KR101358589B1 - Process for oligomerizing dilute ethylene - Google Patents

Abstract

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

Description

PROCESS FOR OLIGOMERIZING DILUTE ETHYLENE}

FIELD OF THE INVENTION The field of the present invention is an apparatus and method for converting diluted ethylene into hydrocarbons in a hydrocarbon stream. These heavy hydrocarbons can be used as motor fuel.

Dry gas is the generic name for an offgas stream from a fluid catalytic pyrolysis unit containing all gases having a boiling point lower than ethane. The offgas stream is compressed to remove as much of C3 and C4 gas as possible. Sulfur is also mostly adsorbed from offgas streams in scrubbers using amine adsorbents. The remaining stream is known as FCC dry gas. Typical dry gases include 5-50% by weight of ethylene, 10-20% by weight of ethane, 5-20% by weight of hydrogen, 5-20% by weight of nitrogen, 0.1-5.0% by weight of carbon monoxide and carbon dioxide, and 0.01% by weight, respectively. It contains less than% hydrogen sulfide and ammonia and the remainder of methane.

Currently, FCC dry gas streams are transported to the burners as fuel gas. The FCC unit, which handles 7,949 kiloliters (50,000 barrels) per day, will burn 181,000 kg (200 tons) of dry gas with 36,000 kg (40 tons) of ethylene as the daily fuel. Attempts to recover such ethylene appear economically advantageous because there is a high price difference between the fuel gas and the motor fuel product or pure ethylene. However, dry gas streams contain impurities that can contaminate the oligomerization catalyst and are too diluted so that ethylene recovery is not economically justified by the gas recovery system.

Oligomerization of the concentrated ethylene stream into the liquid product is a known technique. However, oligomerization generally involves the use of butylene or propylene, especially from dehydration feedstock or liquefied petroleum gas (LPG), to produce gasoline range olefins. Ethylene is rarely used as oligomerization feedstock because of its fairly low reactivity. There is a need to utilize dilute ethylene in the refinery stream.

We have found that ethylene in a dilute ethylene stream, such as an FCC dry gas stream, can be catalytically oligomerized with Group VIII and / or VIB metals on amorphous silica-alumina catalyst with heavy hydrocarbons. Heavy hydrocarbons can be separated and mixed in gasoline and diesel pools. The inventors have found that zeolite catalysts suitable for oligomerization of ethylene are rapidly deactivated in the presence of impurities such as carbon dioxide, ammonia and hydrogen sulfide. The impurities do not substantially affect catalysts comprising Group VIII and / or Group VIB metals on amorphous silica-alumina supports. In conclusion, the diluted ethylene in the FCC dry gas stream can be oligomerized into a liquid fuel product that is easy to separate from the unconverted gas stream. The unconverted gas can then be combusted as fuel gas, but more valuable ethylene is removed as heavy hydrocarbons.

Advantageously, the method and apparatus make it possible to utilize ethylene in a dilution stream, in which there are feed impurities which can be catalytically present.

Additional features and advantages of the invention will be apparent from the description, drawings, and claims of the invention provided herein below.

1 is a schematic diagram of an FCC unit and an FCC product recovery system.
2 is a graph depicting ethylene conversion over time on the stream for Examples 6-8.
3 is a graph depicting ethylene conversion over time on the stream for Example 9. FIG.
4 is a graph depicting ethylene conversion over time on the stream for Example 10.

The present invention can be applied to any hydrocarbon stream containing ethylene, preferably dilution ratio ethylene. Suitable, dilute ethylene streams may generally comprise from 5 to 50% by weight of ethylene. The FCC dry gas stream is a suitable dilute ethylene stream. Other dilute ethylene streams such as coker dry gas streams may also be used in the present invention. Since the present invention is particularly suitable for FCC dry gas, the present application describes the use of ethylene from the FCC dry gas stream.

Referring to FIG. 1, like numbers refer to like components, and FIG. 1 depicts an essential oil complex comprising a FCC unit portion 10 and a product recovery portion 90 generally. The FCC unit section 10 includes a reactor 12 and a catalyst regenerator 14. Process variables generally include pyrolysis reaction temperatures of 400 ° C. to 600 ° C. and catalyst regeneration temperatures of 500 ° C. to 900 ° C. Both pyrolysis and regeneration occur at absolute pressures (72.5 psia) below 506 kPa.

1 shows a typical FCC reactor 12 in which a heavy hydrocarbon feed or crude oil stream in distributor 16 is contacted with a regeneration pyrolysis catalyst coming from regenerated catalyst standpipe 18. This contact can occur in the narrow riser 20, which extends upwardly towards the bottom of the reactor vessel 22. Contact of the feed with the catalyst is fluidized by the gas from fluidization line 24. In an embodiment, the heat from the catalyst evaporates the hydrocarbon feed or oil, which is then pyrolyzed to a low molecular weight hydrocarbon product in the presence of the catalyst as both rise up the riser 20 and are transferred to the reactor vessel 22. do. Inevitable side reactions occur in the riser 20 and coke deposits remain on the catalyst to degrade the catalytic activity. The pyrolyzed light hydrocarbon product is then separated from the coked pyrolysis catalyst using a cyclone separator, which may include one or two stages of cyclone 28 and primary separator 26 in reactor vessel 22. The gaseous, pyrolysis product exits reactor vessel 22 through product outlet 31 to line 32 for delivery to downstream product recovery 90. The spent or coked catalyst requires regeneration for further use. The coked pyrolysis catalyst, after separation from the gaseous product hydrocarbons, descends to stripping 34, where steam is injected through a nozzle to purge any hydrocarbon vapors. After the stripping operation, the coked catalyst is transported to the catalyst regenerator 14 through spent catalyst standpipes 36.

1 shows a regenerator 14 known as a combustor. However, other types of regenerators are also suitable. In the catalyst regenerator 14, a stream of oxygen containing gas, such as air, enters through the air distributor 38 and contacts the coked catalyst. The coke is burned from the coked catalyst to provide regenerated catalyst and flue gas. The catalyst regeneration process adds real heat to the catalyst to provide energy to offset the endothermic pyrolysis reaction that occurs in the reactor riser 20. The catalyst and airflow flow upwards along the combustor riser 40 located in the catalyst regenerator 14, and after regeneration are initially released through the disengager 42 and separated. Further recovery of the regenerated catalyst and the flue gas leaving the disengager 42 is performed using first and second stage separator cyclones 44 and 46, respectively, present in the catalyst regenerator 14. The catalyst separated from the flue gas is distributed through the deep legs from the cyclones 44 and 46, while the relatively light flue gas in the catalyst sequentially exits the cyclones 44 and 46 and the flue gas outlets of the flue gas line 48 ( 47 exits the regenerator vessel 14. The regenerated catalyst is conveyed back to the riser 2 through the regenerated catalyst standpipe 18. As a result of the coke combustion, the flue gas vapor exiting the top of the catalyst regenerator 14 in line 48 together contains CO, CO ₂ , N ₂ and H ₂ O together with a small amount of other species. The hot flue gas exits regenerator 14 through flue gas outlet 47 in line 48 for further processing.

The product recovery part 90 communicates downstream with the product outlet 31. "Downstream communication" means that at least some of the materials entering the downstream communicating component can be operatively flowed from the component in communication with it. "Communication" means that material flow is operatively allowed between the listed components. In product recovery 90, gaseous FCC product in line 32 is directed to the bottom of FCC main fractionation column 92. Main column 92 communicates downstream with product outlet 31. Some fractions of the FCC product are recovered at the bottom from heavy slurry oil from line 93, heavy circulating oil stream from line 94, light circulating oil in line 95 recovered at outlet 95a and outlet 96a. A heavy naphtha stream in line 96 may be included and recovered from the main column. Any or all lines 93-96 are cooled and pumped back to main column 92 to cool the main column at a generally higher position. Gasoline and gaseous light hydrocarbons are removed from the main column 92 in the overhead line 97 and condensed before entering the main column receiver 99. The main column receiver 99 is in downstream communication with the product outlet 31, and the main column 92 is in upstream communication with the main column receiver 99. "Upstream communication" means that at least some of the flowing material from the component in upstream communication can operatively flow into the component in communication with it.

The aqueous stream is removed from the boot in the receiver 99. In addition, the condensed light naphtha stream is removed in line 101 while the overhead stream is removed in line 102. The overhead stream in line 102 contains a gaseous light hydrocarbon that may comprise a dilute ethylene stream. Streams in lines 101 and 102 may enter vapor recovery 12 of product recovery 90.

Although steam recovery 120 is shown as an adsorption-based gsmtem, any vapor recovery system may be used, including a cold box system. To achieve sufficient separation of the light gas components, the gas stream in line 102 is compressed in compressor 104. More than one compressor stage can be used, but typically dual stage compression is used. The compressed light hydrocarbon stream in line 106 is combined with the streams in lines 107 and 108, cooled and delivered to high pressure receiver 11. The aqueous stream from receiver 11 may be sent to main column receiver 99. The gaseous hydrocarbon stream in line 112 containing the dilute ethylene stream is sent to the primary adsorber 114 where it is contacted with unstable gasoline from main column receiver 99 in line 101 to separate between C3 + and C2- hydrocarbons. Is carried out. The primary adsorber 114 is in downstream communication with the main column receiver 99. The liquid C 3+ stream in line 107 is returned to line 106 before cooling. The primary offgas stream in line 116 from the primary adsorber 114 comprises a dilute ethylene stream for the purposes of the present invention. However, in order to further concentrate the ethylene stream and recover heavy components, line 116 is optionally directed to a secondary adsorber 118 where a circulating stream of light circulating oil in line 121 is diverted from line 95. Most residual C5 + and some C3-C4 materials are adsorbed in this primary offgas stream. The secondary adsorber 118 is in downstream communication with the primary adsorber 114. Light circulating oil from the bottom of the secondary adsorber in line 119 enriched with C 3+ material is returned to main column 92 via a peripheral pump to line 95. The overhead of the secondary adsorber 118 comprising dry gas, which is mostly C2-hydrocarbons with hydrogen sulfide, ammonia, carbon dioxide and hydrogen, is removed from the secondary offgas stream in line 122 to include the dilute ethylene stream.

Liquid from high pressure receiver 110 in line 124 is directed to stripper 126. Most of C2- is removed from the overhead of stripper 126 and returned to line 106 via overhead line 108. The liquid bottoms stream from stripper 126 is sent to butane removal column 130 via line 128. The overhead stream in line 132 from the butane remover contains C3-C4 olefinic product, while the bottom stream in line 134 containing stable gasoline can be further processed and sent to the gasoline reservoir.

The dilute ethylene stream of the present invention may comprise an FCC dry gas stream comprising 5-50% by weight of ethylene and preferably 10-30% by weight of ethylene. Usually methane is the main component in the dilute ethylene stream at a concentration of 25 to 55% by weight and substantially 5 to 45% by weight of ethane is substantially present. 1-25% and generally 5-20% by weight of hydrogen and nitrogen may be present in the dilute ethylene stream, respectively. Saturated levels of water may also be present in the dilute ethylene stream. If secondary adsorber 118 is used, C3 +, which does not exceed 5% by weight, will be present with generally less than 0.5% by weight of propylene.

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

We have found that many impurities in the dry gas ethylene stream can contaminate the oligomerization catalyst. Hydrogen and carbon monoxide can reduce metal sites inactively. Carbon dioxide and ammonia can attack acidic sites on the catalyst. Hydrogen sulfide may attack the metal on the catalyst to produce metal sulfides. Acetylene may be polymerized to form a viscous rubber on the catalyst or device.

A secondary off gas stream in line 122, including the dilute ethylene stream, may enter the optional amine adsorber unit 140 to remove hydrogen sulfide at low concentrations. A dilute aqueous amine solution, for example comprising monoethanol amine or diethanol amine, enters the adsorber 140 via line 142 and contacts the flowable secondary offgas stream to adsorb hydrogen sulfide, abundant aqueous containing hydrogen sulfide The amine adsorption solution is removed from the adsorption zone 140 via line 143 and recovered and possibly subsequently processed.

The amine-treated dilute ethylene stream in line 144 enters the optional water washing unit 146 to remove residual amines transported from the amine adsorber 140 and ammonia and carbon dioxide in the dilute ethylene stream in line 144. To reduce the concentration. Water enters the water wash section of line 145. The water in line 145 is generally slightly acidified to enhance the capture of basic molecules such as amines. The aqueous stream in line 147 is rich in amines and potentially ammonia and carbon dioxide can be further processed leaving the water wash unit 146.

The optionally amine treated dilute ethylene and possibly the water washed stream in line 148 may then be treated in optional guard layer 150 to remove at least one of impurities such as carbon monoxide, hydrogen sulfide and ammonia at low concentrations. have. The guard layer 150 may include an adsorbent that adsorbs impurities such as hydrogen sulfide, which may contaminate the oligomerization catalyst. The guard layer 150 may contain a plurality of adsorbents to adsorb one or more impurity types. A typical adsorbent for hydrogen sulfide adsorption is ADS-12, an adsorbent for CO adsorption is ADS-106 and an adsorbent for ammonia adsorption is UOP MOLSIV 3A, all available from UOP, LLC. The adsorbents can be mixed in a single bed or arranged in a continuous bed.

The dilute ethylene stream in line 151, which is amine treated if possible, washed with water, and possibly adsorbed, and further removed from hydrogen sulfide, ammonia and carbon monoxide, will generally have at least one of the following impurity concentrations: from 0.1% by weight to Up to 5.0 wt% carbon monoxide and / or 0.1 wt% up to 5.0 wt% carbon dioxide, and / or at least 1 wppm up to 500 wppm hydrogen sulfide and / or at least 1 up to 500 wppm ammonia, and / or at least 5 up to 20% by weight of hydrogen. The type of impurities present and their concentration vary depending on the treatment process and the source of the dilute ethylene stream.

Line 151 conveys the dilute ethylene stream to compressor 152 and is pressurized to the reactor pressure. Compressor 152 is in downstream communication with main column 92, product recovery portion 90, and product outlet 31. The compressed dilute ethylene stream may be compressed to at least 3,550 kPa (500 psia), preferably not more than 10,445 kPa (1500 psia), suitably from 4,930 kPa (700 psia) to 7,687 kPa (110OOpsia). The dilute ethylene stream is preferably pressurized to above the critical pressure of ethylene of 4,992 kPa (724 psia) relative to pure ethylene to avoid rapid catalyst deactivation. Compressor 152 may include one or more stages with interstage cooling. A heater may be needed to raise the compressed stream to the reaction temperature. Compressed dilute ethylene is conveyed to the oligomerization reactor 156 in line 154.

The oligomerization reactor 156 is in downstream communication with the compressor 152 and the primary and secondary adsorbers 114 and 118, respectively. The oligomerization reactor preferably contains a fixed catalyst bed 158. The dilute ethylene feed stream is preferably contacted with the catalyst in a downflow run. However, upflow drive may also be suitable. The catalyst is preferably an amorphous silica-alumina base with Group VIII and / or Group VIB metals of the periodic table using chemical abstract service tin. In one aspect, the catalyst has a Group VIII metal promoted with a Group VIB metal. In one aspect, the catalyst has a silica to alumina ratio of no greater than 30, preferably no greater than 20. As a rule, silica and alumina are present only at the base so that the silica to alumina ratio is the same for the base and for the catalyst. The metal may be impregnated on the silica-alumina base or ion exchanged. Culling is also contemplated. The catalyst for the invention may have a low temperature acidity ratio of at least 0.15, suitably 0.2, and preferably higher than 0.25, as measured by the ammonia temperature elevated desorption method (ammonia TPD) described below. In addition, suitable catalysts have a surface area of 50 to 400 m 2 / g measured by nitrogen BET.

Preferred oligomerization catalysts of the present invention are described below. Preferred oligomerization catalysts include amorphous silica-alumina supports. One component of the catalyst support used in the present invention is alumina. The alumina can be any of a variety of hydrous aluminum oxide or alumina gels such as alpha-alumina monohydrate of boehmite or pseudo-boehmite structure, alpha-alumina trihydrate of gibbsite structure, beta-alumina trihydrate of bayerite structure, and the like. Particularly preferred aluminas are commercially available from Sasol North America Alumina Product Group under the trade name Catapal. This material is an ultra high purity alpha-alumina monohydrate (pseudo-boehmite) which has been found to produce high purity gamma-alumina after high temperature calcination. Another component of the catalyst support is amorphous silica-alumina. Suitable silica-alumina with a silica to alumina ratio of 2.6 is available from CCIC, a subsidiary of JGC (one minute).

Another component used in the preparation of the catalyst used in the present invention is a surfactant. The surfactant is preferably mixed with the alumina and silica-alumina powders described above. The final mixture of surfactant, alumina and silica-alumina is formed, dried and calcined as described below. Calcination effectively removes the organic components of the surfactant by combustion but only after the surfactant has faithfully performed its function according to the invention. Any suitable surfactant can be used in accordance with the present invention. Preferred surfactants are those selected from a series of commercially available surfactants sold under the trade name "Antarox" by Solvay S.A. "Antarox" surfactants are generally characterized as modified linear aliphatic polyethers and are low foaming biodegradable detergents and wetting agents.

Suitable silica-alumina mixtures are prepared by mixing proportional volume of silica-alumina and alumina such that the desired silica-to-alumina ratio is obtained. In an embodiment, 85 wt% amorphous silica-alumina having a silica to alumina ratio of 2.6 and 15 wt% alumina powder will provide a suitable support. In embodiments, ratios other than 85-to-15 amorphous silica-alumina to alumina ratios may also be suitable, unless the final silica-to-alumina ratio of the support is suitably greater than 30 and preferably greater than 20. have.

Any convenient method may be used to introduce the surfactant with the silica-alumina and the alumina mixture. The surfactant is preferably mixed during the mixing and formation of the alumina and silica-alumina. The preferred method is to mix a blend of alumina and silica-alumina with an aqueous solution of the surfactant before the final formation of the support. It is preferred that the surfactant is present in the paste or dough in an amount of 0.01 to 10% by weight, based on the weight of the alumina and silica-alumina.

Monoprotic acid such as nitric acid or formic acid can be added to the mixture in aqueous solution to peptize the alumina in the binder. Additional water is added to the mixture to provide sufficient wettability such that the dough is extruded or spray dried to a sufficient consistency.

Pastes or doughs may be prepared in the form of shaped particulates, and in a preferred method, a dough mixture of alumina, silica-alumina, surfactants and water is extruded through a die equipped with openings of appropriate size and formation, followed by extrusion The molded material is separated into an extrudate of desired length and dried. Additional calcination steps may be applied to provide additional strength to the extrudate. Generally, calcination is carried out in a dry air stream at temperatures between 260 ° C. (500 ° F.) and 815 ° C. (1500 ° F.).

The extruded particles may be of any suitable cross-sectional shape, such as symmetrical or asymmetrical, but are usually symmetrical cross-sectional shapes, preferably spherical, cylindrical or multileafed. The cross-sectional diameter of the particles can be as small as 40 μm; Generally, 0.635 mm (0.25 inch) to 12.7 mm (0.5 inch), preferably 0.79 mm (1/32 inch) to 6.35 mm (0.25 inch), most preferably 0.06 mm (1/24 inch) to 4.23 mm (1/6 inch). Among the preferred catalysts, the conformation is a cross-sectional shape similar to a three leaf clover, for example as shown in FIGS. 8 and 8A of US Pat. No. 4,028,227. Preferred clover-shaped particles are defined as a circle of 270 ° arc in which each "leaf" in cross section has a diameter of 0.51 mm (0.02 inch) to 1.27 mm (.05 inch). Other preferred particles are those having a four-lobed cross-sectional shape, including symmetrical shapes, and asymmetrical shapes as shown in FIG. 10, etc. of US Pat. No. 4,028,227.

Typical features of the amorphous silica-alumina support as used herein are the total pore volume, average pore diameter and surface area large enough to provide substantial space and area for depositing the active metal component. The total pore volume of the support, as measured by conventional mercury porosimetry, is generally 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. In general, the pore volume of the support having a pore diameter greater than 100 ohms is less than 0.1 cc / gram, preferably less than 0.08 cc / gram, most preferably less than 0.05 cc / gram. B.E.T. The surface area, measured by the method, is generally at least 50 m 2 / gram, for example at least 200 m 2 / gram, preferably at least 250 m 2 / gram, most preferably from 300 m 2 / gram to 400 m 2 / gram.

To prepare the catalyst, the support material is hybridized by single impregnation or plural impregnation of the calcined amorphous refractory oxide support particles with at least one precursor of at least one metal component from Group VIII or Group VIB of the Periodic Table. The Group VIII metal, preferably nickel, should be present at a concentration of 0.5 to 15% by weight and the Group VIB metal, preferably tungsten, should be present at a concentration of 0 to 12% by weight. Impregnation can be carried out by any method known in the art, for example, by spray impregnation, which sprays a solution containing a metal precursor in dissolved form onto the support particles. Another method is a multi-immersion method in which the support material is repeatedly contacted with the impregnation solution with or without intermittent drying. Another method involves immersing the support in a large amount of impregnation solution or circulating the support in the impregnation solution; alternatively, the pore volume for introducing the support particles into an impregnation solution of exactly sufficient volume to fill the pores of the support, or It is a porous saturation method. Sometimes the pore saturation method can be modified to use an impregnation liquid having a volume of 10% less to 10% more than the volume to fill the pores accurately.

When the active metal precursor is introduced by impregnation, subsequent or secondary calcination at high temperatures such as, for example, 399 ° C. to 76 ° C. (750 ° F. to 1400 ° F.) converts the metals into their individual oxide forms. In some cases, calcination may be followed by each impregnation of the individual active metal. Subsequent calcination results in a catalyst containing the active metal in their individual oxide form.

Preferred oligomerization catalysts of the present invention have an amorphous silica-alumina base impregnated with 0.5-15% by weight of nickel in the form of 3.175 mm (0.125 inch) extrusions with a density of 0.45-0.65 g / mL. It is also contemplated to introduce the metal into the support via other methods such as ion exchange and co-mulling.

Other catalysts suitable for the present invention utilize co-gelled silica-alumina supports prepared by known oil droplet methods that can utilize macrosphere supports. For example, an alumina sol, used as an alumina source, is mixed with an acidic water glass solution as a silica source and this mixture is further mixed with a suitable gelling agent, for example urea, hexamethylenetetraamine, or mixtures thereof. This mixture is still released below the gelling temperature, through a nozzle or a rotating plate, into a hot oil bath maintained at the gelling temperature. The mixture is dispersed in the oil bath as droplets formed into spherical gel particles while passing through the oil bath. The alumina sol is preferably prepared by mixing aluminum pellets with a fixed amount of treated or deionized water and adding thereto an amount of hydrochloric acid sufficient to decompose a portion of the aluminum metal and form the desired sol. Appropriate reaction rates are carried out at the reflux temperature of the mixture.

The spherical gel particles prepared by the oil droplet method are aged and finally water washed for at least 10 to 16 hours, generally in an oil bath, and then in a suitable alkali or basic medium for at least 3 to 10 hours. Proper gelation of the mixture in the oil bath and subsequent aging of the gel spheres are not readily performed below 48.9 ° C. (120 ° F.), and at 98.9 ° C. (210 ° F.), the rapid release of gas will rupture the spheres or May weaken. By maintaining sufficient superatmospheric pressure during the formation and maturation steps to maintain water in the liquid layer, high temperatures can often be applied, with improved results. If the gel particles are aged at superatmospheric pressure, an alkaline aging step is not necessary.

Windy, the spheres are water washed with water containing a small amount of ammonium hydroxide and / or ammonium nitrate. After washing, the spheres were dried at a temperature of 93.30 ° C. (200 ° F.) to 315 ° C. (600 ° F.) for a period of 6 to 24 hours or longer, followed by 426.67 ° C. (800 ° for 2-12 hours or longer). F) to 760 [deg.] C. (1400 [deg.] F.).

The Group VIII component and Group VIB component consist of the silica-alumina carrier material co-gelled by any suitable coprecipitation method. Thus, the carrier material may be dipped, immersed, suspended or otherwise impregnated with an aqueous impregnation solution containing soluble Group VIII salts and soluble Group VIB salts. One suitable method involves impregnating the carrier material into the impregnation solution and evaporating it to dryness in a rotary steam dryer, the concentration of the impregnation solution being such that the final catalyst composite has an atomic ratio of 0.1 to 0.3 of nickel to nickel and tungsten. Decide Another suitable method involves immersing the carrier material in the aqueous impregnation solution at room temperature until the solution is completely infiltrated. After absorbing the impregnation, the carrier drains the free surface liquid and dries on a mobile belt calciner.

The catalyst composite is generally dried at a temperature of 93.3 ° C. (200 ° F.) to 260 ° C. (500 ° F.) for 1 to 10 hours before calcination. According to the invention, calcination is carried out in an oxidizing atmosphere at temperatures of 371 ° C. (700 ° F.) to 650 ° C. (1200 ° F.). Other gases containing molecular oxygen may be applied, but the oxidative atmosphere is suitably air.

Another suitable catalyst is a mineralized silica-alumina spherical support having a diameter of 3.175 mm (0.125 inch) impregnated with 0.5-15 wt% nickel and 0-12 wt% tungsten. Other metal introduction methods may be appropriate and may be considered. Suitable density ranges for other catalysts are 0.60 to 0.70 g / mL.

The diluted ethylene feed may be contacted with the oligomerization catalyst at a temperature of 200 ° C to 400 ° C. The reaction takes place primarily in the gas phase of GHSV 50 to 1000 hr-1 on an ethylene basis. Surprisingly, the inventors have found that at least 40% to 75% by weight of ethylene in the feed stream is converted to heavy hydrocarbons, despite the presence of impurities in the feed that contaminate the ethylene dilution in the catalyst and feed. Ethylene is first oligomerized to heavy olefins on the catalyst. Some of the heavy olefins may be cyclized on the catalyst, and the presence of hydrogen may promote the conversion of olefins to paraffins, which are all heavier hydrocarbons than ethylene.

The catalyst remains stable despite impurity supply but can be regenerated upon deactivation. Suitable regeneration conditions include, for example, applying hot air at 500 ° C. for 3 hours with the catalyst in situ. The activity and selectivity of the regenerated catalyst is comparable with fresh catalyst.

The oligomerization product stream from the oligomerization reactor in line 160 may be transported to oligomerization separator 162, which may be a simple flash drum to separate the gas stream from the liquid stream. The oligomerization separator 162 is in downstream communication with the oligomerization reactor 156. The gaseous product stream in overhead line 164 comprising light gases such as hydrogen, methane, ethane, unreacted olefins and light impurities may be transported to combustion unit 166 to produce a stream in line 167. Alternatively, gaseous products in overhead line 164 are combusted to provide a source of flue gas to ignite a heater (not shown) and / or to turn on a gas turbine (not shown) for power generation. Overhead line 164 is in upstream communication with combustion unit 166. The liquid bottoms stream containing heavy hydrocarbons in line 168 from oligomerization separator 162 may be lowered to a valve and recycled back to product separation 90. Recirculation line 168 is in downstream communication with bottom 169 of oligomerization separator 162. As a result, main column 92 communicates downstream and upstream with oligomerization reactor 156. The bottoms stream is preferably recycled via recycle line 168 to main column 92 at a location between heavy naphtha outlet 96a and light circulating oil outlet 95a. Alternatively, recycle line 168 supplies light circulating oil line 95 or heavy naphtha line 96. The recycle line communicates downstream with the oligomerization reactor 156 and upstream with the main column 92. Alternatively, the oligomerization product in line 160 or 168 may or may not be saturated and may be transported to the fuel tank without recycling to product separation zone 90.

Example

The usefulness of the present invention is demonstrated by the following examples.

Example 1

Nickel and tungsten on amorphous silica-alumina-deposited spherical base were synthesized through the procedure provided above for the alternative catalyst of the present invention. The metal constitutes 1.5 wt% nickel and 11 wt% tungsten in the catalyst. The spherical base is 3.175 mm in diameter. The catalyst has a silica to alumina ratio of 3, a density of 0.641 g / mL and a surface area of 371 m 2 / g.

Example 2

Extruded amorphous silica-alumina was synthesized by combining amorphous silica-alumina having a silica to alumina ratio of 2.6 provided by CCIC, and pseudo-boehmite sold under the trade name Catapal and having a weight ratio of 85-to-15. Pseudo boehmite was peptized with nitric acid prior to mixing with amorphous silica-alumina. Sufficient water to wet the dough and surfactant provided under the trade name Antarox were added to the mixture. The catalyst dough was extruded through a 1.59 mm opening in a cylindrical die plate, broken into pieces and then calcined at 5500 ° C. The final catalyst consists of 85 wt% silica-alumina and 15 wt% alumina with a silica-to-alumina ratio of 1.92 and a surface area of 368 m 2 / g.

Example 3

3.37 grams of Ni (NO ₃ ) ₂ .6H ₂ O was dissolved in 32.08 grams of DI water. The extruded amorphous silica-alumina of Example 2 was contacted with the nickel solution by adding the nickel solution four times and vigorously shaking between additions. A light green extrudate was obtained. The nickel metal was then converted to the oxide form by drying the extrudate at 110 ° C. for 3 hours, then heating to 500 ° C. at 2 ° C./min, maintaining it at 500 ° C. for 3 hours and then cooling to room temperature. . The light gray extrudate was found to contain 1.5 wt% nickel.

Example 4

Samples of MTT zeolites with a silica-to-alumina ratio of 40 were obtained from Zeolyst Corporation. MTT zeolite is combined with pseudo-boehmite and extruded through the 3.175 mm opening of the cylindrical die plate before calcination to 550 ° C. The final catalyst consists of 80 wt% MTT zeolite and 20 wt% alumina.

Example 5

The catalyst of Example 1 was tested for olefin oligomerization at 28O [deg.] C., 6,895 kPa (1000 psig), 58 olefin gas hourly space velocity (OGHSV) with a fixed bed run on 10 mL of catalyst. The feed consists of 30 wt% C ₂ H ₄ and 70 wt% CH ₄ . The results are shown in Table 1.

Example 6

The catalyst of Example 2 was tested for olefin oligomerization at 28O < 0 > C, 6,895 kPa (1000 psig), 586 OGHSV with fixed bed operation on a 10 mL catalyst. The feed is 23 wt% C ₂ H ₄ , 14 wt% C ₂ H ₆ , 35 wt% CH ₄ , 13 wt% H ₂ , 13 wt% N ₂ , 1 wt% CO, 1.5 wt% CO ₂ , 10 wppm Consisting of H ₂ S and saturated with water vapor at 25 ° C. and 3,447 kPa (500 psig) before feeding the oligomerization reaction. The results are shown in Table 1 and FIG.

Example 7

The catalyst of Example 3 was tested for olifen oligomerization at 28O < 0 > C, 6,895 kPa (1000 psig), 586 OGHSV with fixed bed operation on 10 mL of catalyst. The feed is 23 wt% C ₂ H ₄ , 14 wt% C ₂ H ₆ , 35 wt% CH ₄ , 13 wt% H ₂ , 13 wt% N ₂ , 1 wt% CO, 1.5 wt% CO ₂ , 10 wppm Consisting of H ₂ S and saturated with water vapor at 25 ° C. and 3,447 kPa (500 psig) before feeding the oligomerization reaction. The results are shown in Table 1 and FIG. For 27-44 hours on stream, 1 ppm NH ₃ was also added to the feed. There was no change in conversion or selectivity.

Example 8

The experiment of Example 7 was repeated except that the concentration of H ₂ S in the feed was 50 wppm instead of 10 wppm. Table 1 and Figure 2 show the results.

2 is a graph showing the C ₂ H ₄ conversion versus time on the streams of Examples 6-8. Nickel on the amorphous silica-alumina catalyst of Example 3 operates better than the silica-alumina base of Example 2 in terms of ethylene conversion. The catalysts of Examples 2 and 3 were also less affected by feed impurities.

Example 9

The catalyst of Example 4 was tested for olefin oligomerization at 28O < 0 > C, 6,895 kPa (1000 psig), 586 OGHSV with fixed bed operation on 10 mL of catalyst. The feed is 23 wt% C ₂ H ₄ , 14 wt% C ₂ H ₆ , 35 wt% CH ₄ , 13 wt% H ₂ , 13 wt% N ₂ , 1 wt% CO, 1.5 wt% CO ₂ , 10 wppm Consisting of H ₂ S and saturated with water vapor at 25 ° C. and 3,447 kPa (500 psig) before the oligomerization reaction feed. The results are shown in Table 1 and FIG.

3 is a graph of C ₂ H ₄ conversion over time on stream for Example 9 showing the effect of impurities that may contaminate the MTT zeolite catalyst of Example 4. FIG. After 20 hours of reaction, the conversion fell below 10% by weight while the conversion with the catalyst of Example 3 in Examples 7 and 8 was maintained at approximately 60% by weight.

Example 10

The catalyst of Example 4 was tested for olefin oligomerization at 28O < 0 > C, 6,895 kPa (1000 psig) kPa, 613 OGHSV with fixed bed operation on a 10 mL catalyst. The feed consists of 30 wt% C ₂ H ₄ and 70 wt% CH ₄ . At 21 hours on the stream, hydrogen was added to obtain a feed consisting of 27 wt% C ₂ H ₄ , 63 wt% CH ₄ and 10 wt% H ₂ . At 45 hours on the stream, 500 wppm NH ₃ in H ₂ was added to obtain a feed of 27 wt% C ₂ H ₄ , 63 wt% CH ₄ , 10 wt% H ₂ and 50 wppm NH ₃ . The results are shown in Table 1 and FIG.

FIG. 4 is a plot of C ₂ H ₄ conversion over time on a stream from Example 10 showing the effect of impurities H ₂ and NH ₃ on the MTT zeolite catalyst of Example _4. FIG. As can be seen in FIG. 3, the ethylene conversion dropped when hydrogen was introduced into the feed at 20 hours on the stream. In addition, upon introduction of ammonia at 45 hours on the stream, ethylene conversion rapidly decreased significantly.

Example Conversion Rate C ₂ -C ₄ selectivity C ₅ -C ₁₀ selectivity C ₁₀ + selectivity Example 5 92 30 53 17 Example 6 38 8 34 58 Example 7 75 9 18 73 Example 8 74 9 17 74 Example 9  5 hours on stream 77 One 31 68  20 hours on stream 7 10 35 55 Example 10  10 hours on stream 90 One 31 68  35 hours on stream 65 4 31 65  45 hours on stream 63 4 31 65  70 hours on stream 30 6 24 70

Example 11

Ammonia temperature elevated adsorption (ammonia TPD) involves first heating a 250 mg catalyst sample to a temperature of 550 ° C. at a rate of 5 ° C./min under a helium atmosphere of 100 ml / min flow rate at 20% by volume oxygen. After 1 hour hold, the system is washed for 15 minutes with helium and the sample is cooled to 150 ° C. The sample is then saturated with a pulse of ammonia in helium at 40 ml / min. The total amount of ammonia used significantly exceeds the amount needed to saturate all acidic sites on the sample. Samples are purged with 40 ml / min of helium for 8 hours to remove physically adsorbed ammonia. While continuing helium purge, the temperature is raised to a final temperature of 600 ° C. at a rate of 10 ° C./min. The amount of desorbed ammonia is monitored using a calibrated thermal conductivity detector. The total amount of ammonia is determined by integration.

The total acidity is calculated as the ratio of the total amount of desorbed ammonia to the dry weight of the sample. As used herein, the total acidity value is given in millimoles of ammonia per gram dry sample. The catalyst active for oligomerization of the dilute ethylene stream is acidic, ie has an overall acidity of at least 0.15, preferably at least 0.25, as measured by ammonia TPD.

The low temperature peak is calculated as the ratio of the total amount of ammonia desorbed from the sample to the dry weight of the sample before reaching 300 ° C. temperature. As used herein, low temperature peak values are given in millimoles of ammonia per gram of dry sample. Catalysts active for oligomerization of dilute ethylene streams have a low temperature peak, ie a low temperature peak measured by ammonia TPD, of at least 0.05, preferably at least 0.06.

The ratio of cold peaks to total acidity is given in unit-less ratios known as cold acidity ratios. Catalysts resistant to feed impurities in the dilute ethylene stream, active for oligomerization of the dilute ethylene stream, have a low temperature acidity ratio, measured by ammonia TPD, of at least 0.15, suitably at least 0.2, preferably greater than 0.25.

catalyst Total acidity, mmol / g Cold peak, mmol / g Cryogenic acidity rate Example 1 0.323 0.116 0.36 Example 2 0.285 0.092 0.32 Example 3 0.264 0.084 0.32 Example 4 0.311 0.037 0.12

It can be seen from the examples that the zeolite catalysts are significantly less effective in terms of ethylene conversion by feed impurities while the catalysts of the present invention remain effective ethylene oligomerization catalysts despite the presence of impurities in the feed, which are typical catalyst poisons. The catalyst of the present invention maintains an ethylene conversion of at least 40% by weight, generally 60% by weight, preferably at least 70% by weight.

Without further effort, one skilled in the art, using the foregoing description, believes that the present invention may be utilized to its fullest extent. The particular preferred embodiments described above are, therefore, merely illustrative and are not intended to limit the remainder of the disclosure in any way.

In the foregoing description, unless stated otherwise, all temperatures are stated in degrees Celsius, and all parts and ratios are by weight.

From the foregoing description, those skilled in the art can readily identify the essential features of the present invention, and various changes and modifications can be made to the present invention for various uses and conditions without departing from the scope of the present invention.

Claims (10)

As a method of oligomerizing ethylene,
Providing a feed stream comprising from 5 to 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 Doing;
Contacting said feed stream with a catalyst comprising an amorphous silica-alumina base and a metal selected from the group consisting of Groups VIII and VIB of the Periodic Table; And
Converting at least 40% by weight of ethylene in the feed stream to heavy hydrocarbons
Oligomerization method comprising a. The method of claim 1, wherein the catalyst has a silica to alumina weight ratio of no greater than 30 in the catalyst. The method of claim 1, wherein the catalyst is an amorphous silica-alumina base impregnated with 0.5-15 wt% nickel. The method according to claim 1, wherein the catalyst has an acidity ratio measured by ammonia temperature elevated desorption test is 0.15 or more as measured at 300 ℃. The process of claim 1, wherein the feed stream comprises less than 0.5 weight percent propylene. The method of claim 1, wherein the contacting step is performed at a pressure between 4,992 kPa and 10,445 kPa that is above a critical pressure for ethylene. Contacting the pyrolysis catalyst with a hydrocarbon feed stream to pyrolyze the hydrocarbon into a pyrolysis product hydrocarbon and coke is deposited on the pyrolysis catalyst to provide a coked pyrolysis catalyst;
Separating the coked pyrolysis catalyst from the pyrolysis product;
Adding oxygen to the coked pyrolysis catalyst;
Regenerating the pyrolysis catalyst by combusting coke on the coked pyrolysis catalyst with oxygen;
Separating the pyrolysis product to obtain a dilute ethylene stream comprising 5-50% by weight of ethylene;
Compressing the diluted ethylene stream to a pressure of 4,826-7,584 kPa; And
Contacting the diluted ethylene stream with an oligomerization catalyst
≪ / RTI > 8. The process of claim 7, wherein said contacting step is carried out in a fixed bed of said oligomerization catalyst. A fluid catalytic pyrolysis reactor in which a pyrolysis catalyst is contacted with a hydrocarbon feed stream to pyrolyze the hydrocarbon feed into a pyrolysis product and coke is deposited on the pyrolysis catalyst to provide a coked pyrolysis catalyst;
A product outlet for discharging said pyrolysis product from said reactor;
A regenerator for combusting coke from the coked pyrolysis catalyst in contact with oxygen;
A product recovery portion in communication with the product outlet for separating the pyrolysis product into a plurality of product streams comprising an ethylene containing stream;
A compressor in communication with the product recovery section for compressing the ethylene containing stream; And
Fixed bed oligomerization reactor in communication with the compressor for oligomerizing ethylene into bihydrocarbons in the ethylene containing stream
/ RTI > 10. The apparatus of claim 9, further comprising a main column receiver in communication with the product outlet, a primary adsorber in communication with the main column receiver for providing a primary offgas stream comprising the ethylene containing stream.

Images