁻¹, even more preferably between 1 and 25 h⁻¹, and most preferably between 1 and 10 h⁻¹. A further preference is between 1 and 5 h⁻¹. The hourly space velocity by weight in stage (c) is between 0.5 and 100 h⁻¹, more preferably between 1 and 100 h⁻¹, even more preferably between 1 and 50 h⁻¹, and most preferably between 2 and 30 h⁻¹ or between 2 and 20 h⁻¹. Preferably, the hourly space velocity by weight in stage (a) is greater than the hourly space velocity in stage (c). Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ weight in stage (c). The partial pressure of the hydrocarbon is preferably below 1 bar, more preferably below 0.5 bar and even more preferably below 0.2 bar. The partial pressure of the hydrocarbon excluding aromatic compounds in stages (a) and (c) is preferably below 1 bar, more preferably below 0.5 bar and even more preferably below 0.2 bar. The temperature of the hydrocarbon mixture optionally in a mixture with a recycle stream to stage (a) has a temperature between 450 and 750 °C, preferably between 450 and 650 °C, more preferably between 500 and 650 °C. The temperature of the first high-boiling fraction optionally mixed with a recycle stream in stage (c) is between 400 and 750 °C, preferably between 450 and 700 °C, more preferably between 450 and 650 °C and even more preferably between 500 and 650 °C. The reactor in stage (a), stage (c), and / or stage (f) is preferably a multi-reactor configuration. For example, such a configuration could be a set of reactors operated in parallel. These reactors could be the reactors listed above or combinations thereof. Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ As described above, the applicant discovered that the presence of aromatic compounds is advantageous for achieving longer cycles, stabilizing activity, and helping to maximize the conversion of paraffinic / naphthonic mixtures. For this reason, the invention also applies to the following processes. A process for preparing propylene from an initial hydrocarbon feed comprising paraffins, naphthenic compounds, aromatic compounds and optionally up to 10 wt% of olefins by adding aromatic compounds to the initial hydrocarbon feed resulting in an enhanced feed containing between 10 and 70 wt%, preferably between 20 wt% and 50 wt% and even more preferably between 25 and 40 wt% of aromatic compounds, wherein the aromatic compound content in the initial feed is below the lower end of these ranges and wherein the enhanced feed is catalytically cracked in the presence of an acid cracking catalyst to propylene and other reaction products. The aromatic compounds added to the initial hydrocarbon feed can be any aromatic compound, including those that boil within the gas oil boiling range. The aromatic compounds described above are appropriately added to the initial hydrocarbon feed. The catalyst of Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ Acid cracking can be as described in this application. The reactor and conditions can be those described for this invention. Alternatively, the reactor can be a fluidized bed. The aromatic compounds added to the initial hydrocarbon feed can be aromatic compounds such as those separated from the reaction products of this process and reused in the process. It appears that the aromatic compounds are not cracked to any significant degree into other products and that the amount of total aromatic compounds in the enhanced feed and the reactor effluent is approximately the same. Such a recycling process is described in the following process according to the invention. A process for preparing propylene from a hydrocarbon feed comprising paraffins, naphthenic compounds, aromatic compounds and optionally up to 10% by weight of olefins, wherein the process comprises the following steps: (aa) feeding the feed mixed with a recycle stream and having a temperature between 450 and 700 °C, preferably between 550 and 700 °C, to a continuously operated reactor comprising a high-density acid cracking catalyst where the mixture is contacted with a high-density acid cracking catalyst at a hydrocarbon partial pressure excluding aromatic compounds below 3 bar, more preferably below 1 bar, Q7C7 iη / ZZΖΠZ / E / YILI even more preferably below 0.5 bar and even more preferably below 0.2 bar and at a weight hourly space rate of between 1 and 30 h-1, preferably between 2 and 30 h-1, (bb) isolating propylene and optionally other low-boiling compounds from the effluent of stage (aa) in which high-boiling fractions remain, (cc) recycling part of the high-boiling fraction to the reactor of stage (aa) wherein the total content of aromatic compounds in the combined mixture as fed to the reactor in stage (aa) is maintained between 5 and 50% by weight, preferably between 10 and 40% by weight, and even more preferably between 20 and 30% by weight, optionally further feeding to the reactor a hydrocarbon mixture comprising compounds aromatic. The high-density acid catalyst in step (aa) can be as described above for step (c). The preferred conditions and catalyst for operating this process are the same as described above for steps (c) through (e). It is also preferred to add an aromatic compound conversion using the high-boiling fraction in a step (dd) with hydrogen in the presence of an aromatic conversion catalyst, such as that found in a reactor, to obtain a fraction rich in aromatic compounds. Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ where all or part of the aromatic-rich fraction is recycled to stage (aa) as an additional hydrocarbon mixture. The conditions and catalyst are as described above. The feed for this process can boil more than 90% by volume between 35 and 280 °C and preferably more than 90% by volume between 35 and 240 °C. The feed appropriately has a dilution content of less than 10% by weight, more preferably less than 5% by weight, and even more preferably dilution-free. Examples of such mixtures are the naphtha fractions obtained in a refinery hydroprocessing process, such as hydrocracking or hydrotreating. The invention also relates to a process configuration suitable for preparing propylene from a hydrocarbon mixture comprising definites, which comprise: (i) one or more first reactors operated in parallel comprising an amorphous heterogeneous cracking catalyst or a heterogeneous cracking catalyst comprising a medium or large pore zedite having a silica-to-alumina ratio of between 1 and 1000, (ii) first distillation and / or flash separation units fluidly connected to the outlet of one or more first reactors operated in parallel having at least Q7C7 iη / ZZΖΠZ / E / YILI an outlet for a fraction comprising propylene 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 units to the inlet of one or more first reactors operated in parallel, (iv) one or more second reactors operated in parallel comprising a heterogeneous cracking catalyst comprising up to 80 wt% of ZSM-5 having a silica-to-alumina ratio of between 2 and 1000, preferably between 25 and 50, between 1-20 wt% of P2O5 and a binder and wherein the inlet of the second reactors is fluidly connected to the outlet for the high-boiling compounds of the first distillation and / or flash separation unit,(v) second distillation and / or flash separation units fluidly connected to the outlet of one or more second reactors of (iv) having at least one outlet for a fraction comprising propylene 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 units to the inlet of one or more of the first reactors operated in parallel and to the inlet of one or more second reactors operated in, Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ parallel . The first and second reactors may be a fixed-bed reactor, a fluidized-bed reactor, a bubbling-bed reactor, a boiling-bed reactor (e.g., a radial-bed reactor), or a moving-bed reactor, or combinations thereof. Preferably, the first and second reactors are fixed-bed reactors, and in one embodiment of this invention, the first and / or second reactors are not fixed-bed reactors. The first and second reactors may be equipped with internal tubes to allow the flow of superheated steam or other superheated medium. This steam will add energy by indirect heat exchange to the endothermic reaction taking place in steps (a) and / or (c), enabling the conversion of the reactant to higher conversion levels within the reactors. Preferably, the process configuration further comprises inlet means (vii) for an additional hydrocarbon feed fluidly connected to the inlet of one or more second parallel reactors. The flash separation units may be suitably combined with an extraction stage to recover any compounds higher than Cs that may be present in the gaseous effluent. Preferably, the process configuration further comprises (viii) one or more parallel-operated aromatic conversion reactors fluidly connected to the Q7C7 iη / ZZΖΠZ / E / YILI outlet for high-boiling compounds from the second distillation and / or flash separation unit and means for recycling part of the effluent from the aromatic conversion reactors to the inlet of one or more first reactors, to the inlet of one or more second reactors and the inlet of the aromatic compound conversion reactors. The aromatic conversion reactors suitably have a hydrogen inlet and have a bed of a heterogeneous catalyst comprising ZnO, a medium-pore zeolite, and a binder as also described above. The aromatic compound conversion reactor can be a fixed-bed reactor, a fluidized-bed reactor, a bubbling-bed reactor, a boiling-bed reactor, a radial-bed reactor, or a moving-bed reactor. The one-stage process or stage (c) of the two-stage process according to this invention is an energy-intensive process that can produce large quantities of light olefins, such as ethylene, propylene, and butylene. The hydrocarbon feed can be heated to a reactor inlet temperature of 450°C or higher via a feed / effluent heat exchanger followed by a natural gas-fired heater. The reactor effluent can then be further reduced in temperature by a combination of a cooled heat exchanger Q7C7 Ln / Zznz / E / YIAI is heated by air through a water-cooled heat exchanger, with the target temperature being appropriately between 25 and 30 °C. The low-boiling fraction can then be separated through a single-stage equilibrium distillation operation, in which the overhead vapors and low-boiling products are further separated in a product recovery unit. The unreacted, high-boiling liquid hydrocarbon can be advantageously recycled. This recycled material, which has a temperature, for example, between 25 and 30 °C, is combined with a fresh hydrocarbon feed and reheated to reactor inlet temperatures of 450 °C or higher. A less energy-intensive alternative process modifies the basic process described above as follows. The reactor effluent exchanges heat with the feed in a feed / effluent heat exchanger network and is subsequently fed to a distillation column, such as a debutanizer distillation column. The operating temperature of the reactor effluent leaving the feed / effluent heat exchanger can range from 200 to 300 °C. Butylene and the lighter components of the reactor effluent are withdrawn from the top of the distillation column as low-boiling compounds. Pentane and the heavier components of the reactor effluent are the bottom products of the Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ distillation column, i.e., high-boiling-point compounds. The bottom product is treated as a recycle as described in the basic configuration above. The operating temperature of the recycle stream exiting the bottom stream of the debutanizer distillation column can be 250–350 °C. A preferred recycle-to-fresh-feed ratio ranges from 2.0 to 4.0 (recycled mass / fresh-feed ratio). For example, 67% by weight of the reactor feed consists of recycle. The alternative process described above eliminates the need to impart an energy equivalent of up to 420 °C to the reactor feed for this recycle stream, resulting in significant energy savings.A negative consequence of the alternative process flow scheme is that the pressure drop in the debutanizer distillation column can be 0.21 bar (3.0 psig). This is higher than the pressure drop of approximately 0.14 bar (2.0 psig) that results from cooling the reactor effluent in air-cooled heat exchangers followed by water-cooled heat exchangers. This additional pressure drop of 0.068 bar (1.0 psig) can be added to the reactor inlet operating pressure when the alternative process is used. Because increasing the reactor inlet pressure can negatively affect the selectivity of Q7C7 ii / ZZPZ / E / YILI propylene, it is preferred that the debutanizer distillation column operate under partial vacuum to eliminate the required increase in reactor inlet pressure. The upper accumulator drum of the debutanizer distillation column will function as the drum section for a centrifugal compressor. The centrifugal compressor can preferably produce a full vacuum of approximately 0.69–1.03 bar (10–15 psia). This will eliminate the need to increase the reactor inlet pressure while still obtaining the benefit of the alternative process flow. The reactor inlet pressure for the alternative process can operate, for example, at 1.27 bar (18.5 psia) versus 2.56 bar (37.1 psia) for the base process, resulting in a further increase in propylene selectivity. The discharge from the centrifugal compressor can pass through a chilled water heat exchanger to reduce the operating temperature to approximately 30°C. The cooled hydrocarbon stream is then routed to a high-pressure separator for the efficient removal of higher molecular weight compounds. The overflow from this separator is routed to a second boost centrifugal compressor to provide pressure for product separations. The liquid hydrocarbon stream from the high-pressure separator can then Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ can be distilled in the product recovery section with a lower energy requirement as a result of operating at a higher pressure. Example 1 A pilot plant fixed-bed reactor containing 1.5 grams of a fixed-bed catalyst at a WHSV of 30 h-1 was fed with FCC naphtha with a boiling point between 20 and 206 °C and having the composition shown in Table 1. The temperature in the reactor was 600 °C. ZSM-5 crystal with SAR 30 (CBV 3024E, from Zeolyst) is milled in a 55 / 45 wt / w blend with alumina (ex Sasol) and extruded to prepare a formed mass. The extruded mass was dried at 120 °C overnight and calcined for 3 hours in a stream of air at 600 °C. The calcined extruded products were impregnated to incipient moisture with phosphoric acid and then dried at 120 °C overnight and calcined for 3 hours at 600 °C in circulating air. Table 1 Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ Total normal paraffin 4.4% weight Total isoparaffin 31.1% weight Total saturated naphthene 6.9% weight Total unsaturated naphthene 3.9% wt. Total normal olefin 12.3% wt. Total isolefin 19.8% wt. Total diolefin 0.2% wt. Total aromatic compounds 21.4% wt. Total 100.0% wt. Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ The composition of the reaction products is listed in Table 2. Example 2 Example 1 was repeated except that 20 wt% of the feed was replaced with toluene. This resulted in the overall conversion (defined as: (mass production of H2, C1-C4 hydrocarbons, and delta aromatic compounds) / (mass feedstock)*100%) decreasing from 22 wt% to 19 wt%, and the conversion of the FCC naphtha itself increasing from 22 wt% to 24 wt%. The addition of toluene illustrates the advantageous effect of a recycle containing aromatic compounds in the cracking reactor. In Example 1, the conversion of the feed as such was 22%. In Example 2, the conversion was the absolute conversion based on the total feed, which was 19%, or the 24% conversion of the FCC naphtha portion of the feed. The product selectivities were not affected by the addition of aromatic compounds, as shown in the results reported in Table 2. Table 2 Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ Example 1 Example 2 Feed FCC Naphtha FCC Naphtha plus 20% toluene Reaction Products (% by weight of aromatic compounds in the delta fraction of Cl-C4) CH4 1% 1% C2+ 1% 1% C2= 16% 16% C3+ 4% 3% C3= 42% 43% iC4+ 1% 1% nC4+ 1% 1% iC4= 9% 9% nC4= 15% 15% Aromatic Compounds 10% 10% Total Coke Yield 0.13% by weight 0.08% by weight Example 3 Example 1 was repeated for approximately 3000 minutes (50 hours), except that the feed was now hexane and the WHSV was 60 h-1. The feed conversion over time is shown by black circles (no aromatic compounds) in Figure 1. Example 4 Example 3 was repeated except that 20 wt% of the feed was replaced with toluene. The feed conversion over time is represented by open circles (w / ring) in Figure 1. Figure 1 shows that the conversion of the crackable portion of the feed (hexanes) was initially lower when toluene was added. Example 4 with toluene showed substantially improved stability and catalyst deactivation over time in the stream because catalyst coking was significantly reduced. The addition of toluene did not negatively affect the product selectivities of propylene (around 35% for both experiments) and butylenes (around 20-24% for both experiments). Example 5 A pilot plant fixed-bed reactor containing 3 grams of a fixed-bed catalyst described in Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ Example 1 at a WHSV of 10 h-1, a non-olefinic feed was fed having the composition indicated in Table 3 and boiling between 20 and 220 °C for about 1200 minutes. The temperature in the reactor was 600 °C. The conversion over time is shown as the black circles (first stage) in Figure 2. Table 3 Q7C7 ίη / ΖΖΠΖ / Ε / ΥΙΛΙ Compounds % by weight Total normal paraffins 23 Total isoparaffins 44 Total naphthalenes 23 Total olefins 0 Total aromatic compounds 10 Xylenes 3.3 C9 aromatic compounds 6.7 Total 100 Example 6 Example 5 was repeated, in which part of the liquid effluent from the reactor was recycled back into the reactor, thus replacing part of the feed. The feed then consisted of 80% by weight of the recycle and 20% by weight of the fresh feed. The recycle contained 2% by weight of olefins. The conversion over time is shown as open circles (recycling) in Figure 2. Figure 2 shows that the overall conversion (of the total combined feed to the reactor) is higher when a portion of the liquid effluent is recycled to the reactor. The selectivities for the desired C3 and C4 olefins were not significantly affected when comparing Examples 5 and 6. It is hereby stated that, as of this date, the best method known to the applicant for putting the aforementioned invention into practice is the one that is clear from the present description of the invention.

Claims

1. A process for preparing propylene from a hydrocarbon mixture having a diffin content of between 5 and 50% by weight and boiling to more than 90% by volume between 35 and 280 °C or from a hydrocarbon feed comprising paraffins, naphthenic compounds and / or aromatic compounds and optionally up to 10% by weight of diffins, characterized in that it comprises the following steps: (a) feeding the hydrocarbon mixture optionally mixed with a recycle stream and having a temperature between 450 and 750 °C to a reactor where the feed is contacted with a low-density acid cracking catalyst at a hydrocarbon partial pressure of less than 3 bar and a weight space velocity of between 0.5 and 100 Ir1, (b) isolating propylene and, optionally, other low-boiling compounds from the effluent of stage (a) in which first high-boiling fractions remain, (c) feeding all or part of the first high-boiling fraction, optionally mixed with a recycle stream and having a temperature between 400 and 750 °C, to a reactor where the first high-boiling fraction is contacted with a high-density acid cracking catalyst at a hydrocarbon partial pressure of less than 3 bar and a weight-hour space velocity of between 0.5 and 100 h-1 and wherein the temperature of the hydrocarbon mixture optionally mixed with a recycle stream as fed to the reactor in stage (a) is lower than the temperature of the first high-boiling fraction, optionally mixed with a recycle stream as fed to the reactor in stage (c), (d) isolating propylene and optionally other low-boiling compounds from the effluent of stage (c) in which a second high-boiling fraction remains, and (e) recycling all or part of the second high-boiling fraction to stage (a) and / or to stage (c) as the optional recycle stream.

2. The process according to claim 1, characterized in that the space hour weight velocity of stage (a) is greater than the space hour weight velocity of stage (c).

3. The process according to any of claims 1-2, characterized in that part of the first high-boiling fraction and / or all or part of the second high-boiling fraction is contacted in a step (f) with hydrogen in the presence of an aromatic conversion catalyst such as that 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 an optional recycle stream and / or to step (f).

4. The process according to claim 3, characterized in that the contact in step (f) is carried out at a temperature of between 400 and 700 °C, at a weight space velocity of between 0.1 and 50 h-1, a partial pressure of the hydrocarbon below 10 bar and a partial pressure of hydrogen below 10 bar.

5. The process according to any of claims 3-4, characterized in that the aromatic conversion catalyst is a heterogeneous catalyst comprising ZnO, a medium pore zeolite and a binder.

6. The process according to any of claims 1-5, characterized in that the low acid density catalyst is an amorphous catalyst.

7. The process according to any of claims 1-5, characterized in that the low acid density catalyst in step (a) is a heterogeneous Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ catalyst comprising a medium or large pore zeolite having a silica to alumina ratio of between 2 and 1000.

8. The process according to claim 7, characterized in that the heterogeneous catalyst comprises up to 70% by weight of ZSM-5, between 1-20% by weight of P2O5 and a binder.

9. The process according to claim 8, characterized in that the heterogeneous catalyst comprises between 25 and 80% by weight of ZSM-5.

10. The process in accordance with any of claims 1-8, characterized in that the space speed hourly weight in stage (a) is between 0.5 and 50 h'1.

11. The process according to any of claims 1-9, characterized in that the high acid density catalyst in step (c) is a heterogeneous catalyst comprising up to 80 wt% of ZSM-5 having a silica to alumina ratio of between 2 and 1000, between 1-20 wt% of P2O5 and a binder.

12. The process according to any of claims 7-10, and claim 11, characterized in that the silica to alumina ratio of the high acid density catalyst is less than the silica to alumina ratio of the low acid density catalyst.

13. The process according to claim Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ 12, characterized in that the heterogeneous catalyst comprises between 35 and 50% by weight of ZSM-5.

14. The process in accordance with any of claims 1-13, characterized in that the space velocity hourly weight in stage (c) is between 1 and 50 h-1.

15. The process according to any of claims 1-14, characterized in that the partial pressure of the hydrocarbon excluding aromatic compounds in steps (a) and (c) is less than 1 bar.

16. The process according to any of claims 1-15, characterized in that the hydrocarbon mixture comprises a fraction as isolated from the effluent of a fluid catalytic cracking process and / or isolated from the effluent of a steam cracking process.

17. The process according to any of claims 1-16, characterized in that the content of aromatic compounds in the hydrocarbon mixture, including optional recycle streams fed to the reactor in stage (c), is between 10 and 80% by weight.

18. The process according to any of claims 1-17, characterized in that a portion of the second high-boiling fraction obtained in step (d) is recycled to step (c) as the recycle stream. Q7C7 ίη / ZZΖΠZ / E / YΙΛΙ 19. The process according to claim 18, characterized in that in step (d) propylene and other low-boiling compounds are isolated from the effluent of step (c) in a debutanizer distillation column operating under partial vacuum and wherein the second high-boiling fraction is obtained as a bottom product of the debutanizer distillation column.

20. A process for preparing propylene from an initial hydrocarbon feed comprising paraffins, naphthenic compounds and / or aromatic compounds and, optionally, up to 10 wt% of olefins by adding the aromatic compounds to the initial hydrocarbon feed, resulting in an enhanced feed containing between 10 and 70 wt%, and characterized in that the enhanced feed is catalytically cracked in the presence of an acid cracking catalyst to give propylene and other reaction products.

21. The process according to claim 20, characterized in that the improved feed contains between 20 and 50% by weight of aromatic compounds.

22. A process for preparing propylene from a hydrocarbon feed comprising paraffins, naphthenic compounds and / or aromatic compounds and optionally up to 10 wt% of olefins, characterized in that it comprises the following steps: (aa) feeding the feed mixed with a recycle stream and having a temperature of between 450 and 700 °C to a continuously operating reactor comprising a high-density acid cracking catalyst wherein the mixture is contacted with a high-density acid cracking catalyst at a hydrocarbon partial pressure, excluding aromatic compounds, of less than 3 bar and a weight space rate of between 1 and 30 h'1, (bb) isolating propylene and optionally other low-boiling compounds from the effluent of step (aa) wherein high-boiling-point fractions remain,(cc) recycling part of the high-boiling fraction to the reactor of stage (aa) wherein the total content of aromatic compounds in the combined mixture as fed to the reactor in stage (aa) is maintained between 5 and 50% by weight, optionally further feeding to the reactor a hydrocarbon mixture comprising additional aromatic compounds.

23. The process according to claim 22, characterized in that in step (bb) propylene and other low-boiling compounds are isolated from the effluent of step (aa) in a distillation column operating under partial vacuum and wherein the high-boiling fraction is obtained as a bottom product of the distillation column.

24. The process according to any of claims 22-23, characterized in that the high acid density catalyst in step (aa) is a heterogeneous catalyst comprising up to 80 wt% ZSM-5 having a silica to alumina ratio of between 2 and 1000, between 1-20 wt% P2O5 and a binder.

25. The process according to any of claims 22-24, characterized in that a portion of the high-boiling fraction is contacted in a step (dd) with hydrogen in the presence of an aromatic conversion catalyst as present in a reactor to obtain an aromatic-rich fraction and wherein all or a portion of the aromatic-rich fraction is recycled to step (aa) as the additional hydrocarbon mixture.

26. The process according to claim 25, characterized in that the contact in step (dd) is carried out at a temperature of between 400 and 550 °C, at a weight space velocity of between 0.5 and 5 h-1, a hydrocarbon partial pressure below 10 bar and a hydrogen partial pressure below 10 bar.

27. The process according to any of claims 25-26, characterized in that the aromatic conversion catalyst is a heterogeneous catalyst comprising ZnO, a medium pore zeolite and a binder.

28. A process configuration suitable for preparing propylene from a hydrocarbon mixture comprising olefins, characterized in that it comprises (i) one or more first reactors operated in parallel comprising an amorphous heterogeneous cracking catalyst or a heterogeneous cracking catalyst comprising a medium- or large-pore zeolite having a silica-to-alumina ratio of between 1 and 1000, (ii) first distillation and / or flash separation units fluidly connected to the outlet of one or more first reactors having at least one outlet for a propylene-containing fraction and one outlet for high-boiling compounds, (iii) means for recycling the high-boiling compounds from the outlet of the distillation and / or flash separation units to the inlet of one or more first reactors operated in parallel,(iv) one or more second reactors operated in parallel comprising a heterogeneous cracking catalyst comprising up to 80 wt% ZSM-5 having a silica-to-alumina ratio of between 2 and 1000, between 1-20 wt% P2O5 and a binder and wherein the inlet of the second reactors is fluidly connected to the outlet for high-boiling compounds of the first Q7C7 ίη / ZZΖΠZ / E / YILI distillation and / or flash separation unit, (v) second distillation and / or flash separation units fluidly connected to the outlet of one or more second reactors of (iv) having at least one outlet for a fraction comprising propylene and one outlet for high-boiling compounds,(vi) means for recycling high-boiling-point compounds from the outlet of the second distillation and / or flash separation units 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.

29. The process configuration according to claim 28, characterized in that it further comprises inlet means (vii) for an additional hydrocarbon feed fluidly connected to the inlet of one or more second parallel reactors.

30. The process configuration according to any of claims 28-29, characterized in that it further comprises (viii) one or more parallel-operated aromatic compound conversion reactors fluidly connected to the outlet for high-boiling compounds of the second distillation and / or flash separation units and means for recycling a portion of the effluent from the aromatic compound conversion reactors to the inlet of one or more of the first reactors, Q7C7 Ln / Zznz / E / YIAI to the inlet of one or more of the second reactors and to the inlet of the aromatic compound conversion reactors.

31. The process configuration according to claim 30, characterized in that the aromatic conversion reactors have a hydrogen inlet and have a bed of a heterogeneous catalyst comprising ZnO, a medium pore zeolite and a binder.