TWI412585B - Process for oligomerizing dilute ethylene - Google Patents

Abstract

The process and apparatus converts ethylene in a dilute ethylene stream that may be derived from an FCC product to heavier hydrocarbons. The catalyst may be an amorphous silica-alumina base with a Group VIII and/or VIB metal. The catalyst is resistant to feed impurities such as hydrogen sulfide, carbon oxides, hydrogen and ammonia. At least 40 wt-% of the ethylene in the dilute ethylene stream can be converted to heavier hydrocarbons.

Description

Method for oligomerizing ethylene oxide

The field of the invention relates to an apparatus and method for converting dilute ethylene to heavier hydrocarbons in a hydrocarbon stream. These heavier hydrocarbons can be used as an engine fuel.

The flow of exhaust gas from a fluid catalytic cracking unit (which contains all gases having a boiling point lower than ethane) is referred to as a dry gas. The exhaust gas flowing through the compression, in order to remove a maximum of C ₃ and C ₄ gas as possible. Sulfur from the exhaust stream is also mostly absorbed in a scrubber using an amine absorbent. The residual gas flow is called FCC dry gas. A typical dry gas stream comprises 5 to 50% by weight of ethylene, 10 to 20% by weight of ethane, 5 to 20% by weight of hydrogen, 5 to 20% by weight of nitrogen, 0.1 to 5.0% by weight of each carbon monoxide and carbon dioxide and less than 0.01 Weight % hydrogen sulfide and ammonia balanced with methane.

Currently, the FCC dry gas stream is sent to a burner as a fuel gas. A 7,949-kiloliter (50,000-barrel) FCC unit can burn 181,000 kg (200 tons) of dry gas containing 36,000 kg (40 tons) of ethylene per day. Because of the large price difference between fuel gas and engine fuel products or pure ethylene, attempts have been made to recover this ethylene based on its economic advantages. However, the dry gas stream contains impurities which can poison the oligo-polymerization catalyst, and the ethylene concentration is too lean, so that it is not economical to recover ethylene by the gas recovery system.

An oligomeric polymerization process for concentrating an ethylene stream to a liquid product is a known technique. However, oligomeric polymerization typically involves the use of propylene or butenes, especially from liquefied petroleum gas (LPG) or dehydrogenated feedstocks, to produce gasoline-based olefins. Ethylene is rarely used as an oligomeric feedstock oil because of its extremely low reactivity.

Diluted ethylene is required in refinery logistics.

It has been discovered that ethylene in a dilute ethylene stream, such as a FCC dry gas stream, can be catalyzed by a Group VIII and/or Group VIB metal on an amorphous meteorite-alumina catalyst to form heavier hydrocarbons. The heavier hydrocarbons can be separated and incorporated into gasoline and diesel pools. It has been found that zeolite catalysts suitable for the oligomerization of ethylene can be rapidly deactivated in the presence of impurities such as carbon oxides, ammonia and hydrogen sulfide. These impurities do not substantially affect the catalyst comprising a Group VIII and/or VIB metal on the amorphous vermiculite-alumina support. Thus, the dilute ethylene in the FCC dry gas stream can be oligomerized into a liquid fuel product that is easily separated from the unconverted gas stream. The unconverted gas can then be combusted as a fuel gas, but the higher value ethylene of the heavier hydrocarbons has been removed.

The method and apparatus utilize ethylene in the presence of a dilute stream and in the presence of impurities as a raw material for the catalytic poison.

Additional features and advantages of the invention are apparent from the description of the invention, the appended claims and claims.

The invention is applicable to any hydrocarbon stream containing ethylene and preferably a dilute portion containing ethylene. A suitable dilute ethylene stream typically comprises between 5 and 50% by weight ethylene. The FCC dry gas stream is a suitable stream of dilute ethylene. Other dilute ethylene streams, such as coked dry gas streams, can also be used in the present invention. Since the invention is particularly suitable for FCC dry gases, the subject application is described with respect to utilizing ethylene from the FCC dry gas stream.

In FIG. 1, as indicated by the numerals, FIG. 1 illustrates a refinery assembly 6 that typically includes an FCC unit section 10 and a product recovery section 90. The FCC unit section includes a reactor 12 and a catalyst regenerator 14. Processing variables typically include a cracking reaction temperature of 400 to 600 ° C and a catalyst regeneration temperature of 500 to 900 °C. Both cracking and regeneration occurred at an absolute pressure of 506 kPa (72.5 psia).

1 shows a typical FCC reactor 12 in which a heavy hydrocarbon feedstock or feedstock stream system in splitter 16 is contacted with a regenerated cracking catalyst flowing from a regenerative catalyst riser 18. This contact can occur in the narrow riser 20 and extend up to the bottom of the reactor 22. The contact of the feedstock and the catalyst is fluidized by the gas from the fluidization line 24. In one embodiment, the heat from the catalyst evaporates the hydrocarbon feedstock or oil, and then, in the presence of a catalyst, the hydrocarbon feedstock is cracked into a lighter molecular weight hydrocarbon product, both of which are transferred to the upper end of the riser 20 Reactor 22. An unavoidable side reaction occurs in the riser 20, and coke deposits remain on the catalyst, reducing catalyst activity. The cracked light hydrocarbon product is then separated from the coked cracking catalyst using a cyclone separator that may include a primary separator 26 and one or two cyclones 28 in the reactor 22. The gaseous cracked product exits reactor 22 through line 31 to line 32 for transport to downstream product recovery section 90. The spent or coked catalyst needs to be regenerated for further use. After the coked cracking catalyst is separated from the gaseous product hydrocarbons, it falls into the stripping section 34 where the gas stream is injected through the nozzle to purge the residual hydrocarbon vapor. After the stripping operation, the coked catalyst is sent to the catalyst regenerator 14 via the consumption catalyst riser 36.

Figure 1 depicts a regenerator 14 (also referred to as a combustion chamber). However, other types of regenerators are also suitable. In the catalyst regenerator 14, an oxygen-containing gas stream, such as air, is introduced through the air splitter 38 in contact with the coked catalyst. The coke is burned from the coked catalyst to provide regenerated catalyst and exhaust. The catalyst regeneration process adds a significant amount of heat to the catalyst to provide energy to counteract the endothermic cracking reaction that occurs in the reactor riser 20. The catalyst and air together flow upwardly along the burner riser 40 located in the catalyst regenerator 14 and, after regeneration, are first separated by release by the selector 42. The additional recovery of the regenerated catalyst and the off-gas exiting the classifier 42 is obtained using the first and second cyclones 44, 46 in the catalyst regenerator 14, respectively. The catalyst separated from the exhaust gas enters through the cyclone legs of the cyclones 44, 46 while the relatively light exhaust gases in the catalyst sequentially exit the cyclones 44, 46 and pass through the exhaust gas outlets 47 in the exhaust gas line 48. Leave the regenerator 14. The regenerated catalyst is transported back to the riser 20 by the regeneration catalyst riser 18. As a result of coke combustion, in line 48, the off-gas from the top of the catalyst regenerator 14 contains CO, CO ₂ , N _{2 ,} and H ₂ O in addition to a small amount of other materials. The hot exhaust gases exit the regenerator 14 through the exhaust gas outlet 47 in line 48 for further processing.

The product recovery section 90 is connected downstream to the product outlet 31. "Downstream connection" means that at least a portion of the material that can be manipulated to the downstream connected component is from a component associated therewith. "Connected" means that the stream of material can be manipulated between a number of listed components. In the product recovery section 90, the gaseous FCC product in line 32 is directed to the lower region of the FCC main fractionation column 92. The main column 92 is connected downstream to the product outlet 31. Several portions of the FCC product can be separated and taken from the main column comprising the following components: heavy oil slurry from the bottom of line 93, heavy cycle oil stream in line 94, line 95 from outlet 95a Light cycle oil and heavy naphtha stream in line 96 obtained from outlet 96a. Any or all of the lines 93-96 can be cooled and pumped back to the main column 92 to cool the main column, which is typically at a higher position. The gasoline and gaseous light hydrocarbons in the overhead line 97 from the main column 92 are removed and condensed prior to entering the main column receiver 99. The main column receiver 99 is connected downstream to the product outlet 31, which is connected upstream to the main column receiver 99. "Upstream connection" means that at least a portion of the material from the upstream connected components can be manipulated to flow to the components associated therewith.

The aqueous stream is removed from the feed hopper in the receiver 99. In addition, the condensed light naphtha stream in line 101 is removed while the overhead stream in line 102 is removed. The overhead stream in overhead line 102 comprises gaseous light hydrocarbons, which may include a dilute ethylene stream. The streams in lines 101 and 102 can enter vapor recovery section 120 of product recovery section 90.

The vapor recovery section 120 is shown as an absorption based system, but any vapor recovery system can be used, including a cryostat system. In order to obtain a sufficiently separated light gas component, the gaseous stream in line 102 is compressed into compressor 104. More than one compressor can be used, but two-stage compression is usually used. The light hydrocarbon stream compressed in line 106 is connected by a stream in lines 107 and 108, cooled and passed to high pressure receiver 110. The aqueous stream from receiver 110 can pass through main column receiver 99. The gaseous hydrocarbon stream containing a dilute ethylene stream in line 112 through the primary absorber 114, wherein the gasoline unstable contact with the main column line 101 from the receiver 99, the resulting C ₃ + and C ₂ - hydrocarbons between Separation. The primary absorber 114 is connected downstream to the main column receiver 99. Liquid in line 107 C ₃ + stream in line 106 prior to cooling return. The primary exhaust stream from line 116 of primary absorber 114 contains dilute ethylene for the purposes of the present invention. However, in order to further concentrate the ethylene stream and to recover the heavier components of line 116, it may be directed to secondary absorber 118, wherein the recycle stream of light cycle oil in line 121 from line 95 absorbs the primary waste stream. Most of the residual C ₅ + and some C ₃ -C ₄ materials. The secondary absorber 118 is connected downstream to the primary absorber 114. C ₃ + higher light cycle oil content of the material from the bottom of the absorber in line two by the circulating pump 119 is returned to line 95 of column 92 competent. Two exhaust gas stream is removed in line 122 contains primarily C ₂ - hydrocarbons of the overhead dry gas and hydrogen sulfide, ammonia, carbon oxides and hydrogen of two absorber 118 distillate, to dilute ethylene stream comprising.

The liquid from the high pressure receiver 110 in line 124 is sent to stripper 126. Most of the C ₂ - in the overhead of stripper 126 is removed and returned to line 106 via overhead line 108. The liquid bottoms stream from stripper 126 is passed via line 128 to debutanizer column 130. The overhead in line 132 from debutanizer distillate stream containing C ₃ -C ₄ olefinic product, containing both a line of stable gasoline bottoms stream 134 for further processing and delivery to the gas storage vessel.

The dilute ethylene stream of the present invention may comprise a FCC dry gas stream having between 5 and 50 weight percent ethylene and preferably 10 to 30 weight percent ethylene. Methane having a concentration between 25 and 55% by weight is usually the major component of the dilute ethylene stream, and the substantial content of ethane is typically between 5 and 45% by weight. The dilute ethylene stream may comprise between 1 and 25% by weight and usually between 5 and 20% by weight of hydrogen and nitrogen, respectively. The saturated ethylene content can also be included in the dilute ethylene stream. The use of two absorbers 118, the C ₃ + content of no more than 5 wt%, and typically comprises 0.5 wt% or less of propylene.

In addition to hydrogen, other impurities such as hydrogen sulfide, ammonia, carbon oxides and acetylene may also be present in the dilute ethylene stream.

It has been discovered that many impurities in the dry gas ethylene stream may poison the oligomeric catalyst. Hydrogen and carbon monoxide may reduce the metal sites to inactivate them. Carbon dioxide and ammonia attack the acid sites on the catalyst. Hydrogen sulfide can attack the metal on the catalyst to form metal sulfides. Ethylene can be polymerized and glued to the catalyst or equipment.

The secondary exhaust stream in line 122 containing the dilute ethylene stream can be directed to an optional amine absorber unit 140 to remove hydrogen sulfide to reduce concentration. A dilute aqueous amine solution (such as comprising monoethanolamine or diethanolamine) is introduced via line 142 into absorber 140 and contacted with a flowing secondary exhaust stream to absorb hydrogen sulfide, while a higher concentration contains hydrogen sulfide. The aqueous amine absorption solution is removed from the absorption zone 140 via line 143 and recovered and may be further processed.

The amine treated dilute ethylene stream in line 144 can be directed to a water cleaning unit 146 as desired to remove residual amines from the amine absorber 140 and reduce ammonia in the dilute ethylene stream in line 144. The concentration of carbon dioxide. Water is directed to a water scrubber in line 145. The water in line 145 is typically slightly acidified to enhance retention of basic molecules such as amines. The aqueous stream in line 147 rich in amines and possibly in ammonia and carbon dioxide is retained in water wash unit 146 and may be further processed.

The amine treated dilute ethylene and possibly water-washed stream in line 148 may then be treated in a temporary guard bed 150 to remove one or more impurities (such as carbon dioxide, hydrogen sulfide, and ammonia). Lower concentration. The guard bed 150 may contain an adsorbent to adsorb impurities (such as hydrogen sulfide) that may poison the oligomeric catalyst. The guard bed 150 can comprise a plurality of adsorbents that absorb more than one impurity. Typical absorbents for absorbing hydrogen sulfide are ADS-12, CO-absorbing ADS-106, and ammonia-absorbing UOP MOLSIV 3A (both available from UOP, LLC). The adsorbents can be mixed into separate beds or can be arranged in a continuous bed.

The dilute ethylene stream in line 151 which may be amine treated, possibly water washed, and possibly subjected to adsorption treatment to remove more hydrogen sulfide, ammonia and carbon monoxide typically has at least one of the following impurity concentrations: 0.1% by weight up to 5.0% by weight carbon monoxide. And/or 0.1% by weight up to 5.0% by weight of carbon dioxide, and/or at least 1 wppm up to 500 wppm of hydrogen sulfide and/or at least 1 to up to 500 wppm of ammonia, and/or at least 5 to up to 20% by weight of hydrogen. The presence of such impurities and their concentrations may vary with the treatment and source of the dilute ethylene stream.

Line 151 carries the dilute ethylene stream to compressor 152 for pressurization to reactor pressure. The compressor 152 is downstream connected to the main column 92, the product recovery section 90, and the product outlet 31. The compressed dilute ethylene stream can be compressed to at least 3,550 kPa (500 psia) and perhaps no more than 10,445 kPa (1500 psia) and preferably between 4,930 kPa (700 psia) and 7,687 kPa (1100 psia). The dilute ethylene stream is preferably pressurized to a critical pressure above ethylene (which is 4,992 kPa (724 psia) of pure ethylene) to avoid accelerated catalyst deactivation. Compressor 152 can include one or more devices with interstage cooling. A heater may be required to raise the compressed stream to the reaction temperature. The compressed dilute ethylene in line 154 is conveyed to oligomerization reactor 156.

The oligomerization reactor 156 is connected downstream to the compressor 152 and the primary and secondary absorbers 114 and 118, respectively. The oligomerization reactor preferably comprises a fixed catalyst bed 158. The dilute ethylene feed stream is preferably contacted with the catalyst during a downward flow operation. However, an upflow operation may be appropriate. Preferably, the catalyst has an amorphous vermiculite-alumina matrix from a metal of Groups VIII and/or VIB of the Periodic Table Notation using Chemical Abstracts Service. In one aspect, the catalyst uses a metal of the VIII and is supplemented with a Group VIB metal. In one aspect, the catalyst has a vermiculite to alumina ratio of at most 30 and preferably at most 20. Typically, since vermiculite and alumina are only in the matrix, the vermiculite to alumina ratio is the same for the catalyst and the substrate. The metal can be impregnated onto or ion exchanged with a vermiculite-alumina substrate. Co-milling is also considered. The catalyst of the present invention may have a low temperature acidity ratio of at least 0.15, preferably 0.2 and preferably greater than 0.25 as determined by an ammonia temperature programmed desorption method (Ammonia TPD) as described below. Furthermore, suitable catalysts may have a surface area between 50 and 400 m ² /g as determined by nitrogen BET.

Preferred oligomeric catalysts of the invention are described below. The preferred oligomeric catalyst comprises an amorphous vermiculite-alumina support. One of the components of the catalyst carrier used in the present invention is alumina. The alumina may be a different hydrated alumina or alumina gel (such as a-alumina monohydrate of pseudo-acre or pseudo-agar stone structure, alpha-alumina trihydrate of water lead-oxygen structure, Bayer stone) Structure of β-alumina trihydrate and the like). A particularly preferred alumina system is available from Sasol North America Alumina Product Group under the trade name Catapal. This material is a very high purity alpha-alumina monohydrate (pseudo-boehmite) which has been shown to produce high purity gamma-alumina after calcination at elevated temperatures. Another component of the catalyst carrier is an amorphous vermiculite-alumina. A suitable vermiculite-alumina having a vermiculite to alumina ratio of 2.6 was purchased from CCIC, a subsidiary of Japan JGC.

Another component used in the preparation of the catalyst in the present invention is a surfactant. The surfactant is preferably mixed with the alumina and vermiculite-alumina powders described above. The resulting mixture of surfactant, alumina and vermiculite-alumina is then shaped, dried and calcined as described below. The calcination process can effectively remove the organic component of the surfactant by combustion, but only after the surfactant faithfully performs its function in accordance with the present invention. Any suitable surfactant can be utilized in accordance with the present invention. Preferred surfactants are selected from a series of surfactants sold under the trademark "Antarox" by Solvay S.A. The "Antarox" surfactant is generally characterized by a modified linear aliphatic polyether and is a low foaming biodegradable cleaner and wetting agent.

A suitable vermiculite-alumina mixture is prepared by mixing vermiculite-alumina and alumina in a volume ratio to obtain the desired vermiculite to alumina ratio. In one embodiment, an 85 wt% vermiculite to alumina amorphous alumina-alumina and 15 wt% alumina powder having an alumina ratio of 2.6 provides a suitable support. In one embodiment, an amorphous meteorite-alumina to alumina ratio other than 85 to -15 may be suitable, as long as the final vermiculite to alumina ratio of the support is at most 30 and preferably at most 20 can.

The surfactant can be incorporated into the vermiculite-alumina and alumina mixture using any suitable method. Preferably, the surfactant is incorporated during the mixing and formation of alumina and vermiculite-alumina. A preferred method is to incorporate an aqueous solution of the surfactant into the blend of alumina and vermiculite-alumina prior to forming the final support. Preferably, the surfactant is in the form of a paste or agglomerate in an amount of from 0.01 to 10% by weight based on the weight of the alumina and vermiculite-alumina.

A monoprotic acid such as nitric acid or formic acid can be added to the mixture in the aqueous solution to peptize the alumina in the binder. Additional water can be added to the mixture to provide sufficient moisture to form a dough of sufficient consistency for extrusion or spray drying.

The paste or dough can be prepared in a particulate form, wherein the preferred method is to have a desired mixture of alumina, vermiculite-alumina, surfactant, and water. The open mold is extruded and the extruded object is then divided into extrudates of the desired length and dried. Another calcination step can be utilized to provide additional strength to the extrudate. Typically, the calcination is carried out in dry air at a temperature of from 260 ° C (500 ° F) to 815 ° C (1500 ° F).

The extruded particles may have any suitable cross-sectional shape, i.e., symmetrical or asymmetrical, but most often have a symmetrical cross-sectional shape, preferably spherical, cylindrical or multilobal. The particles may have a cross-sectional diameter as small as 40 μm; however, they are typically 0.635 mm (0.25 inches) to 12.7 mm (0.5 inches), preferably 0.79 mm (1/32 inches) to 6.35 mm (0.25 inches), and The optimum is 0.06 mm (1/24 inch) to 4.23 mm (1/6 inch). The preferred catalyst configuration is similar to, for example, the cross-sectional shape of the clover shown in Figures 8 and 8A of US 4,028,227. Each "leaf" of the preferred clover-shaped particle section is defined by a 270[deg.] arc that is between 0.51 mm (0.02 inch) and 1.27 mm (.05 inch) in diameter. Other preferred microparticles have a quadrilobal cross-sectional shape (including asymmetric shapes and symmetrical shapes (such as in Figure 10 of US 4,028,227)).

The typical characteristics of the amorphous meteorite-alumina support used herein are that the overall pore volume, average pore diameter and surface area should be large enough to provide a large amount of space and area for the deposition of the active metal component. The overall pore volume of the support as determined by conventional mercury porosimetry methods is typically from 0.2 to 2.0 cc/gram, preferably from 0.25 to 1.0 cc/gram and most preferably from 0.3 to 0.9 cc/gram. Typically, the number of void volumes of the support having a pore diameter greater than 100 angstroms is less than 0.1 cc/gram, preferably less than 0.08 cc/gram and most preferably less than 0.05 cc/gram. The surface area determined by the BET method is typically above 50 m ² /g, for example above 200 m ² /g, preferably at least 250 m ² /g, and optimally 300 m ² /g to 400 m ² / Gram.

When preparing a catalyst, the support material is a single impregnated or multiple impregnated calcined amorphous refractory oxide support microparticle (having one or more precursors from at least one metal component of Group VIII or VIB of the periodic table) ) and combine. The Group VIII metal (preferably nickel) should be present in an amount of from 0.5 to 15% by weight, and the Group VIB metal (preferably tungsten) should be present in an amount of from 0 to 12% by weight. The impregnation process is accomplished by any method known in the art, for example, a spray dipping process in which a solution containing a metal precursor in dissolved form is sprayed onto the carrier particles. Another method is a multiple impregnation process in which the support material is repeatedly contacted with an impregnation solution with or without batch drying. Another method involves immersing the support in a large amount of impregnation solution or circulating the support therein, and yet another method is a pore volume or pore saturation technique in which the carrier particles are introduced into an impregnation solution having a capacity sufficient to fill the pores of the support. The pore saturation technique can sometimes be modified to utilize an impregnation solution that is between 10% and 10% less than the capacity of only filling the pores.

If the active metal pre-system is incorporated by impregnation, the subsequent or second calcination process will convert the metal to its temperature at elevated temperatures (eg, between 399 and 760 ° C (750 ° and 1400 ° F)). The respective oxide form. In some cases, calcination may be carried out after each active metal is separately impregnated. Successive calcination processes can produce catalysts containing active metals in their respective oxide forms.

The preferred oligomeric catalyst of the present invention has an amorphous state vermiculite-alumina substrate impregnated with 0.5 to 15% by weight of nickel in the form of an extrudate of 3.175 mm (0.125 inch) and a density of 0.45 to 0.65 g/ml. Also included are metals incorporated into the support by other methods, such as ion exchange and co-milling.

Another catalyst suitable for the present invention utilizes a co-gelatinized vermiculite-alumina support made by the well-known oil droplet method which allows the use of a support in the form of a large sphere. For example, an alumina sol used as a source of alumina is combined with an acidified water glass solution as a source of vermiculite, and the mixture is further combined with a suitable gelling agent (for example, urea, hexamethylenetetramine or a mixture thereof) )merge. The mixture, still below the gelation temperature, is discharged into a hot oil bath maintained at the gelation temperature by means of a nozzle or a rotating disc. The mixture was dispersed as a drop into an oil bath to form spherical gel particles during passage through the channels. The alumina sol is preferably prepared by combining alumina particles with a quantity of treated or deionized water and adding sufficient hydrochloric acid to digest a portion of the aluminum metal and form the desired sol. A suitable reaction rate is due to the reflux temperature of the mixture.

The spherical gel particles prepared by the oil drop method are usually aged in an oil bath for at least 10 to 16 hours, then in a suitable alkaline medium for at least 3 to 10 hours, and finally washed with water. Proper gelation of the mixture in the oil bath and subsequent gelation of the gel ball is not easy to accomplish at less than 48.9 ° C (120 ° F), and at 98.9 ° C (210 ° F), the rapid emission of gas easily breaks and weakens the sphere. . Using a higher temperature often improves the results by maintaining sufficient superatmospheric pressure to maintain water in the liquid phase during the forming and curing steps. If the gel particles are aged under superatmospheric pressure, an alkaline ripening step is not required.

The ball system is washed with water, preferably with water containing a small amount of ammonium hydroxide and/or ammonium nitrate. After washing, it is dried at a temperature of 93.3 ° C (200 ° F) to 315 ° C (600 ° F) for 6 to 24 hours or longer, and then calcined at a temperature of 426.67 ° C (800 ° F) to 760 ° C (1400 ° F) 2 to 12 Hours or more.

The Group VIII component and the Group VIB component are comprised of a co-gelatinized vermiculite-alumina support material by any suitable co-impregnation technique. Thus, the support material can be submerged, impregnated or suspended, or immersed in an aqueous impregnation solution containing a soluble Group VIII salt and a soluble Group VIB salt. A suitable method comprises immersing the support material in an impregnation solution and evaporating to dryness in a rotary steam dryer, the concentration of the impregnation solution being such that the atomic ratio of nickel to nickel plus tungsten contained in the final catalyst composite is 0.1 to 0.3. Another suitable method involves immersing the support material in an aqueous impregnation solution at room temperature until the solution completely penetrates the support. After absorbing the impregnation solution, the carrier drains the free surface liquid and is dried in a moving belt calciner.

The catalyst composite is typically dried at a temperature of 93.3 ° C (200 ° F) to 260 ° C (500 ° F) for 1 to 10 hours prior to calcination. According to the invention, the calcination is carried out in an oxidizing gas at a temperature of from 371 ° C (700 ° F) to 650 ° C (1200 ° F). Although other gases containing molecular oxygen may be used, the oxidizing gas is suitably air.

Another suitable catalyst is an oil droplet treated vermiculite-alumina spherical support impregnated with 0.5 to 15 weight percent nickel and 0 to 12 weight percent tungsten having a diameter of 3.175 mm (0.125 inch). Other suitable metal incorporation methods are contemplated. Suitable densities for other catalysts can range between 0.60 and 0.70 g/mL.

The dilute ethylene feed can be contacted with the oligomeric catalyst at temperatures between 200 and 400 °C. This reaction mainly occurs in the ethylene-based gas phase of GHSV 50 to 1000 hr ^-1 . Surprisingly, it has been found that at least 40% by weight and up to 75% by weight of ethylene is converted to heavier hydrocarbons in the feed gas stream despite the presence of impurities in the feedstock which poison the catalyst and thin the ethylene. The ethylene is first oligomerized on the catalyst to a heavier olefin. Certain heavier olefins may be cyclized on the catalyst and, in the presence of hydrogen, will promote the conversion of the olefins to paraffins, which are all hydrocarbons heavier than ethylene.

Although the raw material is not pure, the catalyst remains stable, but it can be regenerated upon deactivation. Suitable regeneration conditions include, for example, treating the catalyst in situ for 3 hours in hot air at 500 °C. The activity and selectivity of the regenerated catalyst is comparable to that of fresh catalyst.

The oligopolymerization product stream from the oligopolymerization reactor in line 160 can be passed to oligopolymerization separator 162, which can be a simple cylinder to separate the gaseous stream from the liquid stream. The oligopolymerization separator 162 is connected downstream to the oligopolymerization reactor 156. A gaseous product stream comprising light gases (such as hydrogen, methane, ethane, unreacted olefins, and light impurities) in overhead line 164 can be passed to combustion unit 166 to produce a stream in line 167. Alternatively, the gaseous product in overhead line 164 can be combusted to ignite the heater (not shown) and/or to provide a source of fuel gas that is diverted to the gas turbine (not shown) to generate electricity. The overhead line 164 is connected upstream to the combustion unit 166. The liquid bottoms stream from the heavier hydrocarbons in line 168 of the oligopolymer separator 162 can be lowered onto the valve and recycled back to the product separation section 90. Recirculation line 168 is downstream connected to bottom residue 169 of oligopolymer separator 162. Thus, the main column 92 is connected downstream and upstream to the oligomerization reactor 156. The bottoms stream is preferably recycled through recycle line 168 to a column 92 located between the heavy naphtha outlet 96a and the light cycle oil outlet 95a. Alternatively, a recirculating line 168 is fed to the light cycle oil line 95 or the heavy naphtha line 96. The recycle line is connected downstream of the oligomerization reactor 156 and upstream to the main column 92. Alternatively, the oligomeric polymerization product in line 160 or 168 may or may not be saturated and transported to a fuel tank that does not need to be recycled to product separation zone 90.

Instance

The use of the invention can be demonstrated by the following examples.

Example 1

Nickel and tungsten supported on a spherical matrix of amorphous vermiculite-alumina oil droplets are synthesized using the processes described herein above for other catalysts of the present invention. The metal comprises 1.5% by weight of nickel and 11% by weight of tungsten of the catalyst. The spherical substrates have a diameter of 3.175 mm. The catalyst has a ratio of vermiculite to alumina of 3, a density of 0.641 g/mL and a surface area of 371 m ² /g.

Example 2

The extruded amorphous amorphous vermiculite-alumina synthesis system will be amorphous amorphous vermiculite-alumina (the ratio of vermiculite to alumina provided by CCIC is 2.6) and the pseudobore provided under the trade name Catapal Stone, according to the weight ratio of 85- to -15 combination. The pseudo-acre stone is firstly peptized with nitric acid and then mixed with the amorphous vermiculite-alumina. A surfactant known under the trade name Antarox and water sufficient to wet the dough are added to the mixture. The catalyst mass was extruded from a 1.59 mm opening in a cylindrical template and calcined at 550 ° C after being divided into segments. The final catalyst consisted of 85% by weight vermiculite-alumina and 15% by weight alumina with a vermiculite to alumina ratio of 1.92 and a surface area of 268 m ² /g.

Example 3

3.37 grams of Ni(NO ₃ ) ₂ ‧6H ₂ O was dissolved in 32.08 grams of deionized water. The nickel solution was added in four portions and violently shaken between additions to bring the nickel solution into contact with the extruded amorphous meteorite-alumina of Example 2. A light green extrudate is formed. Then, the extrudate was dried at 110 ° C for 3 hours, then heated at 2 ° C / min to 500 ° C for calcination and then maintained at 500 ° C for 3 hours and then cooled to room temperature to convert the nickel metal into an oxide. form. The light grey extrudate was found to contain 1.5% by weight nickel.

Example 4

A sample of MTT zeolite having a vermiculite to alumina ratio of 40 was purchased from Zeolyst Corporation. The MTT zeolite was combined with pseudo-boehmite and extruded through a 3.175 opening in a cylindrical form and then calcined to 550 °C. The catalyst was composed of 80% by weight of MTT zeolite and 20% by weight of alumina.

Example 5

The effect of the catalyst of Example 1 on olefin oligomerization was tested at 280 ° C, 6,895 kPa (1000 psig), 586 OGHSV (olefin gas hourly rate), in a fixed bed operation on 10 mL of catalyst. The feed consisted of 30% by weight of C ₂ H ₄ and 70% by weight of CH ₄ . The results are shown in Table 1.

Example 6

The effect of the catalyst of Example 2 on olefin oligomerization was tested in a fixed bed operation at 280 ° C, 6,895 kPa (1000 psig), 586 OGHSV, on 10 mL of catalyst. The feed consisted of 23% by weight of C ₂ H ₄ , 14% by weight of C ₂ H ₆ , 35% by weight of CH ₄ , 13% by weight of H ₂ , 13% by weight of N ₂ , 1% by weight of CO, 1.5% by weight of CO ₂ , 10 The wppm H ₂ S composition was saturated with water vapor at 25 ° C and 3,447 kPa (500 psig) before entering the oligomerization reaction. The results are shown in Table I and Figure 2.

Example 7

The effect of the catalyst of Example 3 on olefin oligomerization was tested in a fixed bed operation at 280 ° C, 6,895 kPa (1000 psig), 586 OGHSV, on 10 mL of catalyst. The feed consisted of 23% by weight of C ₂ H ₄ , 14% by weight of C ₂ H ₆ , 35% by weight of CH ₄ , 13% by weight of H ₂ , 13% by weight of N ₂ , 1% by weight of CO, 1.5% by weight of CO ₂ , 10 The wppm H ₂ S composition was first subjected to water vapor saturation at 25 ° C and 3,447 kPa (500 psig), followed by oligomerization. The results are shown in Table I and Figure 2. 1 ppm NH _{3 was also} added to the feed during the 27 to 44 hours of operating the stream. No change was found in conversion or selectivity.

Example 8

The experiment of Example 7 was repeated except that the H ₂ S concentration in the feed was changed to 50 wppm instead of 10 wppm. The results are shown in Table I and Figure 2.

Example 2 based C ₂ H ₄ conversion of 6-8 versus the operating time of the stream. Based on the ethylene conversion, the nickel on the amorphous vermiculite-alumina catalyst of Example 3 exhibited better conversion than the vermiculite-alumina substrate of Example 2. The catalysts of Examples 2 and 3 were also affected by the impurities of the raw materials.

Example 9

The effect of the catalyst of Example 4 on olefin oligomerization was tested in a fixed bed operation at 280 ° C, 6,895 kPa (1000 psig), 586 OGHSV on 10 mL of catalyst. The starting material consisted of 23% by weight of C ₂ H ₄ , 14% by weight of C ₂ H ₆ , 35% by weight of CH ₄ , 13% by weight of H ₂ , 13% by weight of N ₂ , 1% by weight of CO, 1.5% by weight of CO ₂ , 10 wppm The H ₂ S composition was first subjected to water vapor saturation at 25 ° C and 3,447 kPa (500 psig), followed by oligomerization. The results are shown in Table I and Figure 3.

C ₂ H ₄ conversion system of FIG. 3 of Example 9 versus the operating time of the stream, which shows the impurity effect can MTT 4 of Example zeolite catalyst poisoning. After 20 hours of reaction, the conversion decreased to less than 10% by weight, while the conversion of the catalyst from Example 3 in Examples 7 and 8 was maintained at about 60% by weight.

Example 10

The effect of the catalyst of Example 4 on olefin oligomerization was tested in a fixed bed operation at 280 ° C, 6,895 kPa (1000 psig), 613 OGHSV, on 10 mL of catalyst. This material consisted of 30% by weight of C ₂ H ₄ and 70% by weight of CH ₄ . At the time of 21 hours of operation, hydrogen gas was added to obtain a raw material composed of 27% by weight of C ₂ H ₄ , 63% by weight of CH ₄ and 10% by weight of H ₂ . At 45 hours of operation, H ₂ containing 500 wppm of NH _{3 was} added to obtain a raw material containing 27% by weight of C ₂ H ₄ , 63% by weight of CH ₄ , 10% by weight of H ₂ and 50 wppm of NH ₃ . The results are shown in Table I and Figure 4.

Example 4 based ₂ H ₄ C 10 conversion rate of the stream versus the operation time, showing the effects of H ₂ and NH ₃ were impurities of MTT zeolite catalyst of Example 4. Referring to Figure 3, the ethylene conversion is reduced after the introduction of hydrogen into the feedstock for 20 hours of operation. In addition, the ethylene conversion rate decreased significantly rapidly after the introduction of ammonia at 45 hours of operation.

Example 11

The ammonia temperature programmed desorption test (ammonia TPD) consisted of first heating 250 mg of the catalyst sample to 550 ° C at a rate of 5 ° C/min in a helium atmosphere containing 20 ml/min of oxygen at a flow rate of 100 ml/min. After one hour of maintenance, the system was rinsed with helium for 15 minutes and the sample was cooled to 150 °C. The sample was then saturated with a 40 ml/min helium gas pulse containing ammonia. The total amount of ammonia used is significantly greater than the amount required to saturate the acid sites on all samples. The sample was washed with 40 ml/min of helium for 8 hours to remove physically adsorbed ammonia. Continuous cleaning with helium gas was carried out to raise the temperature to a final temperature of 600 ° C at a rate of 10 ° C / min. The amount of desorbed ammonia was monitored using a standard thermal conductivity detector. Integration of points determines the total amount of ammonia.

The total acidity is obtained from the ratio of the total amount of ammonia desorbed to the weight of the sample. The total acidity value as used herein is expressed in millimole units of ammonia per gram of dry sample. The catalytic acidity which is active for the oligomerization of the dilute ethylene stream has an acidity determined by ammonia TPD of at least 0.15, and preferably at least 0.25.

Before the temperature reached 300 ° C, the ratio of the total amount of ammonia desorbed from the sample to the dry weight of the sample gave a low temperature peak. The low temperature peak as used herein is expressed in units of ammonia milligrams per gram of dry sample. The catalyst active for the oligomerization of the dilute ethylene stream has a low temperature peak which is determined by the ammonia TPD to have a low temperature peak of at least 0.05, and preferably at least 0.06.

From the ratio of the low temperature peak to the total acidity, the unitless ratio is obtained, which is called the low temperature acidity ratio. Catalysts which are resistant to feedstock impurities in the dilute ethylene stream and which are reactive toward the oligomerization of the dilute ethylene stream have a low temperature acidity ratio of at least 0.15, suitably at least 0.2, and preferably greater than 0.25, as determined by ammonia TPD.

It can be seen from the examples that the zeolite catalyst is affected by the raw material impurities and the efficiency of ethylene conversion is greatly reduced, and the catalyst of the present invention is an effective ethylene oligomerization catalyst despite the presence of impurities typical of the catalytic poison in the raw material. . The ethylene conversion of the catalyst of the present invention is maintained at least 40% by weight, typically 60% by weight and preferably 70% by weight or more.

If not described in further detail, those skilled in the art will be able to utilize the present invention to the maximum extent possible. Therefore, the particular embodiments that have been described above are merely illustrative and are not intended to limit the disclosure in any respect.

In the above, all temperatures are expressed in degrees Celsius unless otherwise indicated, and all parts and percentages are by weight.

From the above description, those skilled in the art can readily determine the essential characteristics of the invention, and various changes and modifications can be made to the various uses and conditions without departing from the spirit and scope thereof.

6. . . Refinery complex

10. . . FCC unit section

12. . . FCC reactor

14. . . Catalytic reactor

16. . . Splitter

18. . . Regeneration catalyst riser

20. . . Riser

twenty two. . . reactor

twenty four. . . Fluidized pipeline

26. . . Primary separator

28. . . Cyclone separator

31. . . Product export

32. . . Pipeline

34. . . Stripping section

36. . . Consumption of catalyst standpipe

38. . . Air splitter

40. . . Burner riser

42. . . Sorter

44. . . Cyclone separator

46. . . Cyclone separator

47. . . Exhaust gas outlet

48. . . Exhaust line

90. . . Product recovery section

92. . . Fractionation column

93. . . Pipeline

94. . . Pipeline

95. . . Pipeline

95a. . . Export

96. . . Pipeline

96a. . . Export

97. . . Tower top pipeline

99. . . Main column receiver

101. . . Pipeline

102. . . Pipeline

104. . . compressor

106. . . Pipeline

107. . . Pipeline

108. . . Tower top pipeline

110. . . High voltage receiver

112. . . Pipeline

114. . . Primary absorber

116. . . Pipeline

118. . . Secondary absorber

119. . . Pipeline

120. . . Steam recovery section

121. . . Pipeline

122. . . Pipeline

124. . . Pipeline

126. . . Stripper

128. . . Pipeline

130. . . Debutanizer column

132. . . Pipeline

134. . . Pipeline

140. . . Absorber unit

142. . . Pipeline

143. . . Pipeline

144. . . Pipeline

145. . . Pipeline

146. . . Cleaning unit

147. . . Pipeline

148. . . Pipeline

150. . . Protective bed

151. . . Pipeline

152. . . compressor

154. . . Pipeline

156. . . Oligopolymerization reactor

158. . . Fixed catalyst bed

160. . . Pipeline

162. . . Oligomerization separator

164. . . Tower top pipeline

166. . . Combustion unit

167. . . Pipeline

168. . . Recirculation line

169. . . Bottom residue

Figure 1 is a schematic diagram of an FCC unit and an FCC product recovery system;

Figure 2 is a graph of ethylene conversion of Examples 6 to 8 versus operating time of the stream;

Figure 3 is a graph of the ethylene conversion of Example 9 versus the operating time of the stream;

Figure 4 is a graph of ethylene conversion from Example 10 versus run time for the stream.

6. . . Refinery complex

10. . . FCC unit section

12. . . FCC reactor

14. . . Catalytic reactor

16. . . Splitter

18. . . Regeneration catalyst riser

20. . . Riser

twenty two. . . reactor

twenty four. . . Fluidized pipeline

26. . . Primary separator

28. . . Cyclone separator

31. . . Product export

32. . . Pipeline

34. . . Stripping section

36. . . Consumption of catalyst standpipe

38. . . Air splitter

40. . . Burner riser

42. . . Sorter

44. . . Cyclone separator

46. . . Cyclone separator

47. . . Exhaust gas outlet

48. . . Exhaust line

90. . . Product recovery section

92. . . Fractionation column

93. . . Pipeline

94. . . Pipeline

95. . . Pipeline

95a. . . Export

96. . . Pipeline

96a. . . Export

97. . . Tower top pipeline

99. . . Main column receiver

101. . . Pipeline

102. . . Pipeline

104. . . compressor

106. . . Pipeline

107. . . Pipeline

108. . . Tower top pipeline

110. . . High voltage receiver

112. . . Pipeline

114. . . Primary absorber

116. . . Pipeline

118. . . Secondary absorber

119. . . Pipeline

120. . . Steam recovery section

121. . . Pipeline

122. . . Pipeline

124. . . Pipeline

126. . . Stripper

128. . . Pipeline

130. . . Debutanizer column

132. . . Pipeline

134. . . Pipeline

140. . . Absorber unit

142. . . Pipeline

143. . . Pipeline

144. . . Pipeline

145. . . Pipeline

146. . . Cleaning unit

147. . . Pipeline

148. . . Pipeline

150. . . Protective bed

151. . . Pipeline

152. . . compressor

154. . . Pipeline

156. . . Oligopolymerization reactor

158. . . Fixed catalyst bed

160. . . Pipeline

162. . . Oligomerization separator

164. . . Tower top pipeline

166. . . Combustion unit

167. . . Pipeline

168. . . Recirculation line

169. . . Bottom residue

Claims (10)

A method of oligomerizing ethylene, comprising: providing a feed stream comprising: between 5 and 50% by weight of ethylene, and at least one impurity selected from the group consisting of at least 0.1% by weight of carbon monoxide, at least 1 wppm of hydrogen sulfide, at least 1 wppm of ammonia, at least 5% by weight of hydrogen, and at least 0.1% by weight of carbon dioxide; the feed stream and the amorphous meteorite comprising a metal having a group selected from Groups VIII and VIB of the Periodic Table - The alumina substrate is contacted; and at least 40% by weight of the ethylene in the feed stream is converted to heavier hydrocarbons. The method of claim 1, wherein the catalyst has a meteorite to alumina ratio of at most 30. The method of claim 1, wherein the catalyst is impregnated with an amorphous state vermiculite-alumina matrix of from 0.5 to 15% by weight of nickel. The method of claim 1, wherein the catalyst has a low acidity ratio determined by the ammonia temperature programmed release desorption test of at least 0.15. The method of claim 1 wherein the feed stream comprises up to 0.5% by weight propylene. The method of claim 1, wherein the contacting step is performed at a pressure above a critical pressure of ethylene. A method comprising: contacting a cracking catalyst with a hydrocarbon feed stream to crack the hydrocarbon into a cracked product hydrocarbon having a lower molecular weight and depositing coke on the cracking catalyst to produce a coked cracking catalyst; Separating the coked cracking catalyst from the cracked product; adding oxygen to the coked cracking catalyst; burning the coke on the coked cracking catalyst with oxygen to regenerate the cracking catalyst; separating the cracked product Obtaining a dilute ethylene stream comprising from 5 to 50% by weight of ethylene; compressing the dilute ethylene stream to a pressure between 4,826 and 7,584 kPa; and contacting the dilute ethylene stream with an oligo-polymerization catalyst. The method of claim 7, wherein the contacting step is performed in a fixed bed of the catalyst. A device comprising: a fluid catalytic cracking reactor for contacting a cracking catalyst with a hydrocarbon feed stream, cracking a hydrocarbon feedstock stream into a cracked product having a lower molecular weight, and depositing coke on the cracking catalyst to obtain coking cracking a catalyst; a product outlet for flowing the cracked product from the reactor; a regenerator for contacting the coke with oxygen to burn on the coked cracking catalyst; a product recovery section connected to the product outlet, the product is recovered The section divides the cracked product into a plurality of product streams, including a stream comprising ethylene; a compressor connected downstream of the product recovery section for compressing the ethylene-containing stream; and a fixed connection to the compressor In a bed oligopolymerization reactor, the ethylene oligos in the ethylene-containing stream are polymerized into heavier hydrocarbons. The device of claim 9, further comprising: connected to the product outlet A main column, a primary absorber connected to the main column receiver, provides a primary exhaust stream comprising the ethylene-containing stream.