DIRECT OLEFIN REDUCTION OF THERMALLY CRACKED HYDROCARBON STREAMS

Abstract

A process that catalytically converts olefinic (Alkenes, typically liquid at standard temperature and pressure) material in thermally cracked streams to meet olefin content specifications for crude oil transport pipelines. A thermally cracked stream or portion of a thermally cracked stream is selectively reacted to reduce the olefin content within a reactor operating at specific, controlled conditions in the presence of a catalyst and the absence of supplemental hydrogen. The process catalyst is comprised of a blend of select catalyzing metals supported on an alumina, silica or shape selective zeolite substrate together with appropriate pore acidic components.

Claims

1.-20. (canceled)

21. A process for producing an upgraded hydrocarbon product, comprising: supplying an olefin-containing bitumen stream to a catalytic reactor for contacting a catalyst material without the addition of supplemental hydrogen to convert olefins and produce a treated hydrocarbon stream with a reduced olefin content, the catalyst material comprising: a support material; and a catalytic metal material comprising: an olefin cracking metal catalyst to crack olefins into smaller hydrocarbon components; and a reforming metal catalyst for converting the smaller hydrocarbon components into longer-chain hydrocarbons by reaction pathways that include polymerization, cyclization, aromatization; withdrawing the treated hydrocarbon stream from the catalytic reactor.

22. The process of claim 21, further comprising cooling the treated hydrocarbon stream after withdrawal from the catalytic reactor to produce a cooled treated hydrocarbon stream, and separating the cooled treated hydrocarbon stream into a vapour stream and a liquid stream.

23. The process of claim 21, further comprising adding a supplementary stream comprising low carbon number molecules to the olefin-containing bitumen stream prior to supplying to the catalytic reactor.

24. The process of claim 23, wherein the low carbon number molecules comprise olefins.

25. The process of claim 23, wherein the low carbon number molecules comprise methane, ethane, ethylene, propane, propylene, butane or butylene or a combination thereof.

26. The process of claim 21, wherein the catalytic reactor comprises a vessel sized for flows between liquid hourly space velocities of 0.1 h.sup.−1 and 2 h.sup.−1.

27. The process of claim 21, wherein the catalytic reactor is operated between atmospheric pressure and 70 bar.

28. The process of claim 27, wherein the olefin cracking metal catalyst comprises silver and the reforming metal catalyst comprises gallium.

29. The process of claim 21, wherein the catalytic reactor is operated between 70 bar and 140 bar.

30. The process of claim 29, wherein the olefin cracking metal catalyst comprises silver and the reforming metal catalyst comprises platinum or palladium.

31. The process of claim 21, wherein the catalytic reactor is operated at temperatures between 300° F. and 662° F.

32. The process of claim 21, wherein the olefin-containing bitumen stream is in liquid phase when entering the catalytic reactor.

33. The process of claim 21, wherein the catalytic reactor comprises: a main catalytic bed comprising the catalyst material; and an upstream pre-treatment unit configured to remove contaminants, the upstream pre-treatment unit comprising a catalytic bed or an absorbent bed and being configured to remove at least sulfur-based molecules having deleterious effects on the catalyst material.

34. The process of claim 21, wherein the olefin cracking metal catalyst comprises at least one noble metal.

35. The process of claim 34, wherein the reforming metal catalyst comprises at least one platinum group metal or at least one post-transition metal, or a combination thereof.

36. The process of claim 35, wherein the at least one platinum group metal is selected from the group consisting of palladium and platinum.

37. The process of claim 35, wherein the at least one post-transition metal is gallium.

38. The process of claim 35, wherein the support material has acidic activity.

39. The process of claim 35, wherein the support material comprises alumina-based material, silica-based material or zeolite material or a combination thereof.

40. The process of claim 39, wherein the support material is formed as an extruded structure.

41. The process of claim 35, wherein the catalytic metal material is present in an amount of at least 0.1 wt % and less than 10 wt % on a total weight basis of the catalyst material.

42. The process of claim 21, wherein conversion of the olefins in the catalytic reactor is at least 75 wt % based on the total amount of olefins in the olefin-containing bitumen stream supplied into the catalytic reactor.

43. The process of claim 21, wherein conversion of the olefins is performed without supplemental hydrogen donor compounds added to the catalytic reactor.

44. The process of claim 21, wherein the catalytic reactor comprises: an inlet to introduce the olefin-containing bitumen stream; a reactor body in fluid communication with the inlet to receive the olefin-containing bitumen stream, the reactor body containing a flow distribution assembly and a fixed reactor bed comprising the catalyst material for flowing the olefin-containing bitumen stream to contact the reactor bed; and an outlet in fluid communication with the reactor body for removal of the treated hydrocarbon stream.

Description

BRIEF DESCRIPTION OF DRAWINGS

[0044] FIG. 1 depicts the integration of the direct olefin removal unit with a heavy hydrocarbon processing unit.

[0045] FIG. 2 depicts the streams and equipment associated with the direct olefin removal unit.

[0046] FIG. 3 depicts an embodiment of FIG. 2 with addition of a promoting stream to the direct olefin conversion reactions.

[0047] FIG. 4 depicts the use of two reactors in parallel to permit continuous operation.

DETAILED DESCRIPTION

Introduction

[0048] Olefins (alkenes) are readily produced by β-scission reactions during the pyrolysis of hydrocarbons as occurs in thermal cracking reactors. This produces a product with relatively high concentrations of olefins in the lighter hydrocarbon fraction (IBP-650° F.). Olefins, particularly in the naphtha fraction (IBP-350° F.), are an undesirable component of a crude feed to a refinery because of their high reactivity, which leads to polymerization, producing gums and sludge that can foul equipment. Because of the high reactivity and tendency to polymerize, pipelines have placed limits on olefin concentration and now require essentially zero olefin content of pipeline products (non-detectable which is <0.5 wt % with current measurement techniques).

[0049] Conventional practices to remove olefins from pyrolysis products, such as coker distillate, rely on hydrotreating the olefins to saturate the double bond. Hydrotreating is an oil refinery catalytic process in which hydrogen is contacted with petroleum intermediate product streams to remove impurities, such as oxygen (O), sulfur (S), nitrogen (N), or unsaturated hydrocarbons (olefins). Fixed bed hydrotreating is the industrial standard for reducing S and N in upgraded/refined products. To achieve the S and N reduction, this process basically converts all of the olefins in the hydrocarbon stream to a saturated product. However, this is an expensive approach to reduce the olefin content only of the naphtha and light distillate fractions. Hydrogen is expensive to either purchase or produce. The reactors are high pressure/high temperature vessels that are expensive, catalyst is expensive, and the operating costs are high. In addition, hydrogen generation for hydrotreating bitumen in upgrading and refining is a large contributor to the cost of the upgrading process and to the generation of greenhouse gas (GHG) emissions. The industry in general and bitumen upgraders specifically, would benefit from a non-hydrogen based technology that would efficiently remove olefins at a lower cost. Direct olefin reduction (DOR) is a process intensification technology concept advancement using the cyclization reaction pathway, as an example, that can both reduce costs and GHG emissions by removing the need to generate hydrogen.

[0050] The objective of the DOR is to reduce the cost of converting olefins through a different approach than hydrotreating cracked hydrocarbons generated during the heavy oil conversion process, by limiting the reaction to just olefins while minimizing sulfur and nitrogen species reactions and thus reducing the size and cost of the sulfur handling facilities. Olefin treatment is necessary as long as there is a restrictive olefin specification (<0.5 vol %) on transport pipelines.

[0051] Cyclization reactions are organic chemical transformations that yield cyclic products via conversion of molecules such as olefins, as an example, to produce mono- or polycyclic pipeline acceptable products. Because they are intramolecular transformations, they are often very rapid and selective. Selective reaction chemistry can be achieved at carbons bound to a variety of functional groups, and mechanisms such as catalysis to effect cyclization generation are numerous. The cyclization step usually involves the cleavage of a multiple bond. The two ends of the multiple bond constitute two possible sites of reaction. The broken bond provides an attractive location for another bound atom or unbound atom to bind as a cyclic product. Five- and six-membered rings are the most common products; formation of smaller and larger rings is rarely observed.

[0052] The carbon-carbon double bond of an olefin appears to be a stronger bond than a standard carbon-carbon or carbon-hydrogen bond with a high total bond energy (˜613 kj/mol) but this comprises of two single carbon-carbon bonds which alone have a bond energy of ˜347 KJ/mol. The bond energies from each single carbon-carbon bond are not additive when creating the double bond (<697 KJ/mol). As a result, the bond energy to break one of the carbon double bonds is ˜307 kJ/mol (on average) which is less than a single carbon-carbon bond and less than a carbon-hydrogen bond (˜413 KJ/mol). Selective breaking of the carbon-carbon double bond can be achieved with the appropriate process design (ex. catalyst type and pore size) and conditions (eg. temperature, pressure and flow rate). The reaction may be tailored by selection of catalysts and controlling energy input to the process.

[0053] Overall, three conditions must be met for efficient cyclization to take place: A method must be available to generate olefin bond breakage selectively on the catalyst substrate, atoms are attracted to generate non-multiple bonds, and all steps must be faster than any undesired side reactions such as olefin recombination or reaction with sulfur and/or nitrogen bearing bonds. Cyclization reaction conditions are often mild and functional group tolerance is high. However, the catalytic process needs to ensure that the relative rates of the various potential side reactions must be carefully controlled so that cyclization is favored. Side reactions need to be mitigated.

[0054] Formation of aromatics during the pyrolysis of hydrocarbons has been observed since the beginning of the utilization of pyrolysis of hydrocarbons. Nohara and Sakai have elucidated a proposed reaction mechanism to explain the observed chemistry as is shown in Note 1. The reaction is a Diels-Alder cyclization reaction involving olefins and diolefins produced during pyrolysis. Although the reaction does occur, it is not a primary pyrolysis reaction and is considered a secondary reaction. As a result, the conditions required to produce the aromatics are relatively extreme and the yield of aromatics from this reaction are low. However, the reaction does proceed at a measureable rate, but has been considered a novelty.

##STR00001##

[0055] Interestingly, this is the same fundamental chemistry that is the basis for the refinery process of naphtha reforming. Naphtha reforming is a standard refinery process used to convert low molecular weight alkanes into higher value aromatics for octane enhancement and feedstocks for the petro chemical industry. The process occurs in the presence of a noble metal catalyst and a high partial pressure of hydrogen. Three main reactions occur that generate hydrogen creating the high hydrogen partial pressure, the dehydrogenation of naphthenes, isomerization of paraffins and dehydrogenation/aromatization of paraffins. A fourth reaction, the hydrocracking of paraffins also occurs due to the presence of generated hydrogen. Typically the feed to refinery reformers are low in sulfur, with a desulfurization step upstream of the unit, and contain very low concentrations of olefins.

[0056] The invention of this application can use similar catalyst but differs in that the target feedstock contains large concentrations of olefins (˜20 wt %) and the process may tolerate higher sulfur and nitrogen feedstocks.

[0057] Understanding and application of the reforming chemistry was further expanded during the 1980's. Tetra-ethyl lead was removed from gasoline as an octane booster for environmental reasons and without tetra-ethyl lead, the octane pool was not sufficient to produce the needed gasoline. Efforts were made to develop technologies to produce an octane booster, in particular benzene, toluene and xylene (BTX). A readily available feedstock for aromatization was liquefied petroleum gas (LPG). During the 1980's a great deal of research was conducted to understand the chemistry of the aromatization reactions and the effectiveness of various catalysts.

[0058] This chemistry has resulted in the development of three commercial processes that are indicated to be applicable to both naphtha reforming and LPG aromatization. The commercial processes are: [0059] M-2 Forming (Mobil) [0060] Cyclar (BP-UOP) [0061] Aroforming (IFP-Salutec)

[0062] The chemistry of the aromatization of alkanes was elucidated with model compounds and since the emphasis at that time was on aromatization of LPG, most of the work was performed on propane, butane and hexane because they are representative of the components in LPG. However, this chemistry is also applicable to the higher molecular weight components in naphtha. The chemical mechanism for the aromatization is shown below in Note 2.

[0063] The reaction begins with step 1, dehydrogenation of an alkane (in the case of Note 2—propane) to form an olefin. Unlike pyrolysis chemistry, which is free radical chemistry, this dehydrogenation occurs by carbenium ion chemistry rather than by β-scission. Steps 2 and 4 in the reaction mechanism involve transient intermediates that first react the olefins to yield oligomers, quite possibly the reaction involves some di-olefins. The oligomers undergo cyclization reactions in step 4 to produce alicyclics like cyclohexane. The alicyclics undergo dehydrogenation in step 5 to yield the desired aromatics.

##STR00002##

[0064] Step 3 is a form of the reverse reaction of step 2 and is a cracking type of reaction. However, the reaction is not limited to producing only C.sub.3 olefins as other carbon numbered olefins can also be produced. Step 2 will convert these olefins back into the oligomers so the forward reaction (4) is favored. It is interesting to note that the dehydrogenation reaction (step 5) is not reversible. Therefore, the production of the aromatics is not equilibrium controlled.

[0065] The reaction mechanism presented in Note 2 uses the olefin (propene) as the intermediate to yield the oligomers for cyclization and aromatization. This being the case, then substituting an olefin in place of an alkane as the feedstock should favorably help the reaction and be more applicable to the problem at hand. When propene is used as the feedstock in place of propane the conversion of propene is essentially complete and the reaction is about 20 times faster than with propane. Substituting hexene for hexane in the reaction results in a rapid and selective conversion of hexene into aromatic products with practically no cracking products. Therefore, since the current application is targeting the conversion of olefins in the feedstock with up to 20 wt % concentration in the feed, the chemistry is favorable to the goals of this invention, and differs from the goals and design of refinery reforming where olefins in the feed are discouraged.

[0066] In addition, using olefins, rather than alkanes, is energetically more attractive. Table 1 provides the temperature for conversion of several alkanes to either alkenes or aromatics without a catalyst. As can be observed from the results in the Table 1, the temperature to convert any of the alkanes to an olefin is much higher than the temperature required to convert the alkane to benzene. In addition, as the number of carbon atoms in the alkane increase, the required temperature decreases.

TABLE-US-00001 TABLE 1 Temperature requirements to convert selected alkanes to aromatics or olefins (Scurrell, M. S., 1988, “Factors Affecting the Selectivity of the Aromatization of Light Alkanes on Modified ZSM-5 Catalysts”, Applied Catalysis, 41 pg. 89-98.). Temperature required for ΔG.sub.t.sup.0 = 0 (° C.) Alkane Conversion to benzene Conversion to alkene Methane 1075 1350 (Ethene) Ethane 575 774 (Ethene) Propane 450 655 (Propene) n-Hexane 320 575 (2-Methylpropene)

[0067] In terms of the interest to convert olefins to aromatics, the results in Table 1 may be interpreted to mean that the conversion of olefins to aromatics should proceed more easily than conversion of the alkane to the olefin. The results discussed above show this to be the case and the results in Table 1 suggest that the high energy step in the reaction shown in Note 2 is the formation of the olefin in step 1 (high activation energy) and is likely the rate limiting step in the overall reaction.

[0068] The aromatization reactions are performed over a catalyst, often ZMS-5 zeolite catalyst with exchange heavy metal cations (zinc, platinum or gallium). The zeolite catalysts are known to be acid catalysts acting primarily as a Brönsted acid. In the aromatization of alkanes, the acid functionality of the zeolite catalyst is required to form an olefin from the alkene through the carbenium ion chemistry. The acidic activity is native to the catalyst and as a result using the zeolite as the basis for the catalyst will impart some acidic activity.

[0069] The exchange metal cations are bi-functional and participate in the dehydrogenation of the alkane (step 1 in Note 2) and the dehydrogenation of alicyclic intermediates to yield aromatics (step 5 in Note 2). Studies comparing the activity of ZSM-5-H with ZSM-5-Zn show that the addition of the Zn is advantageous for the reaction to yield high levels of aromatics. The general consensus is that zinc (Zn) functions to provide the activity to initiate reaction step 1 and does quite well providing the activity for dehydrogenation to yield aromatics.

[0070] Platinum (Pt) has also been shown to have a high degree of dehydrogenation activity. However, it also appears that Pt also has significant cracking and de-alkylation activity and yields of cracked products (step 3 in Note 2) are increased. In addition, the yield of benzene is increased at the expense of C8 aromatics (xylenes). The activity of Pt appears to catalyze significant retrograde reactions that equate to carbon rejection for the process. For this reason, Pt is not frequently used in commercial aromatization processes. Other platinum group metals such as ruthenium, rhodium, palladium, osmium, or indium may provide less cracking and could be used for the aromatization reactions. Gallium (Ga) has also been used as the exchange metal cation in the zeolite catalysts. The activity of Ga has been shown to be similar to that of Zn and Pt in that it promotes step 1 (Note 2) as well as being active in dehydrogenation of the alicyclic components (step 5 in Note 2). However, the alicyclic dehydrogenation activity of Ga appears to be higher than that observed for the other two metals. Gallium is not plagued with volatility or de-alkylation issues and is a metal preferred for use in the commercial catalyst. Other post-transition metals such as tin, aluminum or lead could provide similar benefits and could be substitutes to gallium.

[0071] Certain noble metals (silver, gold), and transition metals (cobalt, nickel, copper, zinc) have been used as catalysts for various degrees of cyclization, aromatization and isomerazation reactions and could be candidates for direct olefin conversion.

[0072] The conversion of a number of C6 isomers has been studied over the aromatization catalysts and provides some insight into the application of this chemistry to conversion of olefins. The results from this study found that the decreasing order of reactivity of the isomers was: n-hexane>3-methylpentane>2,2-dimethylbutane. This order of reactivity is the reverse of that expected based on classical carbonium ion chemistry. However, the order of reactivity does imply limitations in the access of the reactant to the catalyst.

[0073] Catalyst activity resides on both the surface of the catalyst and within the pore structure. When the catalyst is “fresh”, the reactivity of the branched isomers is high, indicating probable conversion on the surface. As the catalyst ages, the reactivity decreases as expected for surface activity. However, the activity does not cease and suggests a high degree of diffusion control of the reaction based on the capability of the branched isomers to enter the pore structure of the catalyst. This suggests that control of the pore size of catalysts selected for olefin aromatization will need to be tailored because deactivation at the surface will occur before deactivation within the pore structure. Then the reaction rate will be diffusion controlled. Another consideration for pore size in this application is that the product molecules will tend to have different shapes and could be larger than the feed molecules and could be trapped in the catalyst if the pore size selection does not account for the geometry of the reaction product molecules.

[0074] Cyclization reactions, such as the Diels-Alder reaction mechanism, and aromatization reactions, which convert alkanes or olefins to cyclo-paraffins and aromatic species respectively have been observed and studied in pyrolysis of hydrocarbons. This chemistry has been extended to petroleum refining and aromatization is the fundamental chemistry utilized for naphtha reforming and conversion of liquefied petroleum gas (LPG) to benzene, toluene, and xylene (BTX). The aromatization chemistry has been successfully applied to these processes. However, to date, reference to the cyclization and aromatization chemistry for the reduction of olefin content in olefin rich naphtha and kerosene (light distillate) has not been found.

[0075] Another instance of olefin conversion can result from isomerization reactions which are a more selective form of cyclization (i.e. cycloisomerization) in which cracked hydrocarbon molecules are transformed into molecules that have the same number of atoms but the atoms have a different arrangement. The resulting arrangement does not include any double bonds (no olefins) and the molecules can be linear in nature or cyclical.

[0076] Table 2 shares the average heat of reactions for the typical hydrogen-based catalytic conversion reactions with the reactions organized from “least energy” required to the “most energy” required. Olefin saturation reactions require the least energy so a simple reaction process strictly for olefin removal such as cyclization or cycle-isomerization can be targeted in lieu of commercial hydrotreating.

TABLE-US-00002 TABLE 2 Average Heat of reaction for hydrotreating reactions (Tarhun 1983, Gary & Handwerk 2001) Average Heat of reaction Reaction J/mol Olefin Saturation −340,000 Mild Hydrodesulfurization −251,000 (mercaptans, Sulfides) Hydrogenation (di, tri- aromatic rings) −125,500 Medium hydrodesulfurization −95,000 (disulfides, thiophenes) Hydrodemetallization −72,000 Hydrodeoxygenation −68,000 Hydrodenitrogenation −64,650 Severe hydrodesulfurization −58,000 hydrodearomatization −53,000 (mono, di, tri aromatic rings) Hydrocracking −41,000

[0077] By only promoting olefin conversion reactions (breaking the carbon-carbon double bond), while having no effect on the other hydrocarbon molecules, less byproducts such as hydrogen sulfide (H.sub.2S) are generated compared to hydrotreating processes, thereby decreasing the complexity and cost of the facility that uses an olefin removal unit instead of a hydrotreater. In most cases, refiners do not pay the full value for sulfur and nitrogen removal in their purchased heavy hydrocarbon feeds so hydrotreating prior to transport to the refinery does not appear to provide the necessary benefit for the cost. Heavy hydrocarbon conversion facilities such as partial and conventional upgraders would realize significant cost savings and process simplification from a targeted direct olefin removal technology to meet crude pipeline specification along with exhibiting a lower greenhouse gas footprint.

DESCRIPTION OF VARIOUS EMBODIMENTS

[0078] The detailed description set forth below in connection with the appended drawings is intended as a description of various embodiments of the present invention and is not intended to represent the only embodiments contemplated by the inventor. The detailed description includes specific details for the purpose of providing a comprehensive understanding of the present invention. However, it will be apparent to those skilled in the art that the present invention may be practiced without these specific details.

[0079] Direct olefin reduction is a process intensification/simplification of hydrotreating, removing the hydrogen production and consumption steps, that is used to remove olefins and other molecules from hydrocarbons that were generated during hydrocarbon processing. Instead of using hydrogen and a catalyst to remove olefins as in conventional hydrotreating, a specific catalyst is only used on the target hydrocarbon stream to preferentially convert olefins to non-olefinic hydrocarbons such as cycloparaffins, napthalenes and aromatics. FIG. 1 shows a specific application for the direct olefin removal unit 20, as a supporting technology to a heavy hydrocarbon conversion unit 10. Examples of unit 10 that can be integrated with unit 20 are heavy oil and bitumen upgraders, heavy oil and bitumen partial upgraders and any crude refining entity interested in a lower cost solution with the objective of olefin conversion only.

[0080] A heavy hydrocarbon stream, 1, is sent to a heavy hydrocarbon processing unit, 10, which can comprise of one or a plurality of conversion and/or separation operating units, where two product streams are generated. The heavier hydrocarbons exit unit 10 as stream 13 and are re-blended with the olefin reduced stream 29 to make a final product, stream 30 for pipeline transport to a downstream hydrocarbon processing customer (refinery). The lighter hydrocarbons with a concentration of olefinic material exit unit 10 as stream 12. Stream 12 is directed to the direct olefin removal unit 20. As shown in FIG. 2, if required, stream 12 passes through a heat exchanger, 21, to obtain the necessary operating temperature for the feed, stream 22, then to the direct olefin removal reactor 23. Reactor 23 contains a fixed bed with one or more of: platinum group metals (ruthenium, rhodium, palladium, osmium, indium and platinum), noble metals (silver, gold), transition metals (cobalt, nickel, copper, zinc) and post-transition metals (gallium, lead, tin, aluminum), the catalysts formed with silica or shape selective substrate (ex. Zeolite class) support. The olefin-rich stream interacts with the catalyst so that the olefinic bonds are selectively converted reducing the concentration of olefins from stream 22 to a point that when output stream 29 is blended with stream 13, this results in a final product stream 30 which satisfies pipeline specification for olefins. The non-olefinic molecules present in the reactor are not reacted or are minimally converted from passing through the fixed catalyst bed. For example, since the non-hydrogen addition cyclization reaction is endothermic (hydrotreating olefin saturation is mildly exothermic), the reactor may have supplemental heat added to maintain a desired operating temperature. In practice, two to three reactors in parallel can be provided to facilitate regeneration of catalyst by switching between reactors so continuous operation can be maintained. Product stream 24 exits reactor 23 with a significantly reduced concentration of olefins. Stream 24 is cooled in heat exchanger 25 creating a specific two-phased flow stream 26 ready for downstream separation. Stream 26 is sent to a separation vessel where the hydrocarbons that are suitable for the final blend stream 30 to meet pipeline specification are condensed to the liquid phase to become stream 29. Stream 29 blends with stream 13 in FIG. 1 to become the final product stream 30, a crude product that can be sent via pipeline to customers. The light hydrocarbon gases generated from the reactor 23 are removed as vapour from vessel 27 as stream 28.

[0081] In instances when the heavy hydrocarbon stream contains contaminants that can adversely affect the performance of the catalyst, such as sulfur-based, nitrogen-based, inorganic-based and/or other deactivating molecules, an absorbent or catalytic bed (e.g. guard bed) can be placed upstream of the main catalytic bed. The guard bed can be housed in the same reactor vessel as the main catalyst or placed in a separate vessel upstream for ease of regeneration.

[0082] Low carbon number molecules (ex. methane, ethane, propane, butane) are generated in the thermal cracking step and form part of the olefinic feedstock to the direct olefin conversion unit. These low carbon number molecules can enhance the cyclization reaction. In another embodiment, as shown in FIG. 3, a dedicated stream, 31, of a mixture of low carbon number molecules such as methane, ethane, ethylene, propane, propylene, butane and butylene are added to the feedstream 22, to create a stream 32 feeding reactor 23 to supplement the existing low carbon number molecules in stream 22 to enhance the olefin conversion reactions in reactor 23.

EXAMPLE 1

[0083] A fluid stream 12 as referenced in FIG. 1, comprising naphtha and distillate boiling range material with 20 wt % olefin content was placed in a fixed bed reactor with catalyst containing 0.1 wt % of a silver and gallium mix on a zeolite substrate. With operating conditions at 350° C. and 15 bar, and a residence time of 1 hour, the olefin conversion was over 75 wt % with a 98 wt % mass liquid yield while the demetallization, desulfurization and denitrification reactions were negligible. The product liquid from this fixed bed reactor representing stream 29 in FIG. 1 when mixed with stream 13 in FIG. 1 would result in a stream 30, that meets pipeline olefin specification.

EXAMPLE 2

[0084] A fluid stream 12 as referenced in FIG. 1, comprising naphtha and distillate boiling range material with 20 wt % olefin content was placed in a fixed bed reactor comprising catalyst containing 0.1 wt % of platinum on an extruded alumina oxide cylinder. With operating conditions at 300° C. and 70 bar, and a residence time of 1 hour, the olefin conversion was over 98 wt % with a 98% mass yield while the demetallization, desulfurization and denitrification reactions were negligible. The product liquid from this fixed bed reactor representing stream 29 in FIG. 1 when mixed with stream 13 in FIG. 1 resulted in stream 30 that met pipeline olefin specification.