Improved Paraffin Utilization in the Production of Linear Alkylbenzenes

Abstract

The present invention provides a process for producing linear alkylbenzenes from an alkane feed with improved yields due to the use of a novel catalyst, the process having a reduced selectivity to produce undesirable aromatic compounds. In some embodiments, the process can now be operated without a unit for removing such aromatic compounds, allowing an operator of the plant to modify the process to increase product yields.

Description

[0001] Priority Statement

[0002] This application claims priority to U.S. Application No. 16 / 818,085, filed on March 13, 2020, the entire disclosure of which is incorporated herein by reference. Technical Field

[0003] The present invention relates to the production of monoalkylated aromatic compounds (including linear alkylbenzenes). Specifically, the present invention relates to the conversion of alkanes, which reduces the production of undesirable aromatic compounds. Background Art

[0004] The alkylation of benzene with olefins produces various alkylbenzene compounds with various commercial uses. Examples include the alkylation of benzene with olefins having 8 to 16 carbons to produce intermediate compounds for the manufacture of detergents. Alkylbenzenes are sometimes referred to as phenylalkanes and are produced as commodities in large-scale facilities throughout the world, where production rates range between 50,000 and 200,000 metric tons per year. Alkylation processes involve reacting benzene with olefins in the presence of a catalyst at elevated temperatures. The catalyst can be a homogeneous or heterogeneous catalyst, such as hydrogen fluoride, aluminum chloride, silica-alumina, or zeolite catalysts.

[0005] The desired alkylated compound is a monoalkylated aromatic compound. Monoalkylated aromatic compounds include linear alkylbenzenes (LABs) used to form linear alkylbenzene sulfonates (LABS), which are compounds commonly used in the manufacture of detergents. Two common reactions for producing monoalkylated aromatic compounds are the alkylation of aromatic compounds (such as benzene) and the transalkylation of polyalkylated aromatic compounds.

[0006] Today, linear alkylbenzenes are produced from C9-C14 linear alkane streams by first dehydrogenating to form monoolefins and then passing them to an alkylation catalyst with benzene. The dehydrogenation process results in a mixture of species (including diolefins, triolefins, and aromatics). These species then must be removed by additional processing steps and at the expense of valuable alkanes. The selectivity to these undesirable by-products increases at higher conversions using conventional catalysts. Specifically, aromatic species also attenuate the activity of solid acid alkylation catalysts. Due to the inherently lower alkylation selectivity to monoalkylated benzene of solid bed alkylation (SBA) relative to hydrofluoric acid (HF) alkylation, HF alkylation producers are not attracted to switching to SBA because they must add and accept lower alkane utilization in additional processing steps. More efficient methods have now been developed in which reduced amounts of aromatic compounds are produced in the dehydrogenation step. Summary of the Invention

[0007] The present invention provides a process for producing monoalkylbenzene, the process comprising passing a C9-C14 paraffin stream and a paraffin recycle stream derived hereinafter to a dehydrogenation zone maintained under dehydrogenation conditions, the dehydrogenation conditions comprising a dehydrogenation catalyst to produce an effluent stream comprising light hydrocarbons, hydrogen, feed paraffins, corresponding monoolefins and diolefins produced relative to the feed paraffins and C9-C14 alkyl aromatics. The effluent stream comprises less than 1.4 wt% C9-C14 alkyl aromatics, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition, the layered catalyst composition comprising a core, an outer layer bonded to the core, the outer layer comprising δ and / or θ alumina and having a layer thickness of less than 100 microns, on which at least one platinum group metal and at least one promoter metal are uniformly dispersed; sending at least a portion of the effluent stream to a selective hydrogenation reactor to convert diolefins to monoolefins and produce a treated effluent stream; sending the treated effluent stream to an alkylation zone; sending an aromatic stream comprising benzene to the alkylation zone operating under alkylation conditions to produce a process stream comprising paraffins, benzene, monoalkylbenzene and heavy alkylbenzene (HAB); separating the process stream in a first separation unit into a first stream comprising benzene and a second stream comprising alkylbenzene and paraffins; passing the second stream to a second separation unit to produce a third stream comprising paraffins and a fourth stream comprising alkylbenzene, the third stream being recycled back to the dehydrogenation zone; passing the fourth stream to a third separation unit to produce a fifth stream comprising monoalkylated benzene, a sixth stream comprising heavy alkylbenzene (HAB); passing the sixth stream to a fourth separation unit to produce a seventh stream comprising low molecular weight HAB and an eighth stream comprising high molecular weight HAB; and optionally passing the eighth stream to a transalkylation zone; passing the benzene stream to a transalkylation zone operating under transalkylation conditions to produce a transalkylation effluent stream comprising monoalkylbenzene; and passing the transalkylation effluent stream to the first separation unit.

[0008] In another embodiment, a process for producing monoalkylbenzene is provided, the process comprising passing a C9-C14 paraffin stream and a paraffin recycle stream derived hereinafter to a dehydrogenation zone maintained under dehydrogenation conditions in the presence of a dehydrogenation catalyst to produce an effluent stream comprising light hydrocarbons, hydrogen, feed paraffins, the corresponding monoolefins and diolefins relative to the feed paraffins and C9-C14 alkyl aromatics, wherein the effluent stream comprises less than 1.4 wt% of the C9-C14 alkyl aromatics, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition comprising a core, an outer layer bonded to the core, the outer layer comprising delta and / or theta alumina and having a layer thickness of less than 100 microns, on which at least one platinum group metal and at least one promoter metal are uniformly dispersed; sending at least a portion of the effluent stream to a selective hydrogenation reactor to convert diolefins to monoolefins and produce a treated effluent stream; sending the treated effluent stream to an aromatic separation zone to remove at least a portion of the C9-C14 alkyl aromatics and produce a treated effluent stream; sending the treated effluent stream to an alkylation zone; sending an aromatic stream comprising benzene to the alkylation zone operating under alkylation conditions to produce a process stream comprising paraffins, benzene, monoalkylbenzene and heavy alkylbenzene (HAB); separating the process stream in a first separation unit into a first stream comprising benzene and a second stream comprising alkylbenzene and paraffins; passing the second stream to a second separation unit to produce a third stream comprising paraffins and a fourth stream comprising alkylbenzene, the third stream being recycled back to the dehydrogenation zone; passing the fourth stream to a third separation unit to produce a fifth stream comprising monoalkylated benzene and a sixth stream comprising heavy alkylbenzene (HAB); passing the sixth stream to a fourth separation unit to produce a seventh stream comprising low molecular weight HAB and an eighth stream comprising high molecular weight HAB; and optionally passing the eighth stream to a transalkylation zone; passing the benzene stream to the transalkylation zone operating under transalkylation conditions to produce a transalkylation effluent stream comprising monoalkylbenzene; and passing the transalkylation effluent stream to the first separation unit. BRIEF DESCRIPTION OF THE DRAWINGS

[0009] Figure 1 Shows a flow scheme for producing linear alkylbenzene in the absence of an adsorbent unit to remove aromatic compounds, but with units to remove heavy alkylated benzene and a transalkylation unit to improve yield.

[0010] Figure 2 Shows a flow scheme for producing linear alkylbenzene with an adsorbent unit to remove aromatic compounds and a transalkylation unit.

[0011] Figure 3 Shows a flow scheme for producing linear alkylbenzene in the absence of an adsorbent unit to remove aromatic compounds.

[0012] Figure 4 shows the X-ray diffraction pattern of the alumina outer layer in the catalyst. Detailed implementation

[0013] Figure 1 shows a flow scheme that does not include an adsorbent unit to remove aromatic compounds. Since the generation of aromatics in the present invention is reduced, this step of the prior art method is allowed to be eliminated. A new heavy alkylbenzene distillation column and an alkyl transfer unit are added to increase the production of monoalkylbenzene. An integrated method for producing LAB using a solid alkylation catalyst has been developed. An example of the integrated method is shown in Figure 1 which shows an integrated method that includes a dehydrogenation process, followed by a selective catalytic hydrogenation process, and an alkylation process. The alkane feed 105 is mixed with hydrogen 110 and sent through a heat exchanger 115 and a feed heater 120. The heated stream 125 is sent to the dehydrogenation zone 130. The dehydrogenation effluent 135 exchanges heat with the feed 105 and hydrogen 110 in the heat exchanger 115. Then the dehydrogenation effluent 135 is sent to a separator 140 and separated into a hydrogen stream 145 and a liquid stream 150. The liquid stream 150 is mixed with hydrogen 155 and sent to a selective hydrogenation zone 160, where diolefins are hydrogenated. The effluent 165 from the selective hydrogenation reactor 160 is sent to a stripper 170, where light ends 175 are removed. The bottom stream 180 from the stripper 170 is sent to an alkylation zone 185, where it is mixed with a benzene stream 190. The effluent 195 from the alkylation zone 185 is sent to a benzene distillation column 200. The benzene top stream 205 can be mixed with fresh benzene 210 to form the benzene stream 190. The bottom stream 215 from the benzene column 200 is sent to an alkane distillation column 220. The alkane top stream 225 is mixed with the alkane feed 105 and sent to the dehydrogenation zone 130. The bottom stream 230 from the alkane column 220 is sent to a heavy alkylbenzene distillation column 235, where it is separated into a top stream 240 containing monoalkylbenzene and a bottom stream 245 containing heavy alkylbenzene (e.g., dialkylbenzene). If necessary, the top stream 240 can be further processed in, for example, a finishing column (not shown).

[0014] The bottom stream of the heavy alkylbenzene can be further processed, for example, as shown in a heavy alkylated benzene distillation column 250, to produce a top stream 255 of low molecular weight heavy alkylbenzene and a bottom stream 260 of high molecular weight heavy alkylbenzene mixed with benzene. The high molecular weight heavy alkylbenzene is sent to an alkyl transfer unit 270 to produce an alkyl transfer effluent stream containing benzene, monoalkylbenzene, and unreacted high molecular weight heavy alkylbenzene. The alkyl transfer effluent stream 272 is sent to the benzene distillation column 200 to recover monoalkylbenzene.

[0015] Figure 2 shows a similar toFigure 1 The flow scheme is different in that it includes an adsorbent unit 182 for removing aromatic compounds 183, and the treated stream is sent to the alkylation zone 185. In this particular flow scheme, there is no heavy alkylated benzene distillation column. Except for these exceptions, its function is the same as that of Figure 1 the same.

[0016] Figure 3 It is also similar to Figure 1 and Figure 2 , except that in this flow scheme, there is no adsorbent unit to remove aromatics, no heavy alkylated benzene distillation column, and no transalkylation unit.

[0017] The aliphatic feedstock used in the alkylation process of the present invention contains aliphatic monoolefins having 8 to 20, or 8 to 18, or 8 to 17 carbon atoms per molecule. The feed is usually limited to a range of 4 to 6 carbon numbers at any given time. The aliphatic olefins are usually a mixture of olefins having different molecular weights. The olefins can be α-olefins or a mixture including olefin isomers. In most cases, the location of the olefinic bond in the molecule is not critical because it has been found that most solid alkylation catalysts promote the migration of the olefinic bond.

[0018] For commercial processes, other components can be present in the aliphatic feedstock together with the olefin-containing aliphatic compound. These other components can include, for example, alkanes having 9 to 17 carbon atoms per molecule, which can act as heat sinks to maintain the desired temperature in the alkylation reaction zone, as disclosed in US 9,174,891B2. However, such an amount of alkanes is not critical for the process of the present invention, and an aliphatic feedstock having no alkanes present can be used. If no alkanes are present, then another component that can act as a heat sink and remain unreacted under the processing conditions is required, if needed for a particular application, to maintain LAB linearity and 2-phenyl content.

[0019] Hydrocarbons that can be dehydrogenated include hydrocarbons having 2 to 30 or more carbon atoms, including normal alkanes, isoparaffins, alkyl aromatics, cycloalkanes, and olefins. A preferred group of hydrocarbons is the group of normal alkanes having 2 to 30 carbon atoms. Particularly preferred normal alkanes are those having 9 to 16 carbon atoms. Other particularly preferred alkanes are monomethyl alkanes and dimethyl alkanes having 9 to 16 carbon atoms. Each of the foregoing hydrocarbons can be present alone or in the form of a mixture with one or more of any other of the foregoing hydrocarbons.

[0020] The dehydrogenation conditions include a temperature of 400 °C to 900 °C, a pressure of 1 kPa to 1013 kPa, and 0.1 hr -1 to 100 hr -1Liquid hourly space velocity (LHSV). As used herein, the abbreviation "LHSV" means liquid hourly space velocity, which is defined as the volumetric flow rate of the liquid per hour divided by the catalyst volume, where the liquid volume and the catalyst volume are in the same volume units. Generally, for paraffins, the lower the molecular weight, the higher the temperature required to achieve a comparable conversion. The pressure in the dehydrogenation zone is maintained at the lowest feasible level consistent with equipment limitations to maximize the chemical equilibrium advantage.

[0021] The effluent stream from the dehydrogenation zone will typically contain unreacted dehydrogenatable hydrocarbons, hydrogen, and the products of the dehydrogenation reaction. These products include the desired olefins and undesired light ends, aromatics, etc. This effluent stream is typically cooled and passed to a hydrogen separation zone to separate the hydrogen-rich gas phase from the hydrocarbon-rich liquid phase. Generally, the hydrocarbon-rich liquid phase is further separated by a suitable selective adsorbent, selective solvent, one or more selective reactions, or by a suitable fractionation scheme. The unreacted dehydrogenatable hydrocarbons are recovered and recycled to the dehydrogenation zone. The products of the dehydrogenation reaction are recovered as the final product or as an intermediate for the preparation of other compounds. It is necessary to minimize the formation (or yield) of aromatics in the dehydrogenation zone such that the aromatic concentration does not accumulate in the recycle stream and does not increase the size (and cost) of the selective adsorbents and solvents to remove it.

[0022] A catalyst that has been found to be particularly effective in producing reduced amounts of unwanted aromatic compounds is a catalyst composite having a layered composition that includes a core and an outer layer bonded to the core, the outer layer comprising one or more transition aluminas having at least two diffraction angle peaks between 32.0° and 70.0° 2θ, where the first diffraction angle peak in this range is at 32.7 ± 0.4° 2θ, the second diffraction angle peak is at 50.8 ± 0.4° 2θ and has a thickness of less than 100 microns, and on which at least one platinum group metal and at least one promoter metal are uniformly dispersed, and having a certain concentration of at least one platinum group metal, the platinum group metal having a concentration of 0.00006 to 0.0005 grams per square meter of outer layer surface area in elemental form, the layered composition further having at least one modifier metal dispersed thereon, the core and the outer refractory inorganic oxide being different materials.

[0023] The novel layered catalyst in the present invention provides a lower formation of aromatics in the dehydrogenation zone. The combination of a platinum group concentration of 0.00006 to 0.0005 per square meter of the outer surface area and a composition including δ and / or θ alumina in the outer layer results in the formation of fewer aromatics. Although, as described in US6,756,515, the catalyst activity is still maintained by loading a large amount of the active metal platinum per cubic centimeter of the catalyst or per kilogram of the outer layer, a lower formation of aromatics can be achieved by having an outer layer of δ and / or θ alumina with an average pore size greater than the average pore size of the γ alumina layer. Aromatic products are formed by the successive dehydrogenation of the desired olefins. The large alumina pores allow the olefins to diffuse out more quickly and thus do not undergo the undesired dehydrogenation to aromatics.

[0024] The dehydrogenatable hydrocarbon can be mixed with a diluent material before, simultaneously with, or after flowing into the dehydrogenation zone. The diluent material can be hydrogen, steam, methane, ethane, carbon dioxide, nitrogen, argon, etc. or a mixture thereof. Hydrogen is the preferred diluent. Generally, when hydrogen is used as the diluent, its amount should be sufficient to ensure that the molar ratio of hydrogen to hydrocarbon is from 0.1:1 to 40:1, and the best results can be obtained when the molar ratio ranges from 1:1 to 10:1. The diluent hydrogen stream passed into the dehydrogenation zone will typically be the recycled hydrogen separated from the effluent from the dehydrogenation zone in the hydrogen separation zone.

[0025] For example, water or a substance that decomposes under dehydrogenation conditions to form water, such as an alcohol, aldehyde, ether, or ketone, can be added to the dehydrogenation zone continuously or intermittently in an amount that provides a hydrocarbon feed stream of less than 10000 weight ppm based on equivalent water calculation, preferably less than 5000 weight ppm, more preferably less than 3000 weight ppm, and possibly even less than 1000 weight ppm. The method of the present invention can be operated without adding water or a substance that decomposes to form water to the dehydrogenation zone.

[0026] In some embodiments, a multi-bed alkylation reaction zone and a separation feed stream for controlling the 2-phenyl content in the alkylbenzene product stream are employed. This arrangement structure is described in U.S. Patent No. 8,389,787. The separation bed design of U.S. Patent No. 8,389,787 is optimized for raw material utilization and energy consumption.

[0027] In one embodiment of the present invention, an aromatic compound and an olefin are reacted under alkylation conditions in the presence of a solid alkylation catalyst. These alkylation conditions generally include a temperature in the range between 80 °C and 200 °C, as described above. Typically, as the catalyst ages, the temperature of the alkylation is increased to maintain the desired activity. The alkylation is an exothermic reaction and thus in a substantially adiabatic reactor, the effluent is at a higher temperature than the feed. A substantially adiabatic reactor is a reactor in which the increase in the temperature of the effluent over the feed accounts for at least 75% of the heat generated by the reaction in the reaction zone.

[0028] Typically, the temperature within the reaction zone is maintained within a suitable range by providing a large excess of the aromatic compound to the reaction zone to absorb heat. In the case where the aliphatic feedstock contains paraffins, the paraffins are also used to absorb heat from the exothermic reaction. The high exothermic temperature during alkylation can have a negative impact not only on catalyst deactivation but also on the loss of linearity of LAB due to an increase in the isomerization of olefins to non-linear olefins, which results in product quality deterioration.

[0029] Since alkylation is generally carried out in the presence of a liquid phase and preferably under all-liquid or supercritical conditions, the pressure must be sufficient to maintain the reactants in the liquid phase. The necessary pressure inevitably depends on the olefin and the temperature but is generally in the range of 1.300 MPa(g) to 7.000 MPa(g).

[0030] In some embodiments, the alkylation of benzene with an olefin is carried out in a continuous manner using three or more catalyst beds in series in flow. For the purposes of this document, a catalyst bed is referred to as a reaction zone, whether in the same or a separate vessel as another bed. Each reaction zone has an inlet region and an outlet region. The reactants can be made into a mixture before entering the inlet region of the reaction zone or can be introduced separately and mixed in the reaction zone.

[0031] The catalyst can be used as a fixed bed, a moving bed or a slurry bed. The feed to the reaction zone can be passed upflow or downflow, or even horizontally as in a radial bed reactor; however, the flow of the aromatic compound and the olefin is co-current. In a desired variant, the olefin can be fed to several discrete points within the reaction zone. The feed mixture to the reaction zone, i.e., the aromatic compound and the aliphatic feedstock, is generally provided at a total liquid hourly space velocity (total LHSV) between 0.3 and 6 or 10 h -1 and most commonly between 0.4 and 6 h -1 depending, for example, on the alkylation temperature and the activity of the catalyst. The total LHSV is determined by the LHSV of each bed. The reciprocal of the total LHSV is the sum of the reciprocals of the LHSV of each bed in series.

[0032] It is generally desirable to use a sufficient residence time in the reaction zone such that at least 90, or at least 95, or at least 98, and generally at least 99.5 weight percent of the olefins fed to the reaction zone react in the reaction zone.

[0033] Any suitable solid alkylation catalyst can be used in the present invention, provided that the requirements of conversion, selectivity and activity are met. Generally, the catalyst is acidic. Preferred alkylation catalysts include zeolites having a zeolite framework type selected from the group consisting of: FAU, MOR, MTW and NES. Suitable zeolites include mordenite, ZSM-4, ZSM-12, ZSM-20, ZSM-38, perovskite, beta-spodumene, NU-87, UZM-8, MCM-22, MCM-36, MCM-49, zeolite Y, zeolite X and gottardiite. MOR, MWW, FAU, NES and other zeolite framework types are described in Ch. Baerlocher, W.M. Meier and D.H. Olson, "Atlas of Zeolite Framework Types", 5th Edition, Elsevier: Amsterdam, 2001, which is incorporated herein by reference. Another class of acidic solid catalysts is acidified refractory oxides such as chlorinated, fluorinated or sulfated alumina, gallium oxide, boron oxide, molybdenum oxide, ytterbium oxide, titanium oxide, chromium oxide, silicon oxide, zirconium oxide, etc. and combinations thereof. Clays and amorphous catalysts can also be used for practical applications. Further discussion of alkylation catalysts can be found in U.S. Patent Nos. 5,196,574; 6,315,964 B1 and 6,617,481 B1.

[0034] In this method, newer alkylation catalysts can also be used. For example, one such catalyst comprises a mixture of two types of zeolite materials, where the zeolites are mixed and produced to have two zeolites within a single catalyst pellet. With the novel catalyst, the first zeolite is also characterized by its acidity, where the acidity is characterized by having less than 70% of NH3 desorbing from the zeolite at a temperature greater than 400 °C. The NH3-TPD experimental procedure includes: calibrating the NH3-TPD system by injecting 0.2 cc of NH3 pulses five times every 2 minutes into a 40 cc / min UHP-grade helium stream. The data collected from the thermal conductivity detector is integrated and used to calibrate the detector's response to a known amount of NH3. The equilibrated sample is weighed at 250 mg and placed in the reactor. The sample is pretreated in a 20% O2 / He UHP-grade stream at a rate of 100 cc / min and ramped up in temperature at 10 °C / min to a maximum temperature of 650 °C. The sample is held at this temperature for one hour, then purged with UHP-grade helium for 15 minutes and cooled to the saturation temperature. The pretreatment is used to remove water and residual contaminants. The sample is saturated with anhydrous NH3 using multiple NH3 pulse injections at 150 °C into He flowing at 40 cc / min. The minimum amount of NH3 used to saturate the sample is 50 cc. The excess ammonia is purged from the sample in a flowing (40 cc / min) UHP-grade helium stream for 8 hours. NH3 desorbs from the sample in a UHP-grade helium stream (40 cc / min), where the temperature ramp is 10 °C / min to a final temperature of 605 °C. All gases have been purified using appropriate gas purifiers. The desorbed NH3 is detected using a thermal conductivity detector. Using the detector response obtained at the start of the experiment, the detector response is converted to the number of moles of NH3. The integrated results are reported by integrating the temperature range of interest and reported as millimoles NH3 / g sample. An example of the first zeolite is UZM-8.

[0035] The second zeolite has a silica to alumina molar ratio of less than 8 and includes rare earth elements incorporated into the zeolite framework in an amount greater than 16.5 wt%. The amount of the first zeolite component is between 10 wt% and 90 wt% of the catalyst, and the amount of the second zeolite component is between 10 wt% and 90 wt%. The zeolites are intertwined into individual catalyst particles. Examples of the second zeolite are rare earth substituted X zeolite, Y zeolite or zeolites having an EMT / FAU co-crystallite. The rare earth exchanged ions are incorporated into the zeolite at a low ratio which reduces the acidity due to the increase in the amount of framework alumina at low ratios and also reduces the geometric space in the supercages. The reduced acidity and reduced space significantly suppress the isomerization and cracking pathways while leaving the primary alkylation reaction unaffected. This reduces the unwanted side reactions that reduce the amount and quality of the LAB product. This is contrary to what is expected because it has been found that incorporating or leaving some alkali metal or alkaline earth metal cations in the catalyst significantly improves the catalyst performance. This is especially true for the performance around the linearity of the alkylbenzene and maintaining the linearity as the operating temperature increases. Typically, the alkali metal or alkaline earth metal cations are removed because without rare earth exchange, the alkali metal or alkaline earth metal cations are detrimental to the catalyst life and regenerability.

[0036] The alkylation reaction zone can contain at least 2 or at least 3, most commonly between 3 and 10 reaction zones in series, into which a portion of the aliphatic feedstock is fed. There is usually a finishing alkylation reaction zone after the series to react the residual olefins in the effluent from the last reaction zone in the series. The reaction zones can be in a common vessel or separate vessels. The reaction zones can be of the same or different sizes. Additional reaction zones can be used in parallel.

[0037] In a common commercial configuration of alkylbenzene, the finishing assembly includes a distillation assembly that recovers substantially all of the benzene from the alkylation effluent and provides a relatively pure benzene stream as the overhead distillate. The bottoms stream from this distillation assembly is then passed to a distillation assembly to be separated into an overhead distillate, paraffins and unreacted olefins, and the bottoms from this second distillation assembly will be fed to a heavy distillation assembly where the alkylbenzene product is contained in the overhead distillate. If desired, a finishing column can be used to further purify the alkylbenzene, especially after clay treatment to remove color bodies.

[0038] Although at least one exemplary embodiment has been presented in the foregoing detailed description of the present invention, it should be understood that a vast number of variations exist. It should also be understood that one exemplary embodiment or a plurality of exemplary embodiments are merely examples and are not intended to limit in any way the scope, applicability, or configuration of the present invention. On the contrary, the foregoing detailed description will provide those skilled in the art with a convenient roadmap for implementing an exemplary embodiment of the present invention. It should be understood that various changes can be made to the functions and arrangements of the elements described in the exemplary embodiments without departing from the scope of the present invention as set forth in the appended claims.

[0039] The following examples are provided in illustration of the present invention and are not intended to unduly limit the generally broad scope of the invention as shown in the appended claims.

[0040] Example 1

[0041] Aluminum oxide spheres are prepared by the well-known oil-drop method, which is described in U.S. Patent No. 2,620,314. This method involves forming an aluminum hydrosol by dissolving aluminum in hydrochloric acid. Hexamethylenetetramine is added to the sol to gel the sol into spheres when the droplets are dispersed into an oil bath maintained at 93°C. The droplets remain in the oil bath until they solidify and form hydrogel spheres. After removing the spheres from the hot oil, they are pressure-aged at 135°C, washed with dilute ammonium hydroxide solution, dried at 110°C, and calcined at 650°C for 2 hours to obtain γ-aluminum oxide spheres. The calcined aluminum oxide is then extruded into a fine powder with a particle size of less than 200 microns.

[0042] Next, a slurry is prepared by mixing pseudoboehmite and deionized water and stirred to evenly distribute the tin component. The above-prepared aluminum oxide powder and a 50% aqueous solution of tin(IV) chloride are added to this mixture, and the slurry is ball-milled for 240 minutes to reduce the maximum particle size to less than 50 microns. The slurry is sprayed onto a cordierite core with an average diameter of 1.6 mm using a granulation and coating device to obtain an outer layer of 25 or 65 microns. At the end of the process, some slurry remains that has not coated the core. This layered spherical carrier is calcined at 600°C to 900°C in order to convert the pseudoboehmite and γ-aluminum oxide in the outer layer to δ-aluminum oxide and convert the tin(IV) chloride to tin oxide.

[0043] Based on the weight of the support, the calcined layered support was impregnated with lithium and platinum using a rotary impregnator by contacting the support with an aqueous solution containing lithium chloride and chloroplatinic acid (1:1 solution: support volume ratio). The impregnated composite was heated using a rotary impregnator until no solution remained, dried at 315 °C, calcined at 540 °C, and reduced in hydrogen at 500 °C. The resulting catalysts prepared in this example contained 0.1 wt% - 0.2 wt% of platinum, 0.1 wt% - 0.2 wt% of tin, and 0.1 wt% - 0.2 wt% of lithium relative to the entire catalyst. These catalysts were identified as catalysts A, B, C, and D. The properties of catalysts A, B, C, and D are summarized in Table 1.

[0044] Table 1

[0045]

[0046] Example 2

[0047] The dehydrogenation activity of catalysts A, B, C, and D from Example 1 was tested in a laboratory-scale plant. In a 1.27 cm reactor, 5 cc of catalyst was placed, and a hydrocarbon feedstock composed of 8.8 wt% - 9.3 wt% n-C 10 、40.0 wt% - 41.8 wt% n-C 11 、38.6 wt% n-C 12 、8.6 wt% - 10.8 wt% n-C 13 、0.3 wt% - 0.8 wt% n-C 14 and 1 wt% - 1.4 wt% non-normal substances flowed through the catalyst at a pressure of 138 kPa (or 20 psig), a hydrocarbon molar ratio of 4:1, and a liquid hourly space velocity (LHSV) of 28 h -1 . By adjusting the reactor temperature, the total normal olefin concentration (%TNO) in the product was maintained at 10 wt%.

[0048] Hydrogen and hydrocarbon feeds were combined upstream of the reactor to form a combined feed, and the combined feed was vaporized before entering the reactor. In this example, the catalysts were tested at a water concentration of 2000 weight-ppm based on the weight of the hydrocarbons in the combined feed. The results of the product liquids collected after 48 hours of production for the four listed catalysts are presented in Table 2.

[0049] The aromatic content of the feed and product streams was analyzed using backflushing on a high-performance liquid chromatography (HPLC) system using n-hexane solvent as the mobile phase and a refractive index detector.

[0050] n - paraffin conversion % = (n - paraffin in feed - n - paraffin in product) x 100% / n - paraffin in feed. Aromatic selectivity % = (aromatic in product - aromatic in feed) x 100% / n - paraffin conversion.

[0051] Even though catalyst B shows a lower aromatic selectivity than catalyst A, its aromatic selectivity is still higher than that of catalysts C and D. Thus, the combination of one or more transition alumina phases (δ and / or θ alumina) in the layer with a high Pt density (grams of Pt in the alumina layer / layer surface area (g / m 2 )) results in the lowest aromatic formation.

[0052] Table 2

[0053] Sample ID Aromatic selectivity (%) Catalyst A 2.18 Catalyst B 1.80 Catalyst C 1.60 Catalyst D 1.62

[0054] Example 3

[0055] As Figure 4 shown, the alumina layers of the four catalysts A, B, C, and D of the present disclosure were analyzed by X - ray diffraction to study the presence of δ or θ alumina. An attempt was made to remove only the outer layer material by placing 1 cc of calcined caustic in a small grinding container without a grinding medium. Depending on the rate of powder generation from the outer layer, the samples were ground for different amounts of time. The powders from the four catalytic composites were labeled as samples A, B, C, and D.

[0056] The X - ray patterns of samples A, B, C, and D were obtained using standard X - ray powder diffraction techniques. The irradiation source was a high - intensity x - ray tube operating at 40 kV and 44 mA. The diffraction patterns from the copper K - α irradiation were obtained by a suitable computer - based technique. A flat, compressed powder sample was scanned continuously from 8° to 90° 2θ. The interplanar spacing (d) was obtained from the position of the diffraction peaks represented as 2θ, with the unit of angstroms, where 2θ is the Bragg angle as observed from the digitized data. As will be understood by those skilled in the art, the determination of the diffraction angle (2θ) is subject to both human and mechanical errors, and the combination of these errors can impart an uncertainty of ±0.4° 2θ to each reported 2θ value.

[0057] The outer layers of catalysts C and D essentially consist of δ and / or θ alumina. Their X - ray diffraction patterns contain at least two diffraction angle peaks between 32.0° and 70.0° 2θ, where the first diffraction angle peak within the range is at 32.7 ± 0.4° 2θ and the second diffraction angle peak is at 50.8 ± 0.4° 2θ. Additionally, the X - ray diffraction pattern has at least 2 peaks and / or shoulders between 43 ± 0.4° and 49 ± 0.4° 2θ.

[0058] Example 4

[0059] In Figure 1 、 Figure 2 and Figure 3 's flow process scenarios, the fresh feed / LAB ratio = the flow rate of stream 105 / the flow rate of stream 240. For a given LAB production rate, a lower ratio is preferred due to lower feed consumption. Reported in Table 3 are the Figure 1 、 Figure 2 and Figure 3 calculated fresh feed / LAB ratios for the selected dehydrogenation zone and alkylation catalysts. The molecular weight (MW) of the paraffin feed used in these calculations ranges from 155 g / mol to 165 g / mol.

[0060] The novel layered catalyst in the present invention provides lower aromatic formation in the dehydrogenation zone. A portion of these undesired aromatic compounds in the dehydrogenation zone effluent reacts in the alkylation unit to form low molecular weight heavy alkylated benzenes, thereby further increasing the fresh feed / LAB ratio and thus increasing the cash cost incurred. The combination of the layered catalyst and Figure 1 the flow scenario has approximately the same feed efficiency as the current prior art using HF, while eliminating the safety risks associated with HF and the need for additional processing steps to remove undesired aromatics prior to alkylation with a solid acid catalyst.

[0061] Table 3

[0062] Flow scheme Dehydrogenation zone catalyst Alkylation catalyst Fresh feed / LAB Figure 1 Catalyst A SBA 0.759 Figure 1 Catalyst C SBA 0.753 Figure 2 Catalyst A SBA 0.744 Figure 2 Catalyst C SBA 0.739 Figure 3 Catalyst A HF 0.755 Figure 3 Catalyst C HF 0.749

[0063] Specific implementation scheme

[0064] While the following is described in conjunction with specific embodiments, it should be understood that the description is intended to illustrate and not limit the scope of the foregoing description and the appended claims.

[0065] A first embodiment of the present invention is a process for producing monoalkylbenzenes, the process comprising passing a C9-C14 paraffin stream and a paraffin recycle stream derived hereinafter to a dehydrogenation zone maintained under dehydrogenation conditions in the presence of a dehydrogenation catalyst to produce an effluent stream comprising light hydrocarbons, hydrogen, feed paraffins, corresponding monoolefins and diolefins relative to the feed paraffins and C9-C14 alkyl aromatics, wherein the effluent stream comprises less than 1.4 wt% C9-C14 alkyl aromatics, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition comprising a core and an outer layer bonded to the core, the outer layer comprising an external refractory inorganic oxide having a layer thickness of less than 100 microns, on which at least one platinum group metal and at least one promoter metal are uniformly dispersed; sending at least a portion of the effluent stream to a selective hydrogenation reactor to convert diolefins to monoolefins and produce a treated effluent stream; sending the treated effluent stream to an alkylation zone; sending an aromatic stream comprising benzene to the alkylation zone operating under alkylation conditions to produce a process stream comprising paraffins, benzene, monoalkylbenzenes and heavy alkylbenzenes (HAB); separating the process stream in a first separation unit into a first stream comprising benzene and a second stream comprising alkylbenzenes and paraffins; passing the second stream to a second separation unit to produce a third stream comprising paraffins and a fourth stream comprising alkylbenzenes, the third stream being recycled back to the dehydrogenation zone; and passing the fourth stream to a third separation unit to produce a fifth stream comprising monoalkylated benzene and a sixth stream comprising heavy alkylbenzenes (HAB). One embodiment of the present invention is one, any or all of the first embodiment to the previous embodiments of this paragraph, wherein the mass ratio of the C9-C14 paraffin stream to the mass of the fifth stream is less than 0.75. One embodiment of the present invention is one, any or all of the first embodiment to the previous embodiments of this paragraph, wherein the mass ratio of the C9-C14 paraffin stream to the mass of the fifth stream is less than 0.74. One embodiment of the present invention is one, any or all of the first embodiment to the previous embodiments of this paragraph, wherein the mass ratio of the C9-C14 paraffin stream to the mass of the fifth stream is less than 0.73. One embodiment of the present invention is one, any or all of the first embodiment to the previous embodiments of this paragraph, wherein the alkylation zone comprises a catalyst selected from mordenite, ZSM-4, ZSM-12, ZSM-20, ZSM-38, perovskite, beta-magnesite, NU-87, UZM-8, MCM-22, MCM-36, MCM-49, zeolite Y, zeolite X, gottardiite, MOR, MWW, FAU, RE-Y, NES, fluorinated-ASA or combinations thereof.One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, wherein the alkylation zone comprises a hydrofluoric acid catalyst. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition having an outer layer comprising one or more transition aluminas, the transition alumina having at least two diffraction angle peaks between 32.0° and 70.0° 2θ, wherein the first diffraction angle peak within the range is at 32.7 ± 0.4° 2θ and the second diffraction angle peak is at 50.8 ± 0.4° 2θ. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition having an outer layer comprising one or more transition aluminas, the transition alumina having at least 2 diffraction angle peaks and / or a shoulder between 43 ± 0.4° to 49 ± 0.4° 2θ. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition having a certain concentration of at least one platinum group metal, the concentration of the platinum group metal being 0.00006 to 0.0005 grams per square meter of outer surface area in terms of the element. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, wherein the dehydrogenation zone effluent stream comprises less than 1.2 wt% of C9-C14 alkyl aromatics. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, wherein the dehydrogenation zone effluent stream comprises less than 1.0 wt% of C9-C14 alkyl aromatics. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, further comprising transferring a fifth stream to a sulfonation unit to convert monoalkylated benzene to monoalkylated benzene sulfonate. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, further comprising sending the treated effluent stream to an aromatic separation zone to remove at least a portion of the C9-C14 alkyl aromatics, and then sending the remaining portion of the treated effluent stream to the alkylation zone. One embodiment of the present invention is one, any, or all of the first embodiment to the previous embodiments of this paragraph, further comprising transferring a sixth stream to a fourth separation unit to produce a seventh stream comprising low molecular weight HAB and an eighth stream comprising high molecular weight HAB.Embodiments of the present invention are one, any, or all of the existing embodiments in this paragraph to the first embodiment in this paragraph, and further include delivering an eighth stream to the transalkylation zone; delivering a benzene stream to a transalkylation zone operating under transalkylation conditions to produce a transalkylation effluent stream comprising monoalkylbenzene; and delivering the transalkylation effluent stream to a first separation unit.

[0066] Although no further detailed description is provided, it is believed that those skilled in the art can make the most of the present invention by using the foregoing description and can easily determine the basic features of the present invention without departing from the essence and scope of the present invention to make various changes and modifications and adapt it to various uses and conditions. Therefore, the foregoing preferred specific embodiments should be understood as merely illustrative and not limiting the remainder of the disclosure in any way, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0067] In the foregoing, all temperatures are shown in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

Claims

1. A method for producing monoalkylbenzene, comprising: a. passing a C9-C14 paraffin stream and a paraffin recycle stream derived hereinafter to a dehydrogenation zone maintained under dehydrogenation conditions in the presence of a dehydrogenation catalyst to produce an effluent stream comprising light hydrocarbons, hydrogen, feed paraffins, corresponding monoolefins and diolefins relative to the feed paraffins and C9-C14 alkyl aromatics, wherein the effluent stream comprises less than 1.4 wt% of the C9-C14 alkyl aromatics, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition comprising a core and an outer layer bonded to the core, the outer layer comprising delta and / or theta alumina having a layer thickness of less than 100 microns, on which at least one platinum group metal and at least one promoter metal are uniformly dispersed, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition having a certain concentration of at least one platinum group metal, the concentration of the platinum group metal being from 0.00006 to 0.0005 grams per square meter of the outer layer surface area in elemental form; b. sending at least a portion of the effluent stream to a selective hydrogenation reactor to convert the diolefins to monoolefins and produce a treated effluent stream; c. sending the treated effluent stream to an alkylation zone; d. sending an aromatic stream comprising benzene to the alkylation zone operating under alkylation conditions to produce a process stream comprising paraffins, benzene, monoalkylbenzene and heavy alkylbenzene HAB; e. separating the process stream in a first separation unit into a first stream comprising benzene and a second stream comprising alkylbenzene and paraffins; f. passing the second stream to a second separation unit to produce a third stream comprising paraffins and a fourth stream comprising alkylbenzene, the third stream being recycled back to the dehydrogenation zone; and g. passing the fourth stream to a third separation unit to produce a fifth stream comprising monoalkylated benzene and a sixth stream comprising heavy alkylbenzene HAB.

2. The method according to claim 1, wherein the mass ratio of the C9-C14 paraffin stream to the mass of the fifth stream is less than 0.

75.

3. The method according to claim 1, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition having an outer layer comprising one or more transition aluminas having at least two diffraction angle peaks between 32.0° and 70.0° 2θ, wherein the first diffraction angle peak within the range is at 32.7 ± 0.4° 2θ and the second diffraction angle peak is at 50.8 ± 0.4° 2θ.

4. The method according to claim 1, wherein the dehydrogenation zone catalyst comprises a layered catalyst composition having an outer layer comprising one or more transition aluminas having at least 2 diffraction angle peaks and / or shoulders between 43 ± 0.4° and 49 ± 0.4° 2θ.

5. The method according to claim 1, wherein the dehydrogenation zone effluent stream comprises less than 1.2 wt% of the C9-C14 alkylaromatics.

6. The method according to claim 1, further comprising passing the fifth stream to a sulfonation unit to convert the monoalkylated benzene to a monoalkylated benzene sulfonate.

7. The method according to claim 1, further comprising sending the treated effluent stream to an aromatic separation zone to remove at least a portion of the C9-C14 alkylaromatics, and then sending the remaining portion of the treated effluent stream to the alkylation zone.

8. The method according to claim 1, further comprising passing the sixth stream to a fourth separation unit to produce a seventh stream comprising low molecular weight HAB and an eighth stream comprising high molecular weight HAB.

9. The method according to claim 1, further comprising passing the eighth stream to a transalkylation zone; passing a benzene stream to the transalkylation zone operating under transalkylation conditions to produce a transalkylation effluent stream comprising monoalkylbenzene; and passing the transalkylation effluent stream to the first separation unit.

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