WO2012071137A1

WO2012071137A1 - Process for hydrocracking butane or naphtha in the presence of a combination of two zeolites

Info

Publication number: WO2012071137A1
Application number: PCT/US2011/058599
Authority: WO
Inventors: Geert Marten Bakker; Saravana Bhavan; Bruce Kaygay; Muralikrishna Venkatacharyulu Khandavilli; Malgorzata Anna Koornneef; Chilkoor Soundararajan Laxmi Narasimhan; Marcello Stefano Rigutto; Ingrid Maria Van Vegchel; Ferry Winter
Original assignee: Shell Oil Company; Shell Internationale Research Maatschappij B.V.
Priority date: 2010-11-01
Filing date: 2011-10-31
Publication date: 2012-05-31

Abstract

The present invention provides a process for preparing a gas cracker feedstock, comprising contacting a feed containing one or more paraffins comprising 4 to 12 carbon atoms with a catalyst comprising at least one zeolite having 10-membered ring channels and at least one group VIb, VIIb and/or VIII metal, in the presence of hydrogen at elevated temperatures and elevated pressures and converting at least 40wt% of the paraffins comprising 4 to 12 carbon atoms based on the total weight of paraffins comprising 4 to 12 carbon atoms in the feed to ethane and/or propane to obtain a hydrocracked gas cracker feedstock comprising ethane and/or propane. In a further aspect the invention also relates to a process for producing olefins.

Description

PROCESS FOR HYDROCRACKING BUTANE OR NAPHTHA IN THE PRESENCE OF A COMBINATION

OF TWO ZEOLITES

Field of the invention

The invention relates to a process for preparing a gas cracker feedstock and a process for producing olefins.

Background of the invention

5 An increased demand from the petrochemistry for

ethylene has led the industry looking for new feedstocks suitable to produce ethylene. Ethylene is typically

produced by thermal cracking of ethane, however can also be produced by thermal cracking of much more abundantly 0 available hydrocarbon fractions comprising predominantly

C4 to C12 paraffins. A drawback of using these higher

hydrocarbon feedstocks is found in the low ethylene yields and high production of heavy by-products such sa pitch.

Therefore, it has been proposed to pre-treat such

5 higher hydrocarbon feedstocks by hydrocracking and to

provide the obtained hydrocracked feedstock to a gas

cracker to produce ethylene.

For instance, in US4137147 it is proposed to subject an oil cut having a distillation range between 0 and 360°C 0 to a hydrogenolysis or hydrocracking step, prior to using

the oil cut as a feed to a steam crack process, i.e. a gas cracking process. The cracking catalyst used in US4137147, is an acid treated small pore Mordenite zeolite. In the examples of US4137147 a conversion of no more than 38.5wt% 5 of the initial charge to ethane and propane was attained.

To attain further conversion a second bed comprising a non-zeolite catalyst was required.

There is a need in the art for a process to prepare

feedstock for a gas cracker to produce ethylene with a 0 high conversion to ethane and/or propane. Summary of the invention

It has now been found that it is possible to

hydrocrack a paraffin-comprising feed to obtain a gas cracker feedstock using specific zeolite catalysts, with a high conversion to ethane and/or propane.

Accordingly, the present invention provides a process for preparing a gas cracker feedstock, comprising contacting a feed containing one or more paraffins comprising 4 to 12 carbon atoms with a catalyst comprising at least one zeolite having 10-membered ring channels and at least one group VIb, Vllb and/or VIII metal, in the presence of hydrogen at elevated temperatures and elevated pressures and converting at least 40wt% of the paraffins comprising 4 to 12 carbon atoms based on the total weight of

paraffins comprising 4 to 12 carbon atoms in the feed to ethane and/or propane to obtain a hydrocracked gas cracker feedstock comprising ethane and/or propane.

Reference herein to a zeolite is to a molecular sieve aluminosilicate material. Reference herein to a zeolite having 10-membered ring channels is to a zeolite or aluminosilicate having 10-membered ring channels in one direction, optionally intersected by 8, 9 or 10-membered ring channels in another direction.

The process according to the present invention can be operated at high feed conversions, reducing the need to provide a second or further cracking process.

In another aspect the invention provides a process for producing olefins, comprising:

i) providing a gas cracker feedstock obtained from a process for preparing gas cracker feedstock according to the invention; and

ii) cracking the gas cracker feedstock under cracking conditions to obtain olefins. Brief description of the figure

In Figure 1 a schematic representation of an

embodiment of a process according to he invention is provided

Detailed description of the invention

In process according to the present invention a gas cracker feedstock is prepared by hydrocracking a feed containing one or more paraffins comprising 4 to 12 carbon atoms. The feed containing one or more paraffins

comprising 4 to 12 carbon atoms is further also referred to as paraffinic feed. The obtained gas cracker feedstock may subsequently be thermally cracked in a gas cracker to at least ethylene in the presence of steam. Although such a paraffinic feed can be used directly as feedstock to a gas cracker without any prior hydrocracking step, this typically results in lower ethylene yields and high by^¬ product formation such as pitch. It is known that high ethylene yields can be obtained when the feedstock to the gas cracker is rich in ethane and/or propane. Therefore, it has been proposed to hydrocrack the paraffinic feed prior to providing the feed to a gas cracker.

Gas crackers designed to operate on ethane and or propane, however are sensitive to the presence of C4 or higher paraffins in the feedstock to the gas cracker. The presence of such paraffins again leads to production of by-products including pitch and methane. Therefore, there is a need to attain a high conversion of the paraffin feedstock, preferably with a high yield in ethane and propane. Otherwise, it may be necessary to treat the hydrocracked feedstock to remove the higher paraffins in order to obtain a suitable gas cracker feedstock.

In the process according to the present invention the paraffinic feed is hydrocracked by contacting the paraffinic feed under hydrocracking conditions with a catalyst comprising at least one zeolite having 10- membered ring channels in the presence of hydrogen.

Zeolite having 10-membered ring channels as used in the present invention are sleeted from:

- one dimensional zeolites or aluminosilicates having only 10 membered ring channels in one direction, which are not intersected by other channels, in particular other 8, 10 or 12-membered ring channels, from another direction.

- multi dimensional zeolites or aluminasilicates having intersecting channels in at least two directions, whereby at least the channels in one direction are 10-membered ring channels, intersected by 8, 9 or 10-membered ring channels in another direction.

Examples of suitable zeolites having 10-membered ring channels for use in the process according to the present invention include, but are not limited to MFI-type, MEL- type, TON-type, MTT-type, ZSM-48-type, ITH-type and FER- type zeolites.

MFI-type zeolites have a three dimensional structure.

One preferred MFI-type zeolite is ZSM-5.

MEL-type zeolites have a three dimensional structure. One preferred MEL-type zeolite is ZSM-11

TON-type zeolites are more particularly described in e.g. US-A-4 , 556 , 477. For purposes of the present

invention, TON is considered to include its isotypes, e.g., ZSM-22, Theta-1, ISI-1, KZ-2 and NU-10. One

preferred TON type zeolite is ZSM-22.

MTT-type zeolites are more particularly described in e.g. US-A-4 , 076 , 842. For purposes of the present

invention, MTT is considered to include its isotypes, e.g., ZSM-23, EU-13, ISI-4 and KZ-1. One preferred MTT type zeolite is ZSM-23. ZSM-48-type zeolites are more particularly described in e.g. US-A-4, 397, 827. For purposes of the present invention, ZSM-48 is considered to include its isotypes. Preferred ZSM-48-type zeolites are EU-2 and ZSM-48.

ITH-type zeolites have a three dimensional structure.

One preferred ITH-type zeolite is ITQ-13.

FER-type zeolites have a two-dimensional structure. A preffered FER-type zeolite is Ferrierite

The above mentioned zeolite types and zeolites are for example defined in Ch . Baerlocher and L.B. McCusker,

Database of Zeolite Structures: h1.1p : / /www . i za- structure . org/databases/ , which database was designed and implemented on behalf of the Structure Commission of the International Zeolite Association (IZA-SC), and based on the data of the 4th edition of the Atlas of Zeolite

Structure Types (W.M. Meier, D.H. Olson and Ch .

Baerlocher) .

TON-type, MTT-type, and EU-2-type zeolites are one dimensional zeolites, i.e. they have a one dimensional channel structure, wherein the 10-membered ring channels of the zeolite do not intersect with 8, 10 or 12-membered ring channels in another direction.

The FER-type zeolites are two dimensional, i.e. they have a two dimensional channel structure. The 10-membered ring channels of the zeolite intersect with 8-membered ring channels in another direction.

The MFI and the MEL-type zeolites are three

dimensional zeolites, i.e. they have a three dimensional channel structure. Both having intersecting 10-membered ring channels from three directions.

Preferably, the catalyst comprises at least one of ZSM-5, ZSM-11, ZSM-22, ZSM-23, ZSM-48, or Ferrierite. More preferably, ZSM-5, ZSM-11 or ZSM-22. The more preferred zeolites combine a high selectivity to ethane and/or propane with a high conversion of the paraffinic feed.

The catalyst may also contain two or more zeolites. In case the catalyst comprises two or more zeolites it preferably comprises two zeolites having a 10-membered ring channel, of which at least one zeolite has an one dimensional channel structure and at least another zeolite has a multi-dimensional channel structure. More

preferably, the catalyst comprises ZSM-22 and ZSM-5 or ZSM-22 and ZSM-11. Without wishing to be bound to any particular theory, it is believed that the use of one zeolite having an one dimensional channel structure and one zeolite having a multi-dimensional channel structure has a synergetic effect wherein the multi-dimensional zeolite particularly benefits the cracking of the more refractory paraffins. This results in a higher ethane yield, which is advantageous as ethane can be converted to ethylene with higher yields in a subsequent cracking process in a gas cracker.

Preferably, the catalyst comprises one dimensional channel structure zeolites, having a 10-membered ring channel, and multi-dimensional channel structure zeolites, having a 10-membered ring channel, in a weight ratio of in the range of 10:1 to 1:10, more preferably 3:1 to 1:3, even more preferably 1.5:1 to 1:1.5.

To provide sufficient acidity for the hydrocracking reaction, it is preferred that the zeolites are at least partly in the hydrogen form, e.g. H-ZSM-5, H-ZSM-11, H- ZSM-22, H-ZSM-23, H-ZSM-48 or H-FER. Preferably, at least 50wt%, more preferably at least 90wt%, still more

preferably at least 95wt% and most preferably 100wt% of the total amount of zeolite used is in the hydrogen form. When the zeolites are prepared in the presence of organic cations, the zeolites may be activated by heating in an inert or oxidative atmosphere to remove organic cations, for example, by heating at a temperature over 500 °C for 1 hour or more. The zeolite is typically obtained in the sodium or potassium form. The hydrogen form can then be obtained by an ion exchange procedure with ammonium salts followed by another heat treatment, for example in an inert or oxidative atmosphere at a temperature over

300 °C. The zeolites obtained after ion-exchange are also referred to as being in the ammonium form.

The silica-to-alumina-ratio (SAR) is defined as the molar ratio of S1O2/ I2O3 corresponding to the composition of the zeolite, i.e. aluminosilicate molecular sieve.

Typically, the zeolite has a SAR in the range of from 8 to 500. Preferably, the zeolite has a SAR in the range of from 10 to 200, more preferably 16 to 150.

It is desirable to provide a catalyst having good mechanical or crush strength, or attrition resistance, because in an industrial environment the catalyst is often subjected to rough handling, which tends to break down the catalyst into powder-like material. The latter causes problems in the processing. Preferably the zeolite is therefore mixed with a matrix and a binder material and then spray dried or shaped to the desired shape, such as pellets or extrudates. Examples of suitable binder

materials include active and inactive materials and synthetic or naturally occurring zeolites as well as inorganic materials such as clays, silica, alumina, silica-alumina, titania, zirconia and zeolite. For present purposes, materials, such as silica and alumina

arepreferred because they may prevent unwanted side reactions .

Preferably the catalyst used in the process of the present invention comprises, in addition to the zeolite, 2 to 90 wt%, preferably 10 to 85 wt% of a binder material. The catalyst comprises one or more metals. The

catalyst comprises at least one group VIb, Vllb and/or VIII metal, wherein the group number refers to a group in the Periodic Table of Elements, preferably the catalyst comprises at least one group VIb and or VIII metals, even more preferably at least one group VIII metal. Next to the hydrocracking activity, the incorporation of one or more metals selected from group VIb, Vllb and VIII, optionally in combination with another metal, may reduce the activity of the catalyst towards hydrogenolysis, in particular at temperature above 350 °C, resulting in a reduced methane make .

One preferred catalyst comprises one or more group VIII metals, more preferably one or more VIII noble metals such as Pt, Pd, Rh and Ir, even more preferably Pt and/or

Pd. The catalyst preferably comprises in the range of from 0.05 to 10wt%, more preferably of from 0.1 to 5wt%, even more preferably of from 0.1 to 3wt% of such metals, based on the total weight of the catalyst.

Another preferred catalyst comprises at least one group VIb, Vllb and/or VIII metal in combination with one or more other metals, i.e. metals which are not from group VIb, Vllb or VIII. Examples of such combinations of a group VIb, Vllb and VIII in combination with another metal include, but are not limited to PtCu, PtSn or NiCu. The catalyst preferably comprises in the range of from 0.05 to 10wt%, more preferably of from 0.1 to 5wt%, even more preferably of from 0.1 to 3wt% of such metals, based on the total weight of the catalyst.

Yet another preferred catalyst comprises a combination of a group VIb and a group VIII metal. Examples of such combinations of a group VIb and group VIII metal include, but are not limited to, CoMo, NiMo and NiW. The catalyst preferably comprises in the range of from 0.1 to 30wt%, more preferably of from 0.5 to 26wt%, based on the total weight of the catalyst.

In case, the feed to the process comprises substantial amounts of sulphur, i.e. sulphur in the form of elemental sulphur or sulphur compounds, it may be preferred to use a catalyst comprising a combination of a group VIb and a group VIII metal. In particular, when the feed to the process comprises more than lOppmw, more in particular in the range of from lOppmw to 5wt%, of sulphur, based on the weight of the sulphur atoms and the total weight of the feed, it is preferred to use a catalyst comprising a combination of a group VIb and a group VIII metal, more preferably CoMo, NiMo and/or NiW. These catalysts are particularly suitable in combination with feedstocks comprising sulphur or sulphur-comprising compounds.

Although essentially not required, it may be preferred to pre-sulphide the metal or metals in the catalyst and convert them in a sulfide and/or sulfidic state. Such pre-sulphide pre-treatments are well known in the art and do not need further explanation.

In case, the feed to the process comprises little or no sulphur, i.e. sulphur in the form of elemental sulphur or sulphur compounds, it may be preferred to use a

catalyst comprising one or group VIII metals, optionally in combination with one or more other metals, i.e. metals which are not from group VIb, Vllb or VIII. In particular, when the feed to the process comprises in the range of from 0 to 200ppmw, more in particular in the range of from 0 to lOOppmw, of sulphur, based on the weight of the sulphur atoms and the total weight of the feed, it is preferred to use a catalyst comprising one or group VIII metals, optionally in combination with one or more other metals, i.e. metals which are not from group VIb, Vllb or VIII . The metals may be incorporated into the catalyst by any suitable method known in the art. Examples of such methods include, but are not limited to, impregnation, ion-exchange or other deposition techniques using

solutions containing metal salts, typically followed by a calcination, reduction step and/or drying step.

The metals may be incorporated directly into the zeolite or into the formulated catalyst. The metal-loaded zeolite may be subjected to further calcinations and/or reduction steps directly or may be first be added to a catalyst formulation, thereafter the catalyst formulation, including the metal-loaded zeolite, is subsequently subjected to further calcination and/or reduction steps.

Alternatively, metals may be incorporated in the catalyst by a co-mulling process, wherein metal compounds are mixed with the other catalyst ingredients during extrusion of the catalyst.

In the process according to the invention, the feed is contacted with the catalyst at elevated temperatures and elevated pressures.

Preferably, the paraffinic feed is contacted with the catalyst at a temperature in the range of from 200 to 650°C, more preferably 250 to 550°C. The exact choice of the temperature will depend on the composition of the feed to the process and the desired product . Preferably, the paraffinic feed is contacted with the catalyst at a pressure of in the range of from 3 to 100 bar (gauge) , more preferably 5 to 60 bar (gauge) .

Preferably, the paraffinic feed is contacted with the catalyst at a weight hourly space velocity (WHSV) of in the range of from 0.1 to 20 hr^_1, more preferably 0.5 to 10 hr^-1.

Hydrogen may be provided in any suitable ratio to the paraffinic feed. Preferably, the hydrogen is provided in a molar ratio hydrogen to paraffin feedstock of 1:1 to

100:1, more preferably 1.5:1 to 50:1, wherein the number of moles of the paraffinic feedstock is based on the average mol weight to the paraffinic feedstock. The process according to the invention to achieve a set conversion of the hydrocarbon feedstock. Preferably, the ratio of hydrogen to paraffinic feedstock is chosen such that under the process conditions achieving the desired conversion, the hydrogen content in the obtained gas cracker feedstock is no more than lwt%, based on the total weight of the obtained gas cracker feedstock. This is beneficial as hydrogen is an undesired component of the gas cracker feedstock.

In the process according to the invention, at least 40wt% of the paraffins comprising 4 to 12 carbon atoms in the paraffinic feed are converted to ethane and/or

propane, preferably at least 50 wt%, more preferably

60 wt% of paraffins in the paraffinic feed are converted to ethane and/or propane. Reference herein to the wt% of paraffins comprising 4 to 12 carbon atoms is to the wt% of paraffins comprising 4 to 12 carbon atoms based on the total weight of paraffins comprising 4 to 12 carbon atoms in the paraffinic feed.

Depending on the desired composition of the gas cracker feedstock, the reaction conditions may be selected to obtain a high conversion of the paraffinic feed to ethane. Preferably, the gas cracker feedstock comprises a high ethane content. More preferably, the gas cracker feedstock comprises ethane and propane in a weight ratio in the range of from 3:2 to 500:1, even more preferably a weight ratio in the range of from 7:3 to 100:1. This may for example be achieved by selecting temperatures at the higher end of the preferred temperature range and/or selecting a WHSV at the lower end of the preferred WVSH range. Ethane is the preferred component in a gas cracker feedstock, however, the high temperature and prolonged residence time may result in an increased undesired methane make .

By decreasing the temperature and/or increasing the

WHSV relative to the temperature and WHSV used to convert the paraffinic feedstock to predominantly ethane, the paraffinic feed is predominantly converted to a mixture of ethane and propane. This may for example be achieved by decreasing the temperature and/or increasing the WHSV relative to the temperature and WHSV used to convert the paraffinic feedstock to predominantly ethane. Although, propane is a lesser preferred component in the gas cracker feedstock, also less undesired methane is produced.

In case it is intended to convert the paraffinic feed to predominantly ethane, the catalyst preferably comprises at least one zeolite having a 10-membered ring channel and a one dimensional channel structure. When it is intended to convert the paraffinic feed to predominantly a mixture of ethane and propane, the catalyst preferably comprises at least one zeolite having a 10-membered ring channel and a multi dimensional channel structure.

The paraffinic feed (i.e. the feed containing one or more paraffins comprising 4 to 12 carbon atoms) may be any feed containing one or more paraffins comprising 4 to 12 carbon atom.

The paraffinic feed may comprise compounds other than paraffins. Preferably, the paraffinic feed comprises at least 10wt% of paraffins comprising 4 to 12 carbon atoms, more preferably at least 50wt%, more preferably at least

60wt% of paraffins comprising 4 to 12 carbon atoms.

Preferably, the paraffinic feed comprises in the range of from 10 to 100wt% of paraffins comprising 4 to 12 carbon atoms, more preferably of from 50wt% to 99.5wt%, more preferably of from 60wt% to 99wt of paraffins comprising 4 to 12 carbon atoms. Reference herein to the wt% of

paraffins comprising 4 to 12 carbon atoms in the

paraffinic feed is to the wt% of paraffins comprising 4 to 12 carbon atoms based on the total weight the paraffinic feed .

Preferably, the paraffinic feed is or is derived from naphtha. The paraffinic feed may comprise straight run naphtha or naphtha fractions derived from natural gas, natural gas liquids or associated gas. The paraffinic feed may comprise naphtha fractions derived from pyrolysis gas. The paraffinic feed may also comprise naphtha or naphtha fraction obtained from a Fischer-Tropsch process for synthesising hydrocarbons from hydrogen and carbon

monoxide.

The feed may also comprise higher paraffins, i.e.

paraffins comprising more than 12 carbon atoms. Cracking such higher paraffins typically requires the use of temperatures and pressures which are at the higher end of the preferred temperature and pressure ranges.

Preferably, the paraffinic feed comprises at least one isoparaffin comprising 4 to 12 carbon atoms. Preferably, the paraffinic feed comprises in the range of from 10 to 90wt% of isoparaffins, more preferably 20 to 60wt% of isoparaffin, based on the total weight of the paraffinic feed .

A preferred paraffinic feed comprises isobutane.

Preferably, the paraffinic feed comprises in the range of from 10 to 90wt% of isobutane, more preferably 20 to 60wt% of isobutane, based on the total weight of the paraffinic feed .

Another preferred paraffinic feed comprises n-butane. Preferably, the paraffinic feed comprises in the range of from 10 to 90wt% of n-butane, more preferably 20 to 60wt% of n-butane, based on the total weight of the paraffinic feed .

Particular preferably, the paraffinic feed comprises at least n-butane and isobutane. More preferably, the paraffinic feed comprises in the range of from 10 to

100wt%, more preferably of from 50 to 99.5wt% of n-butane and isobutane, based on the total weight of the paraffinic feed .

The paraffinic feedstock may comprise olefins.

However, as the olefins are hydrogenated during the hydrocracking process, the presence of olefins results in an undesired increased hydrogen consumption. Preferably, the paraffinic feed comprises in the range of from 0 to 20wt% of olefins, based on the total weight of the

paraffinic feed, more preferably of from 0 to 10wt% of olefins. Optional, the feedstock is subjected to a

hydrogenation treatment prior to being supplied to a process according to the present invention.

In another aspect the present invention provides a process for producing olefins. In this process the gas cracker feedstock is subsequently cracked to obtain olefins, preferably including ethylene. Additionally, propylene may be formed. Other by-products may be formed such as butylene, butadiene, ethyne, propyne and benzene. The cracking process is performed at elevated

temperatures, preferably in the range of from 650 to

1000°C, more preferably of from 750 to 950°C. Typically, the cracking is performed in the presence of water (steam) as a diluent. The conversion of ethane and propane is typically in the range of from 40 to 75 mol%, based on the total number of moles ethane and propane provided.

Preferably, the un-cracked ethane and propane are recycled back to the cracking zone. Cracking processes for

producing ethylene and propylene are well known to the skilled person and need no further explanation. Reference is for instance made to Kniel et al . , Ethylene, Keystone to the petrochemical industry, Marcel Dekker, Inc, New York, 1980, in particular chapter 6 and 7.

Examples

The invention will be illustrated by the following non-limiting examples.

Example 1. Butane hydrocracking

Catalyst preparation

Zeolites were used in the ammonium form (NH₄ ⁺) . The zeolite powders were first pressed into pills, crushed and sieved in a 30-80 mesh sieve fraction. Next, the materials were impregnated with a solution containing Pt (N¾ ) ₄ (NO₃ ) ₂ to arrive at a final Pt loading of 0.7 wt%. The catalysts were calcined at for lh 350 °C and for 1 h at 500 °C .

Catalyst A: 0.7 wt% Pt on zeolite ZSM-22,

SAR (Si02/A1203 molar ratio) = 50

Catalyst B: 0.7 wt% Pt on zeolite ZSM-22/ZSM-5,

ZSM-22 and ZSM-5 were present in a weight ratio of 85:15; SAR = 82

Comparative catalyst C: 0.7 wt% Pt on zeolite Beta,

SAR = 249

Comparative catalyst D: 0.7 wt% Pt on zeolite ZSM-12, SAR = 90

To test the catalytic performance of catalysts A to D, the respective catalyst powder was pressed into tablets and the tablets were broken into pieces and sieved.

Example la. n-Butane hydrocracking

A mixture of hydrogen and n-butane was reacted over the catalysts.

The reaction was performed using a stainless steel reactor tube of 1.8 mm internal diameter. About 80 mg of catalyst was placed in the reactor, dried in a nitrogen flow at 100°C, and then reduced in a mixture of 15% hydrogen and 85% nitrogen at atmospheric pressure and 400°C.

The catalyst samples were the brought to 250°C after which a mixture consisting of 2.7 vol% n-butane, 7 vol% H₂, balanced in argon and N₂ was passed over the catalyst at a total pressure of 29 barg (bar gauge) to determine their selectivity towards ethane and propane. A Gas Hourly Space Velocity (GHSV) of 17, 000 ml . g ( catalyst ) ^"1. h^_1, determined by the total gas flow over the catalyst weight per unit time, was used. The catalysts were tested in the temperature range of 250-550°C.

The effluent from the reactor was analyzed by gas chromatography (GC) to determine the product composition. The composition has been calculated on a weight basis of all hydrocarbons analyzed. The selectivity has been defined by the division of the mass of product by the sum of the masses of all C1-C3 products.

Using the above experimental procedure each of the catalyst A to D was tested:

Experiment 1A and 1AA: Catalyst A

Experiment IB and IBB: Catalyst B

Comparative experiment 1C and ICC: Catalyst C

Comparative experiment ID And 1DD: Catalyst D

In Table 1 the product distribution is given for experiments conducted with n-butane with all examples having a conversion between 44 and 62wt% conversion of the C4+ fraction, experiments 1 A, IB, 1C and ID.

In Table 2, the product distribution is given for experiments with n-butane at similar reaction conditions, i . e . at a temperature of 325°C, experiments 1 AA, IBB, ICC and 1DD.

Example lb. Isobutane hydrocracking

A mixture of hydrogen and isobutane was reacted over the catalysts . The reaction was performed using a stainless steel reactor tube of 1.8 mm internal diameter. About 80 mg of catalyst was placed in the reactor, dried in a nitrogen flow at 100°C, and then reduced in a mixture of 15% hydrogen and 85% nitrogen at atmospheric pressure and

400°C.

The catalyst samples were the brought to 250°C after which a mixture consisting of 2.7 vol% isobutane, 7 vol% ¾, balanced in argon and 2 was passed over the catalyst at a total pressure of 29 barg (bar gauge) to determine their selectivity towards ethane and propane. A Gas Hourly Space Velocity (GHSV) of 17, 000 ml . g ( catalyst ) ^_1. h^"1, determined by the total gas flow over the catalyst weight per unit time, was used. The catalysts were tested in the temperature range of 250-550°C.

Using the above experimental procedure each of the catalyst A to D was tested:

Experiment IE and 1EE: Catalyst A

Experiment IF and IFF: Catalyst B

Comparative experiment 1G and 1GG: Catalyst C

Comparative experiment 1H and 1HH: Catalyst D

Table 3 gives the results from experiments performed with iso-butane at 400°C, experiments 1 E, IF, 1G and 1H and 450°C, experiments 1 EE, IFF, 1GG and 1HH.

Results

As can be seen from Table 1-3, the conversion with the use of 10-membered ring zeolites, i.e. catalyst A and B resulted in higher conversion at the same reaction temperature compared to the comparative catalyst C and D in which 12-membered ring zeolites were used.

The selectivity to ethane were in all cases higher for catalyst A and B compared to the comparative experiments using catalyst C and D. Even at very high conversion

(97wt%) the catalyst A displayed a selectivity of 68wt% to ethane .

C4+ CI C2 C3 Selectivity Selectivity Selectivity experiment Temp conversion yield yield yield to CI to C2 to C3

°C wt% wt% wt% wt% wt% wt% wt%

Exp. 1A 300 61.9 7.4 38.0 16.4 12.0 61.5 26.6

Exp. IB 300 44.1 5.8 23.6 14.7 13.1 53.4 33.4

Exp. 1C* 400 50.1 9.3 19.5 21.3 18.6 38.9 42.5

Exp. ID* 300 50.8 8.7 19.5 22.7 17.1 38.3 44.6

Table 1.

Comparative example

Table 2.

comparative example

Table 3.

C4+ CI C2 C3 Selectivity Selectivity Selectivity

Experiment Temp conversion yield yield yield to CI to C2 to C3

°C wt% wt% wt% wt% wt% wt% wt%

Exp. IE 400 71.1 12.9 31.4 26.8 18.1 44.2 37.7

Exp. IF 400 53.3 10.6 23.9 18.8 19.8 44.9 35.3

Exp. 1G* 400 4.8 1.0 1.2 2.6 20.3 25.9 53.8

Exp. 1H* 400 7.4 1.2 2.2 4.0 15.6 29.5 54.9

Exp. 1EE 450 91.0 29.4 48.8 12.8 32.3 53.6 14.1

Exp. IFF 450 88.0 50.2 37.3 0.6 57.0 42.3 0.7

Exp. 1GG* 450 9.9 8.1 1.6 0.1 82.3 16.2 1.5

Exp. 1HH* 450 7.9 5.6 2.2 0.1 70.6 27.5 1.9 comparative example

Example 2 Naphtha hydrocracking

Catalyst preparation

Catalysts were prepared following the same procedure as described under Example 1.

The following catalyst were prepared:

Catalyst E: 0.7 wt% Pt on zeolite ZSM-5,

SAR=80

Catalyst F: 0.7 wt% Pt on zeolite ZSM-11,

SAR=56

Comparative catalyst G: 0.7 wt% Pt on zeolite ZSM-12,

SAR=47

Comparative catalyst H: 0.7 wt% Pt on zeolite Y,

SAR=70

To test the catalysts from examples E to F for

catalytic performance the respective catalyst powder was pressed into tablets and the tablets were broken into pieces and sieved.

Hydrocracking of Naphtha

A mixture of hydrogen and naphtha vapour was reacted over the catalysts. The naphtha was a hydrotreated

straight run naphtha, containing less than 10 ppm sulphur, and further characterized by the PIONA analysis in Table 4. The reaction was performed using a stainless steel reactor tube of 2.0 mm internal diameter. About 75 mg of dried catalyst was placed in the reactor and reduced in a flow of hydrogen at 30 barg and 350°C.

After reduction, a mixture of vaporized naphtha and hydrogen was passed over the catalyst at a pressure of 30 barg and an initial temperature of 350 °C. Naphtha was fed at a weight hourly space velocity of

1.0 g (naphtha) . g ( catalyst ) ^_1. h^_1 and hydrogen at a rate of 2.25 mol/mol (feed) -h based on an average feed molecular weight of 92.5 g/mol. The catalysts were tested at

temperatures in the range of 350-500°C. The effluent from the reactor was analyzed by online gas chromatography (GC) to determine the product composition. The composition has been calculated on a weight basis of all hydrocarbons analyzed. The selectivity has been defined by the division of the mass of product by the sum of the masses of all Cl-

C3 products.

Using the above experimental procedure each of the catalyst E to G was tested:

Experiment 2A: Catalyst E

Experiment 2B: Catalyst F

Comparative experiment 2C: Catalyst G

Comparative experiment 2D: Catalyst H

Table 5 compares product distributions obtained from experiments 2A to 2D at high cracking conversions, i.e. a C4+ conversion over 75wt% and a C3+ conversion over 36wt%.

Based on the same experiments 2A to 2D, a comparison is made in Table 6, at approximately constant methane makes of approximately 12wt% (11.5 to 14.9wt%), based on the same experiments 2A to 2D.

Results

From Table 5 is can be seen that hydrocracking of naphtha employing 10-membered ring zeolites (catalyst E and F) resulted in appreciably higher selectivity to ethane and lower selectivity to methane at comparable conversion levels (in the range of from 75 to 83wt% for

C4+ and 35 to 41wt% for C3+) compared to the hydrocracking of naphtha employing a 12-membered ring zeolite (catalyst

G and H) .

In Table 6, a comparison is made at approximately constant methane makes of approximately 12wt% (11.5 to

14.9 wt%) . Here, again appreciably higher ethane yields are obtained for the hydrocracking of naphtha using 10- membered ring zeolites (catalyst E and F) than for the 12- membered ring zeolites (catalyst G and H) . Table 4.

Table 5.

* comparison at C3+ conversion levels around 40 w%, however in experiment 2D this level was not reached - instead, data obtained at the highest achieved conversion level is included

* comparative example

5 Table 6.

comparative example

Example 3: Integrated process to produce olefins

Using an internal model for a steam cracker and ASPEN simulations several embodiments of integrated process to produce olefins have been modelled aimed at maximizing the final output from a steam cracker by pre-treating a feed by hydrocracking . In Table 7, optional conditions for the hydrocracking step to suit different feed and steam cracker configurations are provided.

In Figure 1 a schematic representation of an

embodiment of a process modelled to obtain the embodiment as provided in the second row of Table 7 is provided.

In Table 8, an explication of the used symbols and the calculated composition and size of several streams is provided .

Table 7

Table 8

Claims

C L A I M S

1. A process for preparing a gas cracker feedstock, comprising contacting a feed containing one or more paraffins comprising 4 to 12 carbon atoms with a catalyst comprising at least one zeolite having 10-membered ring channels and at least one group VIb, Vllb and/or VIII metal, in the presence of hydrogen at elevated

temperatures and elevated pressures and converting at least 40wt% of the paraffins comprising 4 to 12 carbon atoms based on the total weight of paraffins comprising 4 to 12 carbon atoms in the feed to ethane and/or propane to obtain a hydrocracked gas cracker feedstock comprising ethane and/or propane.

2. A process according to claim 1, wherein the feed comprises at least 10 wt% of paraffins comprising 4 to 12 carbon atoms based on the total weight the feed.

3. A process according to claim 1 or 2, wherein the feed is or is derived from naphtha.

4. A process according to any of the preceding claims, wherein the feed comprises at least n-butane and

isobutane, preferably comprises in the range of from 50 to

99wt% of n-butane and isobutane, based on the total weight of the feed.

5. A process according to any one of the preceding claims, wherein the catalyst comprises at least one zeolite selected from ZSM-5, ZSM-11, ZSM-22, ZSM-48, ZSM-

23, ITQ-13 and Ferrierite, more preferably selected from ZSM-5, ZSM-11 and ZSM-22.

6. A process according to any one of the preceding claims, wherein the catalyst comprises two or more

zeolites.

7. A process according to claim 6, wherein the catalyst comprise at least: one zeolite having a 10-membered ring channel and an one dimensional channel structure; and

one zeolite having a 10-membered ring channel and a multi-dimensional channel structure, preferably ZSM-22 and ZSM-5.

8. A process according to any one of the preceding claims, wherein the feed is contacted with the catalyst at a temperature in the range of from 200 to 650°C,

preferably of from 250 to 550°C.

9. A process according to any one of the preceding claims, wherein the feed is contacted with the catalyst at a pressure in the range of from 3 to 100 bar (gauge) , preferably of from 5 to 60 bar (gauge) .

10. A process according to any one of the preceding claims, wherein the feed comprises in the range of from 0 to 200ppmw of sulphur or sulphur compounds, based on the weight of the sulphur atoms and the weight of the total feed and wherein the catalyst comprises a Group VIII metal, preferably Pt and/or Pd.

11. A process according to any one of the preceding claims, wherein the feed comprises at least lOppmw of sulphur or sulphur compounds, based on the weight of the sulphur atoms and the weight of the total feed and wherein the catalyst comprises a Group VIII metal and a Group VIb metal, preferably CoMo, NiW and/or NiMo .

12. A process according to any one of the preceding claims, wherein the obtained gas cracker feedstock

contains no more than 1 wt% of hydrogen, based on the total weight of the obtained gas cracker feedstock.

13. A process for producing olefins, comprising:

i) providing a gas cracker feedstock obtained from a process according to any one of the preceding claims; and ii) cracking the gas cracker feedstock under cracking conditions to obtain olefins.

14. A process according to claim 13, wherein the obtained olefins comprise at least ethylene.