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Title:
PROCESS FOR THE PRODUCTION OF OLEFINS BY AUTOTHERMAL CRACKING
Document Type and Number:
WIPO Patent Application WO/2006/117509
Kind Code:
A1
Abstract:
The present invention relates to a process for the production of an olefin, said process comprising passing a feedstream which comprises a paraffinic hydrocarbon, hydrogen and an oxygen-containing gas through a catalyst zone which is capable of supporting combustion beyond the fuel rich limit of flammability to produce said olefin, said catalyst zone comprising at least a first catalyst bed which comprises platinum and palladium and wherein the feedstream comprises at least 0.5% by volume of the total feedstream of carbon monoxide.

Inventors:
MESSENGER BRIAN EDWARD (GB)
REID IAN ALLAN BEATTIE (GB)
Application Number:
PCT/GB2006/001449
Publication Date:
November 09, 2006
Filing Date:
April 20, 2006
Export Citation:
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Assignee:
INEOS EUROPE LTD (GB)
MESSENGER BRIAN EDWARD (GB)
REID IAN ALLAN BEATTIE (GB)
International Classes:
C07C5/48; C07C11/04; C10G47/14
Domestic Patent References:
WO2002004389A12002-01-17
WO2004106463A12004-12-09
WO2000014035A22000-03-16
Foreign References:
EP0332289A21989-09-13
Attorney, Agent or Firm:
COMPASS PATENTS LLP (Chertsey, Surrey KT16 8LA, GB)
Download PDF:
Claims:
10054(2) 20Claims:
1. A process for the production of an olefin, said process comprising passing a feedstream which comprises a paraffϊnic hydrocarbon, hydrogen and an oxygencontaining gas through a catalyst zone which is capable of supporting combustion beyond the fuel rich limit of flammability to produce said olefin, said catalyst zone comprising at least a first catalyst bed which comprises platinum and palladium and wherein the feedstream comprises at least 0.5 % by volume of the total feedstream of carbon monoxide.
2. A process according to claim 1, wherein the catalyst zone comprises at least a first catalyst bed and a second catalyst bed, wherein the second catalyst bed is located downstream of the first catalyst bed, is of a different composition to the first catalyst bed and comprises a Group VHI metal.
3. A process as claimed in claim 2, wherein the Group VIII metal of the second catalyst bed in platinum.
4. A process as claimed in claim 2 or claim 3 wherein the Group VIII metal loading for the second catalyst bed is from 0.01 to 10 wt % based on the total weight of the catalyst.
5. A process as claimed in any one of claims 2 to 4 wherein the second catalyst bed comprises an unpromoted Group VHI metal catalyst.
6. A process according to any one of the preceding claims, said process comprising: (i) passing a feedstream which comprises a paraffinic hydrocarbon, hydrogen, an oxygen containing gas and at least 0.5% by volume of the total feedstream of carbon monoxide through a catalyst zone which is capable of supporting combustion beyond the fuel rich limit of flammability, said catalyst zone comprising at least a first catalyst bed which comprises platinum and palladium, to produce a product stream comprising one or more olefins, hydrogen and carbon monoxide, (ii) separating at least a portion of the hydrogen and carbon monoxide in said product stream as a recycle stream, and (iii) mixing said recycle stream with one or more paraffinic hydrocarbons and an oxygen containing gas to form the feedstream passed to step (i).
7. A process according to any one of the preceding claims wherein the feedstream comprises 110% by volume of the total feedstream of carbon monoxide.
8. A process according to any one of the preceding claims wherein the platinum loading of the first catalyst bed is from 0.01 to 10 wt% based on the total weight of the catalyst. 10054(2) 21 .
9. A process according to any one of the preceding claims wherein the palladium loading of the first catalyst bed is from 0.010.5 wt% based on the total weight of the catalyst.
10. A process according to any one or the preceding claims wherein one or more unsaturated hydrocarbons is cofed to the catalyst zone.
Description:
PROCESS FOR THE PRODUCTION OF OLEFINS BY AUTOTHERMAL CRACKING

The present invention relates to a process for the production of olefins from paraffinic hydrocarbons in which the paraffinic hydrocarbons undergo autothermal cracking. Autothermal cracking is a route to olefins in which a paraffinic hydrocarbon feed is mixed with oxygen and passed over an autothermal cracking catalyst. The autothermal cracking catalyst is capable of supporting combustion beyond the fuel rich limit of fiammability. Combustion is initiated on the catalyst surface and the heat required to raise the reactants to the process temperature and to carry out the endothermic cracking process is generated in situ.

The autothermal cracking process is described in EP 332289B; EP-529793B; EP-A- 0709446 and WO 00/14035.

Generally the paraffinic hydrocarbon feed and the oxygen is passed over a single catalyst bed to produce the olefin product. Typically, the catalyst bed comprises at least one platinum group metal, for example, platinum, supported on a catalyst support.

In more recent work, catalyst zones comprising two or more catalyst beds have been described, such as in WO 02/04389 and WO 2004/106463.

WO 02/04389, for example, has shown that the selectivity of a catalyst zone comprising a first catalyst bed can be enhanced by positioning a second catalyst bed downstream of the first catalyst bed, even when said second catalyst bed is a catalyst which is substantially incapable of supporting combustion beyond the fuel rich limit of fiammability (that is, a catalyst which is substantially inactive under autothermal cracking conditions).

It has now been found that a catalyst zone comprising at least a first catalyst bed which comprises platinum and palladium can provide superior performance to catalyst beds comprising platinum alone. hi particular, the presence of palladium in the first catalyst bed has been found to provide improved tolerance of the catalyst to carbon monoxide present in the feed to the autothermal cracking reaction. Accordingly, in a first aspect the present invention provides a process for the production of an olefin, said process comprising passing a feedstream which comprises a paraffinic hydrocarbon, hydrogen and an oxygen-containing gas through a catalyst zone

10054(2) 2

which is capable of supporting combustion beyond the fuel rich limit of flammability to produce said olefin, said catalyst zone comprising at least a first catalyst bed which comprises platinum and palladium and wherein the feedstream comprises at least 0.5% by volume of the total feedstream of carbon monoxide. The "catalyst zone comprising at least a first catalyst bed" may be a catalyst zone comprising a single catalyst bed which comprises platinum and palladium or a catalyst zone comprising more than one catalyst bed wherein the first catalyst bed comprises platinum and palladium.

The first catalyst bed comprises platinum and palladium. As noted above, palladium has been found to provide improved tolerance to carbon monoxide in the feed to the autothermal cracking reaction, hi an industrial process, carbon monoxide can be present as a component of any recycle streams and/or in any hydrogen feed. It has now been found that activity and selectivity of platinum only catalysts can decrease markedly in the presence of carbon monoxide, and can even be extinguished at relatively low levels of carbon monoxide, but that addition of palladium provides tolerance to carbon monoxide. Because of the tolerance to carbon monoxide of the catalyst zone, and hence the process, of the present invention, carbon monoxide maybe "deliberately" fed, either as a separate stream, by use of a lower purity (and hence cheaper) hydrogen source or by operating separation steps that lead to recycle streams with higher tolerance on carbon monoxide levels.

In a preferred embodiment the carbon monoxide in the feedstream derives from a recycle stream comprising hydrogen and carbon monoxide which has been separated from the product stream from the autothermal cracking reaction itself, hi this embodiment it is possible to recycle hydrogen from the product stream without treatment or with reduced treatment to separate the carbon monoxide therein, avoiding downstream separation steps and their associated costs.

Thus, in this embodiment, the present invention provides a process for the production of an olefin, said process comprising

(i) passing a feedstream which comprises a paraffinic hydrocarbon, hydrogen, an oxygen-containing gas and at least 0.5% by volume of the total feedstream of carbon monoxide through a catalyst zone which is capable of supporting combustion beyond the fuel rich limit of flammability, said catalyst zone comprising at least a first

10054(2) 3

catalyst bed which comprises platinum and palladium, to produce a product stream comprising one or more olefins, hydrogen and carbon monoxide, (ii) separating at least a portion of the hydrogen and carbon monoxide in said product stream as a recycle stream, and (iii) mixing said recycle stream with one or more paraffinic hydrocarbons and an oxygen- containing gas to form the feedstream passed to step (i).

Carbon monoxide and hydrogen in the feedstream may be provided solely (or essentially solely) from recycle of carbon monoxide and hydrogen from the reaction product stream. Alternatively, carbon monoxide and hydrogen recycled from the reaction product stream may be supplemented by carbon monoxide and/or hydrogen from other sources, such as "fresh" hydrogen feed.

Typically, the recycle stream may comprise up to 20vol% carbon monoxide in hydrogen.

Typically, the feedstream to the process of the present invention may comprise at least 1% by volume of the total feedstream, for example, 1-10% by volume of the total feedstream, of carbon monoxide.

Typical platinum loadings of the first catalyst bed range from 0.01 to 99.99 wt %, preferably, from 0.01 to 20 wt %, and more preferably, from 0.01 to 10 wt %, for example 1-5 wt% based on the total weight of the catalyst. Relatively low palladium loadings are preferred for the first catalyst bed. Low levels have been found to be sufficient to provide the desired tolerance. Higher levels of palladium, however, can lead to significant levels of undesired side reactions, including increased formation of methane and carbon monoxide. Preferably, therefore, the palladium loading of the first catalyst bed is less than 5 wt%, more preferably less than 1 wt% based on the total weight of the first catalyst bed.

Most preferably, the palladium loading of the first catalyst bed is from 0.01 to 0.5 wt %, such as from 0.01 to 0.3 wt %, for example from 0.02 to 0.2 wt% based on the total weight of the first catalyst bed.

As an alternative to or in addition to having a relatively low palladium loading of the first catalyst bed, it is also possible to mitigate the propensity of the presence of palladium to lead to undesired side reactions, by utilising a catalyst zone comprising a first catalyst bed comprising platinum and palladium with a second, different, catalyst bed downstream

10054(2) 4

of the first catalyst bed. This allows a relatively short first catalyst bed to be utilized to provide the desired tolerance to carbon monoxide in the feedstream with a subsequent catalyst bed or beds selected to provide improved activity or selectivity for olefin production. Accordingly, in a second aspect the present invention provides a process for the production of an olefin, said process comprising passing feedstream which comprises a paraffinic hydrocarbon, hydrogen and an oxygen-containing gas through a catalyst zone which is capable of supporting combustion beyond the fuel rich limit of flammability to produce said olefin, said catalyst zone comprising at least a first catalyst bed and a second catalyst bed, said second catalyst bed being located downstream of the first catalyst bed, wherein the first catalyst bed comprises platinum and palladium and wherein the second catalyst bed is of a different composition to the first catalyst bed and either (i) comprises at least one metal selected from the group consisting of Mo, W, and Group EB, IIB, IIIB, rVB, VB, VIIB and VIII of the Periodic Table, or (ii) has the general formula of: wherein M 1 is selected from groups IIA, IIB, IIIB, IVB, VB, VIB, VDB, lanthanides and actinides, M 2 is selected from groups IIA, IB, IIB, IIIB, IVB, VB, VIB, M 3 is selected from groups IIA, IB, IIB, IIIB, IVB, VB, VIB and VIIIB, a, b, c and z are the atomic ratios of components M , M , M and O respectively, a is in the range of 0.1 to 1.0, b is in the range of 0.1 to 2.0, c is in the range of 0.1-3.0, and z is in the range 0.1 to 9.

(As used herein the groups of the Periodic Table are referenced using the CAS notation, as listed in Advanced Inorganic Chemistry, Fifth edition, 1988, by Cotton and Wilkinson.) hi a preferred embodiment of this second aspect the feedstream comprises at least

0.5% by volume of the total feedstream of carbon monoxide, preferably at least 1% by volume of the total feedstream of carbon monoxide, such as 1 to 10% by volume of the total feedstream, of carbon monoxide. The carbon monoxide may derive from any suitable source, but preferably derives from a recycle stream comprising hydrogen and carbon monoxide which has been separated from the product stream from the autothermal cracking reaction itself as described for the first aspect.

10054(2) 5

The first catalyst bed according to the process of the (first or second aspect of the) present invention may also comprise a promoter. Suitably, the promoter may be selected from the elements of Groups IHA, IVA and VA of the Periodic Table and the transition metals (other than platinum and palladium) and mixtures thereof. Preferred Group IIIA metals include Al, Ga, Li and Tl. Of these, Ga and In are preferred. Preferred Group IVA metals include Ge, Sn and Pb. Of these, Ge and Sn are preferred, especially Sn. The preferred Group VA metal is Sb.

Examples of suitable transition metal promoters include Cr, Mo, W, Fe, Ru, Os, Co, Rh, Ir, Ni, Cu, Ag, Au, Zn, Cd and Hg. Preferred transition metal promoters are Mo, Rh, Ru, Ir, Cu and Zn, especially Cu.

The atomic ratio of platinum metal to the promoter metal(s) may be 1 : 0.1 - 50.0, preferably, 1: 0.1 - 12.0, such as 1 : 0.3 -5.

The second catalyst bed, when present, may be any suitable second catalyst bed which either: (i) comprises at least one metal selected from the group consisting of Mo, W, and Group

BB, IIB, mB, ΓVB, VB, VID3 and VTII of the Periodic Table, (ii) or has the general formula of:

M 1 ^bM 3 C O 2 wherein M 1 is selected from groups IIA, IIB, IIIB, IVB, VB, VD3, VHB, lanthanides and actinides, M 2 is selected from groups IIA, EB, IIB, IIIB, IVB, VB, VIB, M 3 is selected from groups IIA, IB, IIB, IIIB, IVB, VB, VIB and VIIIB, a, b, c and z are the atomic ratios of components M 1 , M 2 , M 3 and O respectively, a is in the range of 0.1 to 1.0, b is in the range of 0.1 to 2.0, c is in the range of 0.1-3.0, and z is in the range 0.1 to 9.

The catalysts of (i) may be any of those of WO 02/04389. For example, the second catalyst bed may comprise one or more dehydrogenation catalysts. Dehydrogenation catalysts are those catalysts which are capable of converting paraffmic hydrocarbons to olefins, but are substantially incapable of causing partial combustion of the paraffϊnic hydrocarbon feed to olefin under auto-thermal conditions. Thus, suitably, the second catalyst bed may comprise at least one metal selected from Fe, Ru, Os, Co, Ir, Ni, Mo, W, and Groups IB, IIB, IIIB, IVB, VB and VIIB of the Periodic

Table. Specific examples of such metals include Cu, Ag, Au, Zn, Cd, Hg, Sc, Y La, Ti, Zr,

10054(2) 6

Hf, V, Nb, Ta, Ni, Co, Lr and mixtures thereof, especially Cu, Co, Ni, Ir and mixtures thereof.

Specific examples of suitable catalysts (and which are substantially incapable of causing partial combustion of the paraffinic hydrocarbon feed to olefin under auto-thermal conditions) for use as the second catalyst bed include Ni/Sn, Co/Sn, Co/Cu, Cu/Cr, Ir/Sn, Fe/Sn, Ru/Sn, Ni/Cu, Cr/Cu, Ir/Cu, Fe/Cu and Ru/Cu.

The catalysts of (ii) may be any of those of WO 2004/106463.

Preferably M 1 is selected from group IIIB, M 2 is selected from group IIA and M 3 is selected from group IB. Most preferably M is yttrium, M is barium and M is copper. Preferably the catalyst (ii) is in the form of a perovskite-type structure. Perovskite- type structures include yttrium-barium-copper oxides YBa 2 Cu 3 O 7-S , lanthanum-strontium- iron oxides La 1-x Sr x Fe0 3- δ , and lanthanum-manganese-copper oxides LaMn 1-x Cu x O 3-δ , wherein x is in the range of 0.1-0.9 and δ is typically in the range of 0.01-1, preferably in the range 0.01-0.25. The catalyst (ii) may be promoted by addition of halide-promoters to yield materials of having the general formula of; wherein M 5 M and M and a, b, c and z are as herein described above, X is a halide, preferably F or Cl, and x is typically in the range of 0.05-0.5. Most preferably, the second catalyst bed, when present, comprises one or more

Group VIII metals. Preferably, the Group VIII metal is platinum.

Typical Group VIII metal loadings for the second catalyst bed range from 0.01 to

100 wt %, preferably, from 0.01 to 20 wt %, and more preferably, from 0.01 to 10 wt %, for example 1-5 wt% based on the total weight of the catalyst. The second catalyst bed, when present, may comprise an unpromoted Group VIII metal catalyst. Alternatively, the second catalyst bed may comprise a promoted Group • VIII metal catalyst. Suitably, the promoter may be selected from the elements of Groups πiA, rVA and VA of the Periodic Table and the transition metals (other than Group VIII metals) and mixtures thereof. Preferred Group IDA metals include Al, Ga, hi and Tl. Of these, Ga and hi are preferred. Preferred Group IVA metals include Ge, Sn and Pb. Of these, Ge and Sn are preferred, especially Sn. The preferred Group VA metal is Sb.

10054(2) 7

Examples of suitable transition metal promoters include Cr, Mo, W, Cu, Ag, Au, Zn, Cd and Hg. Preferred transition metal promoters are Mo, Cu and Zn, especially Cu.

The atomic ratio of Group VIII metal to the promoter metal(s) may be 1 : 0.1 - 50.0, preferably, 1: 0.1 - 12.0, such as 1 : 0.3 -5. Specific examples of promoted Group VIII catalysts for use as the second catalyst bed include Pt/Ga, Pt/In, Pt/Sn, Pt/Ge, Pt/Cu.

In addition to first and second catalyst beds the catalyst zone may comprise further catalyst beds. For example, the catalyst zone may comprise 3 to 10, preferably, 3 to 5 catalyst beds. Where the catalyst zone comprises more than two catalyst beds, the catalyst of the further bed(s) may be the same as the catalysts used for the second catalyst bed or may be alternative second catalyst beds within the definitions (i) and (ii) above.

Each catalyst employed in the catalyst zone may be unsupported or supported. Suitably, an unsupported catalyst may be in the form of a metal gauze. Preferably, at least one catalyst in the catalyst zone is a supported catalyst. Suitably, each catalyst in the catalyst zone is a supported catalyst. The support used for each catalyst may be the same or different. Although a range of support materials may be used, ceramic supports are generally preferred. However, metal supports may also be used.

Suitably, the ceramic support may be any oxide or combination of oxides that is stable at high temperatures, for example, stable between 600°C and 1200°C. The ceramic support material preferably has a low thermal expansion co-efficient, and is resistant to phase separation at high temperatures.

Suitable ceramic supports include cordierite, lithium aluminium silicate (LAS), alumina (alpha- Al 2 O 3 ), yttria stabilised zirconia, aluminium titanate, niascon, and calcium zirconyl phosphate, and, in particular, alumina.

The ceramic support may be wash-coated, for example, with gamma- Al 2 O 3 .

The structure of the support material is important, as the structure may affect flow patterns through the catalyst. Such flow patterns may influence the transport of reactants and products to and from the catalyst surface, thereby affecting the activity and selectivity of the catalyst. Typically, the support material may be in the form of particles, such as spheres or other granular shapes or it may be in the form of a foam or fibre such as a fibrous pad or mat. Suitably, the particulate support material may be alumina spheres.

10054(2) . , 8

Preferably, the form of the support is a monolith which is a continuous multi-channel ceramic structure. Such monoliths include honeycomb structures, foams, or fibrous pads. The pores of foam monolith structures tend to provide tortuous paths for reactants and products. Such foam monolith supports may have 20 to 80, preferably, 30 to 50 pores per inch. Channel monoliths generally have straighter, channel-like pores. These pores are generally smaller, and there may be 80 or more pores per linear inch of catalyst. Preferred ceramic foams include alumina foams.

Alternatively, the support may be present as a thin layer or wash coat on another substrate. Where a supported catalyst is employed, the metal components of the catalyst are preferably distributed substantially uniformly throughout the support.

The catalysts employed in the present invention may comprise further elements, such as alkali metals. Suitable alkali metals include lithium, sodium, potassium and cesium. The catalysts employed in the present invention may be prepared by any method known in the art. For example, gel methods and wet-impregnation techniques may be employed. Typically, the support is impregnated with one or more solutions comprising the metals, dried and then calcined in air. The support may be impregnated in one or more steps. Preferably, multiple impregnation steps are employed. The support is preferably dried and calcined between each impregnation, and then subjected to a final calcination, preferably, in air. The calcined support may then be reduced, for example, by heat treatment in a hydrogen atmosphere.

Where more than one catalyst bed is present, a convenient method of achieving the catalyst zone is to use a single reactor with a space being provided between the beds. The space can be provided by placing substantially inert materials such as alumina, silica, or other refractory materials between the catalyst beds. Alternatively, the space between the catalyst beds is a substantial void. The space between the catalyst beds is not critical in relation to the beds. Preferably, however, the space will be as small as practical. Most preferably, there is no substantial space between the catalyst beds, that is, the beds are directly adjacent to one another. Where the catalyst zone comprises more than two beds, the size of the spaces between the respective beds may vary.

For avoidance of doubt "passing a feedstream which comprises a paraffmic hydrocarbon, hydrogen and an oxygen-containing gas through a catalyst zone" means that

10054(2) 9

hydrogen, oxygen and the paraffinic hydrocarbon to be cracked to produce olefin are all present in the feedstream to the first catalyst bed (as well as carbon monoxide). Thus, where more than one catalyst bed is present, additional hydrocarbon is not added between the catalyst beds. Where there is more than one catalyst bed, the size of the catalyst beds can vary from each other. The preferred relative lengths of the first catalyst bed to subsequent catalyst beds will generally depend on the loading of palladium on the first catalyst bed. In particular, as noted above, it has been found that relatively low palladium loadings on the first catalyst bed are sufficient to provide the desired tolerance, and higher levels of palladium can lead to significant levels of undesired side reactions, including increased formation of methane. For similar reasons, it is generally preferred that the first catalyst bed is relatively short compared to the second catalyst bed (or total of second and subsequent beds if present). Preferably therefore the size of the first catalyst bed to the second catalyst bed is less than 1:1, more preferably less than 1:2, for example in the range 1:3 to 1:10.

The catalyst beds may be arranged either vertically or horizontally. The feedstream which comprises a paraffinic hydrocarbon, hydrogen and an oxygen- containing gas may comprise a single paraffinic hydrocarbon or may comprise more than one paraffinic hydrocarbon. Any paraffinic hydrocarbon(s) may be employed which can be converted to an olefin, preferably a mono-olefin, under the partial combustion conditions employed. Preferred paraffinic hydrocarbons are those having at least two carbon atoms.

In particular, the process of the present invention may be used to convert both liquid and gaseous paraffinic hydrocarbons into olefins. Suitable liquid paraffinic hydrocarbons include naphtha, gas oils, vacuum gas oils and mixtures thereof. Preferably, however, gaseous paraffinic hydrocarbons such as ethane, propane, butane and mixtures thereof are employed.

Any suitable oxygen-containing gas may be used, such as molecular oxygen, air, and/or mixtures thereof. The oxygen-containing gas may be mixed with an inert gas such as nitrogen or argon.

Any molar ratio of paraffinic hydrocarbon to oxygen-containing gas is suitable provided the desired olefin is produced in the process of the present invention. The

10054(2) 10

preferred stoichiometric ratio of paraffinic hydrocarbon to oxygen-containing gas is 5 to 16, preferably, 5 to 13.5 times, preferably, 6 to 10 times the stoichiometric ratio of paraffϊnic hydrocarbon to oxygen-containing gas required for complete combustion of the paraffmic hydrocarbon to carbon dioxide and water. The paraffϊnic hydrocarbon is passed over the catalyst at a gas hourly space velocity of greater than 10,000 h "1 , preferably above 20,000 h '1 and most preferably, greater than 100,000 h '1 . It will be understood, however, that the optimum gas hourly space velocity will depend upon the pressure and nature of the feed composition.

Hydrogen is co-fed with the paraffmic hydrocarbon and oxygen-containing gas into the reaction zone. The molar ratio of hydrogen to oxygen-containing gas can vary over any operable range provided that the desired olefin product is produced. Suitably, the molar ratio of hydrogen to oxygen-containing gas is in the range 0.2 to 4, preferably, in the range l to 3.

Hydrogen co-feeds are advantageous because, in the presence of the catalyst, the hydrogen combusts preferentially relative to the paraffinic hydrocarbon, thereby increasing the olefin selectivity of the overall process.

One or more unsaturated hydrocarbons, such as olefins, dienes and acetylenes may also be co-fed to the catalyst zone in the process of the present invention. The first catalyst bed comprising platinum and palladium has also been found to be tolerant to the presence of such unsaturated hydrocarbons in the feedstream. The unsaturated hydrocarbons may be components of recycle streams from the autothermal cracking reaction itself and/or may be provided as separate feeds.

Preferably, unsaturated hydrocarbons are provided as components of recycle streams from the autothermal cracking reaction itself. This allows a recycle stream to be separated from the product stream of the autothermal cracking reaction and recycled without treatment or with reduced treatment to remove unsaturated hydrocarbons, avoiding downstream separation steps and their associated costs.

Preferably, the feedstream is preheated prior to contact with the catalyst. Generally, the feedstream is preheated to temperatures below the autoignition temperature of the feedstream.

A pre-heated feedstream may be obtained by pre-heating one or more components of the feedstream prior to mixing to form the feedstream and/or by pre-heating the mixed

10054(2) 11

feedstream.

Advantageously, a heat exchanger may be employed to preheat the mixed feedstream prior to contact with the catalyst. The use of a heat exchanger may allow the feedstream to be heated to high preheat temperatures such as temperatures at or above the autoignition temperature of the feedstream. The use of high pre-heat temperatures is beneficial in that less oxygen reactant is required which leads to economic sayings. Additionally, the use of high preheat temperatures can result in improved selectivity to olefin product. It has also be found that the use of high preheat temperatures enhances the stability of the reaction within the catalyst thereby leading to higher sustainable superficial feed velocities. It should be understood that the autoignition temperature of a feedstream is dependent on pressure as well as the feed composition: it is not an absolute value. Typically, in auto-thermal cracking processes, where the paraffinic hydrocarbon is ethane at a pressure of 2 atmospheres, a preheat temperature of up to 450° C may be used.

The process of the present invention may suitably be carried out at a catalyst exit temperature in the range 600°C to 1200 0 C, preferably, in the range 850 0 C to 1050 0 C.

The process of the present invention may be operated at any suitable pressure, such as at atmospheric pressure or at elevated pressure. The process of the present invention may be operated at a pressure in the range atmospheric to 5 barg, but is preferably operated at a pressure of greater than 5barg. More preferably the autothermal cracking process is operated at a pressure of between 5-40barg and advantageously between 10-30barg e.g. 15- 25barg.

The reaction products are preferably quenched as they emerge from the reaction chamber to avoid further reactions taking place. Usually the product stream is cooled to between 750-600 0 C within less than lOOmilliseconds of formation, preferably within 50milliseconds of formation and most preferably within 20milliseconds of formation e.g. within lOmilliseconds of formation.

Wherein the autothermal cracking process is operated at a pressure of 5-20 barg usually the products are quenched and the temperature cooled to between 750-600 0 C within 20milliseconds of formation. Advantageously wherein the autothermal cracking process is operated at a pressure of greater than 20barg the products are quenched and the temperature cooled to between 750-600 0 C within lOmilliseconds of formation.

The invention will now be described with the reference to Figures 1 to 12 and the

10054(2) 12

following examples, wherein:

Figure 1 shows a high pressure autothermal reactor, and

Figures 2 to 8 show results in terms of ethane and oxygen conversion, and ethylene selectivity, from addition of carbon monoxide over a number of catalysts. Figures 9 to 12 show results in terms of ethane conversion, and ethylene, methane and carbon monoxide selectivity from use of a sequential bed catalyst compared to a single bed.

With respect to Figure 1, there is shown a high pressure autothermal reactor (1) comprising a reaction zone (2) surrounded by a pressure jacket (3). The reactor consists of a quartz tubular liner (4) located within a metal holder (5).

Oxygen feed via line (6) and paraffmic hydrocarbon and hydrogen feeds via line (7) are passed to a gas mixing zone (8). The mixed gaseous reactants are then passed to the reaction zone. The reaction zone comprises a first catalyst bed (9) and a second catalyst bed (10). As the reactants contact the catalyst beds (9) and (10) some of the feed combusts.

This combustion reaction is exothermic and the heat produced is used to drive the dehydrogenation of paraffmic hydrocarbon feed to a product stream comprising olefins.

The gaseous product stream from the reaction zone passes into a quench zone (11) comprising a gas injection zone (12) wherein it is contacted with a high velocity nitrogen stream at 25 0 C to rapidly reduce its temperature and maintain the olefin selectivity. A liquid water stream is subsequently injected in a water injection zone (not shown) to reduce the product stream temperature to 100-200°C.

The invention will now be illustrated in the following examples. Single bed catalyst tests - Experiments 1 to 3 Preparation of catalysts

Catalyst A: 3wt% platinum on alumina foam.

Eleven alumina foam blocks (99.5% alumina supplied by Hi-tech Ceramics of Alfred, New York, block size 15mm diameter by 30mm deep, with nominal porosity of 45 pores per inch, total weight 60.42g) were impregnated with a solution containing 3.4 Ig of tetrammineplatinum(II) chloride (ex Johnson Matthey) in 150cm 3 of de-ionised water.

After immersion in the platinum(II) solution for ca. 5 minutes, excess solution was removed from the foam blocks and the blocks were dried in air for 30 minutes at 12O 0 C

1 ΛA

10(b4(zj 13

then calcined at 45O 0 C for ca. 30minutes.

After cooling to room temperature the foams were re-immersed in the Pt-solution and the drying, calcining and re-impregnation processes were repeated until all of the Pt(II) solution was absorbed onto the blocks. A total of five impregnations were required during the preparation process.

The platinum-salt-impregnated foams then received a final calcination in air at 1200 0 C for 6 hours. Catalyst B: 3wt% platinum/ lwt% copper on alumina foam.

Twelve alumina foam blocks (as used for Catalyst A, total weight 67.94g) were impregnated using a solution containing 3.56g of tetrammineplatinum(II) chloride (ex Johnson Matthey) in 100cm 3 of de-ionised water and a solution containing 2.54g of copper(II) nitrate (ex Aldrich) in 100 cm 3 of de-ionised water.

The platinum- and copper-solutions were used alternately in the impregnation process. After immersion in the solution for ca. 5 minutes, excess solution was removed from the foam blocks and the foams were dried in air for 30 minutes at 12O 0 C then calcined at 45O 0 C for ca. 30minutes.

After cooling to room temperature the spheres were re-immersed in the solution and the drying, calcining and re-impregnation processes were repeated until all of the solution was absorbed onto the foams.

A total of four impregnations were required for the Pt-solution and four impregnations for the copper-solution during the preparation process.

The impregnated foams then received a final calcination in air at 600 0 C for 6 hours. Catalyst C: 1.5wt% platinum/0.3wt% palladium on alumina foam. Seven alumina foam blocks (as used for Catalyst A, total weight 45.3g) were impregnated with a solution containing 1.15g of tetrammineplatinum(II) chloride (ex Johnson Matthey) and 0.3Og of tetramminepalladium(II) chloride (ex Johnson Matthey) in 100cm 3 of de-ionised water.

After immersion in the Pt(II)/Pd(II) solution for ca. 5 minutes, excess solution was removed from the foam blocks and the foams were dried in air for 30 minutes at 12O 0 C then calcined at 45O 0 C for ca. 30minutes.

After cooling to room temperature the foams were re-immersed in the Pt/Pd-solution

10054(2) 14

and the drying, calcining and re-impregnation processes were repeated until all of the Pt/Pd-solution was absorbed onto the foams.

A total of four impregnations were required during the preparation process.

The impregnated foams then received a final calcination in air at 1200 0 C for 6 hours. Autothermal Cracking Reaction The apparatus of Figure 1 was used.

For use with catalysts prepared on foam blocks the quartz liner had an inner diameter of 15mm and overall catalyst bed lengths ranged from 60mm to 100mm. The catalyst foam block was wrapped with ceramic paper to minimise the potential for reactant gas by- passing the catalyst bed.

Reactant gases were fed to the reactor using Bronkhorst Hi-Tec thermal mass flow controllers. Gaseous effluent was analysed by gas chromatography, with residual oxygen being measured using a trace oxygen analyzer (Teledyne Analytical Instruments).

For each experiment, the required catalysts were loaded into the reactor and the reactor heated to 200°C under nitrogen at the required pressure.

Ethane, hydrogen and oxygen were then introduced to the reactor at a gas hourly space velocity Of SOO 5 OOO h '1 .

A hydrogen to oxygen volume ratio of 2:1 was used.

The oxygen to ethane ratio was varied within each experiment as shown in the respective data below (whilst maintaining the hydrogen to oxygen ratio at 2: 1). For carbon monoxide co-feed experiments, carbon monoxide was introduced (and maintained) at a level corresponding to 9 vol% carbon monoxide in hydrogen. (As a function of the total feedstream the carbon monoxide comprised about 4 vol%.) Experiment 1 (comparative): 3wt% platinum on alumina foam The results from use of catalyst A at a reaction pressure of 0.8 barg are shown in

Figures 2 and 3.

Li particular, Figure 2 shows the ethane and oxygen conversions obtained versus the oxygen to ethane feed ratio in the presence and absence of carbon monoxide and Figure 3 shows the ethylene selectivity obtained versus the ethane conversion in the presence and absence of carbon monoxide.

Figure 2 shows that the presence of carbon monoxide leads to a significant reduction in both ethane and oxygen conversions at a particular oxygen to ethane feed ratio over

10054(2) 15

catalyst A.

Figure 3 shows that the presence of carbon monoxide also leads to a significant reduction in ethylene selectivity at a particular ethane conversion over catalyst A. For a comparable ethane conversion, the typical reduction in selectivity is about 6 g per 100 g ethane converted.

Experiment 2: 3wt% platinum/1.5wt% copper on alumina foam

The results from use of catalyst B at a reaction pressure of 0.8 barg are shown in Figures 4 and 5.

In particular, Figure 4 shows the ethane and oxygen conversions obtained versus the oxygen to ethane feed ratio in the presence and absence of carbon monoxide and Figure 5 shows the ethylene selectivity obtained versus the ethane conversion in the presence and absence of carbon monoxide.

Figure 4 shows that in the presence of carbon monoxide the ethane and oxygen conversions at a particular oxygen to ethane feed ratio are reasonably well maintained over catalyst B.

Figure 5 shows that the presence of carbon monoxide leads to a reduction in ethylene selectivity at a particular ethane conversion over catalyst B, although the reduction is not as significant as for the case of catalyst A (platinum "only" catalyst). For a comparable ethane conversion, the typical reduction in selectivity is about 1-2 g per 100 g ethane converted.

Experiment 3: 1.5wt% platinum/0.3wt% palladium on alumina foam

The results from use of catalyst C at a reaction pressure of 0.8 barg are shown in Figures 6 and 7.

In particular, Figure 6 shows the ethane and oxygen conversions obtained versus the oxygen to ethane feed ratio in the presence and absence of carbon monoxide and Figure 7 shows the ethylene selectivity obtained versus the ethane conversion in the presence and absence of carbon monoxide.

Figure 6 shows that in the presence of carbon monoxide the ethane and oxygen conversions at a particular oxygen to ethane feed ratio are maintained over catalyst C. Figure 7 shows that in the presence of carbon monoxide the ethylene selectivity is also maintained over catalyst C.

Figure 8 shows the data for selectivity against ethane conversions for catalysts A to C

10054(2) 16

in the presence of carbon monoxide as in Experiments 1 to 3 on a single graph. It can be seen that in the presence of carbon monoxide, catalyst C shows the highest selectivity at a particular ethane conversion.

Tables 1 to 3 below reproduce some of the data from Figures 2 to 7 at particular oxygen to ethane feed ratios. In addition, Tables 1 to 3 show the temperature of the catalyst inlet (Cat. temp 1) is decreased significantly for Catalyst A in the presence of carbon monoxide. The reduction in catalyst temperature at the inlet suggests that some "poisoning" of catalyst A is occurring in the presence of carbon monoxide, consistent with the reduced conversions observed.

A smaller effect is seen for the platinum/copper catalyst (catalyst B) consistent with the results showing a smaller drop in conversions on introduction of carbon monoxide (but, as noted above, for these catalysts some loss of selectivity is still observed). TABLE 1 : Effect of CO addition at oxygen: ethane feed ratio of 0.52:1 v/v

3% Pt 3% Pt-1%Cu 1.5% Pt-0.3%Pd pressure barg 0.79 0.8 0.79 0.8 0.8 0.81 feed temp degC 204 209 208 204 206 215 Cat. temp 1 degC 897 703 588 546 917 - 910 Cat. temp 2 degC 948 913 904 895 912 929 feed ratios CO:C2H6 v/v 0 0.10 0 0.10 0 0.10

C2 conv % 69.77 61.12 69.99 68.85 72.15 73.45 02 conv % 95.90 91.82 96.59 96.12 96.79 97.13 ethylene sel. 68.61 64.85 69.79 68.01 68.94 68.05

TABLE 2: Effect of CO addition at oxygen:ethane feed ratio of 0.56:1 v/v

3% Pt 3% Pt-1 %Cu 1.5% Pt-O .3%Pd pressure barg 0.81 0.81 0.8 0.8 0.8 0.8 feed temp degC 204 209 208 204 207 214 Cat. temp 1 degC 914 725 597 575 933 927 Cat. temp 2 degC 964 914 921 913 929 944 feed ratios CO:C2H6 v/v 0 0.11 0 0.11 0 0.11

C2 conv % 76.03 61.12 75.61 74.72 78.18 79.08 O2 conv % 96.60 91.82 96.96 96.70 97.36 97.56 ethylene sel. 66.29 64.85 66.95 65.49 65.90 65.65

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TABLE 3: Effect of CO addition at oxygenrethane feed ratio of 0.60:1 v/v

3% Pt 3% Pt-1 %Cu 1.5% Pt-O .3%Pd pressure barg 0.81 0.79 0.79 0.79 0.81 0.81 feed temp degC 202 207 207 205 206 212 cat temp 1 degC 926 766 603 552 952 944 Cat temp 2 degC 981 922 940 931 948 963 feed ratios CO:C2H6 v/v 0 0.12 0 0.12 0 0.12

C2 conv % 81.36 70.03 80.61 79.55 83.25 84.31 02 conv % 97.25 92.88 97.40 97.17 97.90 98.05 ethylene sel. 64.51 62.15 64.39 63.80 64.22 63.10

Experiments 1 to 3 show, therefore, that the use of a catalyst comprising platinum and palladium leads to improved tolerance of the presence of carbon monoxide in the feed to an autothermal cracking process.

Sequential bed catalyst tests - Experiments 4 to 6

Preparation of catalysts

Catalyst D: 3wt% platinum on alumina spheres.

Alumina spheres (10Og) ex Condea (1.8mm diameter, surface area @ 210m 2 /g) were impregnated with a solution containing 5.42g of tetrammineplatinum(II) chloride (ex Johnson Matthey) in de-ionised water. Preparation was via incipient wetness technique.

After immersion in the Pt(II) solution for 5 minutes, excess solution was removed from the spheres and the spheres were dried in air at 12O 0 C then calcined at 45O 0 C for ca. 30minutes. '

After cooling to room temperature the spheres were re-immersed in the remaining Pt- solution and the drying and calcination process was repeated.

The spheres then received a final calcination in air at 1200 0 C for 6 hours. After calcination the diameter of the spheres had reduced from ca. 1.8mm to ca. 1.2mm. Catalyst E: 3wt% platinum/0.2wt% palladium on alumina spheres.

The procedure for Catalyst D was repeated using a solution that contained 5.42g of tetrammineplatinum(II) chloride (ex Johnson Matthey) in de-ionised water and 0.49g of tetramminepalladium(II) chloride.

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Autothermal Cracking Reaction The apparatus of Figure 1 was used.

For use with catalysts prepared on alumina spheres the quartz liner had an inner diameter of 10mm diameter and an overall catalyst bed length of 60mm was used. Reactant gases were fed to the reactor using Bronkhorst Hi-Tec thermal mass flow controllers. Gaseous effluent was analysed by gas chromatography, with residual oxygen being measured using a trace oxygen analyzer (Teledyne Analytical Instruments).

For each experiment, the required catalysts were loaded into the reactor and the reactor heated to 200°C under nitrogen at the required pressure. Ethane, nitrogen, hydrogen and oxygen were then introduced to the reactor with an ethane flow rate maintained at 120 nl/min and a nitrogen flow rate of 9 nl/min.

A hydrogen to oxygen volume ratio of 1 : 1 was used.

The oxygen to ethane ratio was varied within each experiment as shown in the respective data below (whilst maintaining the hydrogen to oxygen ratio at 1 : 1). Experiment 4:.60 mm bed of 3wt% platinum on alumina spheres

Experiment 5: 60 mm bed of 3wt% platmum/0.2wt% palladium on alumina spheres Experiment 6: Sequential bed with 10 mm bed of 3wt% platinum/0.2wt% palladium on alumina spheres and 50 mm bed of 3 wt% platinum on alumina spheres The results from use of single beds comprising 3 wt% platinum on alumina spheres and comprising 3wt% platinum/0.2wt% palladium on alumina spheres and of a sequential bed comprising of 3wt% platinum/0.2wt% palladium on alumina spheres followed by 3wt% platinum on alumina spheres at a reaction pressure of 20 barg are shown in Figures 9 to 12.

In particular, the figures show that, although the sequential bed shows a slightly lower ethane conversion compared to a single bed of platinum/palladium catalyst (Figure 9) the sequential bed shows higher ethylene selectivity (at comparable ethane conversion)(Figure 10) and correspondingly lower methane and carbon monoxide selectivities (Figures 11 and 12 respectively).

Experiments 4 to 6 show, therefore, that the use of a sequential bed catalyst comprising platinum and palladium in the first catalyst bed (which leads to improved tolerance of the presence of carbon monoxide in the feed compared to a single bed of platinum) provides an improved catalyst zone compared to a single bed of platinum and palladium.

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Further Experiments

A number of similar experiments were performed at pressures of 10 and 20 barg over single beds comprising l-3wt% platinum on alumina and comprising 3wt% platinum/0.2-

0.5wt% palladium on alumina and over a sequential bed comprising of 3wt% platinum/0.2wt% palladium on alumina followed by 3wt% platinum on alumina in to the feed to which carbon monoxide was introduced at varying concentrations.

In a number of the experiments over a single bed of l-3wt% Pt on alumina the autothermal cracking reaction could not be maintained on the addition of carbon monoxide to the feed, leading to extinguishment of reaction. No extinguishment of reaction was observed over catalysts comprising palladium as the first catalyst bed, either as a single bed, or as a sequential bed.

These experiments show, therefore, that the presence of palladium in the first catalyst bed leads to improved tolerance of the presence of carbon monoxide in the feed compared to a single bed of platinum.




 
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