WO2006000449A1

WO2006000449A1 - Zeolite catalysts

Info

Publication number: WO2006000449A1
Application number: PCT/EP2005/006927
Authority: WO
Inventors: Klaus Jens; Trond Myrstad; Ivar Dahl; Ase Slagtern
Original assignee: Borealis AS
Current assignee: Borealis AS
Priority date: 2004-06-28
Filing date: 2005-06-28
Publication date: 2006-01-05
Anticipated expiration: 2006-12-28
Also published as: GB0414442D0; US20070246400A1; CN1997451A; EP1761333A1

Abstract

A process for cracking C4+ hydrocarbons comprising heating the hydrocarbons at a temperature of 400-800°C in the presence of an H-ZSM-5 catalyst having a silicon to aluminium ratio of 61-1000.

Description

Zeolite Catalysts

This invention relates to the cracking of higher hydrocarbons into alkenes using the hydrogen form of pentasil type zeolite catalysts (HZSM-5). In particular, the invention relates to the use of HZSM-5 catalysts having specific silicon to aluminium ratios in the cracking of hydrocarbons to give high yields of propylene. Zeolites are a well-known class of aluminosilicates containing an (Si, Al)_nO_2n framework with negative charge balanced by cations present in the framework cavities. The nature of the cations used can vary greatly but typical examples include potassium, sodium, calcium, silver and hydrogen. The spaces created within a zeolite framework are relatively large and these so called "cages" can accommodate large molecules. However, the external pores of the framework which allow access to the cages of the zeolite tend to be smaller, allowing control over the sizes of molecules which can enter and leave the zeolite structure. The possibility of size selectivity has made use of zeolite catalysts an attractive option for the skilled chemist when carrying out cracking reactions, i.e. reactions where higher hydrocarbons are broken down into smaller hydrocarbon fragments, as well as other reactions such as isomerisation reactions or aromatisation reactions. The cracking of higher hydrocarbons into smaller hydrocarbons, in particular ethylene and propylene, is well known and has been carried out in the petrochemical industry for many years. For example, the use of a steam cracker to convert higher hydrocarbons by pyrolysis into ethylene and propylene is well known. In general however this reaction produces much more ethylene than propylene. Zeolite catalysis therefore provides a further cracking alternative for the skilled chemist. US 6222087 describes the use of phosphorus doped ZSM-5 and/or ZSM-11 zeolites in the formation of light olefins rich in propylene from a hydrocarbon feed containing C_4-7 olefms/paraffins. With a 1-butene feed the invention realises 37% ethylene and propylene but with a light catalytic naphtha feed only 25% ethylene and propylene is formed. US 5968342 describes the use of various metal doped zeolites as catalysts in the conversion of naphtha to ethylene and propylene. ZSM-5 catalysts in which the proton is exchanged with silver or copper are suggested to ensure that the zeolite is substantially free of protons. The text suggests that HZSM catalysts show poor selectivity for lower olefins. GB 2345294 describes a process similar to that in US 5968342, i.e. one in which a proton free zeolite is employed in a cracking reaction. WO 00/18853 describes ZSM-5 catalysts doped with phosphorus and a further promoter such as tin or gallium for use in cracking naphtha to olefins. With the increasing demand for propylene throughout the world, in particular the increasing demand for polypropylene, it would be useful if more of the heavier petrochemical fractions could be converted efficiently into propylene for polymerisation. Such heavier petrochemical fractions are often rich in aromatic compounds notably benzene. The use of aromatic compounds in fuels is being reduced for environmental reasons resulting in a potential surplus of aromatic compounds. Moreover, whilst there is a market for benzene, the market is relatively small. Thus, it is anticipated that the market for aromatic compounds may become saturated. It would therefore be useful if aromatic compounds such as benzene could themselves be converted in high yield into more marketable olefins such as ethylene and propylene. There remains therefore, the need for further cracking processes to be developed to maximise the yields of ethylene and propylene which can be obtained from hydrocarbon feeds. In particular it would be useful if the amounts of propylene produced could be maximised. It has now been surprisingly found that by using a hydrogen ZSM-5 catalyst with a particular silicon to aluminium molar ratio hydrocarbons, such as hydrogenated aromatic fractions or hydrocarbon fractions containing higher olefins, can be cracked into ethylene and propylene in very high yields. Contrary to the teachings of the prior art, these catalysts show excellent olefin selectivity at the Si/Al molar ratios claimed and can be used to crack a variety of feeds to give rise to olefins such as ethylene and propylene in high yield. Thus, viewed from one aspect the invention provides a process for cracking C₄+ hydrocarbons comprising heating the hydrocarbons at a temperature of 400- 800⁰C in the presence of an H-ZSM-5 catalyst having a silicon to aluminium ratio of 61-1000. Viewed from another aspect the invention provides the use of an HZSM-5 catalyst having a silicon to aluminium ratio of 61 - 1000 as a catalyst in the cracking of C₄+ hydrocarbons. The process of the invention may be used on its own or in combination with other processes such as alternative olefin producing processes, e.g. pyrolytic steam cracking. By C₄+ hydrocarbons is meant that the hydrocarbons being cracked have at least four carbon atoms. Preferably, the hydrocarbons being cracked will have between 4 and 20 carbon atoms, preferably between 5 and 12 carbon atoms, e.g. 6 to 10 carbon atoms. The hydrocarbon feedstock can be a single pure hydrocarbon but more usually it will be a mixture of various hydrocarbons, e.g. a light or heavy naphtha fraction, different condensate fractions or a hydrogenated pyrolysis gasoline. The hydrocarbons may be saturated or unsaturated, linear, branched or cyclic and preferably non-aromatic. Preferred hydrocarbons however will be saturated paraffins (alkanes such as pentane, hexane, octane, decane, dodecane), olefins of general formula C_nH_2n where n is from 4 to 12 or most preferably saturated cyclic hydrocarbons, e.g. cyclopentane, cyclohexane, methylcyclohexane, dimethylcyclohexane, methylcyclopentane or decalin. The hydrocarbon feedstock may also be a mixture of any of the above. In a highly preferred embodiment the feedstock comprises at least one C4+ olefin, e.g. a C5+ olefin or diene. The hydrocarbon feed may contain aromatic compounds such as benzene. Such aromatics do not, in general, crack in the presence of the HZSM-5 catalyst of the invention and are therefore left unchanged. Aromatic compounds may therefore be used as diluents in the process of the invention. Preferably however, it would be useful to convert aromatic compounds into useful olefins. Many of the potential feeds to the catalytic cracker contain high volumes of aromatic compounds, e.g. pyrolysis gasoline or condensate fractions direct from an oil refinery or an oil producer. In a preferred embodiment therefore, the hydrocarbon feed may be hydrogenated prior to exposure to the zeolite catalyst. Thus for example, a hydrocarbon feedstock containing benzene can be hydrogenated to give cyclohexane prior to cracking. By hydrogenating the feedstock, the majority of the hydrocarbons fed to the cracking reactor will be saturated, i.e. alkanes and cyclic saturated hydrocarbons. This forms a further aspect of the invention. Thus viewed from a further aspect the invention provides a process comprising hydrogenating a C₄+ hydrocarbon feedstock; and heating the resulting hydrocarbons at a temperature of 400-800⁰C in the presence of an H-ZSM-5 catalyst having a silicon to aluminium ratio of 61 - 1000. A potential reactor set up could involve diluent free cracking, e.g. using a hydrogenated pyrolysis gas feed with recycling of C4+ fractions to the hydrogenator. Alternatively, a reactor set up could involve the use of an aromatic compound as a diluent for a hydrocarbon feedstock. Pyrolysis gasoline from a naphtha steam cracker, which contains a large amount of aromatic compounds, would, for example, be suitable as such a diluent. The cracker may then have a recycling system which passes through a hydrogenator to convert the diluent into a crackable hydrocarbon. This may then be fed back into the cracker (along with fresh diluent) so that in a first pass, the aromatic compound acts as a diluent prior to hydrogenation and itself being cracked. Recycling of C4+ fractions could occur simultaneously. Preferably the temperature within the cracking reactor should range from 500-750⁰C, more preferably 550-700⁰C, especially 550 to 650°C, e.g. 570° to 650°C or 570 to 630⁰C such as about 600⁰C. The residence time of the feed over the catalyst should be long enough to give substantial conversion of the feed but not long enough to give rise to the production of large percentages of aromatic compounds. It has been surprisingly found that when the process described above is employed the yield (calculated as weight of carbon) of ethylene and propylene can exceed 40 wt % carbon. Thus, 40 % by weight of the carbon in the hydrocarbon feedstock is recovered as ethylene and propylene. Preferably at least 50 wt % carbon is recovered as ethylene and propylene. Whilst the ratio of produced propylene to ethylene can vary, e.g. more ethylene than propylene or a 1 : 1 ratio, preferably, the bulk of the produced product is propylene e.g. at least 25 wt% C, especially 30 wt % C. It is also preferred if the ratio of produced propylene to ethylene is greater than 1.2: 1 , e.g. greater than 1.5:1. An important factor in ensuring high percentage yields of ethylene and propylene is ensuring high conversion of initial feedstock. It is preferred if at least 50% of the initial feedstock is cracked, especially at least 60%, e.g. 70%, most especially at least 80% or at least 90%. At higher temperatures, conversion can near 100%. The higher the cracking temperature the higher the conversion. Also at higher temperatures, ethylene selectivity increases. There is, however, a trade off between temperature and hence high conversion/desired selectivity and potential for formation of aromatic compounds. The most preferred temperature to maximize conversion but minimise benzene formation is in the range 570 to 63O⁰C. Viewed from another aspect the invention provides a process for cracking C₄+ hydrocarbons comprising heating the hydrocarbons at a temperature of 400- 800⁰C in the presence of an H-ZSM-5 catalyst having a silicon to aluminium ratio of 61-1000 and recovering the cracked hydrocarbons; wherein at least 40% wt carbon of the cracked C₄+ hydrocarbons is recovered as ethylene and propylene. Any non-converted feedstock can of course be recycled back into the cracker along with any unwanted side products. The cracking reaction obviously yields hydrocarbons other than ethylene and propylene. Other products may include ethane, propane, methane, butane, butene and amounts of C5 and C6+ fractions. Catalytic cracking reactors of use in the invention are known and can operate under the temperatures discussed above using pressure if necessary, e.g. from 0.1 to 10 atm, preferably 0.3 to 2 atm. The catalytic process can be carried out in, for example, a fixed bed, moving bed, or fluidised bed reactor and the hydrocarbon flow can be cocurrent or countercurrent to catalyst flow. The catalyst may be formed from fine solid particles having a size range of from about 0.01 to 10 mm, e.g. 0.2 to 5 mm. Diluent such as an inert gas (nitrogen), methane or aromatic compounds can be employed as is known in the art, e.g. to carry the hydrocarbon gas stream into the reactor. The ratio of inert gas (if used) to hydrocarbons may range from 0: 1 to 1000: 1. Careful selection of diluents may allow further control over the ratio of products formed. Comprehensive discussions of suitable reactor set up can be found in the prior art cited above and are known in the art. The catalyst employed in the invention, i.e. a HZSM-5 catalyst having a particular Si/ Al molar ratio is obtainable from commercial sources such as Sud- Chemie. Silicon to aluminium ratios given in the text are molar ratios, i.e. by a Si/ Al ratio of 100 is meant that the molar amount of Si is 100 times the molar amount of Al. The manufacture of ZSM-5 catalysts is described in US 3702886. Preferred silicon to aluminium ratios are 80-400, especially 100-300, e.g. 85 to 200. It has been surprisingly found that at these high ratios, the amounts of formed lower olefins are high, in contrast to products obtained when using HZSM-5 catalysts having lower Si/ Al ratios. The catalyst is preferably medium pore (e.g. 10 rings). Comprehensive discussions of the manufacture and use of ZSM-5 catalysts can be found in the prior art cited above. The amount of catalyst employed will vary depending on the size of the reactor and the size of the feed but will be readily determined by the artisan. Catalyst can be continuously added to the cracker if necessary. It may also be necessary to regenerate the catalyst using known conditions. It has however, been surprisingly found that the catalyst of the invention only exhibits slow loss of activity during the cracking reaction and hence regeneration may be required only infrequently. It is common for the zeolites to be calcined prior to use, however it is preferred if the HZSM-5 catalyst used in the present invention is not calcined. Moreover, many zeolite catalysts are aged prior to use by exposing them to hydrothermal treatment, e.g. heating at a temperature of 500 to 800⁰C in the presence of steam . The zeolites used in the present invention should not be heated in this fashion prior to use since this may affect their light olefin selectivity, i.e. it is not necessary to expose the zeolite catalysts of the invention to elevated temperatures and 100% steam before use. The catalyst of use in the invention is an HZSM-5 species as hereinbefore described. Hence, the cations within the zeolite should be protons. There is the possibility however that the zeolite may be contaminated with other cations, e.g. sodium or potassium ions. It is preferred if the zeolite of use in this invention comprises primarily protons with low amounts of other cations, e.g. exclusively protons and no other cations. For example, the level of cations (other than protons) should be less than 0.05% wt, preferably less than 0.01%wt. It may be necessary to control the contact time of the hydrocarbon materials with the catalyst by taking pyrolysis properties of the hydrocarbons and reaction temperature into consideration. Contact times, measured as 1/GHSV (gas hourly space velocity) should be from 0.00002 h to 0.002h. The products of the cracking reaction can be fractionated using known techniques. Any unwanted products may be recycled to the reactor so as to increase yield of the desired product(s). Larger side products may be channeled into a condenser prior to recycling. A suitable reactor set up is now described. The setup used in the first Example is shown schematically in Figure 1 and represents an alternative for carrying out a small scale cracking reaction. The skilled chemist/chemical engineer can take the principles described herein for use in an industrial cracking reactor. Fluid hydrocarbons enter the apparatus via liquid pump (1) and enter evaporator (3). An inert gas, e.g. nitrogen, is fed to the evaporator via mass flow controller (2). Gaseous hydrocarbons formed in the evaporator enter reactor (6) containing catalyst (10) in oven (5) via flow switch (4). After cracking, the material from the reactor is transferred to condenser (7) where heavy ends (C9+) may be collected. In Figure 1, C 1-8 components are fed back through flow switch (4) and into a mass spectrometer (8) and/or gas chromatograph (9) for analysis. A larger scale apparatus is shown in Figure 8. The feed is passed into catalytic naphtha cracker (11) and heated by burning fuel gas. The cracked feed is transferred to separator (12) where the components are separated. The separator may also take a feed from a naphtha steam cracker (13). Fractions can then be isolated or recycled by to the cracker (1 1) for further cracking to occur. In Figure 9, a reactor set up for the cracking of pyrolysis gasoline is described. To cracker (14) is fed pyrolysis gasoline and a C4 feed along with hydrogen. The cracked product is transferred to separator (15) where heavy material and light hydrocarbons are separated. The bottoms stream containing the heavier components is passed through hydrogenator (16) to a further separator (17) where methane and hydrogen are separated and passed back to the cracker or isolated. The bottoms stream is also recycled back to the cracker. The top stream from separator (15) is itself passed to a further separator (19) via pump (18) where C 1-3 fractions are separated from C4/5 fractions. These latter fractions are recycled to the cracker and the light fractions isolated as the desired product. The invention will now be described with reference to the following non- limiting examples and Figures. Figure 1 shows a potential catalytic cracking reactor set up. Figure 2 shows the composition in the gas phase in the reactor effluent from catalyst A. Figure 3 shows the composition in the gas phase in the reactor effluent from catalyst B. Figure 4 shows the composition in the gas phase in the reactor effluent from catalyst C. Figure 5 shows the composition in the gas phase in the reactor effluent from catalyst D. Figure 6 shows conversion of cyclohexane in the presence of HZSM-5 Si/Al=200. Figure 7 shows selectivity to ethylene and propylene for HZSM-5 Si/Al=200 fresh and regenerated. Figure 8 shows a potential catalytic cracking reactor set up. Figure 9 shows a potential catalytic cracking reactor set up. Figure 10 shows conversion of cyclohexane over HZSM-5 Si/Al=200 at 400-650°C using either pure nitrogen or nitrogen and benzene as diluent. Figure 11 shows the conversion of C4 and the yield to C2=/C3= and aromates for catalytic testing of C4 mix over HZSM-5 from Sϋd Chemie AG (ZPO31400) with a Si/Al=200. Figure 12 shows carbon distribution in the gas effluent after testing the light fraction of pyrolysis gasoline at increasing temperatures over HZSM-5 from Sϋd Chemie AG (ZPO31400) with a Si/Al=200. Experimental The reactor set up depicted in Figure 1 was employed. The condenser was maintained at 75⁰C and all transfer lines and switches post evaporator were maintained at a temperature of at least 6O⁰C. The gas chromatography used was an Agilente micro GC with four columns. Total analyses of permanent gases and Ci-Cβ took 240 seconds. The gas chromatograph was calibrated with a Ci-C₄ gas mixture (Standard 1) and a mixture of N₂, H₂ and Ci-6 (Standard 2). Cyclohexane (CH) was calibrated by bypass analysis. All Cs compounds were assumed to have the same calibration factor and all Ce compounds were assumed to have the same calibration factor.

Example 1

Reactions were conducted in the temperature region 400-600⁰C with Ig of catalyst, 50 ml/min N₂ flow through evaporator and 0.1 ml/min hydrocarbon flow corresponding to a WHSV [weight hourly space velocity] of 4.7 gCH/(g catalyst*h) with cyclohexane used as feed. The catalysts were pressed to tablets and then crushed into particles with particle size 0.2 to 0.5 mm before testing. The reactor was first heated to 400⁰C under nitrogen flow before switching to feed from the evaporator. Samples were taken from the reactor at 4-15 minute intervals. N₂ was then flushed through the reactor and the reactor heated to 45O⁰C before feed from the evaporator was again admitted and the process above repeated. This process was repeated at 500⁰C, 55O⁰C and 600⁰C before the reactor was cooled to 400⁰C or 45O⁰C for gas chromatographic runs to be performed to check for any catalyst deactivation which may have occurred over the course of the experiment.

Four catalysts were examined using the above protocol:

Catalyst A: Kristal 232 ST (Grace Davison) - a standard cracking catalyst based on rare earth exchanged Y zeolite which has wide pores (12 rings)

Catalyst B : Valfoor CP 811 BL-25 (PQ) a Beta zeolite with wide pores (12 rings)

Catalyst C: A HZSM-5 with Si: Al =28 (Sud Chemie) Medium pores (10 ring)

Catalyst D: A HZSM-5 with Si: Al =85 (ZPO 31170) (Sud Chemie) (10 ring) Catalysts A to C are comparative, catalyst D exemplifies the invention.

Cracking Catalyst A

The results for this catalyst are shown in Figure 2 and in Appendix Table 1. The catalyst mostly gave high conversion but deactivated rapidly. Once the temperature returned to 400⁰C after testing at 55O⁰C, activity was almost non-existent. The carbon mass balance results indicate that a lot of the carbon forms products that are not analysed in the sampling system, i.e. coke and C6+ hydrocarbons, The high H₂ yield points to coke formation. Within the C3 fraction, the majority was propane as opposed to propene.

Catalyst B

The results for catalyst B are presented in Figure 3 and Appendix table 2. The catalyst gave rise to products in the C5+ range evidenced by the carbon balance results.

Catalyst C

The results are presented in Figure 4 and Appendix table 3. Much C3 product is formed by this catalyst but the fraction is primarily propane. The conversion is nevertheless high at all temperatures

Catalyst D

The results for catalyst D are presented in Figure 5 and Appendix table 4. At 600⁰C less than 2% ofthe cyclohexane is unconverted and there is observed a large selectivity for propene. The combined ethylene and propylene selectivity among the gas phase molecules is over 60% at 600⁰C. The carbon mass balance at the higher temperature is approximately 80% which means approximately 50% ofthe formed product is ethylene and propylene. The catalyst showed no deactivation on cooling.

Examples 2 to 8 General conditions Exp Feed Flow N₂ flow Amount GHSV Temp . (ml/min) (ml/min) of cat . (h-¹) ⁷ (⁰C) ⁽g⁾ \... Ex 2 CH 0 . 1 50 1 4338 400 - 650 Ex 5 CP 0 . 1 50 1 4542 400 - 650 Ex 6 MeCH 0 . 1 50 1 4131 400 - 650 Ex 3 CH 0 . 1 50 1 4338 600 Ex 4 CH 0 . 1 50 1 4338 600 Ex 7 HPyglf 0 . 1 50 1 4286 400 - * 650 Ex 8 HPyg** 0 . 1 50 1 4110 400 - 650 *Hydrogenated pyrolysis gasoline light fraction <125°C. ** Hydrogenated pyrolysis gasoline.

Example 2.

HZSM-5 from Sϋd Chemie AG (ZPO31400) with a Si/Al=200 was pressed into tablets and then crushed into particles with particle size 0.2-0.5 mm. One gram of catalyst (Ig) particles was tested at 400-650⁰C in a quartz fixed bed reactor with on¬ line GC analysis. A cyclohexane:N₂ molar ratio of 1 : 2.2 and a GHSV of 4338 per h was used. Liquid cyclohexane (CH) from Merck (99.6%) was evaporated at 75⁰C into the N₂ stream before entering the reactor.

The conversion is calculated on basis of unconverted feed (nitrogen is used as internal standard) and the carbon selectivity* in the effluent to propene and ethylene is shown in Table 1.

Table 1. Conversion of CH and C2= and C3= selectivity over HZSM-5

*Carbon selectivity calculated based on the analysis of C1-C6 products Example 3

The catalyst as described in Example 2 was tested with cyclohexane using the same experimental procedure as described in Example 2. The catalyst was tested for 7105 minutes at 600⁰C. The conversion as a function of time on stream is shown in Figure 6. From the results it may easily be derived a half life time of the catalyst of 3.5 days.

The carbon selectivity to ethylene and propene (based on analysis of products Cl- C6) is given in Figure 7.

Example 4.

The catalyst tested in Example 3 was regenerated in 5% oxygen in He. After regeneration the catalyst was tested with CH using the same experimental procedure as described in Example 2. The catalyst was tested for 1397 minutes at 600⁰C. The conversion as a function of time is shown in Figure 6 and selectivities to ethylene and propylene in figure 7.

Example 5. The catalyst as described in Example 2 was tested with cyclopentane (CP) using the same experimental procedure as described in Example 2. A CP:N₂ molar ratio of 1 : 1.9 and a GHSV of 4542b:¹ was used. Liquid CP from Janssen Chimica 16.775.91 (98%) was evaporated at 45⁰C into the N₂ stream before entering the reactor. The conversion calculated on basis of unconverted feed (nitrogen is used as internal standard) and the carbon selectivity* in the effluent to propene and ethylene is shown in Table 2.

Table 2. Conversion of CP and C2= and C3= selectivity over HZSM-5

*Carbon selectivity calculated based on the analysis of C1-C6 products Example 6. Methylcyclohexane (MCH) as feed The catalyst as described in Example 2 was tested with MCP using the same experimental procedure as described in Example 2. A MCP:N₂ molar ratio of 1 :2.7 and a GHSV of 413 Ih^"1 was used. Liquid MCH from Venton 12548 (99%) was evaporated at 90⁰C into the N₂ stream before entering the reactor.

The conversion calculated on basis of unconverted feed (nitrogen is used as internal standard) and the carbon selectivity* in the effluent to propene and ethylene is shown in Table 3.

Table 3. Conversion of MCH and C2= and C3= selectivity over HZSM-5

*Carbon selectivity calculated based on the analysis of C1-C6 products

Example 7 Cracking of hydrogenated pyrolysis gasoline (light fraction) Hydrogentated pyrolysis gasoline was obtained from Statoil. Some of the hydrogenated pyrolysis gas was distilled to boiling point 125°C. This corresponds to 69.8wt% of the feed. 98.7wt% mass balance was obtained during the distillation (this may indicate that 1.3wt% of the lightest components have been lost during the distillation). The total hydrogenated pyrolysis gasoline (Example 8) and the light fraction < 125⁰C (Example 7) were used as feeds. Piona analyses of the two feeds were performed and the results are given in Tables 5 and 7 below. The total feed contains 80.1wt% naphthenes and the light feed contains 83.2wt% naphthenes. The density of the two feeds was measured by weighing 25 ml of the samples. The light fraction <125°C had a density of 0.744 g/ml and the total hydrogenated pyrolysis gasoline had a density of 0.776 g/ml. The analysis of the effluent was performed with an Agilente micro GC with 4 columns. All detectors were TC detectors Column A : 5 A mol sieve Run at 30.7 psi and 85°C with Ar as a carrier gas. Used for analysis of hydrogen and nitrogen Column B: Poraplot U . Run at 20.7 psi and 70°C with He carrier gas. Used for analysis of methane, ethane and ethylene Column C: Alumina Plot Run at 34.3 psi and 115°C with He carrier gas Used for analysis of propene, propane and C4's Column D: OV-I In order to analysis more of the heavy compounds on the GC, the temperature of the D column was increased to 150°C, compared to 70⁰C in earlier analysis.. Pressure was 29.6 psi. He carrier gas. This column was used to analyse C5-C9 compounds. The catalyst as described in Example 2 was tested with the light fraction boiling <125°C of a hydrogenated pyrolysis gasoline using the experimental conditions as described in Example 1. A Piona analysis of this feed is given in Table 5 below. The yields to ethene and propene are given Table 6.

Table 5 C3 C4 C5 C6 C7 C8 C9 G10 C11 C12

Paraffins 0.020 0.120 10.202 4.420 1.021 0.165 0.016 0.000 0.000 0.000 n-Paraffins 0.020 0.081 5.857 2.272 0.475 0.082 0.016 0.000 0.000 0.000 iso-Paraffins 0.039 4.345 2.148 0.546 0.083 0.000 0.000 0.000 0.000

Naphthenes 1.967 51.053 20.296 7.988 1.317 0.617 0.000 0.000 Mono-naphthenes 1.967 51.053 20.296 7.988 0.910 0.000 0.000 0.000 Di-napbthenes 0.000 0.408 0.617 0.000 0.000

Aroma tics 0.597 0.127 0.000 0.000 0.000 0.000 0.000 Benzenes 0.597 0.127 0.000 0.000 0.000 0.000 0.000 Naphthalene 0.000 NaptWOIef-benzenes 0.000 0.000 0.000 0.000 0.000 lndenes 0.000 0.000 0.000 0.000

Olefins 0.000 0.000 0.000 0.073 0.000 0.000 0.000 0.000 0.000 n-Olefins 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 iso-Olefins 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 Naphtheno-olefϊns 0.000 0.000 0.073 0.000 0.000 0.000 0.000 0.000 Di-olefins 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 Other olefins 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000 0.000

Sum 0.020 0.120 12.169 56.071 21.517 8.152 1.333 0.617 0.000 0.000 Table 6. Conversion of pentanes and cyclohexane (CH) and yields to ethene and propene.

The sum of the C2= and C3= yields are about 43 % at 650 ⁰C. After running the experiment for approximately 8 h, the catalyst has been exposed to feed for approximately 115min. The conversion at 450 ⁰C was about the same at the end of the experiment as in the start of the experiment, which points to no significant deactivation.

Example 8 Hydrogenated Pyrolysis Gasoline (total feed)

The catalyst as described in Example 2 was tested with a hydrogenated pyrolysis gasoline using the experimental conditions as described in Example 1. A Piona analysis of this feed is given in Table 7 below. The yield to ethene and propene are given Table 8. Table 7 Piona analysis of the hght fraction from hydrogenated pyrolysis gasoline

Values are given in weight %

C3 C4 C5 C6 C7 C8 C9 C10 C11 C12

Paraffins 0 045 0 165 13 405 4594 0747 0 155 003 0 096 0 026 0 02 n-Paraffins 0 045 0 111 7 616 1 754 0 345 0 084 0 03 0 019 0 026 0 02 iso-Paraffins 0 053 5 789 284 0 402 0 071 0 0 077 0 0

Naphthenes 2 581 36 766 14657 4032 8 305 1 1 689 2 059 0 Mono-naphthenes 2 581 36 766 14657 3 876 6 484 0 43 0 024 0 Dι-naphthenes 0 156 1 82 11 259 2 035 0

Aromatics 0 425 0 081 0 018 0 0 0 0 Benzenes 0 425 0 081 0 018 0 0 0 0 Naphthalene 0 NapfWOIef-benzenes 0 0 0 0 0 lndenes 0 0 0 0

Olefins 0 0 0 0 051 0 0 0 0 053 0 n-Olefins 0 0 0 0 0 0 0 0 0 iso-Olefins 0 0 0 0 0 0 0 0 0 Naphtheno-olefins 0 0 0 051 0 0 0 0 0 Di-olefins 0 0 0 0 0 0 0 0 0 Other olefins 0 0 0 0 0 0 0 0 053 0

Sum 0045 0165 15987 41785 15536 4206 8335 11785 2138 002

Table 8 Conversion of pentanes and cyclohexane (CH) and yields to ethene and propene

Example 9. Testing with aromatic diluent.

The catalyst described in Example 2 was tested at 400-650°C in a quartz fixed bed reactor with online GC analysis Nitrogen and benzene were used as diluent, along with a cyclohexane N₂ benzene molar ratio of 1 1 3 1, and at a GHSV 4356 per h Liquid 1 1 molar mixture of cyclohexane (CH) from Merck (99 6%) and benzene from Merck (p. a.) was evaporated at 80°C into the nitrogen stream before entering the reactor.

In Figure 10 the conversion of CH is compared with the results from Example 2. In Table 9 the selectivity based on the products are given. In the selectivity calculations, it is assumed that benzene is not a product, which influences the results at temperatures jigher than 600⁰C.

Table 9. Conversion and selectivity overHZSM-5 Si/Al=200

*0.1ml/min CH(Hq) + 50ml/min N2, * 0. lm min CH q + 0.098m m n Bz q +27.7ml/min N2.

Example 10. Testing of a C4 mix HZSM-5 from Sϋd Chemie AG (ZPO31400) with a Si/Al=200 was pressed into tablets and then crushed into particles with particle size 0.2-0.5 mm. One gram of catalyst (Ig) particles was tested at 500-700⁰C in a quartz fixed bed reactor with on-line GC analysis. C4 mix obtained from a commercial steam cracker 19ml/min and 25ml/min nitrogen was used as feed. The C4 mix was contained as a liquid in a container keeping a pressure of 1.5 bar. This was taken from a stream where the original C4 mix had been exposed to a partial hydrogenation of butadiene, and thereafter the isobutene removed by reacting to methyltertbutylether. The analysis of the C4 mix was performed on a separate GC and the analysis is given in Table 10.

Table 10. Analysis of C4 mix obtained from commercial steam cracker

The conversion of the C4's and the yield to C2=/C3= and aromates are given in Figure 11. This example shows that the process is well suited for cracking of C4 mix into light olefins and up to 50% yield to C2=/C3= is achieved at 600-650°C

Example 11. Using light fraction of pyrolysis gasoline as feed

The catalyst (Ig) as described in Example 2 was tested with the light fraction of pyrolysis gasoline as feed. The pyrolysis gasoline was distilled and the fraction boiling at < 125°C was used as feed. This corresponds to 82% of the pyrolysis gasoline sample. Analysis of the light fraction of the pyrolysis gasoline was performed on a separate GC and the analysis is given in Table 11. This light fraction had a density of 0.798g/ml.

The liquid feed was fed with 0. lml/min and evaporated into a 50 ml/min nitrogen stream before entering the reactor.

Table 11. Analysis of "light fraction" of pyrolysis gasoline (not hydrogenated)

Component Wt % Benzene 35.5 Toluene 10.9 Other C₅-C₈, mainly C₅ 49.0 Xylenes + ethylbenzene 4.6

The carbon distribution in the gas effluent is given in Figure 12. This example shows that the process cracks 20% of the light naphta into C2=/C3=. This corresponds to 16.4% of the total pyrolysis gasoline.

Table 1. The cracking catalyst A MoI %i gas phase Product carbon % in gasphase Temp sum C H₂ CH₄ C₂H₄ C₂H₆ C₃H₆ C₃H₈ sum sum sum unconv sum Metan C₂H₄ C₂E₆ C₃H₆ C₃H₈ sum sum sum mol% ia balance C₄ Cj Csprod C*n 1 , C₄ C₅ C₆P io__rtop_ 400 102,9 80,6 0,1 0,0 0,1 0,0 0,0 0,5 0,7 0,2 3,8 22,0 28,5 0,0 0,7 0,1 0,5 5,3 10,0 3,7 79,8 400 103,2 83,8 0,1 0,0 0,1 0,0 0,0 0,4 0,5 0,7 3,5 23,0 27,8 0,0 0,6 0,1 0,5 4,1 7,6 11,9 75,3 oo 450 101,5 77,4 0,4 0,0 0,4 0,0 0,3 2,0 1,8 1,1 5,4 16,3 52,9 0,0 1,4 0,2 1,7 11,4 14,0 10,0 61,5 450 103,5 105,5 0,5 0,2 0,4 0,1 0,3 2,1 2,0 1,3 7,0 21,0 65,1 0,3 1,2 0,2 1,5 9,7 12,5 10,2 64,5 500 101,3 75,8 2,6 0,8 1,0 0,3 1,3 4,9 3,2 1,3 5,7 10,7 75,2 1,1 2,6 0,7 5,1 19,5 17,1 8,6 45,2 500 104,2 130,9 2,6 0,7 0,8 0,2 1,2 4,3 3,2 1,0 9,0 20,0 91,1 0,8 1,9 0,5 3,8 14,1 14,2 5,6 59,0 550 101,6 74,4 8,9 2,3 1,3 0,7 2,4 4,9 2,5 1,0 5,0 8,7 73,0 3,2 3,6 1,9 10,0 20,2 13,7 6,6 40,7 550 123,5 88,3 8,6 1,8 1,1 0,5 2,3 4,0 2,3 1,1 6,3 17,6 76,2 2,4 2,9 1,4 8,9 15,8 12,2 7,1 49,4 550 103,3 100,1 6,8 1,4 0,8 0,4 1,9 2,8 1,6 0,9 4,8 18,3 58,0 2,5 2,9 1,2 9,9 14,5 11,3 7,5 50,1 550 109,7 82,5 6,4 1,3 0,7 0,3 1,9 2,3 1,4 0,8 4,1 18,0 49,9 2,6 3,0 1,2 11,3 14,0 11,2 7,5 49,2 600 103,8 89,2 7,6 1,2 0,7 -Pol — 2,6 1,0 0,9 0,6 2,6-— ^"20,1 35,7 3,4 3,6 1,3 21,5 8,2 9,9 8,5 43,6 600 113,9 82,4 —6,5 0,9 ~ ^' OJ^{" ~} 0,2 2,1 0,7 0,8 0,6 2,3 22,6 30,8 2,9 3,4 1,1 20,8 7,1 11,0 9,0 44,7

O O ΪΛ O O W H U Table 2. Catalyst B Zeolite Beta MoI % in gas Product carbon % in gas phase phase Temp sum C H₂ CH, C₂H₄ C₂H* C₃H₄ C₃H, sum sum sum unconv sum Metan C₂H, C₂H₆ C₃H₆ C₃H, Sum sum sum (⁰C) mol% in balance Q C, Csprod C*n Q C₅ Qp loop (%) 400 98.7 71.7 5.3 0.8 0.5 0.1 0.3 4.4 7.0 4.1 6.9 3.1 106.3 0.8 0.9 0.1 0.8 12.6 26.5 19.5 38.8 400 102.0 79.7 5.3 0.6 0.4 0.0 0.3 2.4 4.2 2.8 9.2 8.1 95.8 0.6^" 0.S 0.1 J.I 7.5 17.6 14.6 57.7 400 99.8 55.7 3.3 0.4 0.3 0.0 0.2 1.5 2.2 1.8 5.9 8.7 58.9 0.7 0.9 0.1 0.9 7.4 14.7 15.0 60.3 400 98.6 45.1 2.6 0.3 0.2 0.0 0.2 1.0 1.7 1.2 4.8 8.0 45.8 0.7 1.0 0.1 1.6 6.8 14.6 12.7 62.5 ON 450 97.9 42.6 9.6 1.0 0.9 0.1 0.6 3.4 3.6 1.8 2.8 4.0 55.0 1.9 3.1 0.4 3.5 18.4 26.1 16.3 30.3 450 96.0 34.6 7.5 0.8 0.6 0.1 0.5 2.0 2.2 1.1 2.2 5.1 36.9 22 3.2 0.4 3.9 16.0 23.6 15.0 35.6 450 96.1 30.3 6.4 0.7 0.5 0.1 0.4 1.5 1.7 0.9 1.9 5.2 30.2 2.3 3.3 0.4 4.3 15.1 22.8 14.2 37.6 500 94.4 30.0 14.2 1.7 1.2 0.3 1.0 3.1 2.6 0.8 1.5 2.0 40.5 4.3 6.1 1.3 7.3 22.8 25.9 9.8 22.5 500 94.6 30.1 13.0 1.6 1.1 0.2 1.0 2.6 2.3 0.8 1.6 2.7 37.8 4.1 6.0 1.3 7.9 20.7 24.8 10.1 25.0 550 93.2 27.5 21.2 2.8 1.6 0.5 1.2 2.6 1.4 0.2 2.6 0.4 40.1 7.1 7.7 2.4 8.7 19.4 14.1 2.4 38.2 550 93.8 29.9 18.9 2.5 1.5 0.4 1.5 2.2 1.6 0.3 2.5 1.3 40.6 6.0 7.3 2.0 10.9 16.1 16.0 4.1 37.4 450 97.6 45.2 5.1 0.5 0.3 0.0 0.3 0.4 0.6 0.4 1.5 12.3 17.1 2.9 3.1 0.3 5.6 7.3 15.1 11.4 54.4

O O O <o O O O O O ΪΛ O O W H U Table 3. Catalyst C, HZSM-5 with SUAl =28 ¹ MoI % in gas Product carbon % in gas phase phase Temp sum C H₁ CR, C₂BU C₁Hs C₃H₆ C₃H₈ sum sum sum unoonv sum Metan C₂H₄ C₂H₄ C₃H₄ C₃H₈ Sum sum sum (⁰C) mol% in balance c. C₅ Qprod Cn Q C₅ Qsp loop (%) 4QO 101.5 48.0 1.2 0.2 0.7 0.3 0.9 15.4 4.3 0.8 0.6 2.9 76.4 0.3 2.0 0.8 3.7 60.4 22.6 5.4 4.9 400 101.0 51.4 1.2 0.2 0.8 0.3 1.0 16.6 4.6 0.9 0.7 2.6 82.1 0.2 1.9 0.8 3.5 60.7 22.6 5.4 4.8 400 100.6 47.8 1.1 0.2 0.8 0.3 0.9 15.7 4.2 0.9 0.7 2.5 77.5 0.2 2.0 0.8 3.7 60.7 21.8 5.7 5.1 400 100.1 46.9 1.1 0.2 0.7 0.3 0.9 15.2 4.3 0.8 0.7 2.5 75.9 0.2 1.9 0.8 3.7 60.3 22.4 5.3 5.3 400 100.3 48.4 1.1 0.2 0.7 0.3 0.9 15.7 4.4 0.8 0.7 2.5 77.9 0.2 1.9 0.8 3.6 60.3 22.6 5.4 5.1 400 99.9 49.7 1.1 0.2 0.8 0.3 1.0 15.8 4.4 0.9 0.7 2.7 78.6 0.2 1.9 0.S 3.7 60.2 22.4 5.5 5.2 400 100.3 48.5 1.2 0.2 0.8 0.3 0.9 15.7 4.4 0.9 0.6 2.5 78.0 0.2 1.9 0.8 3.6 60.5 22.3 5.6 5.0 400 101.8 50.8 1.3 0.2 0.8 0.3 1.0 16.1 4.5 0.9 0.6 3.0 79.8 0.3 2.0 0.8 3.6 60.4 22.6 5.5 4.9 CN 400 100.2 43.6 1.2 0.2 0.7 0.3 0.9 14.2 3.9 0.8 0.6 2.5 70.8 0.3 2.1 0.8 3.9 60.2 22.3 5.4 5.0 400 100.9 44.4 1.2 0.2 0.7 0.3 0.9 14.2 3.9 0.8 0.6 2.9 70.9 0.3 2.1 0.8 3.9 60.0 22.1 5.8 5.0 450 99.5 40.4 3.0 0.6 1.5 0.6 1.8 14.6 3.6 0.5 0.7 0.4 74.8 0.8 4.0 1.6 7.1 58.5 19.0 3.3 5.8 450 100.0 40.1 2.9 0.6 1.5 0.6 1.7 14.7 3.6 0.5 0.7 0.3 75.2 0.7 3.9 1.6 7.0 58.5 19.1 3.4 5.7 450 100.4 44.3 3.2 0.6 1.5 0.7 1.8 15.9 3.9 0.6 0.8 0.3 81.2 0.7 3.8 1.6 6.8 58.8 19.0 3.6 5.6 500 99.6 40.6 6.3 1.9 2.8 1.2 2.6 12.5 2.6 0.2 1.0 0.1 72.9 2.6 7.8 3.4 10.5 51.3 14.2 1.5 8.6 500 98.9 37.2 6.2 1.7 2.7 1.1 2.4 11.5 2.4 0.2 1.0 0.0 68.2 2.5 8.0 3.4 10.8 50.7 14.2 1.5 9.0 550 94.7 27.5 10.3 0.0 _ 4.2 1.6 2.3 6.1 0.7 0.0 1.5 0.1 49.0 0.0 17.1 6.3 14.3 37.1 5.9 0.4 18.9

O O O <o O O O O O IΛ O O

H U Table 4 Catalyst D, HZSM-5 with Si/Al =85 MoI % in gas Product carbon % iα gas phase Bhase Temp C H₂ CEU C₂H, C₂H₆ C₃H₄ C₃H, sum sum sum unconv sum Metan C₂HL₁ QH₆ C₃H₄ C₃H₈ Sum sum sum (⁰C) mol% in balance c_< C Cβprod C*n C₄ C₅ C«p loop (%) 600 103.5 86.0 9.2 2.2 16.1 2.0 11.5 4.2 4.3 1.1 1.0 0.3 114.8 1.9 28.0 3.5 30.2 11.1 15.0 4.8 5.4 600 97.7 81.9 8.9 1.9 13.8 1.0 11.6 4.2 4.3 0.7 1.0 0.3 106.1 1.8 26.0 1.8 32.8 11.9 16.3 3.4 5.9 600 96.7 82.2 8.8 1.8 13.5 1.0 11.6 4.3 4.3 0.7 1.0 0.3 105.3 1.7 25.6 1.8 33.0 12.1 16.5 3.3 5.9 600 96.0 79.3 8.7 1.7 13.1 0.9 11.4 4.1 4.2 0.7 1.0 0.3 102.8 1.7 25.5 1.8 33.2 12.0 16.4 3.4 6.0 600 96.2 81.9 8.7 1.7 13.2 1.0 13.6 2.3 4.3 0.7 1.0 0.3 104.9 1.6 25.1 1.8 38.8 6.7 16.6 3.5 5.9 (N 600 96.4 79.4 8.6 1.7 13.1 1.0 11.4 4.1 4.2 0.7 1.0 0.3 103.3 1.7 25.4 1.8 33.2 12.0 16.4 3.4 6.0 550 97.0 80.9 6.7 0.8 8.8 0.6 9.9 6.0 5.3 1.1 0.9 2.7 99.3 0.8 17.7 1.2 29.9 18.1 21.1 5.6 5.5 550 96.9 78.9 6.5 0.8 8.8 0.6 9.8 5.9 5.2 1.1 0.9 2.7 98.0 0.8 17.9 1.2 29.9 18.0 21.0 5.6 5.6 550 96.5 80.7 6.5 0.8 8.8 0.6 9.9 6.0 5.2 1.1 0.9 2.7 99.0 0.8 17.8 1.2 29.9 18.1 21.1 5.6 5.4 550 96.9 81.9 6.5 0.8 8.8 0.6 10.0 6.1 5.3 1.1 0.9 2.7 100.3 0.8 17.5 1.2 30.0 18.2 21.2 5.6 5.5 550 102.1 81.2 6.8 1.1 9.9 0.7 10.0 6.2 5.4 1.1 0.9 3.0 104.0 1.1 19.1 1.4 28.9 17.9 20.7 5.5 5.3 550 98.0 80.7 6.5 0.9 9.0 0.7 9.9 6.0 5.3 1.1 0.9 2.8 100.0 0.9 18.0 1.3 29.7 18.0 21.1 5.6 5.4 500 105.5 101.8 3.6 0.7 5.4 0.4 6.7 7.1 5.9 3.0 1.6 10.3 102.1 0.8 11.7 0.9 21.6 22.9 25.1 4.9 9.5 450 101.6 90.5 1.3 0.5 2.0 0.1 2.6 4.6 2.8 0.7 0.2 19.6 42.7 1.2 9.3 0.6 18.2 32.2 26.7 8.3 3.4 400 100.4 96.2 0.3 0.4 0.5 0.0 0.7 1.9 1.0 0.3 0.6 26,2 17.6 2.0 5.3 0.3 11.6 32.1 21.6 7.7 19.4 400 100.7 103.4 0.3 0.2 0.5 0.0 0.7 2.0 1.0 0.3 0.6 27.6 18.5 1.3 5.2 0.3 11.1 31.9 21.5 7.7 21.0

O

Claims

1. A process for cracking C₄+ hydrocarbons comprising heating the hydrocarbons at a temperature of 400-800⁰C in the presence of an H-ZSM-5 catalyst having a silicon to aluminium ratio of 61 - 1000.

2. A process as claimed in claim 1 wherein the C4+ hydrocarbons comprise at least one C4+ olefin.

3. A process as claimed in claim 1 or 2 wherein the C₄+ hydrocarbons are hydrogenated prior to cracking.

4. A process as claimed in claim 1 to 3 for cracking C₅+ hydrocarbons.

5. A process as claimed in claim 1 to 4 wherein the hydrocarbons comprise saturated paraffins, olefins of general formula C_nH_2n where n is from 4 to 12 or saturated cyclic hydrocarbons or mixtures thereof.

6. A process as claimed in claim 5 wherein the hydrocarbons comprise cyclohexane.

7. A process as claimed in claim 1 to 6 wherein the C₄+ hydrocarbons are derived from pyrolysis gasoline or hydrogenated pyrolysis gasoline.

8. A process as claimed in claims 1 to 7 wherein the hydrocarbons comprise aromatic compounds.

9. A process as claimed in any one of claims 1 to 8 which takes place in a cracker.

10. A process as claimed in claim 9 wherein at least a portion of the cracked hydrocarbons is recycled back to the cracker and undergoes further cracking.

11. A process as claimed in claim 10 wherein said cracked hydrocarbons are recycled through a hydrogenator before being returned to the cracker.

12. A process as claimed in any one of claims 1 to 1 1 wherein the temperature is 570 to 65O⁰C.

13. A process as claimed in any one of claims 1 to 12 wherein the ratio of Si: Al is 80 to 400.

14. A process as claimed in any one of claims 1 to 13 wherein the ratio of Si: Al is 85 to 200.

15. A process as claimed in any one of claims 1 to 14 wherein the amount of ethylene and propylene formed is at least 40 wt% carbon.

16. A process as claimed in any one of claims 1 to 15 wherein the ratio of formed propylene to ethylene is greater than 1.2: 1.

17. A process as claimed in any one of claims 1 to 16 wherein the amount of propylene formed is at least 30 wt% carbon.

18. A process as claimed in any one of claims 1 to 17 wherein at least 60% of the C₄+ hydrocarbons are cracked.

19. A process as claimed in any one of claims 1 to 18 wherein said HZSM-5 catalyst is not hydrothermally treated.

20. A process as claimed in any one of claims 1 to 19 wherein the HZSM-5 catalyst contains less than 0.05%wt cations (other than protons).

21. The use of an HZSM-5 catalyst having a silicon to aluminium ratio of 61 - 1000 as a catalyst in the cracking of C₄+ hydrocarbons.