WO2006070007A1

WO2006070007A1 - Hydrogenation catalyst and use thereof for hydrogenating fischer-tropsch endproducts

Info

Publication number: WO2006070007A1
Application number: PCT/EP2005/057199
Authority: WO
Inventors: Focco Cornelis Bijlsma; Jan Lodewijk Maria Dierickx; Arend Hoek; Frans Joris Antonius Kellendonk; Anna Elisabeth Maria Oud
Original assignee: Shell Internationale Research Maatschappij B.V.
Priority date: 2004-12-31
Filing date: 2005-12-28
Publication date: 2006-07-06
Also published as: MY170240A

Abstract

A hydrogenation catalyst particularly suitable for hydrogenating oxygenates in a hydrogenation unit of a Fischer-Tropsch plant is disclosed. Said catalyst comprises a metallic active portion in which the metal is a non-noble Group VIII metal and a support, characterised in that the support does not catalyse an acid catalysed reaction and wherein over 90% of the pores within the support are sized between 100 A-400 A. A preferred embodiment comprises more than 5 % and less than 20 % nickel based on a wide pore alumina support. The catalyst successfully hydrogenates oxygenates which otherwise tend to poison a catalyst in a hydroconversion unit downstream. Moreover, the temperature at which the unwanted hydrogenolysis of long chain paraffins to methane occurs is higher for one catalyst disclosed herein than a comparable known catalyst. This allows the hydrogenation plant to operate at a higher temperature.

Description

HYDROGENATION CATALYST AND USE THEREOF FOR HYDROGENATING FISCHER- TROPSCH ENDPRODUCTS

This invention relates to a hydrogenation catalyst and particularly to a hydrogenation catalyst for hydrogenating oxygenates which are present in a Fischer- Tropsch wax stream in a hydrogenation unit used in a Fischer-Tropsch plant.

The Fischer-Tropsch process is often used for the conversion of hydrocarbonaceous feed stocks into liquid and/or solid hydrocarbons. The feed stock (e.g. natural gas, associated gas, coal-bed methane, residual oil stream, biomass, and/or coal) is converted in a first step into a mixture of hydrogen and carbon monoxide (this mixture is often referred to as synthesis gas) . The synthesis gas is then converted in a second step over a suitable catalyst at elevated temperature and pressure into paraffinic and olefinic compounds ranging from methane to high molecular weight molecules comprising up to 200 carbon atoms, or, under particular circumstances, even more.

The obtained product (Heavy Paraffin Synthesis (HPS) product) may be fed to a hydroisomerisation/hydrocracking unit, but is preferably first fed to a hydrogenation unit where the olefins and oxygenates are hydrogenated. In the hydrogenation unit there is no or substantially no hydroisomerisation and/or hydrocracking. Some hydrogenated product may be removed at this point for sale but most of the hydrogenated products proceed to a hydroconversion unit, especially the C5+ fraction, in which hydroisomerisation as well as hydrocracking occurs _ O —

required expanded inner diameter, is pumped, pushed or pulled, sometimes in combination with rotation, through the tubular element.

In one such method the expander is pumped through the tubular element whereby the interior of the expanded portion of the tubular element is pressurised to a maximum pressure at which the expander starts moving through the tubular element. However it has been experienced that the required expansion forces exerted by the expander to the tubular element can be very high, thus potentially leading to damage to the expander and/or the inner surface of the tubular element.

It is therefore an object of the invention to provide an improved method of expanding a tubular element in a wellbore, which overcomes the drawbacks of the prior art.

In accordance with the invention there is provided a method of radially expanding a tubular element extending into a wellbore formed in an earth formation, comprising the steps of: arranging an expander in the tubular element, the expander being operable to exert a radial force to the inner surface of the tubular element; radially expanding the tubular element by pressurising the interior of the tubular element and simultaneously operating the expander to exert said radial force to the inner surface of the tubular element, wherein the expander is operated to exert said radial force to the inner surface of the tubular element independently from pressurising the interior of the tubular element.

By operating the expander independently from pressurising the interior of the tubular element, it is achieved that the interior of the tubular element can be pressurised to a significantly higher pressure so that the required expansion forces exerted by the expander to the tubular element can be relatively low. Suitably the tubular element includes an expanded portion and an unexpanded portion, whereby the interior of the tubular element is pressurised both in the expanded portion and the unexpanded portion. For example, the whole interior of the tubular element can be pressurised.

The invention will be explained hereinafter in more detail by way of example, with reference to the accompanying drawings in which:

Fig. 1 schematically shows a first embodiment of a wellbore provided with a casing expanded according to the method of the invention; and

Fig. 2 schematically shows a second embodiment of a wellbore provided with a liner expanded according to the method of the invention. In the drawings and the detailed description below, like reference numerals relate to like components. Furthermore, it is to be understood that the terms "below", "above", "upward" and "downward" refer to wellbore depths measured along the longitudinal axis of the wellbore and relative to surface.

Referring to Fig. 1 there is shown a wellbore 1 for the production of oil or gas from an earth formation 2. An expandable casing 4 extends from a wellhead 6 at surface to near the lower end of the wellbore 1, whereby the casing 4 is sealingly connected to the wellhead 6. An expander 8 for radially expanding the casing 4, is positioned in the casing 4 whereby an expanded portion 10 of the casing 4 extends below the expander 8, and an unexpanded portion 12 of the casing 4 extends above the expander 8. The lower end of the expanded casing portion 10 is closed by means of a packer 13.

The expander 8 tapers in upward direction from a relatively large diameter corresponding to the inner diameter of the expanded casing 4, to a relatively small diameter corresponding to the inner diameter of the unexpanded casing 4. There are no provisions to seal the outer surface of the expander 8 to the inner surface of the casing 4, so that pressurised fluid can flow between the expander 8 and the casing 4.

The expander 8 is provided with a through-bore 14 providing fluid communication between the interior of the unexpanded casing portion 12 and the interior of the expanded casing portion 10. The expander 8 is connected to a wireline 16 extending through the unexpanded casing portion 12 and the wellhead 6, to a winch 18 at surface. The wireline 16 passes through a through-bore 19 provided in the wellhead 6 in a sealing manner so that the wireline 16 can be axially moved through the through-bore 19 while fluid is prevented from passing through the through-bore 19. Further, a fluid pump 20 is provided at surface for pumping fluid, via a conduit 22 and the wellhead 6, into the casing 4. Referring to Fig. 2 there is shown a wellbore 1 for the production of oil or gas from an earth formation 2. A casing 24 extends from surface to a depth at a selected distance from the wellbore bottom, the casing 24 being fixed in the wellbore by a layer of cement (not shown) . Further, an expandable liner 26 extends in the wellbore below the casing 24, whereby an upper end portion of the liner 26 extends into the casing 24. The liner 26 is suspended on a drill pipe 28, with a side entry sub 30 Typically the method produces a catalyst with metal oxide particles on the support and the metal oxide is reduced in situ before the catalyst is used.

Preferably the metal salt is mixed in a basic solution.

Preferably the catalyst is used to hydrogenate products produced by a Fisher-Tropsch process.

The invention further concerns a process for the hydrogenation of the C^₊ fraction obtained in a process in which a mixture of hydrogen and carbon monoxide is converted over a cobalt containing catalyst, preferably a cobalt manganese catalyst, especially a titania supported catalyst, in which hydrogenation process a catalyst is used comprising nickel on a support that does not catalyse an acid catalysed reaction as measured by the n- heptane test result at 350 ⁰C, the process being carried out at a temperature between 220 ⁰C and 300 ⁰C and a pressure between 8 and 50 bar.

The hydrogenation reactor preferably concerns the C±2+ f^raction of the Fischer-Tropsch process, more preferably the C20+ fraction. Suitably the result in the n-heptane test is above 350 ⁰C, preferably above 360 ⁰C. In particular, the difference between the temperature in ⁰C and the pressure in bar is at least 225, preferably between 230 and 250. The catalyst is a nickel catalyst as further defined above. Preferably the obtained product contains less than 100 ppm wt. oxygen, more preferably less than 50 ppm wt. oxygen, still more preferably less that 20 ppm wt. oxygen. The amount of oxygen in the starting material is suitably between 200 ppm wt. and 2%wt., preferably between 500 ppm wt. % and l%wt., more preferably between 1000 ppm wt. % and 5000 ppm wt. % The Fischer-Tropsch synthesis is well known to those skilled in the art and involves synthesis of hydrocarbons from a gaseous mixture of hydrogen and carbon monoxide, by contacting that mixture at reaction conditions with a Fischer-Tropsch catalyst.

Products of the Fischer-Tropsch synthesis may range from methane to heavy paraffinic waxes. Preferably, the production of methane is minimised and a substantial portion of the hydrocarbons produced have a carbon chain length of a least 5 carbon atoms. Preferably, the amount of C5+ hydrocarbons is at least 60% by weight of the total product, more preferably, at least 70% by weight, even more preferably, at least 80% by weight, most preferably at least 85% by weight. Reaction products which are in liquid phase under reaction conditions may be separated and removed using suitable means, e.g. a gas/liquid separator in the case of a (multi tubular) fixed bed reactor or, when a slurry reactor is used, such as one or more filters. Internal or external filters, or a combination of both, may be employed. Gas phase products such as light hydrocarbons and water may be removed using suitable means known to the person skilled in the art.

The relatively heavy Fischer-Tropsch product which is obtaind using specific Fischer-Tropsch catalysts, e.g. cobalt/manganese catalysts, has at least 30 wt%, preferably at least 50 wt%, and more preferably at least 55% of compounds having at least 30 carbon atoms. Furthermore the weight ratio of compounds having at least 60 or more carbon atoms and compounds having at least 30 carbon atoms of the Fischer-Tropsch product is at least 0.2, preferably at least 0.4 and more preferably at least 0.55. Preferably the Fischer-Tropsch product comprises a C20+ fraction having an ASF-alpha value (Anderson-Schulz-Flory chain growth factor) of at least 0.925, preferably at least 0.935, more preferably at least 0.945, even more preferably at least 0.955, based on the amount of C20 and C30 compounds.

Fischer-Tropsch catalysts are known in the art, and typically include a Group VIII metal component, preferably cobalt, iron and/or ruthenium, more preferably cobalt. Typically, the catalysts comprise a catalyst carrier. The catalyst carrier is preferably porous, such as a porous inorganic refractory oxide, more preferably alumina, silica, titania, zirconia or mixtures thereof.

The optimum amount of catalytically active metal present on the carrier depends inter alia on the specific catalytically active metal. Typically, the amount of cobalt present in the catalyst may range from 1 to 100 parts by weight per 100 parts by weight of carrier material, preferably from 10 to 50 parts by weight per 100 parts by weight of carrier material. The catalytically active metal may be present in the catalyst together with one or more metal promoters or co- catalysts. The promoters may be present as metals or as the metal oxide, depending upon the particular promoter concerned. Suitable promoters include oxides of metals from Groups HA, IHB, IVB, VB, VIB and/or VIIB of the Periodic Table, oxides of the lanthanides and/or the actinides. Preferably, the catalyst comprises at least one of an element in Group IVB, VB and/or VIIB of the Periodic Table, in particular titanium, zirconium, manganese and/or vanadium. As an alternative or in addition to the metal oxide promoter, the catalyst may comprise a metal promoter selected from Groups VIIB and/or VIII of the Periodic Table. Preferred metal promoters include rhenium, platinum and palladium.

A most suitable catalyst comprises cobalt as the catalytically active metal and zirconium as a promoter. Another most suitable catalyst comprises cobalt as the catalytically active metal and manganese and/or vanadium as a promoter, especially on a titania carrier.

The promoter, if present in the catalyst, is typically present in an amount of from 0.1 to 60 parts by weight per 100 parts by weight of carrier material. It will however be appreciated that the optimum amount of promoter may vary for the respective elements which act as promoter. If the catalyst comprises cobalt as the catalytically active metal and manganese and/or vanadium as promoter, the cobalt : (manganese + vanadium) atomic ratio is advantageously at least 12:1.

The Fischer-Tropsch synthesis is preferably carried out at a temperature in the range from 125 to 350 ⁰C, more preferably 175 to 275 ⁰C, most preferably 200 to 260 ⁰C. The pressure preferably ranges from 5 to

150 bar abs., more preferably from 5 to 80 bar abs.

Hydrogen and carbon monoxide (synthesis gas) is typically fed to the three-phase slurry reactor at a molar ratio in the range from 0.4 to 2.5. Preferably, the hydrogen to carbon monoxide molar ratio is in the range from 1.0 to 2.5.

The gaseous hourly space velocity may very within wide ranges and is typically in the range from 1500 to 10000 Nl/l/h, preferably in the range from 2500 to 7500 Nl/l/h.

The Fischer-Tropsch synthesis is suitably carried out in a slurry phase regime or an ebullating bed regime, wherein the catalyst particles are kept in suspension by an upward superficial gas and/or liquid velocity.

Another regime for carrying out the Fischer-Tropsch reaction is a trickle flow regime, especially a fixed bed. A very suitable and preferred reactor is a multitubular fixed bed reactor.

The new catalysts of the present invention are especially suitable for (very) heavy Fischer-Tropsch products, especially the C20"¹" fraction of FF processes having an ASF-alpha value of at least 0.925, preferably 0.935, more preferably 0.945, the ASF value based on the amount of C20 ^and C30 compounds.

Catalyst supports in accordance with the present invention do not catalyse an acid catalysed reaction. In order to determine whether a catalyst catalyses an acid catalysed reaction, the ability of the catalyst to crack n-heptane is analysed at a variety of temperatures.

The n-heptane cracking is measured by first preparing a standard catalyst consisting of the calcined carrier and 0.4 wt% platinum. Standard catalysts are tested as 40-80 mesh particles, which are dried at 200 ⁰C before loading in the test reactor. The reaction is carried out in a conventional fixed-bed reactor having a length to diameter ratio of 10 to 0.2. The standard catalysts are reduced prior to testing at 400 ⁰C for 2 hrs at a hydrogen flow rate of 2.24 Nml/min and a pressure of 30 bar. The actual test reaction conditions are: n-heptane/H2 molar ratio of 0.25, total pressure 30 bar, and a gas hourly space velocity of 1020 NmI/ (g. h) . The temperature is varied by decreasing the temperature from 400 ⁰C to 200 ⁰C at 0.22 °C/minute. Effluents are analysed by on-line gas chromatography. The temperature at which 40 wt% conversion is achieved is reported as the n-heptane test result. Lower n-heptane test results correlate with more active catalysts for acidic reactions. Thus the catalysts in accordance with the present invention should have an n-heptane test result of greater than 350 ⁰C, preferably greater than 360 ⁰C.

The cracking activity of the silica-alumina carrier can be influenced by, for example, variation of the alumina distribution in the carrier, variation of the percentage of alumina in the carrier, and the type of alumina. For alumina supports, the content of silica is preferably limited to a maximum of 1% silica.

Reference in this respect is made to the following articles which illustrate the above: Von Bremer H.,

Jank M. , Weber M. , Wendlandt K. P., Z. anorg. allg. Chem. 505,79-88 (1983) ; Leonard A. J. , Ratnasamy P., Declerck F. D. , Fripiat J. J. , Disc, of the Faraday Soc.1971, 98-108 ; and Toba M. et al, J. Mater. Chem. , 1994, 4 (7), 1131-1135.

A hydrogenation catalyst particularly suitable for hydrogenating oxygenates in a hydrogenation unit of a Fischer-Tropsch plant is disclosed. A preferred embodiment comprises more than 5% and less than 20% nickel based on a wide pore alumina support. The catalyst successfully hydrogenates oxygenates which otherwise tend to poison a catalyst in a hydroconversion unit downstream. Moreover, the temperature at which the unwanted hydrogenolysis of long chain paraffins to methane occurs is higher for one catalyst disclosed herein than a comparable known catalyst. This allows the hydrogenation plant to operate at a higher temperature Method of Manufacture

There are various methods of adding an active element to a support. In the present invention, metal impregnation is preferred. To achieve this, the water pore volume of the support is first measured. A nickel salt, such as nickel nitrate or nickel carbonate, is dissolved in a basic solution such as ammonia in the correct proportion so that on evaporation of the ammonia and water, the required amount of nickel is left on the support.

The ammonia serves to solvate the nickel ions which helps reduce the tendency of the solution to flocculate. The pH of the solution is suitably above 8.5. In acidic solutions H⁺ ions proceed to the surface of the support and repel the Ni^⁺ ions. Monoethylamine may be added to reduce the amount of base that evaporates.

The pores of the support are then filled with the solution, left to dry and then calcined at between 250 and 600 ⁰C, preferably between 450 and 500 ⁰C. This method results in highly dispersed nickel on the surface of the support. The average particle size of the nickel metal is around 25 A (that is 2.5 nm) .

If more than 12.7% nickel is required, the solution will tend to flocculate and so, in such cases, the impregnation must be repeated with solutions of up to around 12.7% nickel to give a final nickel content of greater than 12.7%.

The temperature during impregnation is typically between ambient temperature to 90 ⁰C. If the exotherm caused by impregnation heats the mixture above 90 ⁰C, either more monoethylamine or cooling maintains the base side pH. A support with a wide pore size is preferred. A support made from wide pore alumina is preferred.

One catalyst in accordance with the present invention comprises a wide pore alumina support and 12.0% nickel (Catalyst G) Experiments

To assess the hydrogenation performance of Catalyst G (a trilobe catalyst, 1.6mm) on paraffins resulting from a Fischer-Tropsch reactor utilising a cobalt/manganese catalyst, a series of experiments were carried out and the results are set out and discussed below.

The same feeds for these experiments were also used for other catalysts in order for a back to back comparison. The catalyst was received in pre-reduced form.

Compacted bulk density as loaded was 0.55 g/ml. Diluent was 0.2 mm SiC, 1.24 ml/ml. Catalyst bed height was 0.758 m, and reactor diameter 20 mm. The catalysts were supplied in the reduced and air passivated form and were activated before use in order to remove the oxide layer, thereby obtaining the active nickel phase. A short term activation is required. The catalyst used was activated by passing hydrogen gas over the catalyst bed in the standard manner. During catalyst activation, the preferred gas rate is as high as possible, to keep the steam partial pressure low. For the same reason, the temperature ramp is slowed down when the exotherm of the reduction starts.

Process conditions were 30 bar pure hydrogen, and WHSV (the feed rate over the catalyst) was 1 kg/l/h. Light Product (LP) from an HPS Fixed Bed pilot plant (STY 135, 40 bar) was used. (STY is the Space Time Yield - a measure of the amount (in kg) of product per hour per cubic metre of catalyst at a given temperature and pressure.) This is also the feed for the back checks. The range 160-260 ⁰C was screened in steps of 20 ⁰C per weekday, and the IR oxygenates and olefins measured in each step.

Then, Heavy Product (HP) (STY 115, 40 bar, oxygenates recycle, was co-fed and screened over the same temperature range. This was the blended feed from which a known hydrogenation catalyst did not convert all oxygenates. Distillation and GLC for isoparaffins content are only needed at the lowest temperature that ensures full oxygenates conversion, and at the highest temperature above which the catalyst makes methane. We did both in the light and heavy product run. The results are set out below, starting with a breakdown of the oxygenates in the feed. Oxygenate levels from infra-red absorption analysis are shown in ppmw oxygen.

RESULTS FEED ANALYSIS: Sample: Light Product (LP)

(all values in ppm wt. oxygen)

Aldehyde/ketone 580

Ester 475

Acid/anhydride 170 primary OH 400 secondary OH 700

Total 2325

Sample: Heavy wax (HP)

Aldehyde/ketone 95

Ester 800 Acid/anhydride 40 primary OH 2650 secondary OH 225

Total 3810

RESULTS ADSORPTION EXPERIMENTS Prior to actual testing under hydrogenation conditions, the available catalysts and supports were screened in a diffusion-adsorption test. An excess amount of various catalysts was shaken for 2.5 hours and overnight in molten wax.

Wax mixture: 60% LP/40% HP

"CATALYST A" - Comparative Example Adsorption time: 2.5hr 18 hr

Aldehyde/ketone: 200 175

Ester: 100 <5

Acid/anhydride: 5 <5 primary OH: 100 <75 secondary OH: 125 75

Total: 530 250

Catalyst A has a uniform pore size distribution between 100 A and 750 A. It has more than 20% nickel

(about 25) and pores ranging from 100 to 800 A, and has a support which catalyses acidic reactions. The BET surface area is 162 m^/g. The heptane cracking test result for catalyst A is less than 350 ⁰C. CAT. B

This catalyst has the following pore size distribution:

10% > 400 A 11.5% > 350 A

99% > 100 A 0.25% < 100 A Median pore size = 229.1 A

Adsorption time: 2.5 hr 18

Aldehyde/ketone: 135 110

Ester: 75 25

Acid/anhydride: 5 <5 primary OH: 100 <75 secondary OH: 125 75

Total: 440 210

The heptane cracking test result for this catalyst is higher than 360°. The BET surface area is 100-

120 m²/g.

CAT. C (12.7% Ni)

Adsorption time 2.5 hr 18 hr

Aldehyde/ketone 50 <5

Ester <5 <5

Acid/anhydride <5 <5 primary OH <75 <75 secondary OH 100 <15

Total 150 Not detected

CAT. D (31.5% Ni)

Adsorption time 2.5 hr 18hr

Aldehyde/ketone: 175 105

Ester: 20 60

Acid/anhydride: 20 <5 primary OH: 150 100 secondary OH: 200 200

Total: 750 465

CAT.E (12.0% Ni)

Adsorption time 2.5 hr 18hr

Aldehyde/ketone 115 15 Ester 80 <5

Acid/anhydride <5 <5 primary OH 75 <75 secondary OH 125 <75 Total 395 15

CAT. F (30.2% Ni) Adsorption time 2.5 hr 18hr Aldehyde/ketone: 225 85 Ester: 335 60 Acid/anhydride: 75 <5 primary OH: 250 50 secondary OH: 300 150 Total: 1155 345

Catalysts C, D, E and F were prepared from a support having a BET surface area of 110-115 m²/g. At 220 A pore size the BET surface area is greater than 110 m^/g.

These catalysts show heptane cracking test results of above 360 ⁰C. They also have the following pore size distribution: 5% > 350 A

3% > 400 A

0.2% < 100 A

Median pore size diameter = 220 A

These data show that the catalyst with 12.7% Nickel on the wide pore alumina performs better than catalyst A for hydrogenating oxygenates in heavy wax.

Catalyst B has 28% Nickel and a relatively low surface area. The BET surface area is 100-120 m2/g.

The 12.7% Ni catalyst achieves full conversion of oxygenates at 260 ⁰C in a blend containing 60% HP wax. At these temperatures, isomerisation of detergent feedstocks, and methanation, are negligible. Over 1200 hours of testing, no catalyst deactivation was observed.

Both catalysts, catalyst A and G, completely removed olefins in all hydrogenation (HGU) experiments. Both catalysts completely removed oxygenates from HPS Light Product (LP) , but catalyst A does not completely remove carbonyls (=aldehydes + ketones) , esters, and secondary alcohols, from mixtures containing Heavy HPS product (HP) .

ABSENCE OF UNDESIRED REACTIONS

We checked isomerisation and methanation for the 270, 280, and 290⁰C runs on the same feed on which we tested catalyst A (both light and heavy product) . Typical iso/normals splits are 4.8/95.2. Methanation at the three above temperatures is reflected in 0.4, 0.8-1.0, and 7 %vol in off gas (once through gas phase) . Technically, we draw the line at 1% methane. The methanation limit for the catalyst A is 250 ⁰C. We have found 0.27, 0.33. and 0.71%w gas make (yield in % weight on feed) from three mixtures containing LP and HP a known hydrogenation catalyst at 250 ⁰C. The average 0.44% w of translates to (0.44 kg methane/100 kg feed) /(0.016 kg methane/mol methane) * (22.4 Nl methane/mol methane) / (750 Nl off gas/kg feed) = 0.8% vol in off gas. LIFE TIME EXPECTATION

Our back checks at 24O⁰C showed the following oxygenates contents (carbonyls and esters between brackets) :

Rhr 286.3-294-3 55 ppm (25 + 30) Rhr 370.3-382.8 40 ppm (20 + 20) Rhr 768.3-792.3 40 ppm (20 + 20) Rhr 1377.6 (stream sample) 10 ppm (5 + 5) After 1200 hours, full conversion was achieved at 250 ⁰C. After returning to 240 ⁰C, 10 ppm total oxygenates was measured after stabilisation. The catalyst was stable over the duration of our test. If anything, the catalyst benefits from the high temperature and the hydrogen: its activity increases by 40 ⁰C per 1000 hours. By comparison, catalyst A offers zero cycle length, since it cannot do the job at all below its methanation limit.

CONCLUSIONS

Catalyst G clearly performs better than catalyst A in hydrogenation of oxygenates in LP and HP product. It achieves full conversion of oxygenates in a blend containing 60% HP material. At these temperatures, both isomerisation of detergent feedstocks and methanation are negligible.

Over 1200 hours of testing, no catalyst deactivation was seen. Thus since catalyst A cannot completely remove oxygenates from HP product, catalyst G is the better catalyst for this application. Effect of Pressure

The need to avoid uncontrollable excessive hydrogenolysis of paraffins to methane, or "methanation", sets the upper operation temperature limit for a hydrogenation catalyst. The temperature at which methanation starts, decreases with decreasing hydrogen partial pressure (a negative reaction order) .

The experimental technique of nano-flow test equipment, instead of bench-scale testing was used to test for this. Nano-flow testing offers is a quick test but nevertheless is sufficiently useful to provide accurate and consistent data, which could be related to the bench scale test data.

The nano-flow test equipment set-up is similar to the established equipment and procedures of the heptane- cracking test, and all internal consistency checks and comparisons with bench scale data are positive. The heptane-cracking test measures methane yield as is. The feed system design aims at constant concentration, even at fluctuating pressure. A carrier gas, hydrogen, sweeps through a saturator, in which heptane is kept at 110 ⁰C and 36 bar. In order to increase the operating range of this set-up, we had to load more catalyst for the low-pressure experiments.

We included a check-back to link the data at different amounts of catalyst. Above our expectation, a single check-back provided valuable data on both the order of the hydrogenolysis reaction in heptane, and the effect of the pressure at which the catalyst was reduced. Catalyst samples Three catalyst samples were used in this test run

(Catalyst A, E and G)

Four sets of experimental data at different pressure levels are obtained for the above catalysts. The first set was operated at 31 bar, the second at 26 bar, the third at 16 bar, and the fourth at 8 bar at first, and then without reloading, at 31 bar, for the check-back. The onset temperature of methanation is defined as the temperature, which is required for 5% conversion of heptane. At this temperature level, the only products are methane and hexane. The selectivities follow from the reaction stoichiometry. The methane yield at 5% conversion of heptane is equivalent to 5% * 16/100 = 0.8 weight on feed (%wof) . The temperature of methanation for Catalysts G and E were 20 to 30 ⁰C higher than for Catalyst A at each of the three pressures.

It is interesting to note that the check-backs are all at lower temperature requirement, i.e. at higher activity, than the first data point at 31 bar. This is in agreement with our experience from the bench scale experiments, in which the catalyst reduction kept progressing throughout the whole duration of the experiment, which was a full 1500 hours on one occasion. In the heptane cracking equipment, we reduced the catalyst at the same pressure as during start up. The best reduction is achieved at the lowest practical operating pressure. Consequently, the 8 bar data and the check-back are both measured on the most fully reduced catalyst loads. Connecting these data points, we learn that the effect of pressure on the methanation onset temperature is the same for all catalysts: a decrease of one degree C per bar of hydrogen partial pressure reduction. The check-backs also support the hypothesis that the hydrogenolysis reaction is indeed first order in heptane. For the catalyst that is easy to reduce, Catalyst G, the hydrogenolysis activity is the same for 200 mg catalyst, in the first data point, as for 350 mg catalyst, in the check-back. As the reaction rate constant is calculated from the conversion under the assumption of a reaction order of 1, this equal activity appears to confirm this assumption.

In conclusion, the hydrogenolysis data from the heptane-cracking test are not only internally consistent, but are also in good agreement with the bench scale data.

The measured pressure effect on the onset temperature for hydrogenolysis is a decrease of 1 degree C per bar of hydrogen partial pressure reduction. This deduction is applicable to all the Ni catalysts that have been tested.

A small amount of copper may be added to the hydrogenation catalyst to suppress hydrogenolysis of the catalyst to methane.

Thus embodiments of the present invention provide the benefit that the concentration of oxygenates released from the HGU is below detection level, that is below 5 ppm.

Embodiments of the present invention can also be used at a higher temperature, regardless of pressure, before the hydrocarbon chains are broken apart by methanation. Known prior art catalysts were limited to a temperature of around 265 ⁰C before the unwanted methanation reaction occurs. Embodiments of the present invention can operate at up to 280 ⁰C before the onset of methanation. The activity of the catalyst can be increased. Improvements and modifications may be made without departing from the scope of the invention.

Claims

C L A I M S

1. A hydrogenation catalyst, the catalyst comprising a metallic active portion in which the metal is a non-noble Group VIII metal and a support, characterised in that the support does not catalyse an acid catalysed reaction and wherein over 90% of the pores within the support are sized between 100 A-400 A.

2. A hydrogenation catalyst as claimed in claim 1, wherein the median pore diameter is greater than 170 A, preferably wherein less than 11% of the pore volume is provided by pores with a diameter greater than 350 A.

3. A hydrogenation catalyst as claimed in any preceding claim, wherein the support comprises wide pore alumina.

4. A hydrogenation catalyst as claimed in any preceding claim, wherein the active portion comprises nickel, preferably wherein the catalyst comprises at least 5% but less than 20% nickel.

5. A hydrogenation catalyst as claimed in claim 8, wherein the nickel crystallites are around 2.5nm.

6. Use of a hydrogenation catalyst according to any preceding claim in a hydrogenation unit, preferably for hydrogenating olefins and oxygen-containing compounds.

7. Use as claimed in claim 6 wherein the hydrogenation unit is connected to a Fischer-Tropsch unit for the hydrogenation of a Fischer-Tropsch paraffinic wax stream containing olefins and oxygenates.

8. A method for manufacturing a hydrogenation catalyst according to any of claims 1 - 5, the method comprising: admixing a solution of the metal salt with a support; drying and calcining the mixture.

9. A method as claimed in claim 8, wherein the metal salt is mixed in a basic solution.

10. Process for the hydrogenation of the C5+-fraction obtained in a process in which a mixture of hydrogen and carbon monoxide is converted over a cobalt containing catalyst, preferably a cobalt manganese catalyst, especially a titania supported catalyst, in which hydrogenation process a catalyst is used comprising nickel on a support that does not catalyse an acid catalysed reaction as measured by the n-heptane test result at 350 ⁰C, the process being carried out at a temperature between 220 ⁰C and 300 ⁰C and a pressure between 8 and 50 bar.

11. Process according to claim 10 in which the difference between the temperature in ⁰C and the pressure in bar is at least 225, preferably between 230 and 250.

12. Process according to claim 10 or 11 wherein the catalyst in the hydrogenation process is a nickel catalyst according to any of claims 1 to 5.