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DGMK
German Society for Petroleum and Coal ,. Science and Technology
Tagungsbericht 9903
Proceedings of the DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry” October 13-15, 1999, Erlangen, Germany ,.
(Authors’ Manuscripts)
edited by G. Emig, M. Rupp &J. Weitkamp
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-,- . . . . . ,, I ,.....’....,.- ,. .<,.’:’> CONTENTS ., , ““?:,:’ .:;,.’,’,’. ,< ,, -,.’ ,.,. ,,.” ,.. ~ Page ,6 ., \ -, Aromatics: Bridging Refi~ng and Petrochemicals - 7 a major Challenge for the Future W. J. Petzny, C.-P. Halsig
How will EU Environmental Legislation Impact on the 21 Aromatics Markets? M. Schuller, P. Hodges t The Role of Aromatics in Transportation Fuels 37 K. van Leeuwen
Sulphur Tolerant Dearomatization of Middle Distillates - 45 A Cost Effective Way of Meeting 2005 Specifications ?’ D. Huang, E. Kohler
The Roles of Aromatics and Catalytic Reforming in the 53 2000+ Refinery O. Genis, S. G. Simpson, D. W. Penner, R. Gautam, B. K. Glover
Flexible Upgrading of Light Reformate 65 J. Cosyns, Q. Debuisscheti, B. Didillon, J. L. Ambrosino
Separation Processes for the Recovery of Pure Aromatics 73 ,.. , B. Firnhaber, G. Emmrich, F. Ennenbach, U. Ranke ,-,.,
Developments in Aromatics Separation 99 G. Krekel, J. Eberhardt, T. Diehl, G. Birke, H. Schlichting, A. Glasmacher
The Role of Pyrolysis Gasoline from Steam Crackers 115 beyond the Year 2000 C. Dembny
Maximizing Paraxylene Production with ParamaX 131 J. Rault, P. Renard, F. Alario ,.,~,’,, i .,, ,, .’1 Recent Advances in the Oxidation and Ammoxidation of 139 ,., .1, ;,., ,- ,.,; “., ,’ ., Aromatics ‘ -, .’ ,. , f .2.. , ~~ B. Lticke, A. Martin .’~,,,,,, :,,1~; ,,, , ,., ,., ., Recent Progress and Challenges in the Electrophilic 153 .,-, 1 Substitution of Aromatics ,, .,,.. ,.,1‘J F.-J. Mais
Fundamental Relations between Modification of the External 161 Surface of Zeolites and Catalytic Performance H. P. Roger, K. P. Moller, W. Bohringer, C.T. O’Connor
,’.
Y.,,,< ... ,,7 II Gas Phase Hydrogenation of Benzene with Unusual 169 Cyclohexene Formation “ E. Dietzsch, D. Honicke
SulfurTolerance ofAIkali Exchanged Zeolites for Benzene 177 Hydrogenation L. Simon, J. G. van Ommen, P. J. Kooyman, A. Jentys, J. A. Lercher
Postersession
PolycyclicArenes and Heteroarenes asBackbonesof 185 Diphosphine/amine Ligands forThermostable Homogeneous Catalysts M. W. Haenel, St. Oevers, St. Hillebrand, W. C. Kaska
Low Temperature o-XyIene Oxidation to Phthalic Anhydride 193 on V-Ti-O Catalysts Prepared by Mechanochemistry V.A. Zazhigalov, Al. Kharlamdv, L. Depero, A. Marine, I.V. Bacherikova, J. Stoch, J. Haber
A New Calculation Method for Arenes Solubilities in the Whole 201 Composition Range of Water-Sulfuric Acid System A. L Lutsyk, E. S. Rudakov, V. N. Mochalin
A New Oxidic Catalyst System for Selective CG+Paraffine 207 Aromatization D. L. Hoang, A. Trunschke, A. Bruckner, J. Radnik, H. Lieske
Benzene Oxidation to Phenol by Molecular Oxygen on 215 Medicated MoOS L. V. Bogutskaya, S.V. Khalamejda, V.A. Zazhigalov
Selective Hydrogenation of light Reformate for Production 217 of high-purity Benzene M. Walter, E. Schwab, M. G. Koch, P. Trubenbach, S. Dining
Supported Ru-catalysts for the partial Gas Phase 223 Hydrogenation of Benzene E. Dietzsch, P. Claus, D. Honicke
Heterogeneously catalyzed liquid-phase Hydrogenation 23i of nitro-aromatics using Microchannel Reactors R. Fodisch, W. Reschetilowski, D. Honicke
Safety of large Reactors for the Cumene Oxidation 239 M. Weber Ill
., Base-catalyzed Synthesis of Ethylbenzene by 247 .. Dehydroisomerization of 4-Viny Icyclohexene J. Ackermann, E. Klemm, G. Emig
Catalytic Behaviour of modified ZSM-5 Type Zeolites in the 255 Hydroxylation of Benzene-using nitrous Oxide A. Reitzmann, G, Konig, F. M. Petrat, E. Klemm, G. Emig
263 Effect of ZSM-5 Zeolite Synthesis Following Different Routes on ..’ the Isomerization of Xylenes ,:, R. Monnig, W. Schwieger
.,7 Alkylation of Napthalene on a Zeolite Catalyst in Supercritical 271 ,,. and Gaseous Reaction Phases R. Glilser, J. Weitkamp
On Selectivity Aspects of the Alkylation of Toluene with 279 Methanol over Zeolites M. Rep, A. E. Palomares, G. Eder-Mirth, J. G. van Ommen, J. A. Lercher
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,..,,:, >, ,,~’,,, ..- DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 ., !,.,;:. ,. ,,
W. J. Petzny ‘), C.-P. Halsig 2)” 1,BP Amoco Chemicals, London, United Kingdom 3 Fluo~ Daniel, Haar\em, The Netherlands
Aromatics: Bridging Refining and Petrochemicals - a major Challenge for the Future
‘,
Introduction t,
In the last decade the European and North American Refining and Petrochemical Industries have faced increasing challenges in providing an attractive return on investment. The principal reasons for this situation includes overcapacity, slow down of economic growth, environmental pressurea, environmentally mandated fisel composition changes, ‘dkselisation’, price erosion, and low margins resulting in mergers and plant shut downs. The only potentially beneficial effect of this is less tlagmentation of these European industries.
The marketa for both industries are regarded as cyclic. However, while the petrochemical industry cycle was at least in an upturn in 1988 and 1994/95, the refining industry has been in an economic trough since the mid 80s and still is. It is apparent that both industries are .,, ,.,, .,..,- , continuously searching for opportunities to stay competitive and to improve profitability by ., J.. { adding value to their products and product slates.
Obviously, there is no general advice that one can offer for these challenges. Trends can be recognized and compared with each individual local situation. Newly applied techniques and new technological developments can enable refiners and petrochemical plant operators to increase utilization of existing equipment and, as usual, every situation is ‘special’.
The demand for petrochemical products is now rising at a rate of about twice that of oil products. WMe currently about 8% of the crude oil barrel processed goes into petrochemicals this is predicted to rise clearly above 10% within the next few years, at some locations it is already at 200A.Growth in this area will require fhrther initiatives to search for opportunities for the benefit of both the refiner and petrochemical plants to increase return on existing assets.
. . .
7 OGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 The title of this conference is ‘The Future Role of Aromatics in Refining and Petrochemistry’. As a contribution to this conference this presentation therefore aims to review the trends in refining and petrochemical industries, reviews the aromatics of common interest and the challenges and opportunities for both industries in this area, analysing the bridging opportunities and providing a basis for some of the presentations to follow.
Trends in the Petrochemical and Refining Industries
The considerable overlap of the two industries started with the introduction of the reforming process into the refineries more than 40 years ago, the extraction of aromatics produced, substituting and replacing the aromatics recovered from coke oven gas. . The new gasoline specifications in place or coming in place in North America and in Europe, and finally the rest of the world - demands the reduction of total aromatics and benzene in particular in gasoline more and more will have a significant influence on this relationship. This offers challenges for fiture opportunities but also chances for conflicts.
Though the world wide economy is still a bit in a trough, unprecedented petrochemical development is still taking place in Asia and in the Gulf area in terms of expansion as WCIIas vertical integration from feedstock to product. Though the Asian flue has reduced or delayed some investments many were already on their way when the flue hit. A majority of these new plants are integrated petrochemical - refining complexes with refining capacities between 150,000 and 250,000 bpsd (equivalent to 8-12 m t/year) and ethylene capacities from 0.4 to 1 m t/year of ethylene. These significant investments in integrated petrochemical / refining complexes are driven by the desire to have an optimal operation, benefiting from the synergy between the retincry and petrochemical plant, energy utilisation and security of petrochemical feedstock supply.
Of the more than one hundred refineries in Western Europe about twenty, mostly the larger ones, are integrated with petrochemical operation. Major revamps have been carried out or are in progress at several locations to firther increase utilisation of refinery originating streams.
In Central Europe, most petrochemical facilities are associated with a refinery. In the FSWCIS, refineries and petrochemical plants have historically been separated based on different administrative responsibilities. However, it is evident today that the large, vertically integrated Russian oil companies are foming(or considering the fomation o~ petrochemical ,,’ companies. These are expected to add value to their operations by integrating existing or ,., ,,, planned petrochemical plants into their operating schemes adjacent to the existing refinery ,.-,
facilities. “.,,.“ ‘., . . .
Several new refineries have come on stream in the last few years and the majority of the new capacity is located in Asia. Most of the newly announced refineries are increasingly connected with petrochemical operation as previously mentioned. In Europe, the new MIDER refinery in Leun~ Germany, which started up in late 1997, was the first new refinery built in Europe since 1980 when the Mazeikia Refinery in Lithuania came on stream. The MIDER refinery is ,,, in the midth of the Leuna Regional (Chemical) Network, finally benefiting from the ongoing petrochemical and chemical investments. ., .-., ,:.,
However, in recent decades, a number of refineries have been cIosed and crude distillation ) ,. . . . .: capacity has been taken off stream, primarily in Europe and the USA. Worldwide utilisation ,.”, ,, ,, ,., ,-.’ rates have increased over the last few years. Even so, margins have not increased very much. ,’, ,,. . ‘, ,,. !.,.,,.,. . . . One contributing factor to this situation might be the continuous ‘capacity creep’ due to ,,...... : .,, ,+,. ., .,, . . debottlenmking, improvements in equipment reliability, fewer unplanned shutdowns and ,, ,’.,- ,~! ,. .,, longer cycles between turnarounds. ‘Ilk has resulted in an actual annual throughput increase ;<’,! ,,,‘.. ,~., over the last few years. BP quoted a few years ago that the continuous capacity creep is -, .-, ,. .“, ,. equivalent to ‘four additional world scale refineries going on stream every year’. .,...’., : ,,;.. ,., .,* , .. , “’.,.: :’ As historically developed, the seamless use of crude oils as petroleum products over the entire .-:. -’,,, : :’,..-,’, . .. boiling range of crude oils is being challenged by the new, environmentally driven clean fiels ,’. , :,,
legislation. This is increasingly regulating in detail the composition of fiels and what can and .! what cannot enter the mogas poo[. Product specifications have become so strict that refiners have to find new outlets or conversion for ‘orphan’ streams such as straight-run C5/C6 paraff]ns, the butanes traditionally blended to make higher RVP gasoline, the heaviest portion of FCC gasoline and gas oils. Other streams, like benzene rich naphtha have become unusable for gasoline blending without fiwther processing.
Emphasis has grown on producing light olefins in the FCC/RFCC unit as well as producing FCC gasoline with higher octane and lower sulfir levels in the C5 to 190C cut range.
Today there is actmally a surplus of catalytic cracking capacity in Europe as the demand barrel ,, ., has become out of balance in terms of gasoline to diesel ratio. This imbalance maybe partly caused by tax structures, which still favour diesel over reformulated gasoline as the fiel of the fiture for low-emission light vehicles. But there is considerable debate (e.g. in Germany) about this justification for lower taxes on diesel. Though demand for gasoline has slightly
“. ,.:, ,, ’:, . .’! ,.,./,, > , , ,,,)<:1, ,,,:’,; ,’1,,:., .,, ,.’ : “,.....,. .,: 9 increased over the last few years it is not predicted to rise further, it is split between a small decline in West Europe and a small increase in East Europe. Diesel and jet fuels have significantly increased in demand and a further rise in consumption is predicted.
:. 6. Legislation on firel composition in Western Europe is clear for the year 2000 with already ,. substantial reductions in benzene and total aromatics content in addition to reductions in .“ srrlfir content. Germany is now moving ahead with proposals for more stringent sulfur specifications and earlier ( tax incentivised ) than expected.
Agreed and expected gasoline specifications:
Gasoline Specs EU Current EU 2000 EU 2005 D 2001 D 2003 US 2004
(1) [2) (3) (3) (4) Sulfur wt ppm 500 150 50 50 10 av. 30
Benzene vol % max 5 (av. 2.5) 1 1 1
Aromatics vol% max (aver. 45) 42 35 -25
1) The current gasoline has benzene levels between 0.4 and 4% and aromatic levels between18and 54?6 2) To be finalised in early 2000 3) Voluntary declaration of the German Govensment, German car aad oil industry
4) Proposed by EPA with reducing ma.xS values from 2004 to 2006
It is interesting to look at the production volumes and product values in Western Europe in comparison to the additional aromatic streams potentially to come to get a proper appreciation of the challenges facing the refining and petrochemical industries.
Western Earopean Aromatics Market (1998) m tlycar billion $
Crode consumption 700 65
Gasoline demmrd 120(1) 17
Total aromatics in gasoline 54 Petrochemical aromatics production (only BTX) 10 32.3 Benzene ex pygas 3.5 0.9 -i Benzene production total 6.8
(I) = USA 320 m t/year
There is a factor of almost six in the production volume of aromatics in gasoline and . petrochemical aromatics whereas the price of gasoline and petrochemicals differed by about 90$ /t, creating opportunities for upgrading of selected streams in the petrochemical sector.
10 Based on an average figure of 2.5% berrxene in gasoline and 120 ret/year of gasoline in Western Europe, about 1.8 mtiyear of benzene would have to be taken out to achieve the 1% . ,,,, benzene limit in gasoline valid tlom 1.1.2000 onwards. However, the experience in the US -, and similar trends in Europe have shown, that nearly half of this ‘surplus’ wiIl not be , ;. produced, some part will be hydrotreated / destroyed within the refinery environment and the ..,,. - rest will be extracted to meet the 1‘A benzene level. It is expected that --700 ktiyear of benzene will become available for the petrochemical industry. TMs represents about 11% of . . the European benzene market and with an estimated growth to be in the range of 2.5- 3% / .,, . year Uds volume should be digested in the market without too significant problems in the coming years. A similar calculation for the projected leveI of total aromatics in 2005 with 35 VOFXOwould require the removal of over 12 mt of aromatics from todays 54 mt aromatics in the European gasoline pool. This volume would more than double the current 10 mt Western European aromatics market (BTX only!). This of course does not even take into account the increase in ,. the market when Eastern European countries will join the European Union in the next decade. But it is obvious that thk aspect of oversupply of aromatics looks like a rather uncomfortable prospect for the petrochemical industry. ,., ~., Naturally, not all these aromatics will be extracted, similar to the benzene scenario some of ,<, the aromatics will not be produti anymore. However, aromatics production in a refinery is not only producing a high octane blending component but also the most valuable source of hydrogen in the refinery. And hydrogen will become even more vahrable in the refinery environment with the upcoming of more stringent sulfirr specifications.
Summary of trends: ‘. Demand for petrochemical products increasing globally by more than GDP; coproducts could regionally become an issue Increasing demand for refinery originated petrochemical feedstocks to > 10% Competition for naphtha and in particular low sulfur condensates with refining Highly integrated petrochemical /refining centres developing, in particular in ME & SEA Overcapacity – low margins - capacity creep characterise threats to refining industry Increasing requirements for lighter, cleaner, low sulfhr refinery products Reduction of benzene and aromatics in gasoline, increasing hydrogen demand >,.,,.’
.,, ‘$,
,, 11 Feedstock for the Petrochemical Industry - Interface with Refining Industry
Continuing rapid growth of about 4% / year in petrochemical feedstock demand is predicted. Natural gas and naphtha feedstocks share about equally the worldwide cracking capacity. However, there are significant regional differences. In 1996 95~0 of the feedstock for ethylene-oriented oletin steam crackers in the Middle East was natural gas based. [n the USA over 60’%of the ethylene steam crackers were using natural gas (ethaneJpropane). In Europe, naphtha comprises over 70°/0of the feedstock to olefin steam crackers.
The lowest cost route to ethylene is based on ethandpropane obtained using turboexpander gas plants based on natural gas. However, etbane./propane-based ethylene plants are limited in propylene output and do not make substantial tonnage of ‘co-products’, i.e. butadiene, 1- butene, isobutene and benzene. In contrast, a rather large share of the annual revenues from an olefin steam cracker based on naphtha is derived from these ‘co-products’ plus propylene. The availability of these ‘co-products’ has Ied in Europe to the build up of a sophisticated petrochemical / chemical industry.
For aromatics production via ethylene crackers from the naphtha feedstock a 70C to 145C heart cut is utilised when focussing on BTX production. The benzene-rich BTX portion of the hydrotreated pyrolysis gasoline and the xyiene-nch BTX portion of the reformate from a ‘continuous regeneration’ catalytic reformer may be combined in order to simultaneously recover benzene and p-xylene stt lowest cost. An aromatics complex typically includes a transalkylation and a disproportionation unit to convert toluene and C9 aromatics to benzene and xylenes.
Operators of olefin steam crackers rated for naphthas can often charge substantial percentages of propane and butane if these are of suitable quality. In some instances hydrotreated light paraffinic kerosene and unconverted hydrocracked VGO can be substituted for part of the naphtha. The cracking pattern is similar to naphtha. Interestingly, there has been a notable growth in the finding and exploiting of so-called wet natural gas fields. These fields produce some interesting and often rather unique condensate crudes. These condensates will increasingly become available in the coming years, primarily from the Middle East but also Australia, somewhat easing a bit the naphtha supply problem in Europe.
Naphtha for petrochemical operations is in competition with use of naphtha for both C5/C6 isomerization and catalytic reforming within the fiels refinery. The C5/C6 isomcrisation plant produces a medium octane light mogas blendstock of 100% paraftinicity. Isomeratc becomes attractive within the unleaded gasoline pools if total aromatics and total olcfins and T90 all
12 must be limited. The C5/C6 paraffins have traditionally exhibited very excellent octane uplitl response to lead even though most are ‘normal paraffins’. Without TEL lead the octane ‘as is’ of straight-run C5/C6 paraffins is so low that mogas blending is almost impossible. Isomerisation or the light petro-naphtha options become the only viable alternatives for this ‘orphan’Str_.
The modem catalytic reformer produces both high-octane reformate and as co-product hydrogen. The reformate can be distilled into lower octane C5/C6 paraffins, a benzene rich .,, , heartcu~ a C7-fraction and heavy reformate. The lower octane C5/C6 paraffins can become ,. ;. >,, supplemental light petro-naphtha or be blended into mogas in order to mitigate total aromatics levels. A benzenehohrene rich heart-cut can be sent to a benzeneholuene extraction unit for recovery of the benzene and toluene. Even with a benzendtoluene extraction operation, most recovered tohrene will become a high octane mogas blendstock. Alternatively, a benzene-rich heart cut can be distilled off in the reformate splitter and thk heartcut can be sent to a benzene hydrogenation unit or benzene control can be managed by the isomerisation unit. ,, ?.,’ .,, ...... :;, Aromatics ‘;”:
The aromatics market globally is nearly 70 m t, whereby the market of virgin aromatics is .. ,,,’ ,. about 55 – 60 m tiyear. The demand for aromatics is globally still growing at a multiple of the ,: ... ; world’s economic gro~ however to a lesser degree in West Europe. ;.,:j.;.’,’,,’f.::,. Traditionally, aromatics provide the closest interaction between refining and petrochemical , .:,; . ‘: .,, .,.,, . . operation. Pyrolysis gasoline ex stearncracker was blended into the gasoline pool and BTX %’. . aromatics were extracted from refinery derived reformate. :’ ;( ...... Today, more than 70% of the aromatics are being produced by catalytic reforming of naphtha f..,, ,, .,’ 25V0 ! ,, which provides a mixed stream of benzene, toluene and xylene isomers. About of . ...,,..,- :. ,, aromatics are being produced in olefin plants and are being extracted from pyrolysis gasoline .<, . .,,,’ } .. , “.,’: :. ,.,.. ~. ,, ,$’ which contains pnmaxily benzene and toluene. Due to an increasing use of naphtha and .,. . ,’, ,., .. . .,, ,, heavier hydrocarbons for ethylene production the amount of produced pygas is also due to rise ,. ., ~ and thereby increasing the pressure on the supply side of the aromatics market. Coke oven -! product streams containing primarily benzene and contribute about 2% to the aromatics pool. While the reformate can be fed directly into an aromatics separation unit the streams which originate from thermal processes have fimt to be hydrotreated to convert dienes as well as to remove heteroatom containing impurities. TMs then permits splitting of the BTX cut into a BT
13 The extraction of benzene and toluene from a BT cut is commercially attractive. The tonnage quantities of xylenes-rich cut from olefin steam crackers is fairly modest, but many aromatics producers combine the pyrolysis xylenes cut with cat reformer xylenes as feedstocks to a modem multi-step plant for the production of high purity p-xylene.
Once a market value or production cost is placed on reformate, market forces from the petrochemical industry work to justifi whether the recovery of the individual aromatic compound eg benzene, toluene, xylenes (and higher aromatics) is profitable. Aromatics can be recovered and separated via different routes depending on aromatics content :. and type of aromatics to be separated. Distillation is applied with high aromatics #. ,. concentration, liquid-liquid extraction for streams with medium concentration eg. originating ., from a catalytic reformer.
The majority of the aromatics complexes are located adjacent to refineries as this provides a reliable source of naphtha, provides an outlet for the by-product streams into the gasoline pool and sharing of utilities and offsites.
Benzene and p-xylene are in terms of volume produced the aromatics with greatest market demand which is predicted to rise fin-ther.
Alternative routes to aromatics have been developed based on LPG or light olefins. The Cyclar Process, jointly developed by BP and UOP, enables the conversion of LPG into BTX rich concentrates. A commercial plant is currently being commissioned in Saudi Arabia. Asahi Chemicals and Sanyo Petrochemicals have developed and commercialized the Alpha process which converts oletinic streams into BTX rich concentrates. Feed can be C4/C5 cuts from Pyrolysis Gasoline as well as light FCC Gasoline. Besides producing more than 20 wt% of aromatics combined with a significant reduction in olefins, the process is also claimed to reduce the sulfir content.
Benzene Chain
Benzene; The world market for benzene is about 30 m tiy and is expected to rise to about 40 m tly in 2005, a rise of 3- 4°A/ year. Supply is expected to exceed substantially demand in the
middle and long term, resulting in utilisation rates of around 75°/0. The benzene capacity in West Europe is about 8 m t/y and still rising while the production is totaling about 7 m t/y, demand is expected to rise in West Europe at close to 3 % / year. More than 50~o of benzene in Europe is extracted from pygas, coproduced in the steam cracking of naphtha or gas oils to
make olctins (the dominant fcedstocks in Europe). About 15°/0 of the capacity is based on
14 . rcformate extraction and nearly 20V0on tohrene dealkylation (swing operation). Less than 5’%0 are derived from coal. Already known global capacity additions for benzene and most of its derivatives will most likely exceed demand growth. At the end of this year the new ~ petrochemical complex in Yanbu, Saudi Arabia goes on stream which includes a new benzene plant (350,000 tiy), based on the UOP/BP Cyclar process. Chevron will startup an aromatics complex at Al Jubail/Saudi Arabia with a benzene capacity of 480,000 tiy, based on Chevron’s Aromax technology. ,’ ., More than 50% of chemical grade benzene goes into ethylbenzencktyrene manufacture, which ,, ,<$,,,. ,, .’,’ ,, ,. is a continuously growing markeL 170/ois utilized for the manufacture of cumene, an ,, intermediate to produce phenol. About 15°A of benzene is hydrogenated to produce cyclohexane, base material for the nylon production.
Ethyl Benzene/Styrene. The world market for styrene monomer in 1999 is in the order of 19
m L the Western European market represents about 25~0, similar in size to the US market. Global styrene demand is expected to rise by 3% / year which equals the worldwide GDP growth rate, growth rate in Europe is expected to be equal or slightly below. Styrene monomer as the largest benzene derivative has its own cyclicality based around supply and demand, new investments and the strength and weakness of the consumer market. When styrene is strong benzene usually benefits and when operating rates fall, benzene gets under pressure. The majority (two third) of styrene is used to produce polys~ene, the rest to produce
.,’.) .J ,,. elastomers such as s&rene-butadiene-latex (SBL), styrene-butadiene-rubber (SBR) or ., copolymers such as acrylonitrile-butadiene-styrene (ABS), ,,,,,,’
,, Two routes dominate the production of styrene. In the traditional route chemical grade .,, benzene is alkylated with dilute or polymer-grade ethylene over a fixed-bed zcolite type ,.’ . catalyst in the liquid phase producing ethyl benzene. This is then dehydrogenated to styrene. The second route, the POSM (propylene oxide / styrene monomer) coproduction facilities, ‘, ., have introduced a new dimension to the styrene picture, adding a new variable to the supply/demand psychology. Produced in a molar ratio, per ton of propylene oxide, the leading produc~ 2.2 t of styrene monomer are produced as byproduct. Both, conventional styrene plants and POSM plants are capital intensive and both are ultimately based on benzene.
Significant new styrene capacity, based on the POSM process, is coming on stream in . ,. , Western Europe in 2000 and beyond. It is not clear if the styrene produced via the traditional <.., ethyl benzene dehydrogenation route or via the POSM route will have commercial advantage ;’, , ,.. .,,, besides being produced in new, world scale sized plants. It will most likely, in combination with an oversupply, force smaller and less efficient plants into closure.
Cumene (isopropyl benzene) is produced from chemical grade benzene and chemical grade propylene on zeolitic type catalysts in liquid phase. Typical production capacity is 200 ktiyear of cumene. Essentially all cumene is used for phenol production with acetone as a co-product. Global phenol capacity is close to 8 m t with a predicted growth of about 4% / year. About
ss~. of all phenol produced is being consumed in Europe. Phenol is the building block for a range of plastics and chemicals and by-products loosely grouped under the collective of ‘Phenolics’. This is a real growth area with the major drivers being polycarbonate engineering plastics via bisphenol A and phenolic resins, caprolactam and others such as alkylphenols, aniline (800/0used for MDI production), etc.. A commercialised, alternative route to Phenol is DSM’S Tohrene Oxidation Route, starting from Toluene with benzyl alcohol and benzaldehyde as the main byproducts.
Cyclohexane growth is expected to be around 4% / year, driven by solid Nylon 6 and Nylon 66 demand, which consumes about 90~o of the cyclohexane produced. Adipic acid takes about two third for Nylon 66 production, the rest is taken for the production of Caprolactam, the intermediate for Nylon 6 production. However, new technologies for producing caprolactam starting with butadiene and proceeding through adiponitril may challenge the cylohexane market. The Western European market for cyclohexane appears to be in balance. The price of hydrogen is most critical in the economics for producing cyclohexane.
The other benzene end-users nitrobenzeneJaniline, alkylbenzene and others (maleic anhydride, chlorobenzene, etc.), which are comparatively small consumers, are expected to grow substantially stronger than GDP.
in summary, concerning benzene, it is expected that the market will grow at least up to 2005 by a rate of 3- 4% / year driven by a strong derivative demand, in particular styrene and cumene. Originally raised concern that due to the benzene limitations in gasoline the benzene market would be flooded will not materialize. About I/3 of the total benzene in todays gasoline pool will stay in the pool, another 1/3 will be extracted or destructed and the last 1/3 will not anymore be produced.
Toluene
Toluene is produced in large quantities in BT and BTX extraction plants with primarily pyrolysis gasoline or reforrnate as the major feestock. Toluene extraction capacity in West #.,.
16 Europe is about 2.7 m tly, whereby more than 60’%.is reformate and 30’%.is pygas based. If ‘,’ not back-blended into gasoline or used as a solvent toluene is either eonvertcd to high-purity be~ene via thermal-or hydro- dealkylation (TDA or HDA) or via disproportionation reaction into high-purity benzene and a para rich xylene stream. Thermal dealkylation plants are the first plants to be shut down when the benzene price is dropping. Thermal dealkylation will set the floor price for benzene and is expected to contract by more than 5% / year, in particular in view of the increasing benzene capacities being build. In toluene disproportionation (TDP), toluene is converted to benzene and xylenes with xylene, in particular p-xylene, as the desired product. A new developed technology from Mobil converts toluene to high-purity benzene and a mixed-xylene stream with a para-xylene content of up to 85’%0. Of the three major aromatics, tohrene is the one with the lowest commercial vahre. The price of toluene normally reflects the mogas value. ToIuene extraction capacity in West Europe is about 2.7 m t/y, whereby more than 60% is reformate and 30% pygas based.
., .’ . . XYLENES
Of the Xyfenes only p-xylene is of significant commercial interest. The chemical markets for o-xylene are comparatively small and trivird for m-xylene.
Total p-xylene (PX) output in Europe in 1999 including East Europe is estimated at 2.5 m t, projected to grow to 2.9 m tin 2002. TMs includes the new BP Amoco plant in Geel. The PX demand (predicted growth rates 5 -7%/ year) is driven by the manufacture of PTA (purified terephthalic acid) which goes into polyester fibre, resin and film as well as DMT (dimethyl ., ,,.., terephthalate). About 99~0 of PX goes into PTA and DMT, only - 1’% are used for solvents, speciality coatings, paints, insect repellents and herbicides. ,, ,.., .,, , P-Xylene - world wide 1996 1999 2002 2005 Capacity m tiyear 17.6 19.9 25.3 25.7 ., Demand m t/year 13.1 14.1 16.5 19.2 . . Plant utilisation (%) 75 71 65 76 ,-
The PET (polyethylentherephthalate) – global production production volume 1998 4.6 mtiy - produced from PTA and ethylene glyeol, had fantastic growth rates of 12- 15%/ year in the ., pasL mainly in the plastic bottle business. It is expected to grow again in the coming years by 7- 10%/year, in world-scale plants above 200,000 t / year. ,,.,, , -. t “,, -. . , ,.,,f,’, ,.’ ..,,.. ‘l. ,,J ., ::’.., ;,? ~’,..,.:..,, .t,,~.; 17 :,, ,’ :...-.+’/”’,/ ,,.,; ... ” .,-.’4 Conventional PX technology is based on the isomcrisation of mixed xylenes from reforrnate or steam cracker gasoline. High purity product can be obtained using crystallisation or selective adsorptive separation processes such as UOP’S Parex and IFP’s Ehrxyl processes. Tohrene disproportionation (TDP) offers an alternative route with the more advanced catalysts able to produce a PX rich stream but with coproduct benzene. Todays available technologies enable designing p-xylene production plants with benzene to p-xylene ratios from O to 1.6, depending on local/ regional demand.
Xylene production is an option for a refiner who operates a high octane reformer with continuous catalyst regeneration. Worldscale p-xylene units have at least a capacity of 300,000 t / year, which requires feed from a catalytic reformer of about 30,000 bpsd capacity. High reforming severity is required to improve aromatics yield.
From an economic scale point of view it is diftlcult for a single refiner to justifi the investment into a PX production facility. It is therefore increasingly discussed to build combo units, ie, PX production facilities which are owned by several refineries / petrochemical producers.
In summary, there is globally a healthy outlook concerning PX production. Market demand will steadily rise - and is rising now again with the recove~ of the SE Asian market. However, regional surplus cm be expected and alternative outlets / processing of higher aromatics containing streams, in particular pyrolysis gasoline based, have to be investigated.
The total capacity for o-xylene in Western Europe is ca. 700,000 t/year, dominated by Eni ( 140ktpa), Total Gonfreville(110 ktpa), Shell Godorf and Exxon Botlek with 100 ktpa each. The end users are converting most of it to phthalic anhydride. European demand is forecast to increase slowly in the next five years, operating rates are still declining, expected to increase slowly from 2000 onwards.
Summary
The demand for petrochemical feedstocks derived from retine~ operation will continue to rise from about 8V0today to >10’7. in the next few years.
Legislation is now significantly influencing the composition of most of the refining industry products, the requested general reduction in aromatics content in gasoline has the potential for providing regional oversupplies.
;. 1, ,. 18 t
This is against a background of refining not being a particularly profitable business and investment in refineries is difficult to justify. All this points towards swapping volumes/streams between different refinery sides, debottleneckhg of existing units, increasing further cooperation between refineries and petrochemical companies, creating super-integrated * sites where eventually up to 20°Aof the refinery output will go to higher value petrochemical products as compared to the current 8% average.
Whh the potentially growing surplus of aromatics, in particular higher aromatics in mixed streams such as pygas in Western Europe and the growing demand for ethylene and propylene, considering complete hydrogenation and hydrocracking of heavy pygas to LPG ,,,,...: feed for the steam cracker could provide an interesting alternative. ,~,>, . . Further advances in catalyst and process technologies are required to increase selectivity in xylene isomerisation units, in particular the isomensation of ethylbenzene to xylenes. ., These challenges are also opportunities for the Petrochemical and Refining Industries with the potential of enhancing the competitive position for both, in particular when considering building aromatics extraction or treatment units together for several producers/ consumers.
Expertise, skills and tools are available for the respective industries to assist in the identification and assessment of potentiaI opportunities up to the execution of successful projects.
References:
,.. , Annual Dewitt Petrochemical Review Conference in Houston, March 1998 and 1999 ,., .
Johnson, A.R., Hallee, L.P., Bowen, C.P., ‘Future directions of petrochemical/refinery inter- relationships’, 14th World Petroleum Congress, Stavanger, 1994
.,.’ Graeser, U., Keim, W., Petzny, W.J., Weitkrunp, J., ‘Perspectives in Petrochemistry’; Erdoel, . . . Erdgas Kohle, 11l.Jahrgang, Heft 5, Mai 1995, p.208 -218 -,. ,
A1-Ubaid, A.S., ‘Integrating Refinery and Petrochemical Plants to maximise value added’, Presentation at OAPAC Conference, Cairo, 1996 Crawford, C. ‘Refining for petrochemicals’, Hydrocarbon Engineering, Vol. 2, No. 5, Sept. 1997
Haelsig, C.-P., Taubman, J., Goelzer, A., Crawford, C., ‘ Added Value by Integrating Refining and Petrochemical Operations', Presentation atthe ERTCConference, London, 17.- 19.11.1997
20 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 .,:
-.”! M. Schuller 1),P. Hodges 2) ---- 1)TOTALFjNA, pafis [a D~fense, France ., ,. ~, ? DEWIn & company Incorporated, Wilmslow, Cheshire, United Kingdom ,.’: ;1 ,., .!, ..,.! ,., ,., ,, ,’+.-.,. .,.,,, ,: ,,, ,’:,’, ,,: How will EU Environmental Legislation Impact on the Aromatics Markets? ,..., . ‘,, ,,~.. ! ~., ..$ “,, .:, ,,, ,’;.? .! ,.‘,,.,,.,. ,, ,., ,,, ;.’,’ ,, ., -, .’,;.: ‘,,< 1/A GENERAL MOVE TOWARDS SUCH REGULATIONS
7A/ AN INTERNATIONAL MOVE COMING FROM THE USA :“
,- ,,, The first Clean Air Act in the USA dates back to 1970 and since ‘., . then it has been amended regularly with successful results in ,, terms of air quality improvements. The impact on ozone in the air ,, is one example.
There is no denying ,.’ the air has improved...... :, Worst Daily Ozone Readings in California 700
600
: 500 G . .c1 400 & a 300 z c1 200 E 1111 1“: , , E=py#z:—— — :
1970 1975 1980 1985 1990 1995 1996 [n the USA itself, California has always had since 1943 a leading role in terms of pollution control initiatives.
In the”United States ......
The Clean Air Act dates back to 1970. In 1975, CAFE standards were adopted. Catalytic converters were installed and leaded gasoline was phased out by the early 1980s. Oxygenates emerged as octane enhancers and to aid with emission control by the mid-80s.
The rest of the world and more precisely, the European Union (this is the area we will refer to when talking about Europe) and Japan are following the example developed by American legislators.
The next steps will be in year 2000 and 2005 for the EU. Japan and Korea will also take measures in 2000.
lB/A EUROPEAN MOTIVATION
This policy inspired by the American example has gained strong support from European political authorities. Several surveys have shown that - together with unemployment - environmental concerns are one of the most important issues for the European population (cf: report of “le bureau Europ6en de l’environnement”).
22 -n-EmLEo=mmEma&ATs
120
100 :“
80 . . ,. H 60 ,., , ,’. ,. 40 ,,,
20
0 ,,
LBwJ3mNc= lNG53-E ,, ., ,.. , -. Obviously, these concerns cannever be fully satisfied. Asthe chart shows, the pace of change in the environmental field is constantly accelerating. It took over 25 years for the USA to remove lead from gasoline, but the EU will have accomplished this in most countries in only 15 years.
Human nature being what it is, one has to expect that expectations will constantly be raised as initial targets are met. Fuel ., reformulation is also ultimately limited in what it can achieve alone. ‘,,,. \’- Improvements in vehicle technologies will also continue to be essential. ,, .,, .,’,
lC/ THE DECISIONS: AUTO OIL PACKAGE .,
. . ● What is the Auto-Oil package?
The first Auto-Oil directives proposals were issued in July 1997 by the European commission. It suggested that car engines should ,’ .- be modified in order to obtain the air quality improvements that ; ., were expected. It also suggested for that purpose that fuel should ,. be reformulated. ,, :,,
Following a lengthy decision process, the European council, the ,. ,.. !,, European commission and the European parliament reached an ,, ,., agreement in june 1998 on new fuel specifications to be applied.
,,’
,, 23 ‘, ,,. .,’ ,.-. ● What can be expected?
The results should be very significant. The Auto-Oil package is expected to result in reductions in emissions per vehicle in the order of 70 YO as compared to current standards.
A two-step tightening process has been planned focusing on 2000 and 2005, in order to help ensure that EU Air Quality Objectives for 2010 will be achieved.
Let’s look for a moment at what has been decided for the EU based on a Fuel Quality comparison:
- [ Spec Spec Spec Average I Current \ 2000 2005 I California
I Gasoline I I I I I Sulphur.wrmrn 500 150 50 18-20 Benzen le,voI!10 5 1 0,38-0,60 Aromatil ics,volYO 42 35 23 - .,-10/. 4Q -~,4n Olefins. VUi/0 I I Iv I I 4. Q--?. U 1 Oxygen,vol% I 2,7 2;0-2;2
Diesel Density 0,86 0,845 0,842 Sulphur,wppm 500 350 50 140 Polyaromatics,volYO 11 Cetane 50 51 53,8
Most EU countries will also have to phase out leaded gasoline in year 2000, with derogations available for some Mediterranean countries up to 2005.
The reduction in maximum sulphur contents from 500 ppm down to :. 350 ppm in 2000 and to 50 ppm in 2005 will be a drastic change 1. ,. and address an issue of similar if not more importance than aromatics. .,
In Asia, there is no common legislative body. But Japan will introduce a 1YO benzene limit from April 2000. And Korea will introduce a 1,3Y0 - 2,3°/0 oxygen requirement, and a 2°/0 benzene limit from January 2000.
24
. . ,,
The reaction of the European oil industry has been mixed. The specifications for year 2000 are seen as hard, but also seen as achievable by the refining indusby. But the 2005 specifications ,:, ”, ., ,!, : ., : raise further difficulties since it appears that: ,, ,., ~ .,.,.,, ; , ,’. ,. .,, Either the industiy will have to make some expensive investments that moreover may increase the current surplus of gasoline in Europe Or the industry will not make the necessary investment. will not be able to meet the new limits and will ha~e to negotiate the support of the public authorities for derogations or investment subsidy. Otherwise some sites could face closure decisions.
Unlike the USA, there is currently a 20 million tonne surplus of gasoline in Europe, coupled with a shortage of middle distillate. As shown in the following Chart, refining experts Wood Mackenzie ,. forecast that this situation will continue, $., .
EUROPEAN SUPPLYLI)EMAND IS UNBALANCED ,, ,.’. MT 25/ 20- 8 ■ 1998 15. ~ ■ 2005 — 10.~ 5- ~ o-~ -5- ~ -1o- ~
-15}’ ) GASOLINE . JEVKERO GAS/OIESEL FUEL OIL OIL Source: Wood Mackemie
As a result, refiners face some difficult choices and their options are more limited than might at first appear.
.. 25 .-, 2/ HOW ARE COUNTRIES AND COMPANIES GOING TO ADJUST?
2A/ THE INITIAL SITUATION FOR BENZENE AND AROMATICS VARIES.BY COUNTRY
The current situation of EU countries as far as the quality of gasoline compared to the new regulations are concerned are quite different. Whilst Scandinavian countries are ahead and should be in a position to adjust very easily, Mediterranean are late and will have to make much bigger efforts to match the Auto-Oil requirements. Most other countries are somewhere in an average position.
EuE40GAmLrNEMAEuwlswiv NxlIawxHARAclIiRrslm
Equally, not only is Europe’s gasoline situation significantly different from the US in that we are net exporters of around 20 MT per annum and not importers, but also gasoline markets differ from
,’ . 26 : #, ,.
,. .* .< ,, ...... : <:. ‘ ., ,i .,.. .’ ..;. . .. ‘. . .,. .. -,. - ‘- . ..- ,- -...... , ...,-. ~.,.<,.-:-,.”. ,,, :.,,, /<,-~,‘1 ,’‘ :::.-./,:,. !’,.; ... .,,.,4:,: ,, ‘“!, . .4 ,/- ,“ ,>: ,!. -,, .,,/..,>; -,.,,.$ /,., ,, , , , ,, country to country. The map shows that whilst countries such as . ) ;f... .,:(.’.j ,+i: ,’.,”., ,,.*J the UK and France are net exporters, countries such as Germany ,-;:;,:,;:;;;i:;j+ are importers and others such as Italy are balanced. As a result, /;<{,a.~;:,/ ‘‘$,~;,.,,,( ,ifiofi;,,:,,;.;;, ‘i refiners often start from different assumptions about what is ,,,.~ f.,’f.iff, -,,.;,,v::f,’:1;,)): ,,,:;:\Y/}.,.:!(, possible or desirable. ,, .‘,$~~.’.’$:..I ..’:::$:’:’’,,,,.!(4, ..,...... ,,. \,,./.,.,,.,.1.!,;-,,,;;.-.,, 2B/THE INITIAL SITUATION VARIES ACCORDING TO THE -.:... ,. :,1 ~,., +1,.:.,~-r~ REFINERY SCHEME , .,‘.L.,,,,,-.,! ,., :...;-, : :i;Tf.j ,,.,$,,,.j,y,,: ..~~ ,,,:/} The situation is therefore very different from one scheme to the ~‘ ‘y~:r., ;:j ,, <,,,,,,!.::,. other. . .:,L,,... , :,,:, .’,, ,,,.,.i..,.,,,,.,,,-;;.:. As regards the total aromatics constraints of 42 Y. set for 2000 it is ,,,‘,.,:;:-,‘;.,.,:“:,.: 4;’ ,,(t,;,,’!l: anticipated that refineries having an FCC will have a clear ;,:,/:,’:;‘, f: .’J ,, ‘1,1,,; advantage compared to the others. They will meet this constraint ,.,. , ‘;. -, without major changes. Those that do not have an FCC (for ,:,, ,, ‘/: ,,,-... ,,., example, refineries that have an hydrocracker), will not benefit ,,.’ ,:, ,,, ,,,, from the dilution effects of the FCC gasoline cuts and have to take “,.,,,:..,! ,, measures such as dilution or extraction of an aromatic cut. .,.,,:,. ., ,., ,$,, .,.1 Having said that, let’s not forget, that FCC’s are causing the ,,, , ,,,!,,.: , refinery industry a lot of other problems such as the sulphur and :’,. ...,,,:.,.’ ‘!t, olefin contents or the lack of middle distillates. ,, ,,.4’:-,.,,,; ,$’,, .,J,.),..,.,, + , ‘,.L,“ : ,,j,+ “., ,,,,, Refiners have essentially had to choose from 3 options in order to ,,‘1‘.,~,.:.“.);,’‘,.; . .... ,,,,i~ J.7L meet the 2000 benzene specification: ‘,,, ?,,;,., ,.,:,J!,f ,,,;,,,,..-,. ,’,:, The elimination of benzene precursors by setting a deeper cut point between the light naphtha (sent to isom C5/C6 or sold as light virgin naphtha to the stream crackers) and the heavy naphtha. The utilisation of spare distillation capacity or creation of additional fractionation capacity to produce a benzene heart cut in order to sell high quality benzene. This is the option that has been chosen by TotalFina for benzene. An extrati]on capacity of 250 KT/year will start in Antwerp in January 2000 (which will also produce 500 KT/year of isomerisation grade xylenes). The hydrogenation of the benzene in an isom C5/C6 unit (if available).
,,,,, ,..’’,..,. ,.., . ‘.’. These options mean that the 2000 specification can be met by ,,.,,,..’. .: ,, !:.,,Z ‘,., virtually all refiners. But this is not true of the 2005 specification. ,,.,.,1,,.,..~ ,.:. ,.,. This specification places twin constraints on refiners as the sulphur ‘,:,,, ,‘,, ,.,. limit means that many will need to run their reformer to maximize ,..,,,:, ,’ ..., ,,,“,.1 the production of hydrogen, but at the same time this mode of operation will also maximise total aromatics content, making it more difficult to meet the 35% limit.
2C/ WHAT CAN REFINERS DO IN ORDER TO ADJUST FOR 2005?
In order to reduce total aromatics content, the following solutions are possible:
*Technical alternatives:
● Either building new capacity for blending components (Alkylates/lsomerates/Oxygenates), or importing them, to dilute the gasoline pool and adjust the aromatics content.
● Or building new aromatics capacities to extract the non- desirable molecules (extractive distillation, solvent extraction, xylene fractionation etc).
*Commercial alternatives:
● Segregation and the export of gasoline to areas where regulations are less severe (Lebanon, Nigeria etc).
● Sale/swap of “heart cut” streams either internally or externally.
● Sale of pure aromatics.
*Strategic alternatives : ● Refi~ety mergers in order to increase the size of the refinery system and improve profitability by economies of scale. ● Integration of refinery and petrochemicals operations to improve flexibility, to maximise synergies and optimise aromatics extraction. ● Eventually, dis-investment possibilities such as refinery closures are also a possibility. But these can easily represent a high cost, and whilst reducing the gasoline surplus would leave Europe even shorter of middle distillate.
28 At the end of the day, reality will combine a mixture of all these options driven by economics. ,.,,,. ,~ ., :, ,’., ,> ;, If one takes the example of benzene removal to meet the 2000 ;, ...,,: specification, it seems that about 1/3 of the refiners have chosen ,, .!, -,, , not to invest in removing the benzene precursors. They will ,.,. ,. ,, instead extract benzene heart-cut, with the intention of either ., :’..:’ selling this or swapping it on the market. Europe is short of l’. :., ,.. benzene, and so this additional product can be absorbed by the ,..,.,,::,,:;. ‘., ,, ;..- ., ,., market, at the cost of lowering margins for existing producers. .,’’: .~’.,, ,, .,, But unfortunately, this solution will not work for the 2005 ?,, ,- ., ,, specification, as the markets for toluene and xylene are not big ,. ..; ,.. .,. .,,’ enough, even on a global scale, to absorb the additional quantities that could potentially be produced...... ,.,,,:..,., , ,.,.:.,: ;: ,,,,,.. ,,.., ‘. ,, At the moment, it appears that refiners fall into 5 different types of ,-,,- .,”. . :“, ,7 ,; ,. category in regards to meeting the 2005 specification: . .. . -, Supersites, big and integrated, they have flexibility that provides them with an internal solution. /nvestors, who will do what is necessary to match the limitations. Internal and External Product Exchangers, who will extract some streams with the intention of exchanging them for alternative octane-containing streams either internally with other refineries of the company, or externally in the market. Extractors, who see an opportunity to extract additional aromatics and sell them in the market. ,.. Shutdown candidates, who will face a dificult dilemma.
1’
-!, ,.,
.-,,, ... ,+. ... .,; :: ; ,, . . ,, :. ,., .,.,:,;..: ,, ,, ,. ’+,, ,’), ., ,. . .’! ,,.:,, 29 ,, ::’..,.,, ,. ‘.,,. “,, .1,.,,, ,, .,,,J ,., .,,. ,;,. 3/ THE IMPACT ON AROMATICS MARKETS AND ON PRODUCERS
3A/AN EXCESS OF PETROCHEMICALS PRODUCTS GLmALAR!!cSMARKMs ARESM?LLBYCWPARKN
35
30
25 ❑ Isomers 20 l?ilXylenes 9 15 ■ Toluene IElBenzene 10 E 5 0 1 Asia N America Europe
* Main features of the aromatics markets:
[n 1998 aromatics markets represented about 13 MT/year of product: 6.5 MT benzene, 2.5 MT xylenes, 2 MT paraxylene/orthoxy lene isomers and 2 MT toluene. These markets are growing steadily about 1‘Yo above GDP and expectations for the future are positive. The downstream markets of our products are in the chemicals such as:
30 ,.., .,
- styrene/polystyrene and cyclohexane/nylon ..’
BENZENE’S MAJOR APPLICATIONS
Consumer Dlsposables@ VCR Tape Housing Toys Tyres o~ Nylon Fibres Automotive Resins - Industrial Solvents
PhwoodResi”s ~ .<,, Polycarbonate Resins ,,.,, Nylon ,. , > Detergenk ., ,, Lube Oil Ad I ves @ ,.’.
Polyurethane m Intermediates Dyestuffs
,, - paraxylene/polyester ,.,.- ,:-. :, ’’.:,.1 ;.’: ‘., .. , , I PA1L4XYLItN11’S MAJOR APPLICATIONS I .,’ :-, ., f,’,. ,,. , ‘. ,:. ., ..:.;. ,;
!<’:,, ,,. ,,. .,.,,,f ~,’;.. , .,,. ..-...... ,,‘, . ‘!.’;. ,. ., C1lJ ,, .:, ... ., container ~.,,.,.,: . :. :. Q .,!.,,’,. resin <11> ‘., ; ““”.;,,. ..,,,~’ b . ... , ,?. Px . . ., :... ,.,; ,, ,,’ > staple ‘. ,. c?%
.,,..,,1 ~, ,,; ..,. ,“. ,., In addition, many PX producers such as TotalFina also produce -.... ,,>.’., orthoxylene from xylene feedstock, and this is a major feedstock ,,,, ,, : .<,,, .-. ‘,. < t for the production of plasticisers for the PVC industry and resins for the paint industry.
All these products are essential to our modern way of life and are used extensively in a wide variety of major industries such as automotive, textiles, packaging, construction and pharmaceuticals.
*The theoretical new quantities:
As far as the 2000 specifications are concerned, the major impact i,. . will be on benzene where the theoretical quantity needed to be extracted would be 1,8.MT in order to meet the 1‘A limit. [n reality, estimates today count on 700 KT. These numbers are based on what refiners themselves declared to DeWitt.
This is far less proportionally than in Japan (700 KT also expected, but from a gasoline pool of 55 MT only), but twice the US volume of 350 KT per year which was extracted from a gasoline pool of 350 MT i.e. almost three times bigger than the European gasoline pool.
The move to a maximum 42’%.limit for total aromatics in 2000 is not expected to cause refiners a major problem, as the EU average is not much different. But the 35% limit set for 2005 could theoretically give an additional quantity of up to 10 MT, if all the aromatics were actually to be extracted, based on an expected EU gasoline pool of 150 MT in 2005, when several Central European countries are expected to have joined the EU. This is almost equivalent to the total size of the aromatics market in Western Europe.
It is clearly impossible for the aromatics market to absorb such a volume. Not only would prices fall well below breakeven values. But refiners themselves would be badly hit, as many have large petrochemical operations. h would not make any sense to “solve” a gasoline problem by flooding the petrochemicals markets.
Having said that, it remains to be seen how much aromatics refiners are exactly going to extract. We believe that most of these theoretical quantities will not be produced and will be handled inside refineries as described earlier.
32 *Why would refiners consider to extract such a volume of ,. aromatics?
First, Western Europe has a surplus of gasoline, and so it would be superficially attractive to reduce this through intensive extraction of aromatics. In addition, as gasoline consumption is more or less stable, there is no risk that by adopting such a strategy, Europe would lack gasoline. This is a major difference from the situation that existed in USA.
Second, the refining business is Western Europe is not very profitable. Many refineries are not as big as those in the US, and so they lack the USA’s flexibility and economy of scale. As a consequence refiners are financially constrained and often very reluctant to invest in new units.
*Without major investments, it remains to be seen how much the petrochemicals markets will be able to cope with such new availabilities of aromatics?
As for aromatics in general, the strong growing demand for these products, together with a current positive economic environment will be a suppodive factor.
Having said this, it will not be possible to have the market absorb quantities that it cannot swallow. The expression is famous and does apply here: “You can take a horse to water, but you cannot make him drink”.
The potential volumes available for extraction are such that it would be impossible for the market to absorb them all. Refiners will have to find alterna~ive solutions, such as adapting reformer conditions and/or throughput. Given that the aromatics market in Europe is 13 MT/year and growing about 400 KT/year, it cannot possibly absorb an additional 10 MT in the space of a few years.
In a competitive market, the leading adjustment factor will be the alternative values of the various streams. And not only petrochemical streams will be affected, because the potential for the replacement of existing streams is limited. Both refinery and petrochemicals streams will be affected. If too much aromatics become available, prices will decrease very dramatically. When deciding how to meet the 2005 specification, refiners will have to consider the relative prices of refinery products versus petrochemicals products and versus other investment costs. And they will have to make a choice.
Again, we have to stress that the position with regards to meeting the 2005 specification is different from that surrounding the 2000 specification. With benzene, one of the positive aspects was that there was in Western Europe some “heart cut” extraction capacity available. On paper, “processing” 700 KT/year of new benzene was possible. Equally, the market was net short, relying on imports in order to balance supply with demand.
Nevertheless, these imports came from producers in E Europe, Asia Pacific and Latin America, and these producers will not simply disappear as a result of the new European production. Equally, the European benzene market itself is 6.5 MT/year and growing only by about 2 YO per year i.e. 130KT/year. So the market will not find it easy to absorb the 700 KT of new production.
* The solutions:
● TECHNICAL SWINGS
Many big players in this industry are involved in both crackers and refineries on several sites. This will give them a lot of flexibility and possibilities of adjustments among the various productions. But 6, . . even so, these companies have already made major investments in aromatics production, and will be reluctant to take on the role of ‘swing’ producers of benzene and aromatics.
● COMMERCIAL POSSIBILITIES
Europe May try to export some surplus product to other Regions of the world, but this would only offer limited possibilities as many markets are already suffering from over-capacity.
34
.. ,,. :
● POLITICAL ASPECTS
Some refiners might try to approach the political authorities to request either derogations or tax incentives.
But having considered these options, we feel that the maximum ,,> volumes expected will not all find room on the market. Aromatics ., prices will go down and greater development of internal options for refinery and cracker operation will appear than many players .-.;,., currently expect. .,.
. ,,. 3BIA HEA W COST FOR THE INDUSTRY
The measures to take and that remain to be taken will put a lot of financial pressure on our industry and will not help to improve the already poor state of profitability in the refining sector.
Costs appear at several levels:
From a macroeconomic point of view, most refineries have to ,, make a strategic decision regarding what they would do to meet the limitations: investments must have been decided in terms of ., ,.’ ,{.!,. ;’, elimination or extraction of the new cuts. Some logistic aspects ,,, , ,.,’ were most otten involved, bringing their own additional costs. ,, ,., At TotalFina, we decided to go ahead with the construction of a . benzene super site in Antwerp. We thought it was the best way ,, to consolidate our position in the petrochemicals, to meet the ,, new gasoline specifications and also to maximise our .“ profitability through flexibility. ,., On the contrary, some companies with isolated and unsophisticated refineries ~11 have to face difficult issues such as heavy investments or closure. These may have to choose between running a refinery at a “small” loss (1O M US$ per year) or shut it down which could cost between 50 and 200 M US$ and is a risk taking decision in terms of environment.
CONCLUSION: whether we like it or not, California is leading the way in terms of environmental legislation. The future of MTBE is now at stake and we have to anticipate the next steps to keep ahead and adjust smoothly. We also have to admit that economics and scientific rationale may be less important than public concerns or political factors to understand what will be the next decisions.
There will probably be a time when the car industry will have to contribute more than it does today to these concerns. Nevertheless, it would be unwise for us to plan on the basis of this assumption.
Our challenge is to anticipate what products our markets will require, but also what the environment may impose to us in order to do more than just being obliged to follow.
For the time being, there is no doubt that the new regulations are going to have a major impact on refineries and on the aromatics markets. In petrochemicals, growing markets will help to absorb some additional availabilities of product. But it is nevertheless unrealistic to believe that this growth will be sufficient to enable supply and demand to balance aromatics markets, if many refiners decide to take the extraction option.
We will either face major economic issues or new discussions will have to take place by 2005 in order to contemplate with the public authorities the ways to cope with the new limits.
36
., . . ,-, , . . ...’ : :. ,. .,, ,<,”T , , ,.,.:.,: ,;, DGMK-Conference ‘The Future Role of Aromatics in Refining and Petrochemisby”, Erlangen 1999 ~.,e, ,.’,...,, .. !,’ ...... “ . ,. .,, ,
K. van Leeuwen CONCAWE, Brussels, Belgium
The Role of Aromatics in Transportation Fuels
i Aromatics in transportation fuels Aromatics, in particular benzene, have come under pressure in transportation fuels. ‘ Benzene because it has been classified as a human carcinogen, other aromatics in . gasoline mainly as benzene precursors contributing (to a much lesser degree) via de-alkylation during combustion to exhaust benzene. In this paper we aim to assess : the role of aromatics in the manufacture and performance of gasoline. The average ‘ aromatics content of EU gasoline was at some 38% in 1990 and with mandated 2005 gasoline specifications, the aromatics mntent is expected to reduce to an average level of 33~0 volume. The average benzene content of gasoline in the EU is ., expected to decrease from some 2.3%’o to 0.8% volume by 2000. ,. Traditionally, the main gasoline quality parameter has been octane. The average pool octane in the EU is close to 95 RON (in US terms (M+R)/2 of some 90), with : ULG-95 as the main gasoline grade and with smaller volumes of high octane ULG ,:; 97/98 and low octane regular. The leaded 98 grade is being phased out, with limited derogations beyond 2000. .’, , .-,:’ One of the main effects of the EU specification changes for the years 2000 and 2005 for gasoline is a pressure on the octane level of the gasoline pool. This pressure is a result of the compositional limits imposed on most of the traditional high-octane ‘,’ molecules (benzene, aromatics and oletins) as well as a reduced blending capability for butane in summer to control RVP. Reduction of the sulphur content affects the octane pool since desulphurisation of gasoline blending components may result in a loss of olefins (octane loss) and may also lead to a volume loss e.g. when high- sulphur, high-aromatic heavy CC gasoline is rejected from the gasoline pool.
New components are needed to makeup for the loss of volume and of octane quality in the gasoline pool. Options are isomerates produced from light naphtha ‘. . isomerisation and an increased use of oxygenates or other high-octane components such as alkylate. The combined effect of these changes in composition is that gasoline is expected to become lighter and may cause manufacturing constraints due to a loss of flexibility in gasoline blending by the producer.
The lower summer RVP limit reduces the potential for butane blending in gasoline. ,,, Reduction of (heavy) aromatics in gasoline and the increased use of light isomerates ‘‘, and MTBE also affect the E70 limit. CEN developed new limits for the year 2000 ,, including a slight relaxation of the E70 volatility specification levels for the various
~,. , OGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 37 -. classes - see CONCAWE report 99/51 (proposal for revision of volatility classes in EN228 in light of EU fuels directive, Jan 1999)’.
Therefore, when asking about the role of aromatics in the gasoline, this question would need to be considered in the context of the whole gasoline pool, its average octane requirements, the grades being produced and other limits on composition as well as the refinery type producing the product. With much lower average octane as e.g. in the USA (some 3 points lower), a lower aromatics limit would be feasible but in the EU with its vehicle pare requiring a higher octane level, this issue is a substantial challenge.
When considering potential changes in fuels specifications, it should be realised that the combined specifications aim to sewe aspects such as customer satisfaction, emissions performance and systems durability. In many cases the main emissions improvement is a direct result of new vehicle technology. Such new vehicle technology, as in the case of the catalytic converter introduced in the mid 80’s, may require a new fuel quality (unleaded gasoline) and a new system optimum needs to be established. This octane optimum was investigated and the European standard of ULG-95 was established. Even today the average pool octane in the EU is around 95 RON for the unleaded pool.
Aromatics serve another purpose in the refinery. When producing high octane reformate, the hydrogen by-product is used to supply the ever-increasing refinew hydrogen demand for reducing sulphur in the middle distillate pool.
The current debate on gasoline quality focuses on its sulphur content, driven by the desire to introduce direct injection gasoline engines to reduce C02 emissions. Engine-out NOX emissions of such technology are higher than from conventional gasoline technology and new De-NOX catalytic after-treatment is required. De-NOX catalyst systems are rapidly being developed and are claimed to need very low sulphur fuel. Various technology options are available and it remains to be seen which technology will ultimately be commercialised. Certainly the existing and future fleets do not generally require low sulphur contents as an enabling fuel property that has been specified at 50 ppm for 2005. In case a lower sulphur level is demonstrated as needed to enable a new technology (e.g. de-NOX storage), it should be realised that the current refinery technology available to selectively desulphurise FCC gasoline quickly loses its selectivity at very high sulphur conversion levels, causing olefins conversion and octane loss.
This brings us back to the role of aromatics in gasoline. The essence would be that the gasoline product remains readily producible and that adequate flexibility in product manufacture and blending is maintained. We therefore need to consider always the mix of gasoline grades, the whole of the pool, the potential blending components that can be used and the combined product specifications.
When considering new specifications the debate should be more focused - essentially targeting specifications that have the largest benefits in reducing air toxics - then tighter limits on targeted specifications could more readily be achieved. By selective component blending all refineries could make a limited volume of very low
, 38 sulphur gasoline available - the logistics aspects (cross contamination in transport by tanker and pipeline) and targeting the enabled new technology fleet would then be t issues to resolve.
In diesel there is no total aromatics specification in the EU diesel specification EN- ,7, : 590. The aromatics content varies in the range of some 30-35% and is expected to reduce as a result of the changes in the year 2000 specifications, lowering sulphur ,’, . and density, limiting poly-aromatics content and increasing cetane number. This last “, specification may largely be met by increased cetane improver additive.
Some suggest that total aromatics in diesel should be reduced mainly with an aim to ,, reduce COZemissions from diesel engine-vehicles, however the extra COZ emitted in ,, ., producing hydrogen, required to reduce aromatics in diesel fuel, outstrips the C02 gain achieved in vehicles (basic thermodynamics).
,.,,. EU legislation related to air quality Many-EU legislation initiatives aim to improve air quality for health and to improve environmental protection. The main programmed are in Air Quality Framework ,-,. .., Directive and the Ozone and Acidification Strategies. The implementation measures ---- include setting of targets on Air Quality and limits on acid deposition to protect ,.. ecosystems (National Emission Ceilings Directive) and involve measures such as the ., Auto-Oil Directives for fuels and vehicles, Sulphur in Liquid Fuels Directive and the Large Combustion Plant Directive.
The aim of the current Auto-Oil 2 (following on to AO-I programme) is to identify hither measures that may be required to meet AQ limits in the EU. The programme includes an update of air quality modelling of a number of cities using updated emissions inventories. When considering further measures all potential options should be reviewed. Attributes of good solutions should be that they contribute to solve the stated problem and do so in an EU optimised way.
[n case the remaining problems are mainly city-centre related, these may well be best handled by local measures and city centre related solutions. These may include ,“ .. such options as PM traps on city buses and other city bound vehicles, special fuels ,, for targeted fleets and measures to improve or reduce congestion in the affected areas. .’ ,. By 2010, the emissions contribution from combustion in mobile sources has become a small part of the remaining emissions inventory after the AO-I measures. ,,
,,, . Conclusions ,.. Fuel specifications need periodic review to ensure that fuels and developing engine technologies are properly matched. The examples are numerous, when catalytic ., converters for gasoline vehicles where introduced lead-free gasoline was introduced and the sulphur content of diesel was reduced to enable Euro-2 HD engines. Equally when newer systems with much tighter engineering tolerances were introduced the
39 ... :..,,, companies developed new fuel and lubricant additives that made the new technologies possible. in further developments the link between technology needs and fuel development needs to remain strong, without jeopardizing other aims of society, which include having a reliable supply of low cost fuels.
role of aromatics in gasoline
o role in context of how to produce gasoline o product quality considerations Q customer satisfaction 0 environmental performance o systems durability
aromatics in gasoline m..,”, o octane - main performance parameter o EU octane pool -95 RON clear o US octane pool -92 RON clear (R+MY2=87 o high octane molecules O traditional: aromatics, oletlns, alkylate, C4 added : oxygenates and isomerates o aromates average content reducing from 38% to 33% by 2005
40 !;’,
, .“,}. ,,, ., ,, :., ,,, , ,-. , ,,. ‘ .,, ;....,: ,, : , .:...- m :,,,.:...... ,. gasoline env. specifications !,- .,- ‘:, ... ,,, ;,,,,, . ,., .,,’ o composition effects of directive 98/70/EC <, -.. .:. .,’-. :. 0 reduced summer RVP (C4) .~,,...... ,. >.,:,;>, ,..,,,,... ,, ,,,..,,, 0 reduced suIphur .,- ,’, .,:,-, ,,, CC gasoline- selectivesuIphur removal ,,. , ...’.. with lowolefins(octane)loss o limits on aromatics, benzene, olefins .,,.,,, ., octane replacement needs ,. ,4 ‘, :.., ‘, ..,,; ..’, o producibility - blending flexibility ,’...... ’.’ o reformer as hvdroaen source
AO-1 gasoline/diesel fuel specifications
--- .Ga~ONne. ..-— ...... EN22a. eanofxc 9~n0/Ec ,,, , current 2000 2005 RVP Summer : kPa 70 60 60 EIOO mh _____ :. % vtv.. .40 . 46. 46 E15-o.M@ _... . -.4 Y.vl’ _ ..-. _75 75.. oloflns max , %Vlv - 18 i8 aram.atlcs msg.. i, % vh - 42 35 benzeno max .- .:.; % V&, .5 1-1. oxygenates m/m % Vk 2.5 2.7 2.7 sulphu_r max !t ppmm 500 150 50 le.a~ Fax _.. :..ln!r&.& 0013.-.– . . o.ao5.. _o.oo5 .,– - ‘–Dlhel” “–- ! , EN500 ‘WKMEC 98i70/iC currant 20oO ~ 2005 ,-.., ,., . cetzno nr. mln . - 49 51 51 .-. ~,’,’ dop?!l.. ___ ._. U,. kolms .820~60 845 845 ,. ,,, .,.,. .“ T95 j- d.gr.C 370 360 360 ,,, ” poly-zrom. max_ ~ % m/m ~ - 11 11 ,/; ” . ,, -, sulph.ur max .- [ppm mlm “500 350 :. 50 ,..,-,, AO.2b “mlv, .coIu@B,n.”urj WWS2S0S .{ ’+..! lotus 7/$+ ,~,. , Y, ,’, h,.’. ,,, ,, ,,, ,..; ,,/. :,. ,/’, ? “,{ .,”, ,, ,.,1 :,..:. ,. .’, .’. ,: .’: ’.:, ~,;’” ,, ,, :,,, ,;,:.,- ~ ?“.,,,~,, ‘ :, .,,. , :;, , .. 41 ,’ ~,.,,, , :! ,, ,’ : “.. .’. ,, .,, ,.,; ,., European Air Quality Related Initiatives
ity Framework Ackliyicutka)Jrufzgy and Daugkters ObjectivestOprotect ‘0 ObjectivestoProtect Ecosymem ulplzurIn “ id Fuds Human Health msdPk. Pro,d UpdateoJLCPD ,.
m rw- PrOtmh RAm IPP Irective Progressiveimplementation ofBestAvailableTechniques OnindustrialPlant C3WATSR Salk
Recent legislative developments
o SEVESO-2 o Auto/oil-l, AO-2 o Air Quality Framework Directive (AQFD) o S02, NO,, PM, Pb, CO and Bz 0 Acidification Strategy 0 Sulphur in Liquid Fuels (SLFD) 0 Revised LCPD - Large Combustion installations 0 UN-ECE Protocols (SOX, Nox, VOC, NH3, POPS) 0 Integrated Pollution Prevention (IPPC) 0 lPCC-Kyoto... 0 Ozone Strategy + directive proposal 0 NEC - National Emission Ceilings
42 Auto-Oil II Emissions Basecase : Commission Data ‘aO”w, Bcnzme
0,
commission Data: Auto-Oil 11Emissions Base case V5 , ., Kawe P*rum! *l. -,*< 10mkrms ,. ,,
,. “,’ ,,.,. ,; ,, ,’),,“’.> , ,,.,. , . ..’, ,,., ,., .> ,., :,.. ,,, ; ,: !;,.. , ,’ . , ,.,.,~“ ,:”,,.! -,,,;’,,; : ,, . . . .~’”” “. .,, , ,, .,’ ,. ,;,: ,. 1..,.,,’,,; ‘!, ,...,, >;, ,. ..:, ,, ,,,- (., /, :,( , ,.,-, .; .,, ,’,. ,{ $,. , .;’
,, .’ ~,. 43 -. Auto-Oil II Emissions Basecase : Commission Data CO.c,lve
Nitmgcn oxides
‘-~
conclusions
o fuel qualities need periodic review 0 remain in step with engine technology
o firm link to demonstrated technology o need focus in quality changes
o solve the problem and
o find EU optimum solution
44 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemisfry”, Erlangen 1999
:, D. Huang 1),E. Kohler 2), ‘) United Catalysts inc., Louisville, Kentucky, USA 2)Stid-Chemie AG, Munich, Germany
Sulphur Tolerant Dearomatization of Middle Distillates - A Cost Effective Way of Meeting 2005 Specifications ,,
Two major trends are going to change the refine~ wortd to an extent that has never been observed at ,’ :“. any point in the refining history. First, all over the world, the demand for diesel fuel is steadily on the
increase while markets for residual fuel oil are dramatically shrinking. The imbalance caused by this
trend towards higher quality products can only be compensated for by process technologies which ,. provide upgrading capabilities for low value hydrocarbon fractions. ., Second, tighter specifications on middle distillates will impose further burdens on refiners in terms of
considerable amounts of capital investments. It has become apparent that catalysts and proc&sses
which are currently employed in diesel production will not be able to meet the European 2005 specifications, especially in terms of sulphur content (50 ppm) and polyaromatics (1 wt%), not to
,., , mention the proposed high cetane numbers. FurUrermore, density and end point reduction will call for a ,.’, ,. ,“ shift towards the use of lighter fractions. Alternative feedstreams need to be made available to make ,. ,, ,’,‘ ,;:. ;f) up for the lower boiling fractions going into diesel fuel. First candidate is LCO from FCCU operation. In ~. .,,,’ “,1 ‘ ;,..,,:: ,, :;,: view of the decreased demand for fuel oils where LCO is normally going to, refinera will have to blend ., .’!, ,,,, j ,,1 .,, ,.,, , more LCO into their diesel. However, blendstocks containing large amounts of aromatics such as LCO .. ,., .,.,, . . . . ,,: ,.,’ and visbreaker gas oil are distinguished by a low cetane number, requiring post-processing. , . . ‘(.,,.,,,, .,, ,, ,i The S~D-CHEMIE Group has recently introduced a new highly sulphur tolerant dearomatization ,.’ ,. ,, ,, catalyst (ASATm) providing outstanding capabilities for LCO upgrading. .
1. ASATm - A novel Sulfur tolerant Dearomatisation CataIyst
Efficient low temperature dearomatization of middle distillates is normally accomplished with noble .,
metal catalysts rather than the nickel catalysts which suffer from the affinity to sulfur leading to their
subsequent deactiva~on. A few yeare ago, new catalyst systems with dual noble metal function
emerged in the market which were supposed to be tolerant to much higher levels of sulfur than
standard platinum catalysts. However, none of the units that have been designed to cope with
Californian or Scandinavian diesel specifications is being operated under real sulfur rich conditions.
Thus, the proof is still outstanding and it can be assumed that existing catalysts cannot handle sulfur
levels of about 500 ppm in continuous operation. This means that these catalysts can only be used in .,
45 DGMK-Tagungsbericht 9903, lSBN 3-931850-59-5 ,, <,. .- .,, .,’, grass roots plants which provide pre-treatment by desulfurization down to the 50 ppm level m a first stage reactor.
Therefore, in view of the upcoming specifications and in consideration of the shortcomings of today’s state-of-the-art catalysts, the SUD-CHEMIE Group has recently introduced a new ASATm catalyst (see Tab/e 7 for properties) which is a highly sulfur tolerant dearomatization catalyst providing outstanding capabilities for LCO upgrading [1, 2]. ASATm offers a totally new concept of hydroprocessing using the existing hydrotreaters as a first stage unit without modifications and processing “~tanrjard diesel”, in ~ !300 ppm sulfur as presently produced by MOSt rdiIW&, in a second stage dearomatization (ASAT’”) reactor. This novel concept of a tri-functional catalyst is designed to
. remove polyaromatics completely and reduce total aromatics down to a few percent level
. simultaneously bring the sulfur levels to below 50 ppm
. fully convert the organic nitrogen compounds.
Table 1: ASAP Physical and Chemical Properties
> Carrier Zeolite + Binder
> Noble Metals Proprietary
} Size& Shape 1/16” Extrusions
b CBD 0,5 kgll
From the literature it is known that the addition of palladium to a platinum on alumina catalyst could
improve its sulfur tolerance when used for the hydrogenation of aromatics in diesel [3]. The sulfur tolerance of this type of catalyst, however, could not meet the requirements of the present stringent environmental regulations. Newer technologies [4,5,6] suggest the use of zeolites as catalyst supports to further enhance catalyst’s sulfur tolerance. The tri-functional ASATm catalyst is a hi-metallic noble
metal catalyst supported on zeolite. Its performance benefited from years of efforts in research and development exceeds other catalysts currently on the market. Research efforts have been made concerning the structural effects of alloying and that of the acidic support which relate in part to the catalyst’s sulfur tolerance [7].
46 2. ASATm Performance Features
ASATm has been extensively tested in the pilot plant by S~D-CHEMIE as well as by clients who presently investigate the use of ASATm for their specific processes. The focus of the test work was on
the simulation of a second stage dearomafization reactor as an add-on to the existing hydrotreater
which already produces diesel of <500 ppm sulfur by current specifications. Pra-treated LCO feed of
about 400 ppm sulfur, 43 W % total aromatics and 15 wtYo polyaromatics (for details see Tab/e 3) was
used. Very consistent date have been obtained throughout the test runs which confirm that the
development targets have been met in every aspect. In the following sections the most striking
features of this catalyst are presented and discussed in tight of the future needs.
2.1 Effect of Tempemture on Aromatics Reduction and HDWHDN Activity
Due to its tremendous hydrogenation power, treatment with ASATW can result in a dramatic reduction ,, of total aromatics to a level very close to equitibnum which is timited by temperature. Fig la ,.. ,,,, . . ,’, ., demonstrates that a reduction from 42,5 wt% total aromatics down to 3,3 wt% can be accomplished at .-,:, . ‘,,,,,’ a moderate temperature (327 ‘C). Polyaromatics have vanished completely and are no longer . . ., ,,.,. : ,,., .,:. , ., “,.. : ‘:, detectable under those renditions. The catalyst responds steeply on a marginal increase in ,’,, ,, ,.L, ,, ..,, -’.,, ,.. temperature bringing aromatics and contaminants down to levels far beyond the anticipated ,. .. ”,..:.. “,, , ,; ,, specifications. ,, -! (a) (b)
COncenblban. WI% cmcenmlmll.WvOrm
50 530
~o. . 4C+I ; .— 30 ‘ —Gzzl- 3oo ’- ~,=:- 20. : 2CII
,.. . lCQ ,, -. —...--— , ok n’ .. Teq)efature, “C Feed 307 315 327 .JR Temperature, ‘F Feed 585 604 620
Fig.la and b: Dearomatization (a) and HDWHDN (b) Performance vs Temperature
Conditions: P =62 barg. HJHC = 712 Nm3/m3, LHSV = 1 h’
From Fig la it is obvious that the ASATm catalyst is capable of achieving the 1 wt% polyaromatice
limit according to European 2005 specification at temperatures as low as 307 ‘C (565 “F). This
condition corresponds to 12,5 wt% total aromatics in the product. The catalyst’s ability to hydrogenate
difficult-to-treat-sulfur as well as nitrogen compounds is shown in Fig 7b. The desulfurization .... performance is well balanced with the catalyst’s dearomatization capabilities. The temperature of 307 .>
‘.’ :, . . ,., .,. 47
,., ——.. ‘, “C where future PNA specs of 1 wt% are achieved is exactly the temperature where about 50 ppm of
total sulfur are accomplished. A slightly higher temperature would result in 25 ppm sulfur in the product
which is well within the 2005 specifications for diesel, and temperatures above 315 “C provide
extremely deep HDS to <10 ppm of sulfur in the product. The absence of nitrogen mmpounds at all
conditions is an excellent basis for colour stability.
The product from processing with ASATm has the distinguished properties of a valuable blending
stock. Table 3 summarizes the operating conditions applied to Light Cycle Oil (LCO) and Light Gas Oil
(LGO) as well as the feed and product properties. These two feeds are typical of those frequently
encountered in refineries and would most likely subject to dearomatization. Although ASATm needs
only moderately severe conditions in terms of temperature, pressure and hydrogen supply to get down
to zero level in polyaromatics with the highly aromatic LCO feed, LGO on the other hand can be
effectively treated at temperatures down to 290 “C and pressure as low as 45 barg.
Table 3: Operating Conditions and FeeoYProduct Properties
LCO Case LGO Case
Oneratina Conditions Unit
Pressure [barg] 62 45
Temperature rc] 315 268
H2/011 [Nm3/m3] 712 535
LHSV [hr-1] 1.0 1.2
Properties Unit Feed Prod Feed Prod
LGravity ~APl] 31.6 40.6 39.6 Total Aromati= [Vol%] 42.5 2.5 26 n.d
Polyaromatics [Vol%l 15.1 0 2.4 0
sulfur [ppml 400 9 260 6
Nmogen [ppml 127 0 8 0 Cetane Index [D-1737] 41 51
T95 rc] 359 346 361 356
Gas Make [w%] 02 04
L!quid Yield [,01%] 100 105.5 100 1032
A tremendous improvement is achieved for LCO with regard to gravity (+9 “API) and Cetane Index
(+10). Both can be attributed to the substantial removal of aromatic compounds of low API gravity and
low cetane. Sulfur is below the toughest diesel specifications (< 10 ppm) and nitrogen is totally
converted. A major effect of converting high boiling aromatics, sulphur and nitrogen compounds into
species of lower specific gravity is a shift in the distillation curve. In the LCO case given in Table 3,
ASATW reduces the T95 point by 13 “C by means of hydrogenation of polynuclear aromatics, hence
48 ,
:-, “ — ,,, ,
. ,.’, Product ,., Faed .- w’ F**d y’ I 1 4 / I I I I I ,0, 200 s00 400 100 sw 700 SO’ 2 <00 200 S00 400 S00 600 700 S00 Temperature “F mTemperature “F Fig. 4: Distillation Curves forLGO Fig. 3: Disti//ation Cr.rrvesforLCO .’ ..’ -, :, ’.,’.:
boiling point of di-aromatica from 218 to 186-196 “C and those from tri-aromatics from 340 to 285 “C
[8]. The LGO feed with Iower aromatics and less severe operating conditions results in a less pronounced shiRin T95(50C) andless dvergentboiting cuwesforfeed andpmduct(Fig. 4).
Although thecatalysfs acidic function may have caused ring openings and fhe formation of some
lighter components, there is hardly any gas make, i.e. liquid product loss. There is an appreciable gain
in Iiquid volume yield which is much higher for LCOdueto fhemnveraion of high density aromatics
compounds.
Overall, the results discussed above are clear proofs of ASATm’s supenorily over stata-of-fhe-art
catalysts which can only survive sulfur-upsets” in the range of 500 ppm occurring on a temporary
basis, yet such incidents are devastating to their dearomatiition performance and these catalysta
would need several days to remver the hydrogenation activity [9].
2.2 Design Considerations
The accessibility of hydrogen, as a result of operating pressure and hydrogetioil ratio, and liquid hourly space velocity are the major criteria in designing the catalytic reactors for purposes discussed above.
Ffg5depicS tie ASATm~@lysfs response tochanges in pressure and hydrogen tooilratio ina
setiesof consecutive test inns, allmnducted wititie LCOfeedat LHSVl lf1and315”C/60fYF. The
test resuits also indicate that at a pr~ssure as low as 48 barg, the polyaromatica are still completely
hydrogenated. However, Medesulfutition activi~drops to65ppm so Watasfight increase inH2
partial pressure through the adjustment of hydrogen to oil ratio is recommendable mnsidenng that in ?“,
the range of 50 barg the desulfurization activity depends very much on a favorable hydrogen
environment.
,., 49 ,,, ;.:,.,.,..,.,:.,,, ,4 ., ,, ~).i. ,, ,’.,.; .,,. .’
Concentraka v.&/. Concentration, ppmni
10 100
8 80
6 60
4 40
2 20
0 0 Pressure.barg 62 62 I 55 j 46 62 Pre33ure. barg 62 62 55 48 62 HzMC. Nm21m3 712 535 53s I 535 712 H2MC, Nm3rm3 712 H 535 535 I 535 712
Fig 5: Effect of pressure and hydrogen supply Fig 6: Effect of pressure and hydrogen supply on aromatics reduction on desrdphurization
Investment costs for dearomatization are closely related to the unit size, therefore, the space velocity that the catalyst can work in is of particular importance. Obviously, the allowable space velocity is highly dependent on the hydrogen pressure as well as the hydrogen to oil ratio. Fig. 7 shows the basic behaviour of the ASATTM catalyst in two base cases representing low and medium pressures. It is obvious that PNA’s are not a concern with regard to the proposed 1 $vtY. limit at the medium pressure even at increased space velocities. At the low pressure conditions they are in the low percentage range which would easily meet the less stringent alternative limits as being discussed. Less stringent requirements on total aromatics (15-20 wt%) can also be fulfilled by a low pressure operation or with a unit half in size operating at medium pressure and hydrogen. However, for a 10 wtYo limit on total aromatics according to CARE standards a medium pressure operation will be necessary which would also help to bring the sulfur in the product down to below 50 ppm and completely remove the nitrogen.
., /-+. . r,,.,,,,,..,,,.. , ● Temperature: 280-360 “C d. total aromatics b . Hz ParOal Pressure >45 barg
medium pressure . HJHC ratio: 500-1000 Nm3/m’ & h aro en ● LHSV 1.0 – 2.0 h’ m polyaromatics = —. -----: : -- -- . LHSV
Fig 7: Effect of LHSV on Aromatics in LCO Table 4: Typical ASAP Operafing Conditions (400 ppm S, 127 ppm N, 43/25% Arom/PNA)
Table 4 provides an outline of the ranges of typical operating conditions of ASAT7M based on extended pilot plant testing. These data can serve as a rough guideline for the design of dearomafization reactors with ASAT1~ as the catalyst of choice.
50 It is worth mentioning that ASATm underwent a comprehensive life-test of more than 3000 hours under various operating conditions. Upset conditions were also simulated in this extended test including exposhrg the catalyst to a LCO feed containing 0.5 wt% sulfur which led to a temporary loss of catalyst activity. A simple procedure of hot hydrogen stripping was able to restore the catalyst activity ,. completely. These results provide strong indications that the catalyst will achieve the typical HDS cycle ..’ ,, times.
3. Process Options
Due to the lack of sulfur tolerance of the state-of-the-art catalysts the existing dearomatization technologies are based on multi-stage processing to provide low sulfur feed typically in <50 ppm S to the dearomatization reactor. This means high investment cost both in the case of a grass roots unit and revamps. It creates a hurdle to refiners to adapt their procass potiolio to future requirements. The basic idea behind ASATm is to provide a catalyst for a simple two-stage process whereby the ASATm reactor can be fed with gas oil meeting the present 500 ppm S specification. Thus the addition of a semnd stage HDAr reactor to an existing (conventional) hydrotreater will give product meeting all current and future specifications. This can be accomplished with very inexpensive revamps using existing equipment. The ASATW advantage with comparison to the conventional technology is ., depicted in F/g 8. An interstage clean-up of the process gas to the HDAr reactor is required to achieve ,. ,..,. , optimum dearomatization performance.
,<
G.. 00 Lvend X45mead
L Conventional c.mvmuond Creep Dsep Hos Mm HOA,
0.s 011 81..6 2005Dbsol
sum”<.w mm ‘., , .. ..-> Nwcqm....—-.<5pm_.. ,,, ,, ASAR COn.muonal >.~Tc.+.A,wr,.<10%* ! ,, HDS :: K P PNA.t.Aw. >.S(3<825k#ni > 195 Fig. 8: ASAP vs Conventional Hydmprocessing ASATm can also be used advantageously in multi-reactor systems which incorporate mild hydrocracking and dewaxing, besides hydrodesulfurization. This enables refiners to flexibly adapt their hydroprocessing capabilities to future product specifications and changing feedstock requirements. ASATm is a novel catalyst for hydroprocessing which offers refiners increased options to deal with present and proposed regulatory requirements for the reduction of sulfur and aromatics in diesel fuels. The bimetallic noble metal on zeolite catalyst is an extremely sulfur tolerant, tri-functional catalyst which not only brings a substantial reduction in the product aromatics, but also can achieve a deep desulfurization and complete nitrogen removal to meet the toughest environmental regulations. The application of ASATTMto LCO upgrading manifests the following performance features Z Aromatics conversion close to equilibrium } Increase of cetane Index +10 } Reduction of boiling point >10 ‘C P Considerable gain in liquid product volume Z Low gas-make ‘P Reduction of sulfur from up to 500 ppm to a few ppm E Reduction of nitrogen from up to 200 ppm to zero ppm Current catalyst options require low sulfur content feedstocks with deep desulfurization in the first stage. ASATm provides a simple, cost effective solution as it incorporates the existing HDS reactors without any modification as the tlst stage. This saves the installation of additional HDS capacity or even investment into a grass-roots unit. ASATm not only offers an economically viable solution but 4, provides a flexible tool to refiners to meet all future specifications on middle distillates. References [1] J.C Fletcher, O. Hopkins, D. Huang, E. Kbhler, F. Tungate, NPRA Spring Meeting, 1999 [2] O Huang, E KOhler, PTQ. 3 (4) Winter 1998/99.53 [3] S. M. Kovach and G O. WJson, US Patent 4,049,576 Ashland 011, Sept. 20,1977. [4] S G. Kukes, F. T. Clark and D. Hopkins, US Patent 5,494,870 Amom CorJI., Feb. 27, 1996 [5] B H.C. Winqutst, B 0, Murray, S N. MJam, R.C. Ryan, T.W. Hasbngs, US Patent 5,391,291 Shell 011. Feb. 21, 1995. [6] S G Kukes, F.T. Clark, P D. Hopkins, International Patent WO 9419429 Al Amoco Corp.. Sept 1, 1994 @l D.C. Huang, W.F. Huang, Y. Miao, K.J. Huang and C.F. Li, to be published [8] J. Langston, L Allen, D. Oaw+. PTQ Summer 1999,65 [91 O.R. Lawrence, presented al the JPI Conference, OcJ 2-3, 1996, Toky, Japan. 52 ‘., .,:., ‘:’ ,:.,,,‘, ,:,,.., ,,J DGMK-Conference ‘The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 ,,, 0. Genis ‘), S. G. Simpson ‘), D. W. Penner 2),R. Gautam 2),B. K. Glover 2, 1)UOP Umited, Guildford, England z)UOP LLC, Des P[aines, USA . . The Roles of Aromatics and Catalytic Reforming in the 2000+ Refinery ,, ,. *.,:, ‘. ,1 ,’ .,.-., ;: ; :,y ./ ‘, ..,, Upcoming European environmental legislation will restrict the concentration of benzene and , /. ., ,., ,J ~...’..,, ,,, ,,,, other aromatics in gasoliie. Because catalytic reforming units are the major source of r ,. benzene and aromatics in gasoline as well as refinery hydrogen, the new firel regulations have . ‘,,’- , ‘ , ‘i ,,>.,. .,, (, the potential to significantly affect refinery operation. The marmer in which the new benzene .’ ,’,. :.$ ... ,, ,, ;:,, and aromatics limits are met will be critical to refinery profitability and may change the :, .:,< ..; . availability of refinery-derived aromatics for petrochemical applications. ,.,,,;,’.,,,:, :,, ;,(,.,.’ ,.,, ,., .,, .,..... ,., \,,.‘-1, This paper analyzes the refting and petrochemicrd market factors and technology options ,, ,’./ ~.,,; that will intluence the way refiners and petrochemical producem respond to the new fuels ., ..::0 :.[ ,. ,’ .~ ::.- “:+;,’ legislation. Shortages of benzene, toluene, and xylenes (B’IX) in Errrope for petrochemical ,, ,:..,, ‘,.,, ~ production will provide refiners with an outlet for excess gasoline benzene and aromatics and ‘, .,,;,.,.,, ~ ,.,,-?~,t.< ,;-, ~:, ,,. allow them to improve refinery profitability. .. . . ,’! : ‘,~1:. ,$ J,, 4 ,, ,’ !., ‘, ,,,4 ‘:,,;!”;, ,,. Instead of turning down catalytic reforming units to meet gasoline aromatics limits, refiners ,:y,’, ,::; can increase BTX production by using the proper feedstock and reforming catalysts, thereby ,’. , ,: ,,,~.1’ ~’ improving refinery margins by helping to satis@ the petrochemical market demand for ,, ,, J:’<,;’ ‘,:;:,{,: aromatics. Catalytic reforming will continue to be an important process unit in the 2000+ . .,,’,’, ,~, +, . ,,..,.... ,.,.,.-,,.:, refinery not only for gasoliie production but also for further linking of refinery and ~.., ,.. :f,;:,i.’ .,‘:., petrochemical operations. : .,i,:;,? ;.,>, ~ >..,,:, ., - “,”:j~,., , ‘1 .,, ., ,: 4,$, 4 ;.. ., .,, ,. Legislated changes in European gasoline specifications impose a need for a structural review ..,.‘..’i/,,,., ;, ,! ,: -,;,.,,!,I,,,,.7,- , of the European refting industry. Gasoline production capacity is under-utilized, and a ! !,: -“:, .)..j . ‘. ~,,. $,t,,y; J : , significant increase in gasoline demand is not projected. As a result of new European ,,. , ;,,.,.,/ ;..,.,, ,’;,’.; legislation restricting benzene and aromatics in gasoline, European refiners will be obliged to ,,, , ‘;;;,,;: .,:;.; modi~ their gasoline products. ,.. ,. :,. ,’. , Recent analyses of the European refining sector contirrn that average profitabili~ is low and that the typical return on capital employed is around 4%.]2 However, the highest performing “companies are able to achieve a return on capital approaching 20Y0.Refiners achieving the best performance are controlling costs effectively, differentiating their products, and activeIy searching for diversification opportunities. DGMK-TagungsbericAt 9903, ISBN 3-931850-59-5 Market analysis shows that Europe will be short of aromatics over the next 10 years and will need to import these materials. Refiners have the opportunity to use their surplus refinery- based reforming capacity to satisfy petrochemical aromatics demand at a low marginal cost. However, the refining and petrochemical industries have, with some notable exceptions, tended to function in isolation from each other. The shortage of aromatics for petrochemical use coupled with the surplus of aromatics from gasoline production provides an opportunity to exploit this synergy for mutual benefit. INJROPEAN REFINING MARKET ASSESSMENT New specifications for both gasoline and diesel go into effect in Europe in 2000, and will be followed by more-stringent specifications in 2005. Table 1 compares the new gasoline specifications for 2000 and 2005 with the current EN 228 specifications. The most significant changes are a substantial reduction in sulfur content and restrictions on benzene and aromatics. In 2000, the level of benzene will be limited to less than 1 VOl-O/Owhile total aromatics will be limited to Iess than 42 voI-OA.In 2005, the maximum aromatics content will be reduced to less than 35 vol-%. In addition, sulfur will be reduced in two steps, first to less than 150 ppm in 2000 and then to less than 50 ppm in 2005. Table 1. European Gasoline Specifications I Current New New EN 228 2000 2005 Sulfur, rmm 500 max. 150 max. I 50 max. Oxygen, wt-% Varies 2.7 max. TBD Benzene, VOI-VO 5.0 max. 1.0 max. TBD Aromatics, VO1-’YO NIA 42.0 max. 35.0 max. ,’ . Oletins, vol-~. NIA 18.0 max. TBD #. ,. N/A: Not applicable TBD To be decided ., Catalytic reforming together with fluid catalytic cracking (FCC) provide the backbone of gasoline manufacture. In addition, catalytic reforming provides the key link between refining and aromatics production. The catalytic reforming process, first commercialized by UOP in 1949, increases the octane of heavy naphtha through the selective conversion of paraffins and naphthenes to aromatics over a platinum catalyst. The high concentration of aromatics in the reformate product is responsible for its high octane rating and also makes reforrnate valuable for the production of benzene, toluene, and xylenes. Hydrogen is a valuable byproduct. Refiner-v Confimrrations The new European fuels legislation imposes difficult constraints for the three basic types of refineries. The most common configuration is the FCC-based crackhg refinery, which accounts for about 83°/0of installed European gasoline production capacity. The other two configurations me hydrosklmming and hydrocracking refineries that account for about 100/o and 70/. of European gasoline production capacity respectively. Limits onsulfur andaromatics andtheban onleadin Europe afier2000 will make it more difficult to produce gasoline. Hydroskirnming and hydrocracking refineries will typically not 54 be able to meet the 1 VO1-’YObenzene and 35 vol-% total aromatics specifications mandated from 2005, although FCC based refineries will be close to compliance. Estimates of the average research and motor octane of unleaded gasoline pools indicate that a typical FCC. based refinery will struggle to meet the octane requirements for unleaded premium gasoline (95 research octane), and will have little hope of producing an unleaded super-gmde (98 “ ,., . research octane). Hydroskimming and hydrocracking refineries will have only a little more ,. flexibility in meeting gasoline octane requirements. Estimates of gasoline pool characteristics ., for the three types of European refineries are shown in Table 2. ., ,.’ Table 2. Estimated Gasoline Pool Characteristics of European Refineries ., Refinery Contigurat n FCC I Hydroskimming Hydrocracking No. of Refineries 69 39 11 Capacity, MMTA Crude distillation 550 130 74 Gasoline production 140 18 12 Reformate in gasoline, vol-% 38% 77% 83% Gasoline pool characteristics: Benzene, VOI-% - l% - 4% - 4% :<, ,,,,, Aromatics, VOI-YO -35% -sly. -54~o ,,, , RONC 94 95 96 ,.. MONC 85 88 88 :,,.,, ,, The projected European gasoline production capacity is estimated at 170 million tons per , year. However, European gasoline demand is only 135 million tons annually, or 80% of the overall production capacity.3 If not for a doubling of the net gasoline exports over the last ,’ decade, utilization of gasoline production capacity would have been less than 80’Yo.Gasoline exports currently account for over 10°/0of European gasoline production. The primary reason for the excess gasoline capacity is that the refining industry invested heavily in production in the early 1990s anticipating a growth in demand. However, gasoline demand has not increased in Europe in recent years due to improved engine efficiencies and private motorists switching to diesel-powered automobiles. To meet the growing demand for ;{; diesel, refiners have increased crude runs which also increased gasoline production. The ;,, ., .’, resulting over-supply of gasoline is one reason for the low average profitability in the refining .-. , ,.,. ,. sector. Because gasoline demand in Europe is expected to remain stagnant during the coming . .. . . decade, refiners are reluctant to make major capital investments to meet the new gasoline “.. ,, ,’ .- ,- specifications. .’, .’. , .,’,.,.:.;>. . ..,. ,:. ,- Rcfincrv Benzene and Aromatics YicId ,, .,. ,,, ., !.. ~ .,..,- c - ,. TabIe 3 shows the quantity of benzene and aromatics produced in Europe for each refinery ,.,. -, configuration and the amount that must be removed to meet the legislated specifications. .“.’. .’ “c-, : .,, ., European refiners will be obliged to reduce the benzene content of their gasoline product by ,, ,, ...... , - ‘ ,. over 0.8 million tons per year by 2005. In addition, up to 4.5 million tons per year of other .!” ..~+ , aromatics will need to be removed from the gasoline pool. About 40°/0of tbk will be toluene ,!:.: ‘,:!’ ...... and 50% will be mixed-xylenes. .“’ .’ (’” :’ ., .;,’ ,; :,, 55 Table 3. Potential Excess Benzene and Aromatics in European Gasoline by 2005 Refinery Contribution Total FCC Hyclroskimmirrg Hydrocracking Gasoline source, % of total 83 10 7 100 Gasoline demand, MMTA 112 13 10 135 Benzene production, MMTA ,’ . In gasoline 1.24 0.51 0.42 2.17 1. ,. Excess over specification 0.12 0.38 0.32 0.82 Aromatics production, MMTA In gasoline 39.76 6.58 5.40 51.74 ., Excess over specification 0.56 2.03 1.90 4.49 EUROPEAN AROMATICS MARKET ANALYSIS Reforming has played a significant role in the evolution of the aromatic based petrochemical industry. h economical route to high volume and low cost BTX emerged in the 1950s when catalytic reforming was integrated with efficient aromatic extraction technologies. Since then, steady growth of aromatic petrochemicals and their derivatives has increased the need for BTX-rich feedstock. While little or no growth is expected in European gasoline demand, the demand for aromatic petrochemicals continues to increase. The changing balance between petrochemical and gasoline demand for aromatics requires strategies that will economically satisfi both markets. Over 55% of Europe’s BTX supply is based on reformate. Toluene and xylenes production are dominated by reformate feedstocks (TO~o, 79°Arespectively). Reforming has proven to be the lowest cost technology and will continue to be the dominant source of feedstock for the downstream xylenes markets. New developments in reforming and extraction for more selective BTX operation together with adjustments in reformer feed composition can significantly enhance aromatic yields reIative to conventional, gasoline-based reformer operation. - The driving force behind increased mixed xylene output is the 6’% to EW. amual worldwide growth in p-xylene demand. Consumption of p-xylene in Europe in 1998 was 1.S million tons. This consumption is projected to reach 3.5 million tons per year by 2010. The primary reasons for thk growth are increased consumption of polyester bottle resins in food and beverage packaging and continued penetration of polyester fiber into the apparel market. European producers have not overbuilt capacity to the same extent as in Asia. As a result, Europe continues to be a net importer of both p-xylene and mixed xylenes. Based on known and anticipated capacity announcements, Europe is expected to continue to be short of xylenes into the foreseeable future. . 56 Benzene The supply situation for benzene is more complicated than the supply for xylenes because ,’ benzene is a by-product from a number of processes: gasoline, p-xylene, coke, and ethylene production. Because Europe produces 70% of its ethylene from naphth~ the major source of ., benzene in Europe is from pyrolysis gasoline, a by-product of ethylene crackers. The next ~‘’ major source of benzene in Enrope is tlom catalytic reforming. In 1998, about 35% of ,, benzene originated from reformaie feedstocks. A third source of benzene is from on-purpose benzene production. Historically, on-propose benzene was produced by toluene dealkylation and more recently horn tohrene disproportionation. The latest on-purpose benzene technology is a variation of reforming using a zeolite catalyst that is highly selective to ,,, benzene and toluene. ,“ ,., Two trends in benzene production that influence aromatics phmning have emerged recently. ., l%s~ the growth rate in benzene demand is lower, rdthough benzene demand is significantly higher thanp-xylene demand (7.7 versus 1.7 million tons per year in 1998). Over the next ten years, p-xylene consumption in Europe is projected to nearly double while benzene consumption is expected to grow by about 50Y0.On a global basis, the growth in p-xylene and ethylene demand should generate enough co-product benzene to meet benzene demand. However, Europe will need additional benzene, either from on-purpose benzene production or from increased imports. The second trend is increased benzene production through catalytic reforming instead of the displacement of benzene by dealkylation, which is the highest cost route and therefore the ., price setting mechanism for benzene production. Because catalytic reforming is a ‘..,. . significantly lower cost source of benzene, the long-term trend will be to make more benzene ,, ‘, via reforming rather than dealkylation. Mcctinp Future Aromatics Demand The balance between importing BTX and making it in Europe will depend on the willingness of European refiners to improve profitability by capturing the value of BTX from refinery catalytic reforming units. Although in most cases additional investment will be required to remove BTX from reformate, this investment will provide refiners with added margin from chemicals that have healthy long-term growth prospects. Given the cyclic nature of the petrochemical market and the amount of time needed for project implementation, a technology plan and decision must be made by refiners witidn the next one to two years to capture the full potentird from the next wave in petrochemical demand. OPTIONS FOR MEETING GASOLINE BENZENE AND AROMATICS Lmms ,,., .-.,. , The need to reduce benzene and aromatics in gasoline together with the demand for more :,’, ,,,” BTX aromatics for petrochemicals enhances the opportunity to integrate fiel refineries and petrochemical complexes. Redirecting aromatic components from the gasoline pool to petrochemical facilities mitigates the negative economic effects of gasoline specification changes on fiel refineries. By maximizing capacity utilization of existing reformers, refiners will be able to provide the petrochemical industry with vaIuable interrnedlates. Refinerv Benzene and Aromatics Mmuwzcment Case Studv The effects of the 2005 benzene and aromatics specifications on the refinery economics are quantitatively compared using these three cases . Case 1: Current Refinery Operation ● Case 2: 2005 Refinery Operation—self-contained ● Case 2: 2005 Refinery Operation—petrochemicals integration Case 1: Curren[ Refinerv Operation The current FCC-based refinery configuration, the main producers of gasoline in Europe, is the basis for this case study. An FCC-based refinery (Figure 1) with a capacity of 180,000 BPSD (8.6 million tons per year) processing a mix of Arabian Gulf, Brent, and Urals crude oils has been modeled to illustrate the changing roles of aromatics and catalytic reforminm The reformer is assumed to be a first generation UOPTMCCR PlatformingThi unit with an a~erage reactor pressure of 8.5 kg/cm2g. Figure 1. Current FCC-based Cracking Refinery Naphtha fi.fTPn P ... . ,,,. 1 ~ 1- nsa. Gsoline Brent T TKerosene 1~” “ Arab Mix and Jet 1, Urals Dieseland 1, m I Heating OiI :[1+J“% Fuel Oil ‘@+pti Details of current and target gasoline production are shown in Table 4. The benzene and aromatics levels in the current grades of gasoline exceed the targets for 2005. The diesel product contains a maximum of 500-ppm sulfir in accordance with the current EN590 specification. The reformer generates sufficient hydrogen to supply all the hydrogen consuming units in the refinery. 58 Table 4. FCC-based Refinery: Current and Target Gasoline Production I Premium Super Leaded I Unleaded I Unleaded I Super Current I I 10 o 1.0 I NIA .. Aromatics, VO1-V. I 35.0 I 35.0 NIA ,.,,. , .,. ,,t :: The Case 22005 refinery operation model meets the benzene and aromatics limits by reducing CCR Platforrning unit utilization and installing a new, light-naphtha isomerization unit such as a UOP Penex~-DIH unit. Even with this additional uniL it is still necessary to blend more MTBE into the pool to meet the pool octane requirements. A flow diagram of the modified refinery is showh in F]gure 2. (Case 3 highlights reformate sale to petrocheticaIs.) Case3: 2005 Refinerv Ocrerafion-rrefrochemical integration In Case 3, the economic projections can be improved by maximizing CCR Platforrning unit utilization. The refinexy product slate is changed as a result of increased utilization of the CCR Platforming unit and selling excess reformate for petrochemical use. However, less heating oil is produced in Case 3 because a heavier feed is processed in the CCR Platfonning unit and jet fuel and diesel production are maintained. As a result of selling reformate, gasoline production in Case 3 is 13% lower than the base refinery (Case 1). Figure 2. 2005 FCC-based Cracking Refinery MTf?lI Naphtha A -. —-- 5 ------— - Arab Mix‘mI } Diesel and 1– ] ‘:”..‘:::● Heating Oil .—i I -.. .-e.. — : —-—--—-—.—.- ...... - ● Fuel Oil ... .,’ .. 59 . .- ,.,’ <,-. Case Comrrarison A comparison of these cases demonstrates that it can be economically attractive for refiners to meet the 2005 gasoline specifications by increasing reformer throughput and aromatics production. Refiners that choose this route will need to operate their reforming units as efficiently as possible by taking advantage of the latest process and catalyst developments. Refinery hydrogen production from the Platforming unit is markedly different among the three cases. By maximizing utilization of the Platforming unit to produce reforrnate for petrochemical use, nearly 60V0more hydrogen is produced than in the 2005 self-contained refinery (Case 2), and nearly 20°A more hydrogen than in the base refinery (Case 1). This excess hydrogen is a crucial component in reducing diesel and gasoline sulfur. ,’ A smaller Penex unit is needed in Case 3 to meet the gasoline specifications resulting in a #, ,. reduced investment. The refinery generates a small positive margin over Case 1. This improvement in operating margin provides a mechanism to pay back the investment for reducing both gasoline aromatics and meeting the new diesel and gasoline sulfur specifications. In a sense, the aromatic problem created by the 2005 gasoline specifications can be economically addressed by maximizing production of total aromatics and redirecting some of them to petrochemical applications. The following tables show the relative production rates, gasoline pool properties, capacity utilization, and relative economics of the three cases. The refinery material balance is shown in Table 5. Gasoline properties and gasoline unit capacity are shown in Table 6. Table 5. Refinery Material Balance Case 1 Case 2 Case 3 Current 2005 Refinery 200S Refinery Refinery (self-contained) (petrochemical integration) Refinery capacity, KMTA 8>635 8,635 8,635 Product Sales, KMTA Propylene 104 104 104 LPG 481 475 480 Naphtha 582 635 604 Gasoline Unleaded 95 RON 1802 1929 1856 Unleaded 98 RON 180 193 178 Leaded 98 RON 341 Reformate 479 Jet A-1 420 420 420 Diesel . 1200 1267 1267 Heating oil 1777 1845 1470 Fuel oil 1335 1340 1361 Hydrogen production, KMTA 43.4 31.9 51.2 60 Table 6. Gasoline Production Case 1 Case 2 Case 3 ,. Current 2005 Refinery 2005 Refinc~ Refinery (self-contained) (petrochemical integration) ,, Gasoline Properties (unleaded grades) RONC 95 98 95 98 95 98 ,,, Berrzcne, vol Y. 2.75 2.72 0.58 0.52 0.58 0.57 ,, ,. Aromatics, vol % 41.1 42.0 35.0 35.0 35.0 35.0 . Sulfir, ppm 150 150 50 50 50 50 Oxygen, wt ‘Y. 0.12 1.64 0.15 2.12 0.18 1.89 Gasoline Unit Capacity, KMTA Reforming unit (utiliition) 1079 (85%) 791(62%) 1272 (100%) .’ Isomerisation unit 200 158 ~,. 2303 2303 2303 ,“, ., FCC ., Tbe refinerv–. economics for the three cases are shown in Table 7. For Case 2, the net present value (NW) of the investment (specifically to reduce benzene and total aromatics) is negative. It should be noted that this analysis focuses on investment to meet the benzene and aromatics specifications. Investment for sulfur reduction is not included. Table 7. Refinery Economics Case 1 Case 2 Case 3 ., Current 2005 Refinery 2005 Refinery ,,, ! ,, Refinery (self-contained) (petrochemical integration) ,.,.,,. Cost/Revenue,MM$/yr .’, ,, Product Revenue 1058 1057 1067 Feedstock Purchases . 883 885 885 ,, Gross Margin 175 172 182 Net utility cost 34 33 34 Operating Margin 141 139 148 Economic Indicators Delta investment MM$ Base +21.0 +18.0 Delta Op. Margin, MM$/yr Base -2.0 +7.0 NPV for 20 years @,lO% -31.8 +32.0 Internal Rate of Re-~ % Incrcasin~ Aromatics Yield UOP is continuously working on further developments in Platforming technology to respond to the needs of the refining and petrochemical industries. Excess capacity in catalytic reforming units offers refiners the potential to increase revenue by producing addhional reformate for use as petrochemical feedstock. Modified operations and new developments in Platforming can make tlis an even more attractive prospect. A40dih Feed Properties The first alternative for modified operations is to adjust the cut points of the naphtha feed to improve the BTX content of the ;eformate. An example of this is illustrated in Table 8. The base operation represents a typical gasoline reforming operation in a CCR Platforming unit operating at 8.5 kg/cm2g average reactor pressure and 100 RONC using UOP R- 132Th~ catalyst. The feed is cut to an initial boiling point of 98°C to eliminate most of the CGbenzene precursors. Removing benzene precursors is a typical refining practice for minimizing benzene production in the reformer. The end point of the feed is cut high at 190°C to maximize gasoline production. The second alternative, uses the same naphtha cut at 80°C to include most of the CfI compounds and the endpoint is reduced to 175”C. These changes increase the percentage of BTX precursors in the Platforming unit feed. Along with increased CC!to C8 content in the feed, the reduction in endpoint reduces the catalyst coking rate and enables operation at 101 RONC versus 100 RONC in the base operation. 4 . . The net effect of altering the feed is to increase the yield of BTX horn 35.7 WWO in the base case to 45.1 Wt-O/Oin the modhled feed case. This enhancement of aromatics production can be achieved simply and provides a means for refiners to focus reforming operations toward both aromatics and gasoline production. Table 8. Effect of Reed Boiling Range on Aromatics Production Base Operation, BTX Operation, UOP R-132 UOP R-132 Feedstock cut-point, ‘C 98-190 80-175 Yield, wt-% Hydrogen 2.8 2.8 Benzene 4.0 6.5 Tohrene 13.9 17.4 Cg aromatics 17.8 21.2 Total BTX aromatics 35.7 45.1 62 ,. ,,, “,.,,, - . ..- ,’...... ,, .,..’. , ,.. . . ,’.,<.,.,.:,}.:.. ~’ ,.:. Imnact ofImuroved Catolwts :.,> ,, ,’:7:,:. ..,.<,!.:,’,. ~- In addition to feed cut point adj@rnents, new more selective catalysts provide opportunities >. ’,...- .- ,> to further improve aromatics production. Table 9 compares the yields for BTX operation .,~-,, using UOP R-132 with the recently commercialized UOP R-174~ high yield catalyst. Both operations are at 8.5 kg/cm2g average reactor pressure and 101 RONC. The feed is 80”C to 175°C naphtha BTX feed. In this case, the UOP R-174 catalyst increases BTX yield from 45.1 to 46.8 w-V. while producing 6V0more hydrogen. For higher-seventy operations with leaner feeds, yield benefits could be even higher. No mechanical modifications to the CCR Platforming unit are required to use the R-174 catalyst. Table 9. Catalyst Impact on Aromatics Production UOP R-132 UOP R-174 Feedstock cut-pointj ‘C 80-175 80-175 Yield, wt-% Hydrogen 2.8 3.0 Benzene 6.5 6.6 Toluene 17.4 17.8 Cs aromatics 21.2 22.4 Total BTX flO~tiCS 45.1 46.8 New Directions for Platforming CataIvst Development While current technologies can provide a significant increase in aromatics production, UOP continues to explore avenues for fhrther improvement. Current reforming catrdysts are highly .’ selective for aromatics production in the CS+ range or the C,5to CTrange, but not both. One ... ,/, focus of UOP’S development efforts is on new catalyst systems that exhibit high eonversion efficiency throughout the carbon number range of interest from CGto CS and above. It is anticipated that introduction of this type of improved catalyst system will allow a significant increase in BTX production in Platforming units. The anticipated BTX yield improvement for thk new generation catalyst system is illustrated in Table 10. Table 10. Yield Benefits for Advanced Catalyst System I UOP R-132 I UOP R-174 I Advanced I Catalyst System Feedstockcut-poin~ “C I 80-175 80-175 80-175 Yield, wt-% - Hydrogen 2.8 3.0 3.3 Benzene 6.5 6.6 8.5 Toluene 17.4 17.8 21.5 CXaromatics 21.2 22.4 22.2 T~trd BTX aromatics 45.1 46.8 52.2 The potential benefits, in terms of increased hydrogen and aromatic yield clearly indicate that ,,,,. .,::.: ,, : ,;,.. ~.;.! opportunities remain in the quest for higher BTX yield. Thk and other new imovations will ,, 63 ,. . ensure that the PIatformirrg process continues to be a key technology for gasoline and petrochemical production. CONCLUSIONS Thk paper demonstrates how a simplified but typical refinery can be modified to produce gasoline that meets the European 2005 benzene and aromatics requirements. One possibility for meeting the new specifications may result in more than a 25%’. reduction in reformate for gasoline production, which represents a substantially diminished role for catalytic reforming. The required investment for this option is unlikely to generate a return and, in addition, significantly less hydrogen will be produced. Poor economic projections will make thk choice unattractive for most refiners. Another possibility is an anticipated growth in the demand for benzene and xylenes in petrochemical use to provide an attractive outlet for aromatics produced by catalytic reforming. Refiners who pursue this option can potentially improve asset utilization by operating their reformers at high throughput. Excess reformate can be sold for petrochemical production. The resulting economics are attractive and provide a positive return that can help pay the costs required to meet other firels specifications that go into effect in 2000 and 2005. Gasoline can also be reduced by this option. In addition, substantially more hydrogen is produced as a result of increased reforming unit utilization. This is particularly important since hydrogen will become an increasingly critical feedstock in the future to make lower- sulfur and higher-quality fuels. The role of catalytic reforming in future refineries will change. The process will continue to be important for gasoline production, and its role as a key link for fbrther integration between refining and petrochemicals is Ii.kely to be significantly enhanced. This points the direction for future development in catalytlc reforming technology. REFEIU3NCES 1Roland Berger& Partner, Studv on Oil Retinin~ in the European Communitv, Report prepared for European Commission, DG XVIVB2, page 5, [Dee, 1997]. 2 Wood Mc Kenzie Consultants Ltd., The Future for Eurouean Refining, Muhi Refinery Study, [1999]. 3International Energy Agency, Annual Statistical Smmlement, page 9, [7 August, 1998]. 64 ,, .$ .-, .;!,., DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Ertangen 1999 .,. ,“, J. Cosyns, Q. Debuisschert, B. Didillon, J. L. Ambrosino IFP, Rueil-Malmaison, France Flexible Upgrading of Light Reformate ...,,, ; ,, .,. \;,,!\, ‘~,,’..,:;! ,,.. > !. . --,,f., ,,, Introduction . ‘., ,,, .$ ,.:. Reformate, by far the largest contributor (>80Y0) of benzene to the gasoline pool, is ... . ,<.,,..,; ,, targeted directly by new measures restricting benzene content in gasoline. Various ... reformate upgrading strategies can be employed to meet the new restrictions, one of -: which is to operate or modify the reformer feed fractionator so that benzene precur- .. sors (aliphatic @) are removed frOM the feed to the reformer. The low octane C5-C6 overhead fraction is ideally suited for upgrading either in an isomerization unit to pro- duce high-octane light gasoline or in a steam-cracker to produce petrochemical-value olefins. Adjusting the reformer feed fractionator cut-point may not be sufficient to reduce the benzene content of the gasoline pool below 1Y. volume, the current trend in specifi- cations. in this case, removal of benzene from the reformate product is necessary. ,, :.,!, This can be accomplished through either conversion or extractive distillation tech- ~. .,. :, ,, nologies. Conversion of benzene to cyclohexane via hydrogenation is a possible >“... ,, ‘ ‘i,.~, :, route if overall yields are good and the loss in gasoline pool octane is acceptable. ,. ,: ‘,. ,,,,.. Extractive distillation with purification can be attractive if the refinery is associated ..’./, ,,, ,,., with a petrochemical complex needing benzene; however, this option lowers gasoline production and is often not the best solution for a stand-alone refinery. The options for benzene reduction in refonrrate all require various degrees of frac- tionation and purification. The added investment and operating costs associated with ., the reformate splitter or fractionator must be balanced with individual refinery needs ,. , and requirements, such as gasoline or petrochemicals production, quantity of ben- ! ., zene to be eliminated, added market value of high purity benzene with respect to treatment cost. This article describes two reformate hydrogenation processes that ., are cost-effective in reducing benzene content in the gasoline pool: .,’, ● benzene saturation via reactive distillation: Berrfreem, . benzene or reformate purification for substantial extractive distillation proc- ess improvement Arofir?ing~. Benzene Saturation: Benfree The benzene saturation process (hydrogenation of benzene to cyclohexane) offers two advantages: it reduces the total aromatics content in the gasoline pool and does not require high-purity separation schemes. The apparent drawbacks to benzene saturation are lower hydrogen production and lower reformate octane: benzene (RON/MON = 106/100), cyclohexane (RON/MON = 83177). However, much of the apparent road octane (RON + MON)/2 loss is offset by the gain in product volume . ,, -,’, ,,. . 65 DGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 and the combined effect, octane-barrels, of the process is significantly less pro- nounced as illustrated in the example in Figure 1 where benzene concentration in a reformate has been reduced from its original level to 1 VO170by the Benfree process. Figure 1: Reformate road octane number (Research + Motor)/2 after benzene saturation 102.0 101.0 Product Parameter as Compared 100.0 to Reformate Feed, Per Cent 99.0 98.0 0 2.0 4.0 6.0 8.0 Benzene Originally in Reformate Feed, Volume Per Cent For example, upon treatment by Benfree of a reformate initially containing 5 VOI% benzene, the road octane will be 98.87. of the reformate’s original value. However, because cyclohexane’s density is lower than that of benzene and because hydro- genation adds mass to the product, the volume yield will be 100.9% of the original reformate. The increased volume counterbalances the reduced road octane so that the resulting “octane barrel” (road octane times the volume) is gg.7’~0 of the original reformate value. For reformats having low benzene contents, hydrogenation will often be the most economical means of benzene removal. For higher benzene con- tents, production of high purity benzene by extractive distillation could be an attractive option. :. Aromatics hydrogenation catalysts cannot selectively hydrogenate benzene without t.,. also hydrogenating other aromatic compounds. Therefore, splitting the reformate into ., a lighter fraction containing most of the benzene and a heavier, benzene-depleted fraction is essential. The splitter design will depend on the required degree of the benzene and toluene separation. Typically, if a conventional hydrogenation unit is chosen, the cost of the reformate splitter represents almost 50% of the total invest- ment cost (splitter and hydrogenation unit). Since it could be argued that there is little improvement that can be made in reducing the cost of the splitter, any cost reduction should originate from the hydrogenation unit. With the objective of minimizing investment and operating costs, the Berrfree process was developed and two licenses have recently been awarded. This technol- ogy is based on four key ideas: 1. Installation of a reformate splitter sidedraw (minimal change to the column) and a simple, external fixed-bed reactor that completely hydrogenates benzene in its feed; 2. The reactor operates under essentially stoichiometric hydrogenation con- ditions thus eliminating the need for a hydrogen recycle compresso~ 66 ;:,’ ., ,<,’ .. ‘,’,” ,, . . ,...-’-,.,,!,.. 3. With benzene concentration reduced to low levels, benzene/CT azeotropes ,.: ,..:’; ~., ,,, ,:L ,, are avoided in the splitter ensuring that branched CT alkanes do not enter ,.. . the overhead fraction; >, ‘., .,;,,..’ ,’,, 4. Use of a highly active catalyst operating in the liquid phase With the PO=+ ~ ~ ., -! bility to change out catalysts or isolate the reactor without shutting down . . the-column. Process Description Full range refonnate is fed to the reformate splitter which operates as a dehexanizer, concentrating (% and lower boiling components in the overhead section. Benzene and azeotropic mixtures of benzene and branched CT paraffin isomers are also car- ried into the upper section of the splitter. t Above the feed injection tray, a benzene-rich light fraction is withdrawn from the col- .. umn and pumped to the hydrogenation reactor, see Figure 2. The pump enables the reactor to operate at higher pressure than the column, thus providing increased volu- bility of hydrogen in the feed. Sufficient hydrogen is added to the feed to prevent ~, :, benzene/CT azeotrope formation and to keep the benzene content in the downstream !’ gasoline pool at mandated levels, i.e., below 1.07. vol. for most major markets. The amount of hydrogen is very low, nearly stoichiometric, and essentially all of it is con- sumed. With good feed hydrogen purity, there is practically no gasoline product loss in the column off gas. Benzene conversion to cyclohexane can be easily adjusted to meet changing market demand or regulatory requirements. The reactor effluent, es- sentially benzene free, is re-injected into the column. ,. Figure 2 /3errfree Process Flow Diagram ., Off Gas :-.,., .:. , Cgcg Reformate .,. H2 Light . . . Q- Refo~mate ,’ ,: Heavy m1 I it-l Reformate w ,. The low benzene content in the light fraction above the withdrawal tray disrupts the formation of benzene/iso-CT azeotropes and precludes the presence of CTS in the overhead fraction. This is particularly advantageous when the light reformate is des- ,, tined to be isomerized, because iso-CT paraffins are readily cracked to produce CS ,. and Cq components, thus leading to a loss of gasoline production. ,.. .,’ The f3enfree process is endowed with a highly active catalyst, which is easily acti- ,, vated on site, The high space velocity of the system translates into low catalyst ,, ,.,,,,,,,,,,: ‘., volumes, typically lower than 10 m3 even for high reformer capacity. The unit is de- ,.,, ,.,:),,,, ,....,. -- signed to provide long catalyst run lengths, in excess of three years, which facilitates long periods between refine~ turnarounds. Comparison with Conventional Processes and Catalytic Distillation The Benfree process eliminates several equipment items used by conventional ben- zene hydrogenation processes: feed drum, preheater, separator drum, stripping column and feed pump. The resulting reduction in hydrogenation section investment is approximately 50% with no loss in operating flexibility. The overall investment re- duction for the combined refonnate splitter and hydrogenation section is approximately 250/. with the added advantage that the light fraction contains no iso- CT paraffins. Benfree exhibits several attractive advantages compared to processes where the catalyst is contained within packaging materials that are individually loaded in the distillation column: ● Easy, fast catalyst loading and unloading, ● No need to stop the splitter during reactor shutdown, . No restriction to a single, high-cost specialty catalyst source, ● Reactor design is optimized and independent of the splitter, ● Splitter design is simple and employs standard conditions without a fired heater, ● No excess hydrogen is needed, . No recycle compressor for hydrogen off-gas recycling. Experience. Two Benfree licenses have been recently awarded: both units, repre- senting a combined capacity of 19000 BPSD, are scheduled for start-up in the first quarter of 2000. Five conventional benzene saturation units on light reformate are already in operation. Benzene Purification by Selective Hydrogenation: Arofiningm Olefins and diolefins in reformate. The trend in reforming technology is toward continuous catalyst regeneration (CCR) units. This technology enables operation at low-pressure and high-severity. For thermodynamic reasons, the major reforming reaction, paraffin dehydrocyclization, is enhanced with decreasing pressure. In other words, the lower the pressure, the higher the aromatics and hydrogen yields. At low pressures and in the typical reforming temperature range, dehydrogenation reactions of naphthenes to aromatics are essentially quantitative. However, in addition to these desirable reactions, paraffin and isoparaffin dehydrogenation also occurs resulting in some olefin and diolefin production. Under the most severe conditions (lowest pres- sures, highest temperatures), olefin and diolefin contents increase significantly. The influence of pressure on the olefin and diolefin contents in reformate is exempli- fied in Figures 3 and 4, respectively. The olefin content is measured by the bromine number (BrN). One molecule of bromine reacts for each molecule of either aliphatic or aromatic olefins or diolefins. The olefin content, in grams, is estimated as the number of moles of bromine consumed divided by the average molecular weight of the sample. Diolefins are quantified by measuring the ma/eic anhydride va/ue (MAV) which is expressed in mg of maleic anhydride consumed per gram of reforrnate. Maleic anhydride reacts with conjugated diolefins. 68 .’ <-,> ;, ,., ., ,, ,. -’~...... , . ,$, .,, .-;., Figure 3: Influence of operating pressure on reformate ,’.’ ,,, ., olefins content as measured by the bromine number (BrN) 1 , .’.:,,:,-, ,-, ,’. }, .- . . . . — .— .!.. , ,,! ~, ,,,’.’,- ., —. I ,,,‘, 6 .,, .,,., ,,- ,.,- 5 .,,“, , ,,. . ~~ ; ,. ,.. ., :, Bromine — ,----- ,, 4 ‘,- ,., ,..,., :; , Number, .,, ., :,, ..,, ,,.~,.,, g/loog . . . . 3 .~. ,’ ”.,,,.,.. ‘,; 2 ,, .,...... I I I 1 — .. 0 0 5 10 15 20 ReformerOperating Pressure, bar g. Figure 4 Influence of operating pressure on reformate diolefins content as measured by the maleic anhydride value (MAV) t 1.0 MAV, t, mglg 0.5 Oi ,. o 5 10 15 20 .;, Reformer Operating Pressure, bar g. ,,,,,, Consequences of olefins and diolefins in high severity reformate. When them- ,,.’.,, ,, ,., , ical-grade aromatics are to be produced, diolefins must be completely eliminated and ,, -., olefins must be significantly reduced. The most commonly required specification for the aromatic product is the acid wash co/or (AWC). An example of the sensitivity of .,. the AWC to diolefins and olefins is shown in Figure 5. AWC measurements have ,,-, been performed on a benzene rich cut (exhibiting an AWC = O) after the addition of small amounts of cyclohexadiene or increasing amounts of cyclohexene. An immedi- ate observation is the significant difference beb.veen the two responses cyclohexadiene is two orders of magnitude more reactive to the AWC measurement than cyclohexene. 69 ,’ ... ., Figure 5: Acid wash color of a benzene-rich C6 cut I ‘ PA I . I 0 ‘“ 2 I ““-:: 1’ 0 )f_.._l100 Y 2:yc’”&&’”’& o 2500 5000 7500 10000 Cyclohexene, ppm The cyclohexadiene content must be reduced to less than 5 ppm to meet an AWC value of 1. For cyclohexene, this specification is reached when its concentration is lower than about 300 ppm. The AWC response to branched olefins is significantly lower than the corresponding cyclic or linear olefins as shown in Table 1. Thus, very low AWC values of high severity of reforrnate will require complete removal of diole- fins, a high removal of linear and cyclo-olefins and a partial removal of branched olefins. Table 1: AWC responses to various CGolefins Olefins Olefin content giving an AWC of 1, ppm Cyclohexene 300 n-Hexenes 300-400 Methylpentenes 2500-3000 Attaining chemical grade purity aromatics. Chemical-grade aromatics are refined by extracting the desired BTX components from steam-cracker or reformer effluents, For cracked products, BTX rich cuts are hydrotreated for a complete elimination of diolefins, olefins and sulfur before extraction. BTX rich cuts originating from reforming units are not submitted to hydrotreating because they are largely free of sulfur. Reforming units designed prior to the mid-1980’s operate at significantly higher pres- sures (> 15 bar) than modern units. This results in the production of negligible levels of diolefins and low levels of olefins. BTX extraction processes fed with low severity reformate operate without difficulty however, treatment of the aromatics fractions with an adsorbent is necessary to attain the stringent AWC specifications required for the products. The usable lifetime of the adsorbent material is generally superior to 6 months. With the advent of high severity reforming, adsorbent lifetimes have been dramati- cally reduced to generally less than 3 months. Short lifetimes lead to high adsorbent consumption and high cost associated with adsorbent purchase, change-out, pack- aging, storage, transportation and disposal. In addition, this is a major environmental 70 headache because large amounts of aromatic-rich spent materials must be elimi- nated, generally ending up as toxic landfill. Several commercial benzene extraction processes employ liquid-liquid extraction or extractive distillation. Traces of diolefins, such as cyclohexadiene and methylcyclo- pentadiene, and cyclic olefins such as methylcyclopentene and cyclohexene have a great impact on the extractive distillation process because they tend to remain asso- .. .,. . ciated with the benzene in the extractive solvent. This may require increasing .’ :: significantly the solvent recirculation rate and consequently solvent regeneration duty. Moreover, traces of diolefins in the extracted benzene will negatively impact the AWC value. Therefore an adsorbent tower is often required. The Arofinirrgm process, based on selective hydrogenation, is being employed upstream of the benzene ex- traction unit to eliminate these potential drawbacks. Arofiningm Unlike aromatics-rich steam-cracker streams that contain sulfur, reformer aromatics cuts are sulfur free. Thus, hydrotreating reformate with conventional COMOS hydro- desulfurization catalysts would require excessive process conditions and could add a contaminant (sulfur) to a clean product. On the other hand, conventional hydrogena- tion catalysts, e.g., Pd or Ni, offer milder conditions but have the potential of significant aromatic losses through aromatic ring hydrogenation. ., The Arofining process has been designed to accept aromatics-rich cuts at most bat- ,, tery limit conditions and selectively hydrogenate all of the diolefins and a significant portion of the olefins while avoiding aromatics losses. The catalyst employed for this application matches these requirements perfectly. Commercial results from a unit treating a & benzene concentrate have afforded a product with an undetectable diolefin content and virtually no benzene hydrogenation (cyclohexane make: less than 10 ppm). Arofining operates in the liquid phase in a simple chamber type reactor. The liquid :-’: hourly space velocity is very high, i.e., an order of magnitude higher than that of ad- :,”- .’ -,, sorbent towers. The ready-to-use catalyst is stable, enabling cycle lengths of several ., .,. ’.-.4, ..(: years; thus, no spare reactor is required. .,. ,,..C, ,,, , For CIj cut applications, Arofining uses no excess hydrogen; all of the hydrogen en- . ..., ,. ’-,;. tering the reactor is consumed. Therefore, with the degree of hydrogen purity ,,. available from pressure-swing adsorption (PSA) units, no stabilizer column is re- quired. For this application, the Arofining unit is represented by a feed drum, a feed .’,-, , pump with spare, and the reactor, representing only a few percent of the total cost for ,,. a complete benzene production and purification system (i.e., reformate splitter, Aro- fining section and Extraction system). Industrial experience. Three Arofining licenses have been awarded. The first unit was put on stream in the fourth quarter of 1995 on a CGrich cut upstream of an ex- tractive distillation unit. The initial catalyst load for this unit is still in operation and the AWC of the product is below 1 without treatment in an adsorption bed. Two other units are designed to operate on CJ3cuts. The first of these units will be put on stream in the first quarter of 2000. Conclusion ,. New, low capital cost technological solutions have been developed and commercial- #.,. ized to resolve reformer-related product specification problems. The Benfreelh~ ., process enables the attainment of the gasoline pool benzene specification (< 1 % VOI benzene) with minimum investment and operating costs while maintaining versatility and improving operability (no iso-CT paraffins in the light reformate cut). The Arofininglh’ process is an ultra-low investment, positive environmental impact process that significantly improves the performance of extraction plants, enabling the production of on-specification aromatics (B, T, X and combinations thereof) from high-severity CCR reformers while drastically reducing (and eliminating in some cases) adsorbent purchases and the toxic waste disposal problem generated by used absorbents. . 72 .,, , ‘,, ;., . . f ., ’”; ‘,’ ,(,., DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistsy”, Erlangen 1999 f,..~,, ,., f.,’, ,,. ,,,- ‘, .’,,, :..,: B. Firnhaber, G. Emmrich, F. Ennenbach, U. Ranke !’, , ..’, .,. Krupp Uhde GmbH, Dortmund, Germany ;.,,, ...,.,.,,,:,.::.., ‘,,..-”,., ?.! !, ”’.. ,., , Separation Processes for the Recovery of Pure Aromatics ,,..,,,.,,, ,. ,--, ,.- ,, ...... ,, . ., .$. ,,, . .: ..... : ;: :. .,,- f,, ..,, Introduction ... , ,!...... - >,, .,,, ,, Pure aromatics, benzene, toluene and the three (ortho-, meta-, para-) xylenes (BTX), ‘ .‘ -! ‘ are building blocks for a multitude of chemical products. They are recovered from product streams of the oil processing industry, which contain aromatics in high concentrations. Coalderived aromatics - coal conversion processes were the main aromatics source in the first half of the century-play only a minor role today. i The separation processes have the task of recovering pure aromatics from mixtures with other hydrocarbons which cannot be separated by distillation. The separation of aromatics and non-aromatic hydrocarbons is feasible if a solvent is used, which has :,, ?, a markedly higher volubility for aromatics than for non-aromatics and which is ., applied either in liquid-liquid extraction or in extractive distillation. The status of and the developments in extraction technology are the main subjects of this report. ,, To recover single xylenes, the compounds have to be separated after being extracted from the other C6 aromatics. The boiling points of ethyl benzene, m-xylene ,.,,. and p-xylene are all within a very narrow temperature range. While o-xylene can be .. .’ separated by tlactionation, special physical properties, like molecular structure and freezing point, are used for the recovery of pure m-xylene and p-xylene in sorption ‘ and/or crystallization processes. The important processes are the Parex and the MX ~‘ ,’ h 73 DGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 . . Sorbex processes of UOP and the Eluxyl process of IFP. Details on these processes are not discussed here, Separation technology is today an important part of the petrochemical industry due to the high demand for aromatics worldwide. . The Aromatics Industry (Applications - Demand - Supply) The applications of BTX aromatics are listed in Fig. 1, which includes the percent distribution of benzene consumption in 1997 [1]. Benzene was mainly applied for the production of ethyl benzene (55%), cumene (17%) and cyclohexane (13%). The overall demand for benzene amounted to 27 million tpa, by far the highest amount for a single aromatic. Second largest demand was for p-xylene, totalling 13 million tpa, which is used almost exclusively in polyester production. The total aromatics market currently amounts to nearly 60 million tpa. – 550/0Ethylbanzene -— Styrene — Polystyrene, ASS Resina, SBR v:y,$~rme - “’%c”mene { — ‘::G:%R::’SAdhesives, ASS Resins, Waxes Caprolaclam — Nylon 6 Benzene - 13% Cydohexane Adipic %“d — Nylon E6 f Cydoh6xanme — Nylon 12 – 6% Ndmbenzene –— Aniline r- - 4% .4kylbenzene — U@ - 2% Ch!oro- benzenes I Tc4uerw Oiiaocyanale Polyurethane Fcama I Nitrotoluenes TDP Toluene Explosives, Oyes Solvents O-Xylene Phlhalic .Anhydride Alkyd Resins, Melhylaoylate Xylenea M-Xylene IsWhlhalic Acid Polyesters, A!kyd Resins P-Xylene Terephlhalic Add/ Polyesters Oimethvltharaphthalate Fig. 1: Main Applications of BTX Aromatics 74 ,., , f,, ” . ~~ L..,’,: .- : -,., ,, ;.. ,,.,..; .,.~’ :., ..,<. <,:,’, ,,-~,:.,,,,, ... .- !,, ,, .,;,,: In Fig. 2, the aromatics demand worldwide is shown as a graph for the years 1997 to ,,, $ ,. ,, .’~,.-,, 2003. The forecasts [2] agree that p-xylene will have the highest growth rate, a stunning 8.4 Wa, followed by benzene with a healthy 4.3 Wa. These growth rates are based on the predicted rapid built-up of end-use markets such as polyesters, polystyrene, phenolic resins and nylon. The demand for toluene will remain constant as the use of toluene in chemical products is limited. However, tohrene extraction capacity will increase due to continued conversion of toluene into benzene and t’ xylenes by disproportionation. The differences in demand growth have an effect on . . . the recovery technology, as will be discussed later. . ,, ?’ (MilIion Metric Tone/Year) 401 (MillionMetric Tons/Year) 40 1 . . 35- 35- 30- 30- ,. 25- 25- -i, , ,,, ,. ,. ; ,-...... -, ,, 10- +:.::,:’:::;. , ;:”’- - ““.:- ,., . . ,... :, . . <, ~ ~ .:<,.,!, ..-: ,, .,- ,,, 5- -. .,. 1997 1998 1999 2000 2001 2002 2003 1997 1998 1999 2000 2001 2002 2003 ,.. ,’ -. Fig. 2: Aromatics Demand Worldwide The main sources of aromatics are reformate from catalytic ref0rmin9, PYrolYsis ‘ , gasoline from steamcrackers and coke oven light oil from coke oven plants [3]. Reformate from catalytic reforming processes provides the base SUPPIYfor BTX and ‘ heavier aromatics (Fig. 3-01 ). However, a large part of toluene and most of the .: ,,, heavier aromatics from reformate are converted to benzene and xylenes, mainly for ,,’, ” .,, , the purpose of p-xylene production. The remainder of the aromatics is produced from . pyrolysis gasoline and from coke oven light oil. ,. ,., 75 Most of the benzene is produced from pyrolysis gasoline, followed by reformate (Fig. 3-02). Some benzene is obtained from the hydrodealkylation (HDA) of heavier aromatics and by toluene disproportionation (TDP). About 6 YO of benzene is recovered from coke oven light oil. I Coke Oven HDArrDP Reforrnate Pyrolysis Light Oil 22% 33% Gasoline 4 % Oven Light Oil 6 Y. Reforrnate Pyrolysis I 72% Gasoline 39% Fig. 3-01: Fig. 3-02 world Aromatics SupPIy by Source Worfd Senzene Supply by Source Fig. 3: Aromatics Supply by Source Extraction Processes Generally, two different processes are used for producing pure aromatics from feedstocks rich in aromatics, the liquid-liquid extraction process and the extractive distillation process. Many factors determine which of the two processes is to be used, the main ones being the feedstock and the end products required. The liquid-liquid extraction (LLE) process uses the fact that aromatics and non- aromatics m“th the same boiling point have different degrees of volubility in the solvent. The extraction process (Fig. 4) is effected in several stages by supplying the two liquid phases, i.e. solvent and feedstock, to the extractor in countercurrent flow. The extractor must be designed to ensure a good transfer rate between the two phases while still allowing sufficient phase separation. 76 . I %-’ A. Water ,.. ------~ *—“, ..’ ::.,..., .,,— -.. .“, . 9 ..:. ~ :,..:. ,:., .... + ‘p 1- ...... Aromatica .,’. eam cl= !’.t ‘ l. —:. ..— .—. — –r LeanSOlvenl 1 Fig. 4: Liquid-Liquid Extraction (Morphylex) ,, Increased aromatics concentration in the solvent during the extraction process also ::::: ~,.,:. ‘‘: ,,,.,..,.,., improves the so[ubility of the non-aromatics.The non-aromatics with lower boiling ,.. .,.-,- ( points are more soluble than the higher boiling ones and can be separated in a ‘, ,, . ,,.. ,. .- -.’ primaty distillation stage horn the solvent mixture which is rich in aromatics. The .,, ,,,. ., , overhead product from the distillation process is fed back to the base of the .,....-., : ‘$,:,,;.:,. extractor, a few stages below the feed zone, as a countersolvent. This causes the :.”.,,,> ‘f ,,’ ..$.-~ ,! -“ ,, ..,- ,. ,, higher-boiling non-aromatics to be expelled from the aromatics concentrate. , ,, ,, ...., -, The aromatics content of the overhead product, and consequently the internal recycle rate, can be reduced considerably by feeding in additional solvent at the head of the distillation column several trays above the original feed position, resulting in extractive distillation. The bottom product from this distillation column is passed on to a second distillation column where the aromatics are separated from the solvent. The lean solvent is returned to the cycle. ,“- The solvent must meet a variety of demands, the most important being those made ,, on its solvency and its selectivity, i.e. the difference in the volubility of aromatics and t, non-aromatics. In addition, its density must differ sufficiently from that of the feedstock, its boiling point must be higher than that of the aromatics to be extracted ., 77 and it must display thermal stability as well as some other important operational characteristics. Finally, it must be available at a reasonable cost. Several different solvents have been used in LLE. The first commercial-scale LLE process for pure aromatics recove~ was launched in 1952 by UOP and Dow Chemical, the Udex process using diethylene glycol as solvent. [n the sixties, other processes followed. Shell introduced sulfolane as a superior solvent. UOP later became the exclusive Iicensor of the Sulfolane process and continued its development. Lurgi used a mixture of N-methyl pyrrolidone and ethylene glycol in the Arosolvan process, Krupp Uhde used N-formylmorpholine with a small percentage of water in the Morphylex process and Union Carbide used tetraethylene glycol in the Tetra process. However, UOP succeeded in contracting most aromatics plants with the Suifolane process. Now, in addition to Sulfolane, UOP also offers the Carom process which was commercialized by Union Carbide in 1986 and which uses a mixture of tetraethylene glycol and a proprieta~ cosolvent as solvent. All processes use a water wash column to extract solvent from the raffinate. Various methods are used to process the bottom effluent of this column internally. In Morphylex, the water is added to the solvent feed of the extractor (Fig. 4), in Sulfolane and Carom the water is used for steam stripping, while Arosolvan, which has a water wash for raffinate and extract, uses a separate distillation column. The LLE is predominantly used for separating aromatics within a wide boiling range, such as benzene to xylenes, and for extracting products with comparatively low aromatics contents, thus enabling two phases to be developed. However, as xylenes are not usually recovered from pyrolysis gasoline any more due to the high ethylbenzene content in this feedstock, and as the extraction of xylenes from reforrnate for the production of p-xylene is greatly reduced because of new selective conversion processes, the combined extraction of BTX is gradually being replaced by the extraction of benzene or the combined extraction of benzene and toluene. 78 In the extractive distillation (ED) process, the solvent alters the vapor pressure of the components to be separated, the vapor pressures of the aromatics are lowered .. ,.. \ more than those of the less soluble non-aromatics. The ED process is relatively !., simple (Fig. 5). The solvent is supplied to the head of a distillation column with a ,,, . .,, .4’ central feed inlet. The non-aromatic vapors leave the ED column with some of the ,; solvent vapors. No reflux is required as the internal reflux is controlled by the solvent ., feed temperature. The solvent is recovered from the overhead product in a small ..” column with section reflux, which can be either mounted on the main column or ,’ separate from it. The bottom product of the ED column is fed to a distillation stage, i.e. the stripper, to separate pure aromatics and solvent. After extensive heat ,’ exchange, the lean solvent is recycled to the ED column. Non-arsma!ics . ~-.–- -.. : -- StStipp b —.—-—. — —, 200‘c i Tq d ~ ‘: ‘ I 1 !- I 1 I A - I Fig. 5 Extractive Distillation (Morphylane) ,.,,<,.... ., The product purity is affected by a number of different parameters including the ~”,. ‘,,.>,..,.‘ .,, , . . solvent selectivity, the number of separation stages, the solventlfeedstock ratio and .,,,, ,,. the internal reflux ratio in the ED column. The solvent properties needed for this ,,,...,’.. ..,,, .,$. , ,.’:,:‘,:. process are similar to those required for the LLE process high selectiviut ~e~al ;~:.;,, ~~,, :,, ,’,.,: -’ stability and a suitable boiling point. As the effect exerted by the solvent on the , :,, ,., . . . . vapour pressure of the feed components is limited, feedstocks with defined final 79 ,’- . c, . . boiling points must be used if pure aromatics are to be obtained, e.g., fractions comprising aromatics with one or two carbon numbers only. Aromatics separation by extractive distillation is a German domaine. It was introduced almost simultaneously by Krupp Uhde and Lurgi in the sixties for the recovery of benzene and toluene from coke oven light oil. The need for a new technology arose when aromatics from oil refineries came on to the market. These were separated by LLE and were purer than the products of existing processes. LLE was not suitable for processing coke oven light oil because of its high aromatics content. The Lurgi Distapex process used N-methylpyrrolidone as solvent, while the Krupp Uhde Morphylane process used N-formylmorpholine. Both processes have been quite successful since. A comparison of the ED and LLE processes shows that the ED process design is simpler than that of the LLE, resulting in lower costs for equipment and utilities. ED is also capable of processing all feedstocks, irrespective of the aromatics content. The ED is much more economical than the LLE especially for the recovery of benzene or the combined recovery of benzene and toluene. As the importance of xylene extraction has diminished, ED processes are favored today for new capacities, while LLE still leads in terms of both volume of aromatics produced and number of plants. Process Configurations The main feedstocks for aromatics recovey, Reformate from catalytic reformers, Pyrolysis gasoline from steam crackers and Coke oven light oil from coke oven plants, have very different compositions with regard to the content of paraffins, olefins, naphthenes and aromatics as well as different quantities of impurities, such as compounds containing sulfur, oxygen, nitrogen and chlorine. In Fig. 6, typical 80 ., ;, compositions are listed for hydrorefined coke oven light oil and pyrolysis gasoline as .,\ ‘,,. well as reforrnate. ,. ,!”, ,.. ,, ,:. , % Hydrotreated Hydrotreated Reformate .,. coke oven benzole pyrolysis gasoline ,“, , % Benzene 65 40 3 .’ Toluene 18 20 13 Xylenes 6 4,5 18 Ethylbenzene 2 2,5 5 “. Higher aromatics 7 3 16 Total aromatics 98 70 55 ,“ 2 30 45 .’ Non-aromatics ,,- ,., 100 100 100 ,. .- ,,, ., Fig. 6 Typical Feedstock Compositions for Aromatics Recovery Various pretreatment, aftertreatment and distillation stages lead to a variety of :,’,,’, process ~nfigurations which are determined by the kind of feedstock applied and by ., the aromatics to be produced. Aromatics RecoveW from Reforrnate Using Liquid-Liquid Extraction Reforrnate produced by the catalytic reforming of desulfurized naphtha is fed to the extraction process without any pretreatment (Fig. 7). Small concentrations of olefinic compounds in the extract necessitate clay treatment in order to meet product .. ,.,. ,’- specifications. The treated extract is fed to the aromatics fractionation section where benzene, toluene and the Cs aromatics are separated. The Ca aromatics, ethyl ,:,: .,’ benzene and the three xylenes, can be sold on the market as mixed xylenes or can .; be processed further for production of pure single xylenes. ., Benzene Non-Aromatics I I “i Column Column C61umn Fig. 7 Aromatics Recovery from Reformate Using Liquid-Liquid Extraction The Morphylane Extractive Distillation Process The advantages of the extractive distillation process over the liquid-liquid extraction process were mentioned in detail above with particular emphasis on the simplicity and cost effectiveness of the process when only one or two aromatics are to be produced. The Krupp Uhde Morphylane extractive distillation process, as it stands today, is the result of a continuous development effort over the last 30 years w“th the aim of improving product purity, energy efficiency and construction costs. The basic ;. design (Fig. 6) and the solvent, N-formylmorpholine, are the same for all Morphylane 6, ,. plants [4], irrespective of the feedstock used or the product required. Up to now, Krupp Uhde has been awarded contracts for more than 40 Morphylane plants worldwide for the recovefy of aromatics from coke oven light oil, pyrolysis gasoline and reforrnate. Plant capacities range from 28,000 tpa to over 800,000 tpa. Product purities of 99.99 Y. by wt have been achieved for benzene and toluene. Table 1 lists the operating results of several plants using different feedstocks. 82 ., ,, ,’, , ?’ ,, Tab. 1: Operating results in Morphylane plants using different feedstocks ,., Heat consumption Yield Purity per metric ton of feed kcsl %bywt %’0by Wt ,, .;, Feed from Reformate f ., ,,, Benzene only 219,000 99.40 >99.99 Benzene & Toluene 271,000 Benzene 98.50 Benzene >99.99 ,,, ,.” Toluene 99.80 Toluene >99.99 ;! .+. ,, .,. , Feed from Pyrolysis Gasoline ,’ Benzene only 246,000 99.00 >99.98 .,. Benzene & Toluene 241,000 Benzene 98.00 Benzene 99.99 Toluene 98.70 Toluene 99.85 ,, Feed from Coke Oven tight Oil ,’ Benzene only 207,000 >99.70 >99.95 Benzene & Toluene 483,500 Benzene 99.50 Benzene 99.99 Toluene 99.50 Toluene 99.98 The solvent used in the Morphylane process is a single compound, namely N- formylmorpholine (NFM). No agents or promoters are added to the solvent. The ,, ,, salient features of NFM are summarized as follows [5] . .’ High selectivity and high solvent efficiency. NFM shows the best selectivity ,,. when compared with other solvents, especially in the ED process. .,, NFM exhibits low basicity. Mixing with water at a ratio of 1:1 gives a pH-value ‘‘ ,, of 8.6, which inhibits any corrosion to equipment. Many other solvent; demmpose or hydrolyze to acidic compounds and need a pH-control system < ,, at least. ., NFM is also characterized by a high permanent thermal stability and its extremely low tendency to polycondensation or gumming. This results in low solvent consumption as well as in low costs for regeneration. According to the results of biological tests carried out on mice, rats and rabbits, NFM is practically non-toxic. Krupp Uhde has tested many different solvents and solvent mixtures since it ., ,,,,.,,’ developed its extractive distillation processes. NFM as a single solvent has proved ,, ‘, ,. itself to be the best choice for aromatics extractive distillation. .,, Typical process configurations employing the Morphylane extractive distillation process are described in the following sections chronologically rather than in order of economic importance. Aromatics from Coke Oven Light Oil Coke oven light oil is produced during the coking of hard coal. It has a very high aromatics content with the non-aromatic components being highly unsaturated and having the tendency to polymerize. In addition, compounds containing sulfur, oxygen, nitrogen and chlorine must be removed by hydrotreatment before pure aromatics can be recovered. Non-Aromatics I I Ce Fraction Motor Toluene Off-Gas r I (H2S, NH,, CH4, etc) I I t Crude Hz n Coke Oven Light Oil $gdq;~ + + + i’genation:~ ~ Stabilizer Pr- ToIuene Xylene SoIvent I fmctionator Column CoIumn CoIumn Fig. 8: Aromatics Recovery from Coke Oven Light oil ,. t.. . Hydrogenation is usually carried out at a pressure of 30-50 bar in a two-stage ,, reactor. In the first stage, the compounds which form polymers are hydrogenated at relatively low temperatures w“th a nickel/molybdenum catalyst. After temperature increase, full hydrotreatment is carried out at 320-360”C on a cobaltlmolybdenum catalyst. The feedstock for extraction is obtained by stabilization and distillation steps (Fig. 8). 84 Hydrotreated coke oven light oil is normally used exclusively to recover pure t benzene. The higher aromatics, separated by distillation, have a low non-aromatics . . content and can be marketed as such. Extractive distillation is therefore the most economical process. . ,, ?, Aromaticsfrom Pyrolysis Gasoline Similar to coke oven light oil, pyrolysis gasoline is mainly used to recover benzene or benzene and toluene. Mixed xylenes are not usually recovered from pyrolysis ,. gasoline as the content of xylenes in the feedstock is relatively low and the ;, , .,, economic use of mixed xylenes, when recovered from pyrolysis gasoline, is limited ,,. due to the high ethylbenzene content in pyrolysis gasoline. Again extractive ,,$,..- ““ .’ . distillation is the preferred recovery process. .:. , . . . . .,’. .-, “: Non- Aromatics to Cracker cc to Gasoline Pool ! C6-G ‘Aromatics Exbaotive: * ., orr- DIatillation Pyrolysis Gas Gasoline ;Zd StaP. ?. - ., Hydro- ,’ tradment .’ ., . . t Hz .,’ Pool ,,, :. ,, ,, Fig. 9: Aromatics Recovery from Pyrolysis Gasoline ~,., A typical process route for the recove~ of benzene and toluene, starting from crude ., pyrolysis gasoline, is illustrated in Fig. 9. In the first step, selective hydrogenation, diolefins are saturated at a relatively low temperature in order to avoid ,,, , ., .,-. ,-,’,,: :. polymerization. The C5fraction is usually separated from the selectively ., ,., ,,,,,,,,,} ,-.-,:. hydrogenated pyrolysis gasoline in a depentanizer upstream of the full ,,,,.., ,.,;,, ..> :,:, J,, ,,! !,:., :) 85 ,,: ., .,.. hydrogenation unit and sent to the gasoline pool. In this way hydrogen can be saved and the size of the tlrll hydrogenation unit can be reduced. In the full hydrogenation stage, olefins and impurities such as components containing nitrogen and sulfur are completely hydrogenated. The off-gas is separated in the stabilizer from the fully hydrogenated pyrolysis gasoline and then sent back to the steamcracker. However, if the Cs fraction is sent back to the steamcracker as feedstock, the C> fraction should be fully hydrogenated and separated downstream of the full hydrogenation stage, either in a combined depentanizerk.tabilizer or in the extractive distillation together with the non-aromatics, thus dispensing with the need for a depentanizer system. In order to extract the desired aromatics, a specific aromatics cut has to be separated from the pretreated pyrolysis gasoline. In the case of benzene recove~, a C& cut is separated in a predistillation column and sent to the extractive distillation and a CT.cut is separated in the case of benzene and toluene. The CT+or the CE+ fraction is sent as valuable octane feedstock to the gasoline pool. In some cases it can be economical to convert the C7+aromatics into benzene to maximize the benzene output. In this case, a thermal hydrodealkylation unit is integrated in such a way that the extracted toluene and the xylenes from the predistillation are dealkylated to benzene (Fig. 10). The off-gas produced can then be used as fuel gas. The toluene must not necessarily be extracted before being fed to the hydrodealkylation stage. However, more hydrogen is consumed when non- extracted toluene is used as the cracking of the CT non-aromatics produce a larger amount of off-gas. Depending on the available feedstock and the desired products, individual process configurations and optimizations of heat integration between the individual units are provided taking into account local conditions such as specific utility availability and costs. Consequently the most economic solution regarding investment and operating costs can be provided. ,’- , c , . . ,, 86 ,, ,.. ?, ,, ., NorkAmmatisa Pure to cracker ,. t + ,, c< to ~, J GasotinePool ... . . ~e. ~ Aromatke ~ ~Extraative . orr- orr- D1atIllation ,,$ “,,,, Pyrolysis Gas Gas ,:. , Gasoline c, from Cracker ~1’* tep. w step Fuel Gaa ‘m m ~~u;t ,,. + nent Hydru-. ,,, -, H, % dealkylation 4- : -; Ca+ ‘1; ‘; % u Pod I Fig. 10: Maximized Benzene Recovety from Pyrolysis Gasoline ,. Aromaticsfrom ReformateUsing ExtractiveDistillation ,“ Reformate is the preferred feedstock for producing p-qlene due to i~ relatively low ~, benzene content and the high toluene and xylene content. A process configuration ,, ‘‘ using extractive distillation is presented. Extractive distillation is also used in the reduction of benzene in motor gasoline by recovering pure benzene from reformate. . . In order to comply with governmental regulations with regard to benzene content in :‘ ., . ‘ motor gasoline, the refining industry has to adjust its operations to either convert or recover benzene. Krupp Uhde developed an optimized process configuration in term of investment and operating costs for benzene recovery from reformate (Fig.1 1). The reformate from catalytic reforming is fed to the reformate splitter to separate the C&fraction. This fraction, which also contains Cs hydrocarbons, is sent to a selective hydrogenation stage in order to saturate diolefins which influence the .. acid wash color of the benzene product. The selective hydrogenation stage consists ,, ,.,,.,. of a small reactor which is operated at mild temperatures and pressures in a trickle bed [6]. The small amount of offgas is separated vithin the reactor. This upstream ‘‘: ... . , .<’ ‘: ,’,/-,’ “ ,:J/y ‘ ,’,, hydrogenation of diolefins avoids the necessity of downstream clay treating and : ,, . ,’.’ subsequent product distillation, thus saving considerable investment and operating Fig. 11: Recovery of Pure Benzene from Reformate for Benzene Reduction in Motor Gasoline The pretreated C& fraction is sent to the extractive distillation stage, where high purity benzene is removed from the non-aromatics in a simple two column system. One of the advantages of the Morphylane process in this case is that diluted hydrogen and light hydrocarbons can easily be handled in the extractive distillation column and as such no upstream depentanizer/stabilizer system is required. Other ;. advantages are the very high benzene purity and the very low benzene content in 4 ‘, the non-aromatics. Some years ago Krupp Uhde developed an optimized benzene reduction process for typical refinery applications. The main aim was to reduce the energy requirements when removing the small quantity of benzene from reformats resulting in sophisticated heat integration. Recently a plant where this improved configuration had been implemented was started-up, demonstrating the following excellent performance figures: 88 . .) .“ . . Steam consumption per ton of reformate 0.26 t ., ’,.,. Benzene quality >99.99 %foby wt. ,’ -,’ ~, .,,, .,. , , <0.15 ~o by VOi. .,, Benzene content in gasoline ,,, ., .,,! ,! .,, .’ ‘,.’,.,, ,. Depending on the refiner’s specific needs, other process schemes are also available ,.. .,,: ,-, ,, ~, e.g. a combination of reformate tlactionation into several blending cuts and ,,, .”,- ,’ benzene extractive distillation or a combined benzene and toluene extraction. Krupp Uhde’s so called Octenar process is a special process scheme in which high-purity benzene is recovered alongside a C7+aromatics concentrate in a single extractive distillation with till range reformate as feedstock. The Octenar process [8] is a modified version of the Morphylane process in as far as two extract tlactions are taken from the solvent stripper, thus producing three different products from the reformate Benzene with a purify of B 99.99 Y. by wt., Raffinate w“th a benzene content of <0.1 % and Aromatics concentrate W-th approx. 6% C7+ non-aromatics and an octane number of> 110 RON. Octenar was originally developed with the aim of producing high-octane gasoline. The raffinate, which did not contain any aromatics, was discharged and the aromatics concentrate fed to the gasoline pool. Today’s regulations pertaining to gasoline allow this process to be used for benzene removal and for reducing the overall aromatics content. Concern has been voiced that recovery of benzene from reformate with the aim of reducing its content in motor gasoline would flood the benzene market. In view of the :, high benzene growth rate and of the fact that benzene is still produced by the .... hydrodealkylation of higher aromatics, this concern is not likely to be verified. :-.$, In the case of p-xylene production which is illustrated in Fig. 12, catalytic reformate is first split into a C7. and a C6+fraction. The C7. fraction is sent to an extractive ‘,’ distillation stage where benzene and toluene are separated from the CT. non- ,,,,,!, .,,.,,,’. ,,, ,,., ,,. ,.’ ,,-,,~.,,: 89 ;,,,., , “,, aromatics. A selective hydrogenation stage is added upstream of the ED. This configuration is preferred to the alternative which involves clay treatment after the ED process, Fig. 12: Recovery of Aromatics from Reformate Using Extractive Distillation The CB+ fraction is sent directly to the p-xylene loop without extracting the xylenes. As the non-aromatics content in this fraction is very low and can easily be treated in the p-xylene loop, only benzene and toluene have to be extracted and not BTX. The toluene separated is sent to the toluene disproportionation stage, where the toluene is converted into benzene and xylenes. After fractionation, the benzene produced is sent together with the extracted benzene as a high-purity product to battery limits, whereas unconverted toluene is recycled back to the disproportionation stage. After separation, the mixed xylenes fraction from the disproportionation stage is sent with the xylenes fraction from the reformate splitter and recycle xylenes from the isomerization stage to the p-xylene adsorption, where pure p-xylene is recovered. The remaining xylenes (m-, o-xylene and ethylbenzene) are sent to the isomerization 90 stage to be converted into additional p-xylene. If o-xylene is also required as a product, this is separated by distillation in a separate column system. A toluene and Cw aromatics transalkylation stage is integrated in the process scheme instead of toluene disproportionation, if maximum p-xylene production is desired. In this case, the extracted toluene and the complete Cgt aromatics fraction are sent to the transalkylation stage to be converted to benzene and xylenes. The design of the detailed process configuration and the heat integration between the individual units are based on the specific feedstock composition, the desired products and the availability and cost of utilities. For example, the vapors from the xylenes distillation can be used to heat the benzene and toluene columns, etc. Outtook The separation or production of pure aromatics still has considerable potential for development. Opportunities exist in improving the processes applied today as well as in solving new separation problems. With its extensive knowledge in extraction technology and excellent laboratory facilities, Krupp Uhde is well equipped to participate in this endeavor. Two new processes, described below in more detail, have been developed and demonstrated on pilot scale by Krupp Uhde in recent years an improved Morphylex liquid-liquid extraction process and a process for the recove~ of polymer-grade styrene from crude pyrolysis gasoline using extractive distillation. Current pilot-plant testing aims at further improving the Morphylane extractive distillation process. First results indicate feasible reductions in energy demand by 3070 and in investment COStSby 2070. The improved Morphyfex Liquid-Liquid Extraction Process While the ED process with NFM as a solvent enjoyed great success in its various forms, the LLE process with NFM, the Morphylex process which was introduced in 91 the 60s, was less successful. There were many reasons, the main one being its relatively high energy demand. A e Hea aromaks Washer b A @~ r~ Waler+ Exlradof ,,-. solvent ExtraciJva;Ji!iwater“m Di$tilabon Q . column Water ~~ C&ng Side Feed . Slnpper ~ Aromati e——— “ r 8 8 t- v = , d,~, 1 Steam %lwnkarcmsks I ./ I-WI,,, d.t I I r-l’” I Fig. 13: The New Morphylex Liquid-Liquid Extraction Process The fact that there is still a market for combined extraction of BTX in some parts of the world has led KU to improve its process considerably in the last few years. The process engineering and R&D departments have worked together to develop a process which constitutes a considerable improvement in both investment costs and in energy consumption. The flow diagram in Fig. 13, which shows the fully integrated pilot plant used to demonstrate the process, will help to explain how the process works . The plant now has only 2 columns - each with a smaller attachment - the extractor and a combined ED/stripping column. In the extractor, the solvent, NFM with 4 to 6% water, is fed from top to bottom as a continuous phase. The feedstock, reformate, is fed to the column several stages above the base of the extractor. The difference in density causes the feedstock to bubble upwards in countercurrent to the solvent. During this procedure the aromatics pass into the solvent and the non-aromatics stay in the light phase. Internals, structured packings or sieve trays, ensure that the phases are well distributed throughout the cross-sectional area of the column. The 92 overhead product from the second column, which mainly comprises relatively low- ., /.,, ., boiling non-aromatics, is fed to the base of the extractor as a countersolvent. The ,, head and the base of the extractor act as phase separating vessels. The non- ., ..’, aromatics with a slight concentration of NFM in solution are drawn off overhead, and the solvent containing all the aromatics and some non-aromatics is drawn off at the base. The extractor is operated at near atmospheric conditions, with temperatures of 30 to 50 “C and pressures of 1 to 3 bar. The addition of water reduces the volubility of all hydrocarbons, that of the non-aromatics more than that of the aromatics, so that the selectivity is also increased. The water is also used for washing out the NFM which is dissolved in the overhead product tlom the extraction process. The selectivity of the water for this separation process is extremely high; the NFM can be ,., , ,,.,, ,.,, recovered in just a few stages. ,.- ,, .,:, ,. ..,, Column 2 is divided into 4 sections. The bottom product from the extractor is fed into the column between sections 2 and 3 (from above), and additional solvent is fed in ,, above section 2. Section 1 is a small solvent recovew stage which allows solvent- free water to be recovered from the overhead product. Section 3 is used to strip the non-aromatics from the aromatics/solvent mixture. Some of the vapors produced in the bottom stripper, section 4, are used to heat the ED and some are fed into a small “ lateral column where the pure aromatic product is separated from the solvent. This lateral column does not have a bottom reboiler and consequently the bottom product ,,. /,. still contains some aromatics. For this reason the bottom product is returned to the .’, \.. , . -. .,, ED which is operated at reduced pressure because of the boiling temperature ,, .’, ,- ., ,., threshold. ,.~,,,,’,,:,,.:..,. ,,’ ,. <-: .“,?., ,.-, ...... - ., ,~- The stripped solvent is practically water-free. The water undergoes overhead ,., ,, ,- .-,., ,, azeotropic distillation and is produced as a separate phase in the reflux drums. This water is then fed to the solvent recovery stage of the extraction process. The .. . ..”.. ,: ‘?,::’,’~’ < ,;,-, :’,’ discharge stream from the solvent recovety stage, i.e. water w-th a few percent NFM, : .:,,’,’.. ,,,. , :. ’... .“ is used to strip the heavy aromatics which cannot be separated from the NFM by ,.,; ,., .’ ,,,, distillation. For this purpose, part of the NFM recycled to the extractor is mixed with the water discharge to form two phases. The upper hydrocarbon phase is added to the overhead product from the extraction process and the lower phase is added to ‘.,.,J ;... .,. ,,, ,, :!! ‘t, : the remaining cycle NFM to regulate the required water concentration. This procedure avoids the costly distillation of the reformate beforehand. It has been proven that a constant low concentration of heavy aromatics can be attained in the cycle solvent during continuous operation with a standard reformate which has not been distilled to separate the high boilers. The numbers of Table 2 apply to consumption, production yield and product quality attained when using the modified Morphylex process: Table 2: Performance of the Modified Morphylex Process I Consumption per tonne of feed Steam (20 bar) 0.46 t Cooling Water (At= 10°) 12 m’ Electric Power 18kWh Product Yield Benzene -100 ?tobywt. Toluene 99.7 % by wt. EB, Xylenes 94.0 % by wt. E!@Y Benzene 99.999 % by wt. Toluene ~ 99.9970 by wt. EB, Xylenes >99.99 ~0 by Wt. Recovery of Polymer-Grade Styrene from Pyrolysis Gasoline A process for producing styrene from crude pyrolysis gasoline by subjecting a CE heart cut to an extractive distillation process with dimethylacetamide as solvent was described by Toray Ind. as long ago as 1970 [9]. A commercial unit has not been built despite of the attractiveness of the concept in terms of resource savings. Instead of producing styrene from benzene and ethylene in several process stages, it is extracted directly from crude pyrolysis gasoline. The main reason for this is probably the complexity of the task. Styrene used in polymerization processes must 94 ... ,’ be extremely pure and pyrolysis gasoline contains many compounds which are difficult to separate from the styrene. . . . The new process [10], which has been demonstrated on a pilot-scale, produces styrene with the necessary purity. The individual stages of the process are (Fig. 14): ‘ Two-stage distillation . . ,,, - Selective hydrogenation of phenyl acetylene ,’ ‘, Extractive distillation using NFM as solvent (with ED, stripping and solvent ,., ., regeneration columns) Secondary purification (with acid treatment, washing and drying stages and flash distillation). ,, . . ., , ACM Hn ., ,. ,,.,,.,, ,,, ,, ,,4 ., ,, ., ., Fig. 14 Recovery of Polymer-Grade Styrene from Pyrolysis Gasoline The feedstock is untreated and consequently very reactive pyrolysis gasoline. The ~‘ ,., styrene content of this product is between 3 and 7 YO by wt depending on the pyrolysis and recovery conditions. Once a polymerization inhibitor has been added ; to the feedstock, the benzene/toluene tlaction is distilled overhead in the first column .,;.-.’ ~~ ,. and the Cs aromatics fraction in the second column. The styrene content in the Ca ., .“ ,, ,. fraction is between 35 and 45% by wt. Phenyl acetylene which cannot be separated from styrene by fractionation or extractive distillation must be removed by selective hydrogenation of the C8 cut prior to the ED process. in this process, the acetylene bond is hydrogenated without styrene loss. The feedstock sent to the ED mainly comprises ethyl benzene and xylenes as well as styrene. The extractive distillation section is similar to the standard Morphylane process, except that reduced pressure is applied in both columns and that a solvent regeneration column is added. Both these measures are added as a result of styrene’s tendency to polymerize. The s~rene product from the ED contains <1 ppm NFM, but its yellow coloring makes it unsuitable for polystyrene production. This color is caused by traces of unsaturated hydrocarbons (polyenes) which cannot be ,’- . separated off by distillation or by extractive distillation. They must be removed in a #,. . chemical reaction. The treatment process uses an acid wash. Colorless polymer- ‘.. grade styrene is obtained after water wash, drying and flash distillation stages. The following operating results will be attained with 6 ‘%. by wf sfyrene in the crude pyrolysis gasoline: ~ >96 oh UtiliN consum@ion Rer ton of feedstock Low-pressure steam ( 4 bar g) 0.8 t Medium-pressure steam ( 9 bar g) 0.1 t Cooling water 7.6 m’ Electric power 9.0 kWh Solvent 5.0 g Inhibitor (DNBP) 300 g Product quality Styrene 99.9 % by wt Aromatics c 600 ppm Total sulfur 10-20 ppm Solvent ~ 1 ppm Color ~lOmg Pt/1 96 ,.. ., ,’, , ?’ Summary ,. Aromatics recovery is a large and growing industry. Higher growth rates are forecast “‘ . forp-xylene and benzene than for the petrochemical industry ingenerel. Process .;, ,,, ,. configurations are changing due to a different product mix and to the application of new improved conversion processes. Liquid-liquid extraction, the preferred ,, ~~,, ., separation process for four decades, is gradually being replaced by extractive ,,, ,.’ distillation which has both investment and operational cost advantages...... References ,. [1] M. A. Fisler, Benzene- Boom or Bust in 2000, 1998 World Petrochemical Conference, Houston, April 1998 [2] J. N. Bonarius, Market review- Base Aromatics, Phenol and Styrene, ,’, ,. Ist European Petrochemicals Technology Conference, London, June 1999 ,,- ,., ,.: [3] S. Palmer, Benzene – Where we are now, Dewitt Petrochemical Review 1998 ,., . [4] G. PreuRer, G. Emmrich, Erdol und Kohle 3&, 207 (1983) ,. [5] H.J. Vollmer, G. Emmrich, N-FormylmoTholine - The superior solvent for aromatics recovery, 1OhChisa Congress, Prague, 1990 ,’ [6] P. Polanek, M. Hooper, J. Miller, G. Emmrich, Purification of reformer ,.. ., streams by catalytic hydrogenation, NPRA Annual Meeting, San Antonio, ., March 1996 m G. Emmrich, U. Ranke, Reducing benzene in gasoline case studies, HydrocarbonTechnology International, Spring 1996 [8] G. Emmrich, Octenar - An economical process for producing low benzene gasoline, NPRA Annual Meeting, New Orleans, March 1992 [9] M. Sate, Hydrocarbon Proc. May 1973,141 .. [10] G. Emmrich, B. Firnhaber, H. Gehrke, U. Ranke, Int. J. of Hydrocarbon ‘,... Engng. ~, 62 (1998) ~ ,,., .. .’ ... ! , 97 <~” ---- ,,. . . . y.’. ,, ,~..,. ,:; .,K. ,,. , 6, . . ., 98 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 ,> .:, ; ‘} G. Krekel, J. Eberhardt, T. Diehl, G. Birke, H. Schlichting, A. Glasmacher . . . Lurgi Oel Gas Chemie GmbH, Frankfurt, Germany ,,, .- ,. .,.,. . .. . ‘. . . :. .. ,,. , ‘ !,, .,. ,,> Developments in Aromatics Separation ?’..,,/,.: ..., ., ::. ,. -’ ,,, .“, ,.. !.. , .,.,- ,.r ,, Abstract Extractive distillation and liquid-liquid extraction are the up-to-date used separation technologies for the production of pure aromatics for petrochemical applications. This paper gives a short characterisation of the present aromatics market and discusses developments in the aromatics separation technologies and configuration options in view of different feedstocks and desired products. .’ ., -, Introduction /. Aromatic compounds rang among the most important basic chemicals for the ‘, -.:’.,,, ‘, ,.: ., ,. ,, .,. petrochemical industry. The key compounds of the extended aromatics chemistry :-:, ,,, are benzene, toluene, the isomers of xylene, and ethylbenzene. ~,.,,,, ! ,,. ,:, ,’f - .’f> .. Aromatics are almost exclusively produced from petroleum and coal. However, the aromatics concentrations in these fossil fuels is too low for an economical recovery. Suitable raw materials can be obtained by thermal or catalytic treatment though. Today, aromatics are essentially recovered from the following raw materials: ,., (1) Coke oven light oil from coking of hard coal, , ?“, (2) Reformate from naphtha reforming, (3) Pyrolysis Gasoline from steam cracking. Historically, the aromatics production from coke oven light oil was most important. Due to the decline in coke production, the share of the aromatics production from coke oven light oil has fallen to less than 5 ‘3!. of the total production (Fig. 1). Today, most of the aromatics are recovered from reformate (about 70 Y.) followed by pyrolysis gasoline (about 25 ‘Y.). ,., Regional differences in aromatics production can be observed. In the US the .,,.” aromatics production from reforrnate is most important, since the mayor part of the ., ., steam crackers is operated with gaseous feedstocks. Hence, the yield of pyrolysis gasoline is relatively small. In Western Europe and in Japan aromatics are produced above average from pyrolysis gasoline. .,, , ,/, ., ;:,+-,r ,; Table 1 summarises typical compositions of the raw materials for the aromatics recovety. Reformate has a low concentration of benzene and considerable higher J ,,1,!, ..’,,,. ,.. ,, , ... ,,, , ,,-, :, ,. .. ,, ,., ,,‘,, ,: .!. ,, ,+, ,, , J’.,, ,, , ., .,,,,. , 99 OGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 ,,. . , ,’ -j .,, .,..; Sources for Aromatics Coke Oven Pyrolysis Light Oil Gasoline about 5°A about 25’%0 - Reform ate about 70% Figure 1: Sources for aromatics concentrations of toluene and xylenes. On the other hand, the composition of pyrolysis gasoline is reflected in a high concentration of benzene. Nevertheless, the composition of pyrolysis gasoline can be influenced by the severity of the steam cracking process. The greater part of coke over light oil consists of benzene and toluene showing only a vety low non-aromatics content. As can be infered from Table 1, reformate and pyrolysis gasoline contain non-aromatics with higher concentrations between 25 and 50 ‘A. The non-aromatics consist of a large number of paraffinic and naphthenic components showing very close boiling points to the aromatics and partly forming azeotropes. Using conventional distillation techniques, the separation of pure aromatics from non-aromatics is uneconomical, because it requires large numbers of trays and a high reflux ratio. Alternative methods for separating the close boiling aromatics from non-aromatics were developed during the sixties. It was through the use of liquid-liquid-extraction followed by extractive distillation. Table 1: Typical compositions of raw materials for aromatics recovery Composition, ‘YO Reform ate Pyrolysis Coke Oven Gasoline Light Oil I Benzene 4 38 65 Toluene 17 20 18 Xylenes I 18 5 6 Ethylbenzene/sty rene 5 4 2 Higher aromatics 11 5 7 Non-aromatics 45 28 2 100 ,, .,, Lurgi has almost 40 years of experience in the aromatics separation business. The i; ,”,, first commercial AROSOLVAN@ liquid-liquid extraction unit for the recove~ benzene, ,., toluene, and x#enes (BTX) of high purity went on stream in 1961. Since then 16 AROSOLVAN plants were commissioned in Japan, South America, Western Europe and the US. [n these plants BTX and BT aromatics are extracted mainly from reformate and pyrolysis gasoline feedstocks. Two AROSOLVAN@ plants were built to produce solvents free of aromatics from kerosene and naphtha feedstocks. In the mid-sixties Lurgi developed the DISTAPE~ extractive distillation process for the recovery of aromatics. N-methyl yrrolidone (NMP) is used as the solvent. The first applications of the DISTAPE # process were made on coke oven light oil, because the high aromatic content in this feedstock made the use of an extractive .-, ,, distillation technique economically more favorable compared to a liquid-liquid ,,, :. extraction. Being in the lead of the competitors, the first commercial DISTAPE~ ., .,, plant went on stream in 1966. .,”, $ .’,’”, .,,., ..., , Due to its advantages like flexibility to changes in feed quality and the simple .,, ,, operation, DISTAPEfi extractive distillation was used already in the following year ,’,’, ,., ,- ,. ,,,-. ,, for an application to recover benzene from a pyrolysis gasoline feedstock. ,- .’, ,. ,,. , ,,. , f !,,’ , Up to now, the DISTAPE~ process was continuously improved and developed. The ,,. references include phk fOr recovery of benzene, toluene, Ci3-aromatics, and ethyl r..;:,...:..+,:+ ,:.,,,’~. >,.’,?, \ benzene from different feedstocks like coke oven light oil, pyrolisis gasoline, and ,, ->,.!, , .,:,-,.,. reformate. Lurgi has been awarded contracts for 25 DISTAPE~ plants located in .,. Japan, Western Europe, Australia and Asia. Aromatics producers such as Veba, ,<,, ,,.,.,,.’. ,. BP-Amoco, DSM, DOW, ICI, EC-Erd61chemie, and OMV placed confidence in the .--, ,, DISTAPE#’ process. At present, two DISTAPE~ plants with both, improved .,,,.,., ., ,,.. ,’., technology and process configurations are ready to be commissioned...... ,.:’ ,., ,: .... : :! Lurgi’s processes for the separation of aromatics from non-aromatics will be ...... -.., ,, ,..:, ,, described in detail later along with the properties of the applied solvent NMP. .<”. “’,..:, : ,, ...... The Aromatics Market at a Glance The market for aromatic basic chemicals is above all determined by the aromatics consumption in petrochemical processes and thus by the demand for plastics. The most important basic chemicals for the petrochemical industry are benzene with an expected global demand of a little less than 30 Mio tla in 1999 and p-xylene with about 15 Mio t/a. These aromatics are used directly for the production of petrochemical intermediate products like styrene, phenol, and cyclohexane or plastics like polyesters. In addition, there is also a significant market for toluene and mixed xylenes with expected global demands of about 13 Mio tia and about 22 Mio t/a in 1999, respectively. As already mentioned there are regional differences in aromatics production. The same applies to the aromatics consumption. However, the aromatics market can not be considered for any region in isolation, since aromatics are commodity petrochemicals for many years. 101 1400-, ~ -o- Benzene I Contract, fob 1200- \ NWE, DM/mt I 1000- + Toluene Contract, fob 800- NWE, DM/mt 600- + o-Xylene Contract, fob NWE, DM/mt - p-Xylene . 0-11 I h I I t ,1, 1 1 I I t Contract, fob NWE, DM/mt + Naphtha Quart.Average fob NWE, DM/rr Figure 2: Price development of aromatic chemicals between 1996 and 1999 The mesent situation for the aromatic market is marked bv low r)rices for the aromatics, oversupply and low operating rates. The development if the contract prices for the most important aromatics is given with Figure 2. As can be seen, the prices bottomed out in the first quarter of 1999 and have slightly advanced again since the beginning of 1999. This is partly due to the rising oil price. For comparison, the quarterly average naphtha price is also shown in Figure 1. There are several factors from the demand and production side which influence the market situation. Besides individual influences for the single aromatics components, the present market situation is essentially influenced by the following general effects: ● Built-up of surplus capacity. The production capacity for basic aromatic chemicals experienced a strong expansion since the mid-nineties. However, the capacity expansion exceeded the actual demand resulting in considerable surplus capacity. At the end of the investment cycle, the situation shows low operating rates, fierce competition, and weak margins. ● The Asian currency crisis which began in Southeast Asia in 1997 amplified the difficulties. ● Reduction of aromatics in motor fuels. The regulations of the Auto Oil Program in Western Europe limit the benzene content in gasoline to 1 VOI%and the content of aromatics to 42 VO170by January 2000. A further reduction of the aromatics content to 35 VOFYOis scheduled to come into force in 2005. Similar clean fuel legislation will be found across most of the developed world (e.g. US Clean Air Act Amendments of 1990). This has made an additional contribution to the aromatics supply, especially to that of the benzene. 102 , ,. Liquid-Liquid-Extraction Extractive Distillation Non-aromatics Non-aromatica r SOIven Extractor Sxtractive Distillation Solvent Column ,. -n Light Hydrocarbmra ,.,. ,’ Feed . Feed ● Solvant + Solvent + Aromatlca Aromatica . ,.. , Figure 3: Principles of aromatics separation processes ). ~..’, ,., , ● On the other hand. the alobal Dom.dation is arowina and at the same time an .‘ ‘ increase in per capita co~sumpt~on of plastics ~an b~obsewed. Based on this, a Iong-term increase in the demand for aromatic basic chemicals can be expected. This forecast is impressively underlined by the observation that the demand for the most important petrochemicals benzene and especially p-xlylene is still growing in a recessional environment. ,,, Principles of Aromatics Separation Processes The separating of aromatics from non-aromatics from coke oven light oil, pyrolysis gasoline, and reforrnate is carried out mainly by liquid-liquid-extraction or extractive ,,, ., ,, ,’ ., ‘ distillation. .,. , .,. ....!., On the right side of Figure 3, the principle of the most widely used licwid-licwid- .. ’.,,. ,, .,, extraction (LLE) technique is shown. It combines an extractor with an extractive ,., ., distillation column. ,.. ..-, , .. ,, ,,. , ‘ The hydrocarbon feed is introduced into the lower section of the extractor flowing ,.,, ,,. upward counter-currently against the solvent which is flowing down from the top of ,,..,...::.,.::... ,-. . the column forming a separate phase. The solvent extracts the aromatics from ,, ..-.-,... ., !,,,...,-,.!. hydrocarbon phase. However, the separation is not ideal; light non-aromatic . .,, ‘,, . hydrocarbon impurities are co-extracted with the solvent phase due to the solvent ,.,,. /,., . selectivity favouring light non-aromatic components more than heavy non-aromatic .. --:,,, -., components. The bulk of the non-aromatic hydrocarbons leaves the extractor .,, .,,. ,.. overhead as raffinate. ,, -‘; ‘:. .,.; .:. ,,, ., .. ,,. ,.,..: .:,, : ~,. : ,,’. ..,, - ,, ., .”. . ,’,.,,.,.- ,, 103 ,, ...., ,, .-A_-..— ,! .—-—.....-. ,,. ,,, ,.q.,>.. , ., -:> ,,,.,’,, +.. . .* The rich solvent containing aromatics and light non-aromatics is sent to the extractive stripper. Light non-aromatic impurities are removed overhead and are sent back to the lower part of the extractor, where they displace heavy non-aromatic components from the solvent phase. The bottoms product of the extractive stripper consists of pure aromatics dissolved in solvent. It is sent to a solvent stripper for the separation of the aromatics from the solvent by distillation. The LLE is efficient at recovering aromatics of high purity within a wide boiling range. Typically, benzene to xylenes can be extracted together. Low aromatics contents in the feedstock favour the LLE economically, because it does not require an overhead distillation of the non-aromatic fraction. The process principle of the extractive distillation is shown on the right side of Figure 3. In order to separate close-boiling components, the extractive distillation uses a solvent that creates or enhances the volatility differences between the components in the original feedstock. The solvent is a polar compound having a higher boiling point than the components of the feedstock. The hydrocarbon feed is introduced in the middle part of the extractive distillation column. The solvent is fed at the top of the column. By flowing downward the solvent preferentially extracts the more polar aromatics in the mixture allowing the non- aromatics to be distilled overhead. The solvent with dissolved aromatics is recovered as the bottoms product. The bottoms product is fed into a solvent stripper to separate the solvent from the pure aromatics phase. The extractive distillation requires normally a pre-fractionation, because it becomes less efficient with extremely wide boiling range feedstocks. However, the high purity of the aromatics products, the high flexibility, easy operation, and a higher capacity compared with a LLE of similar size equipment make the extractive distillation attractive. Process Configurations for Aromatics Production from Reformate Production of aromatics for the petrochemical industry As already mentioned, reformate contains high contents of toluene and xylenes, whereas the benzene content is low (see Table 1). For this reason, reformate is a . suitable raw material to produce xylenes that are used as basic chemicals for the production of polyesters, resins, and plasticisers. There are many different configurations for aromatics complexes. Figure 4 shows the process configuration of a highly integrated aromatics complex. The complex is designed for the maximum yield of p-xylene and benzene for petrochemical applications. Hydrotreated naphtha is sent to the continuous catalytic reformer (CCR), where paraffinic and naphthenic components are converted to aromatics. If the reformer is operated at high-severity in order to maximise the aromatics production, the C~+ fraction of the rerformate virtually does not contain any non-aromatic impurties. Therefore, only a C~C7 fraction is taken overhead from the reformate splitter and sent to the aromatics extraction unit. The non-aromatics free bottoms stream from 104 -s + L C& Raffinate ,NaphLha A I Benzene Column I I I 1 1 b I I H, I I I .’. I 4 . Dec, De C, w De C, + 4 p_Xylene 1 C,o+ &vln.9ucs Figure 4 Highly integrated aromatics complex for p-xylene and benzene production the reformate splitter is clay-treated and routed directly to the xylene recovery section of the complex. The raffinate stream from the extraction can be sent to the gasoline pool or other destinations. The extracted BT-mixture is clay-treated to remove trace olefins. Pure benzene and toluene is recovered in the benzene and toluene columns, .. respectively. ,,,., ‘ The C8+ fraction from the reformate splitter is fed into DeCa column. Mixed Xylenes ,., .<, of high purity are recovered overhead and are sent to the p-xylene recovev unit, .,’ where p-xylene is recovered by adsorption or crystallisation processes. The .: remaining C6 aromatics are routed to an isomensation unit, where the equilibrium distribution of the xylene isomers is re-established (i.e. p-xylene is produced from ., other C8 aromatics). After separating a C7+ fraction in the DeCT column, the product of the isomerisation is recycled back to the DeCa column. The bottoms stream of the DeCa column is sent to the DeCg column, where an aromatic Cg fraction is recovered overhead. Together with toluene from the toluene column, the Cg fraction is sent to a transalkylation unit for the production of additional xylenes and benzene. After removing light ends, the product from the transalkylation unit is clay-treated and sent to the BT fractionation columns. Xylenes ‘ are recovered with the bottoms product of the toluene column and are routed to the ‘“. ... ,, xylene recovery section. .,,.,, .,, , ,.-. Benzene Benzene Column c6- Clay Treater Reformate Splitter Raffinate ~—————— Reformate * lsom%~~ation 1- –– I –-l I u L_–”y–_~ / I ! 1–+ : To Gasoline Pool I + + + P Figure 5: Process configuration for benzene reduction in reformate ., For the extraction of the BT aromatics from the reforrnate splitter overhead, a LLE unit is usually employed. However, since the boiling range is limited, alternatively an extractive distillation unit can be selected, taking advantage from its flexibility and from the low investment costs. in case the reformer is operated with lower severity, Ce+ non-aromatics remain in the reformate. For this reason, non-aromatics must be separated, before the CE+ fraction can be routed to the xylene recovery section. In this case, it would be advantageous to sent the reformate completely to an LLE unit, because aromatics with a wide boiling range can be extracted from non-aromatics in one step. A Ce+ fraction containing pure aromatics would be yielded by a distillation from the aromatic extract. Benzene Reduction in Reformate Clean fuel legislation like the European Auto Oil Program limit the benzene content in motor gasoline. The refiner can select between three options to manage the benzene content in the gasoline pool. There is the possibility to avoid the benzene production in the reformer by reducing the severity or by modifying the reformer feedstock (i.e. removing the benzene precursors in the naphtha feed). Furthermore, a benzene rich fraction of the produced reforrnate can be hydrotreated converting benzene to cyclohexane. on the other hand, the benzene content in the reformate can also be reduced in an attractive way by extractive distillation as shown in Figure 5. A CG- fraction is generated from the debutanised refomate in a reformate splitter. A quantity of the 106 ,,, $, ’,,,,, ,{ .,, ; ,<.!,. .;! 1, .:, reformate is bypassed around the splitter and is sent together with the bottoms product of the reformate splitter to the gasoline pool. The amount of the bypass stream is determined by the degree of benzene reduction required in order not to exceed the maximum benzene concentration in the gasoline pool. The benzene rich C6- fraction from the reformate splitter is sent to an extractive distillation unit. The obtained benzene is clay-treated to remove trace of olefines in order to cope with the acid wash colour test. The product of the clay treatment is fractionated in a benzene column to produce high purity benzene. The recovered benzene can be sold on the petrochemical market. The raffinate from the extractive distillation unit can be blended into the gasoline pool. Alternatively, the raffinate can be sent to an isomerisation unit for further improvement of the octane number. Process Configurations for Aromatics Production from Pyrolysis Gasoline Benzene Production from Pvrolvsis Gasoline ,, The amount and composition of the pyrolysis gasoline (pygas) from steam crackers ,, differ extensively with the type of the feedstock and with the severity of the cracking operation. Small amounts of pyrolysis gasoline with high benzene contents, but almost no CBaromatics are produced from light cracker feeds such as liquid natural gas or ethane. Naphtha and heavier feedstocks produce higher quantities of pyrolysis gasoline containing considerable amounts of C8 aromatics. Different from reformate, substantial amounts of aliens, olefins and impurities such as sulphur and nitrogen are found in pyrolysis gasoline, which must be removed before the pygas can be processed in an extractive distillation unit. Figure 6 shows a process configuration to remove benzene from the pyrolysis ,., . gasoline so that a benzene free octane blendstock for the gasoline pool is produced ,’ and pure benzene can be sold on the petrochemical market. “L ,, .,, , Pyrolysis gasoline from the cracker is sent to the selective hydrogenation unit, where .-, diolefines are converted to olefines in order to prevent the diolefines from polymen- C~-toGasolinePool * DeC6 DeC6 Full Hydrogenation selective + + and + Hydrogenation ~ OesuphurkaUon H, u Raffinate T v Ofl.gas H, on-gas v ~+ to Gasoline Pool. Figure 6: Process configuration for benzene recovery from pyrolysis gasoline sation. The C5- fraction from the overhead of the following DeCs column is blended into gasoline. In this way, the second stage hydrogenation unit can be reduced in size. The bottoms stream of the DeCs column is fed in second stage hydrogenation unit. Olefins are saturated with hydrogen and impurities like sulphur are removed by hydrogenation. The stabilised and sulphur free product is sent to the DeCGcolumn, where a benzene rich CIj fraction is obtained overhead. Pure benzene is recovered in the following extractive distillation unit. The non-aromatics from the extractive distillation can be blended into the gasoline pool or may be used as feedstock for the steam cracker. The bottoms product of the DeC6 column is sent to the gasoline pool as octane blendstock. Production of benzene and mixed xylenes from Lwrolvsis qasoline Besides the limitation for benzene, the Auto Oil Program limits also the aromatics content of gasoline to 42 Vol% by 2000 and to 35 VOIVOby 2005, respectively. With a process configuration as described in the previous section, a benzene free gasoline blendstock can be produced from pyrolysis gasoline. However, in most cases the blendstock exhibits aromatics contents which are considerably higher than the limiting values from the Auto 0)1 Program. As shown before, the existing benzene production plant on the basis of pyrolisis gasoline consists of a two stage hydrogenation unit and an extractive distillation unit. A further aromatics reduction can be done by adding a second step extractive distillation unit in order to extract also toluene and xylenes from the gasoline blendstock. The process configuration is shown in Figure 7. The bottoms product from the DeCo column is fed into a DeCB column. The overhead of the DeC8 column represents a C#C13fraction containing TX aromatics and non-aromatics. h is sent to a second step extractive distillation unit. The extract is splitted in a following TX column into toluene and mixed xylenes with high purity. The raffinate stream can be sent to the gasoline pool as aromatics free octane blendstock. High purity toluene and xylenes can be sold on petrochemicals market. Additionally, raffinate free of aromatics can be used as octane blendstock for the gasoline pool or can be sold on the market for refining products. If the pyrolysis gasoline contains only low concentrations of Cfj aromatics, toluene may be recovered alone from a CTfraction of the pryrolsis gasoline. Since there is only a limited demand for toluene from the petrochemical industry, the price of toluene normally is below the prices of benzene and mixed xylenes, respectively. For this reason, it may be attractive to add value on the toluene product by converting it to benzene and mixed xylenes. This can be done by installing a toluene dispropotiioning unit (TDP). 108 Benzene 4 + A Non.b’iromallcs DeC, to gasoline pool Column .,. % * ,, E c5- ,’ Hydro- ToIuene ?’ genaled Dlsproportkming 7 Pygas 0 DeCa Slabllizer Column ~ ~ f .,,. I 1 I MixedXytenes I ,,, ,”, C,+ -Producl ,, I ., Figure 7: Production of benzene and mixed xylenes from pyrolysis gasoline Toluene from the TX column is fed into the TDP unit. The product is stabilised and sent to a benzene column, where produced benzene is taken overhead. The bottoms stream of the benzene column is recycled back to the TX column, where the produced mixed xylenes are separated from the unconverted toluene. Only benzene, mixed xylenes and an highly valuable octane blendstock for gasoline are yielded with this type of aromatics complex. Lurgi’s processes for aromatics separation AROSOLVAN@ - Lumi’s Iicwid-licruid-extraction rxocess The AROSOLVAN@ process works on the principle of liquid-liquid-extraction combined with an extractive distillation as described above. It recovers pure benzene, toluene, xylens, and Cg-aromatics from reforrnate or pyrolysis gasoline at the same time. It is especially economical for feeds with low B and T content, since no pre-distillation is required. The AROSOLVAN@ process mainly consists of three columns the extractor, the extract recycle column, and the solvent stripper (see Figure 8). The feedstock is -,. charged to the feed stage of the extractor. Solvent is fed to the extractor top. Flowing downward the extractor, the solvent is loaded with aromatics and light non- aromatics. The raffinate leaving the extractor top still contains some solvent which is removed by scrubbing with water. The rich solvent from the extractor bottoms containing aromatics and light non- aromatic impurities is sent to the extract recycle column. The column overhead product consisting of light non-aromatics and light aromatics is recycled back to the ,., 109 Raffinate I Recvlce PJ ‘X’rac’ —& Extractor — . Feed 4 iz- Figure 8: AROSOLVAN@ liquid-liquid-extraction process bottom stage of the extractor, where it displaces the higher boiling non-aromatics from the solvent phase. The extract recycle column bottoms product consisting of solvent and the pure aromatics is separated into aromatics and solvent in the stripper under vacuum. The lean solvent is returned to the top stage of the extractor. Recovering its heat for the process, the lean solvent is cooled down by heat exchange in the extract recycle column reboiler and with other process streams. The BT or BTX extract from the stripper overhead is fed to a scrubber, where residual solvent is recovered by adding water. Subsequently, the BT or BTX mixture is sent to a downstream distillation section where pure aromatics are obtained. The solvent loaded water from the scrubber is sent to the water column, where solvent is recovered and recycled to the process. A water stream is obtained as overhead product. It is recycled in a closed loop back to the scrubber. A mixture of approximately 60% N-methylpyrrolidone and 40 % glycol is used as the solvent. This mixture shows an optimum balance of the solvent properties required for the extraction process. ., The AROSOLVAN” process is capable of processing feedstocks with wide boiling ranges. Very high purity aromatics are recovered with excellent yields. The AROSOLVAN@ process can be designed highly flexible regarding the composition of the feedstock and the heat sources available. The extraction process operates at 40 – 60 ‘C and low pressure. Thus, depentanising of the feedstock is not necessary. Table 2 shows typical utility consumption figures of the AROSOLVAN@ process. It is advantageous that only medium pressure, but no high pressure steam is required. 110 Estractive - Distillation Column Solvent Stripper ., 1— ,’ t’ Raffinate Column 4-Pure ArOmatiss- ‘“’ Component Aromatics cut a 4 * ..,.. — + ., solvent + rucmauc +’ -.’, Component ., ,, -., ,., Lean Soh’mt TI .,. Figure 9: DISTAPE#’ extractive distillation process ,,”, DISTAPE~ - Lunai’s extractive distillation process The DISTAPEfl process works on the principle of extractive distillation as already ,, mentioned above. It is especially effective for recovering individual aromatics (i.e. benzene, toluene, xylenes, or ethylbenzene) from a heart cut containing the aromatic compound to be obtained. However, the DISTAPE~ process allows also ,. to recover BT or TX aromatics from broader fractions. In this case, a downstream fractionation of the recovered aromatics extract is required. ,. The DISTAPE# process comprises two columns, the extractive distillation column ~ and the solvent stripper in which the solvent is circulated. An additional small ,., column serves to recover solvent from the separated non-aromatics (see Figure 9). .,. , .,’ The pre-fractionated feedstock, i.e. the heart cut with the aromatic component to be ,,, recovered, is routed to the middle section of the extractive distillation column. In the .“, presence of the solvent the aromatic component and the non-aromatics are separated in the column. The aromatic component passes together with the solvent to the bottom and is ,,,. ~ routed to the stripper. There it is separated from the solvent under vacuum. The -,. overhead aromatic component then leaves the plant as pure product and the solvent ‘.’, ‘. is circulated back to the extractive distillation column. High heat utilisation is obtained by intensive heat exchange of the circulating solvent. The necessary additional heat is supplied by medium pressure steam at 12 -14 bar making the process especially economic. Table 2: Utility Consumption of Lurgi’s aromatics separation processes Utility Consumption AROSOLVANW DISTAPE~ (per metric ton aromatics) (per metric ton benzene) Steam (MP) 0,8 mtl) 0,6 m?’ Cooling water 24 ms 24 ms Electrical power 11 kWh 4 kWh Solvent loss 0,05 kg 0,01 kg Noles 1) Medmm pressure steam of 12-16 bar 2) Medium Pressure steam of 12-14 bar The non-aromatics still containing small quantities of solvent are obtained at the top of the extractive distillation column. The solvent is recovered in the raffinate column and returned to the solvent recycle. The DISTAPE~ process allows to produce high purity aromatics with excellent yields. Aromatics can be recovered vety economically due to the simplicity of the process and the low utility consumption (see Table 2). Only moderate pressures and temperatures are applied in the plant. The working temperature on the product side is 175°C in maximum. This allows the use of reasonable and readily available medium pressure steam as heating medium for the column reboilers. Due to favorable, non-corrosive characteristics of the solvent N-methylpyrrolidone, the overall plant can be constructed of carbon steel. Recent developments of the DISTAPE~ process The DISTAPE#’ extractive distillation process is subject to continuous improvement and optimisation. Recently, considerable progress was made in reducing investment costs, required plot area, and utility consumption figures. The developments can be summarised as follows (see Figure 10): (1) The raffinate column is mounted on the top of the extractive distillation column. As a result, the investment costs and the required plot area can be reduced. (2) Simultaneously, the operating pressure in the combined raffinate and extractive distillation column was increased so that the column overhead pressure is slightly higher than the atmospheric pressure. With it, the actual vapour volume flow in the combined raffinate and extractive distillation column can be decreased. This results in a smaller column diameter saving investment costs. However, the amount of pressure increase is limited by demand for heat integration, since the pressure increase leads to a temperature increase in the column bottoms. The temperature increase can only be allowed to such extent that the latent heat from the lean solvent can be used for reboiler heating in the extractive distillation section of the main column. (3) With the improved design structured packings will be employed in top sections of the combined raffinate and extractive distillation column as well as the solvent stripper. As a result, the column diameters can be reduced saving investment costs. 112 —A C.w. ~ I ➤ Combined Raffinata Rafflnate and EDC Column Solvent Stripper Benzene C.w. Feed + M. P. Steam Condensate Rich solvent I I Lean solvent Figure 10: Advanced DISTAPE)@ process Furthermore, structured packings show a considerable lower pressure drop compared with conventional valve trays. This is a clear advantage regarding the solvent stripper which is operated under vacuum conditions. Due to the lower pressure in the stripper bottoms, the residual benzene content in the lean solvent can be reduced significantly. Consequently, the solvent circulation flow can be reduced. This leads to savings in investment costs for the extractive distillation column, pumps, piping etc. as well as to savings in costs for utilities. (4) Further advances comprise among other things the development of a computer- based process simulation model for the non-ideal extractive distillation system, improvement of heat-integration by application of Pinch-Technology, or the development of an advanced control system. NMP - the solvent N-methylpyrrolidone (NMP) is used in the DISTAPE~ process and AROSOLVAN@ process as solvent. NMP features vefy advantageous characteristics for the application in LLE and the extractive distillation. The essential criteria are: ● NMP exhibits a high selectivity and an excellent capacity that leads to low solvent circulation streams and therefore low utility consumption. ~. ,,‘ ,’ ‘ ,:, “.”, : ,. 113 ,’ ., ,,,. ,, , . ,,,.,,. ‘,. ,,,., i, ,,, ,. ;.,,-, ,, ,, .,’ .,:’ . The low boiling point of NMP leads to low working temperatures in the plants. Medium pressure steam can be used for the solvent recovery. ● No tracing of solvent containing piping is required due to the low solidification point of NMP. ● NMP shows excellent chemical stability and thermal resistance. . Due to the low corrosiveness of NMP, the plants can be constructed of carbon steel. No corrosion problems were observed yet. . NMP is non-toxic. Since a closed circulation system is applied for the solvent, there are no environmental problems. ● NMP is readily available all over the world at reasonable cost. ● Due to its outstanding characteristics, NMP has also been used for many years as solvent for other processes like butadiene extraction or gas desulphurisation. Conclusions Extraction technologies for the separation of aromatics from non-aromatics have been used for nearly 40 years. It can be observed that liquid-liquid-extraction and especially extractive distillation are flexible technologies that can be adapted to challenges like the recent demand for the reduction of benzene and aromatics in gasoline due to clean fuel legislation. Continuos development of the extraction technologies leads to simplified processes, thus saving investment costs and reducing utility consumption. Advances in process simulation techniques allow to design the extraction plant tailor-made for the feed composition and the product specifications. Advanced plant control systems allow to operate the plants with stable process conditions close to the maximum capacity. These improvements give the opportunity to recover aromatics in a cost-effective way. Therefore, it can expected that extraction technologies are also successful in the beginning of the next millennium. 114 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erfangen 1999 ,., ,.. ,., ,, C. Dembny Linde AG, Hoellriegelskreuth, Germany The Role of Pyrolysis Gasoline from Steam Crackers beyond the Year 2000 -,, , ,,’, ,, :,’ !:, -, ,, ’..> 1 Introduction ,., r., , “Pyrolysis Gasoline” (referred to here as “pygas~ is a complex m“~ure of numerous hydrocafions wilh C-numbers ranging between 5-10 consisting mainly of aromatics (BTX) and, to a lesser extent, olelins and paraffins. Pygas occurs as a by-product during the thermal cracking of hydrocarbons to ethylene. In the past, ethylene producers have paid less attention to the potential of an optimized usage of such side-streams. [n general, the pygas product was blended to the motor gasoline pool with the benefit of upgrading the Iattefs quality. This situation will now change because of the legislation to be passed in near future “European Auto Oil Program I + II-. These regulations will tighten the specifications of motor gasoline concerning maximum contents of aromatics, benzene, olefins and other substances. A summary of the expected regulations is given in Figure 1. As a consequence, a mnsiderable surplus of benzene is expected (European Chemical News, 28 June -4 July 1999, p. 11). Due to these legislative constraints, alternatives to the common processing methods will be required. The pygas production in steam crackem is influenced by a number of parameters . Type of cracker feedstock . Origin of liquid feedstocks . Pygas processing units ● Furnace operation parameters. In the followimr emphasis is put on the influence of the cracking process on pygas production. The reactions in the c&cking furnaces, the heart of a steam cracker, themselves depend in their performance on different physical parameters defining the cracked gas composition. Important parameters as . Cracking severity . Residence time . Operating Pressure . Steam-to-hydrocarbon ratio (S/HC ratio) are interconnected and influence each other. In the following, the influence of the above mentioned parameters will be discussed in detail. At the end of the arlicle, an outlook on upcoming developments will be given highlighting various tendencies influencing the furnace operation and Ihe effects on pygas production. 2 Pwras from Steam Crackers The main task of a steam cracker (also called olefins or ethylene plant) is to produce ethylene and ... propylene. From that point of view, all other products are side-streams and by-products such as hydrogen, methane fraclion (mainly used as fuel for the cracking furnaces), C4’S, pyrolysis gasoline >,,, :: 115 DGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 .-,., .-, and pyrolysis oil. The steady pressure 10 improve the performance of steam cracking units forced ethylene producers to find economical ways to handle also these side-streams, especially now in the face of the new regulations. Most interesting is the reduction of aromatics and benzene content in motor gasoline fmm an average content of 2.3 vol.-Yo to 1.0 VOI.-YO.In other words, a surplus of approx. 750.000 mta benzene production in Europe is expected. Ttris surplus is assumed to increase further if in 2005 an even lower limit to benzene should be passed. This has forced a reaction in the refining and steam cracking industries to develop strategies to reduce or minimize side-streams with low revenue on the one hand and - if possible - increase ethylene and propylene production on the other. Figure 2 shows an example derived from cracking a typical full range naphtha feedstock at a cracking severity of P/E = 0.55 kg/kg (P= Propylene; E= Ethylene; definition of cracking severity see section 5, ‘Influence of Cracking Severity”) at furnace outlet. Main components are aromatics (like BTX, styrene and ethylbenzene) which account for more than 60 wf.-~. of pygas. The remaining components belong to the olefins (approx. 25%) and n-ho-paraffins. Within the aromatic fraction, the benzene content can vary considerably depending orI several parameters to be discussed later. In the example of Figure 2 aPPrOximateIY one half of the tOtal aromatics was already conlained in the feedstock and remained unchanged during pyrolysis. The other half was formed by chemical reactions in the cracking furnaces. The following chapter gives a closer insight into the dependence of pygas and benzene production on the processed feedstock. 3 Influence of Feed Characteristics on Pvqas Production Modem steam crackers can be designed to process various feedstocks ranging from ethane up to hydrocracker residue (HCR). However, in most cases - with respect to the investment cost, other economic constraints, basic requirements like feedslock availability and strategic aspects - the design is based to either gas or liquid feedstocks. The pygas production strongly depends on the feedstock being processed in an ethylene plant. In Figure 3 an ovenfiew is presented with typical product composition figures (wt.-%-yields at battety limit) for different feedstocks. [f high ethylene, low pygas and low benzene production is desired, then ethane is the ideal feedstock followed by propane; liquid feeds range in their yields at around 30% ethylene and approx. 20% PY9as. The benzene content of the pygas of liquid feedstocks is less than 1/3. With increasing molecular weight of the feedstock, the main cracked gas components are also shifted to higher molecular masses. As a consequence, heavy feedstocks have significant pyrolysis gasoline and oil production at decreased ethylene/propylene yield. An exception in this summary are hydrocracker residues which can reach ethylene yields similar to naphthas but at significantly higher py oil production. Experience shows that the cracked product distribution is predominantly determined by the feedstock and its composition (Figure 4): n-paraffins are premium feedstock for high ethylene yields, followed by iso-paraffms and naphthenes. Only low yields can be obtained with olefins. Aromatics have practically no ethylene potential. For a better understanding of these differences, a simplified look at the molecular mechanism of thermal cracking is helpful (Figure 5): When heating hydrocarbons to temperatures of approx. 550-650 “C, radicals (high reactive particles) are produced which attack other feedstock molecules under removal of an H-atom. The resulting “feedstock radical” has different possibilities to react furthen 116 ‘., . ,,.,’-!. ‘+”; ,$::!: “,, , ., ;j .,,, “ ,’, ,,, 5 ,,,,’. ~., ,. .... ,, ,., ~ ,,, .<; .,:’,: i ‘,’., ., . In the case of ethane feed the resulting C2H5-radical reacts very fast to ethylene due to the low activation energy of its decay readion. In the case of tiquid feedstocks with longer C-chains (e.g. 5-10 C-atoms per molecule in case of naphtha), chain breaking of the intermediates occurs to predominantly form olefins and other radicals e.g. C3H7, C4H7 and others. Additionally, other reaction types (dehydrogenation, additiorr~ condensation reactions etc.) are taking place which lead to higher olefins and aromatics accounting for the higher pygas yields. Consequently, high ethylene and low pygas yields are produced by: gaseous feedstocks, and if liquid feeds are to be processed: high paraffin and low aromatic content give better performance. Besides ethylene, propylene is a valuable product of steam cracking. As already discussed, tiquid feedstocke produce significant quantities of propylene. Figure 6 indicsles that the European feedstock slate is dominated by naphtha with more than 85% shara approx 8% of the total feedstock is gas oil .>’ and minor portions are gaseous feedstock. [n the global feed slate (Figure 7) the naphtha share is not ,,:,, so pronounced because approx. ss~o of the US crackers are fed with ethanelpropane. ,,, , ,.. , ,. So if being restricted to process tiquid feedstock, naphtha cuts from crudes containing low aromatics -, -,.. (and therefore benzene) should be prefemed. In Figure 8 and Figure 9 an ovewiew of different crudes ,., , is presented to give an insight of the varying contents of especially benzene in the crude and therefore in the naphtha fraction. Crudes originating from Africa, Nearlfvliddle East, Latin and South America in general have benzene contents of up to 0.2 vol.-%, whereas crudes fmm the Far EasffChina, Europe and Russia more often contain up to 0.5 vol.-%. Some Australian crudes reach more than 2 vol.-%. 4 Processing of Pvraas in Steam Crackers Once the cracked gas has been “frozen” by quenching, the yields are practically f~ed. Nevertheless, in ethylene plants, several processing steps can be integrated into the separation train. By this means specific customer demands are met for the desired product distribution at battery timit and the revenue is increased at the same time. Additionally, significant savings in total feedstock consumption are achieved by the recycling of different streams. Figure 10 gives an overview about the possibilities to ..’ .,’ produce chemicals and/or to use certain pygas fractions as recycle streams. In general, the individual ‘,.,.., +,- ‘,. design of pygas processing is mainly depending on: . . /., - ,,, . .,,, . desired products ,,. - ,’ ,, ● costs for feedstock and utilities 9 investment costs In the following the possibilities for pygas processing are discussed in more detail. L“, :,! C5 seDarafion and hvdroaenafion ,, .,.,,, .,. . ,.-, ,“, . . As already indicated, paraffins are a valuable feedstock with regard to ethylene potential. [f no C5 .,,,, ... olefirrs/diolefins recovery is desired, the C5 diolefirrs are usually converted to olefins in a pygas .,....’.. .,.. ,:. selective hydrogenation sIep followed by C5/C6 separation before the C5 fraction is recycled to the .:. ,+,’: .. ~ furnaces. An alternative is an additional 2ti hydrogenation unit to convert C5 olefins partially or ,., . ...,. ;, ...... ‘ :,:..”. ,, completely into tiso-paraffins depending on the ,..J “, ,, ..,. ● remaining available hydrogen production after C2/C3 hydrogenation and .’, .:, ,,. , . economic evaluation of hydrogen. .. -,, When recycling a C5 stream as feedstock to the furnaces it is essential that no more diolefins are present in this fraction, otherwise severe coking and fouling in the furnaces and TLEs occurs. The pygas hydrogenation unit is flexible in that it can be combined with butadiene hydrogenation resulting in a C4/C5-hydrogenation for diene removal. :.:4 ,, , ,,.; ,/ ,. .,; ~, ,:, ,, ,,,‘.’%.,,, ,., i’ ,. 117 . ~.:,.,. ,, ,,. . Another recently reporied option is 10 convert the C4 and C5 fraction by fluid catalytic cracking (FCC) into oletins (mainly propylene). This option will became more viable if propylene shortups influence the economics. C5 dienes recovery lsoprene/Cyclopentadiene/l .3-Pentadiene have a broad spectrum of end-uses including .synlhetic rubber, hydrocarbon resins, thermoplastic elastomers, unsaturated polyester resins, pesticides, flame retardants and other chemicals. The recovery from pygas is achieved using extractive distillation. Benzene segaralion The separation of benzene from pygas and its recovery for other syntheses or hydrogenation is a promising method of reducing benzene in the motor gasoline pool. Analogue to the recycling of C5, the remaining C6 raffinate after benzene extraction is a valuable feedstock to be recycled to the furnaces. Dealkvlation of C7-C8 aromatics If the benzene production from the pygas fraction is to be maximized al the expense of C7-C8 aromatics, a dealkylation (hydroalkyiation) process can be applied. Using thermal or catalytic dealkylation in presence of hydrogen, the aromatics tohrene, xylenes and ethylbenzene are conveded into benzene, a methane/ethane rich overhead stream, oil components and recycfe hydrogen. Before being routed to the catalytic dealkylation process. olefins must be removed from the feedstock. C9/Cl O recovery The pygas fraction contains C9/CIO hydrocarbons consisting of aromatic molecules linked to an oletinic or a cycloolefinic rest; e.g. vinyltoluenes, indene, methyl-indenes. These chemicals are used for the production of resins. The boiting point range of these components is between 150 and 20fYC, just at the borderline between the py oil and pygas thus forming the heaviest fraction of the pygas. The quantity of this type of hydrocarbons strongly depends on the feedstock quality, cracking severity and residence time in the cracking process. Yield prediction is not as accurate as for the main products. The C9/CIO share in the pygas is recovered by distillation from the heavy gasotfne fraction. If no recovery for production of chemicals is desired, alternatively C9/Cl O separation and use as liquid fuel is possible. Besides the light pyrolysis gas oil fraction, the C9/Cl O-fraction can be separated and used as fuel for the cracking furnaces if fuel gas export is desired or there is a shortage. 5 Influence of Furnace Operation Parameters on Pvqas Production During the process design of a steam cracker, the furnaces are optimized for the desired feedslock scenario and all other requirements. For maximized olefins yields, the chemical and physical principles of the cracking process have to be taken into consideration. For a better understanding of furnace operation two thermodynamical statements shall be reviewed: Paraffins are (at ambient temperature) thermodynamically more stable than olefins Above temperatures of 750”C oletins become more stable than paraffins. In other words to transfom paraffins into oletins in a finite time, the feedstock has to be heated to temperatures above 750”C. But, at these temperatures, oletins are very readive and will undergo further reactions (e.g. condensation and addition reactions; see Fig. 5). For very long reaction times, oletins will react to carbon and hydrogen as these elements are more stable than oletins. As a result, the reaction must be kept under “kinetic control” to obtain olefins with reasonable yields. The feedstock must be heated up high and long enough to initiate reactions to oletins, and short enough to prevent the formation of undesired products. 118 ,4,.,,, . In industrial practice with respect to economic aspects these constraints translate into a number of requirements for the operational conditions: . High heat input at high temperatures . Short residence times to minimize by-producls . Limitation of HC partial pressure (dilution effect) . Effective quench of reactor effluents to fix composition. The fundamental physical parameters (which determine the cracked gas composition) for the above mentioned conditions are * Cracking Seventy (Coil Outlet Temperature) . Residence Time . TLE Outfet Pressure (TOP) . Steam/HC-Ratio In the following section the influence of these parameters will be investigated in detail. It must be stressed that in practice, the parameters listed above cannot be varied independently. In other words, it is not possible to tune one parameter with the others E!!I!2the total feed flOW rate being f~ed. The following example (given furnace and coil design; constant total flow rate) explains the correlation: if the TLE outlet pressure is decreased, the residence time decreases too, due to faster fluid velocity. -, Both effects (reduced TLE oullet pressure, shorter residence time) influence the cracked gas yield. For that reason, in the results presented below, the calculations are performed in a way to eliminate . . this mutual dependence. [n this way, effects on yields can be assigned to the respective variables. ,.’ It is impoflant to notice that applying the following results to industrial practice reqUireS aHeIltiOIl alSO to be paid to . cracked gas coil outlet velocity (=> high velocities cause erosion problems) ● coking (=> increased coking reduces the furnace run length) . pressure drop across the furnace (=> high pressure drop may cause feed supply problems) . high tube metal temperatures (=> reduced lifetime of radiant tubes due to increased carburfzation) Commercial scale furnaces operated on a typical full range naphtha were chosen for the calculations. In the following, the influence of parameter variation on ● ethylene yield . pygas yield/production, . benzene content in the pygas fraction is shown. ,.:. Influence of Residence Time ., .,, As indicated eartier, the residence time is an important parameter in determining the cracked gas composition. For a given coil geometry, the residence time is restrided within narrow limits. Figure 11 ., shows the influence of residence time on overall ethylene (overall ethylene defined here as C2H4 + 0.8=C2H6 + 0.9*C2H2) for residence times between 0.1- 0.5s and two different P/E ratios ex furnace. For short residence times at P/E=O.60 kg/kg, the overall yield is approximately 0.3 wt.-% higher, ,’ whereas at P/E=O.40 kg/kg the ethylene yield increases about 1.3 wt.-%. For decreasing residence time, ethylene is increasingly produced via direct conversion of the feedstock to ethylene and less via recycling of ethane, especially at mild conditions (high P/E ratios). ,’ Inversely to the increase in ethylene yield, the pygas production (Figure 12) is reduced by almost 3% at the given conditions (P/E ex furnace = 0.60 kg/kg; S/HC = 0.50 kg/kg) at 0.1 s. Additionally, the -’ benzene content in pygas (shown on the right axis) is reduced from 28 wt.-% to approx. 23 wt.-Yo. That means that a twofold reduction in benzene occurs at reduced residence time due to . ‘. ,. ~ . ‘,;.,+, ~,,’-”. : -. .,. . ., .,,, . . . :; ‘-””v 119 ,’, ,., - ,’. . .J :’(;.,,, * : ,- ;,. ..:.<. ,7A! reduction of total pygas flow rate and reduction of benzene content in the pygas fraction. The two effects add up to a 20% reduction in benzene production. In case of revamping furnaces, a principal decision to have radiant coils with shorter residence time is possible. In this case the installed few original long tubes with large diameters will be replaced by many tubes with smaller diameters and shorter total length. The residence time drops considerably as a result. Influence of Crackirw Severity There are different methods of defining cracking severity: A general parameter is the coil outlet temperature (COT). In the case of gas feedstocks, the term “conversion” (for instance ethane: [C2H6 content in feed minus C2H6 in cracked gas]divided by C2H6 in feed) is used. The cracking severity when processing liquid feeds is often described by the “P/E” ratio giving the wt.-% ratio of Propylene .’. . and Ethylene in the cracked gases. A high cracking severity is therefore related to a low P/E ratio and 6. . . vice versa. ‘., The calculations were performed for P/E ratios (ex furnace) between 0.4 to 0.7 kg/kg. Usually furnaces are designed for a certain band of severities but will not cover the whole range of 0.4 -0.7 kg/kg. If a cracking furnace is designed 10 operate at P/E=O.7 kg/kg, its performance can not be optimal for a P/E=O.4 kgrkg due to: design data (T, p,...) will be most probably exceeded detrimental effects on various parameters like run length, coil outlet velocity, pressure drop increase are to be expected. Nevertheless, to show clearly the influence of different P/E ratios on ethylene, pygas and benzene yield, in Figure 13 and Figure 14 the P/E ratio was varied between 0.40-0.70 kg/kg at a constant residence time of 0.1 8s. Due to the fact that in real furnace operation with increasing severity the furnace load has to be decreased, the residence time would change as well. In the following diagrams this effect is compensated by tuning the total flow rate and therefore a relative scale is chosen for the ethylene yield. A P/E-ratio of 0.6 kg/kg is defined as 100% for reference. A! higher severity the ethylene yield rises, whereas the pygas yield decreases but remains almost constant below P/E=O.5 kg/kg. In contrast to the pygas yield, the benzene content in the pygas increases at high severity from 15 wI.-Y. at P/E=O.7 kg/kg to 38 wt.-% at very severe cracking conditions (P/E=O.4 kg/kg). The benzene production for an decrease in P/E (increase in severity) from 0.6 to 0.4 kg/kg changes as follows under the “theoretical” conditions (constant residence time): +34%: in real furnace operation (with more reduced furnace load) : +20%. Influence of TLE Outlet Pressure fTOP) A low TOP favours trigher olefins yields. For an existing cracker the TOP is a f~ed parameter due to the defined pressure losses between TLE outlet and inlet cracked gas compressor. In case of revamps, such pressure losses can be minimized and the TOP be reduced e.g. from 2.0 bara to 1.6 bara. A lower pressure level is critical with regard 10 the safety margin above ambient pressure at inlet cracked gas compressor. When reducing the TOP, a higher volume flow rate (at constant feed flow rate) results which may require high investment costs for larger cracked gas compressors or rerating of the compressor. 120 ,..,.. ,,> ,“,,. ,,. ,, +,::,!,.!.: ,’.~ .!,... : -’ ,’,, ,1 ,!~ ,, / ,,~?.’. ; ‘ .. ,, ,.-’. .:. .>’:; , ,,, , Figure 15 and Figure 16 show the sensitivity of different TOPS to ethylene and pygas yield. The yield obtained for the parameter’s ,., ‘, ● Residence Time= 0.18s . Cracking Severity P/E= 0.60 kg/kg ,’ c S/HC = 0.50 kglkg . TOP= 1.80 bara (defined as basis 100%). As a result, the overall ethylene yield remains almost constant. At higher TOP more C2H4 is produced indirectly via elhane. The relative pygas yield remains also almost constant but the benzene content in the pygas fradion increases from 22% to 26% with increasing TOP. Within a reasonable variation in pressure, the TOP influencesthe pygas and benzene yield much less than the variation in the cracking severity. ,.,,... ;,,’,: The benzene production change for a TOP rise from 1.60 bare to 2.0 bars is as follows ‘ ‘: , $:,,,’, , :,. , ,. . “,,,,. , :: ! - “theoretica~ (at COflSk3fltresidence time) +34~o; ,, ., in practise (at constant feed flow): +IB’%0. :’1 r’, ,,, .f . -,. . .’. ,, ,:.,,.. ‘, ,,, , ,’ ,,, Influence of Sleam/Hydrocarbon ratio (S/HC ratio) ,; ,: ‘: ,, ( .;,’. :.,,; The S/HC ratio is usually optimized for a given feedstock, lower figures decrease the ethylene yield ,.,,,, ., ,0 ?,,. >1.:’,. / and increase the coking ratq on the other side a to high S/HC ratio is uneconomical due to high , ‘,.,,...,, ,,,,,,,,,,,,: process steam consumption and pressure drop. .,, .;, ,r, ,’ ,, .,..,,, :, ,,. . . . . ! In Hgure 17 and Figure 18 ethylene, pygas and benzene content in pygas are shown for S/fiC ratios ., ... ,;,,, of 0.4- 0.7 (for naphtha feedstock). The same basis as for TOP variation is chosen: ,,’ .,9 ‘. , . .,,,, ,,’. .-~ . Residence Time= 0.18s ,,, ! ‘ ., .,,,, ● Cracking Severity P/E= 0.60 kg/kg ,,. ,,1’” ,, .::: . S/HC = 0.50 kg/kg (defined as basis 100%) ,, !,. ~,., ., ,, :.!/ . TOP= 1.60 bara...... The S/HC ratio influences ethylene and pygas yield in a similar way as already seen for the TOP The ,., -,.’, .,. ,., ... overall ethylene yield remains constant regardless of the S/HC ratiw high steam dilution favours direct : ‘$. ,,:.. ethylene and less ethane production. Pygas is also constant, but the benzene content in pygas ,...,, ,,, ,,,, .. ,“; , decreases from approx. 25Y. to 21 % at high steam dilution (0.70 kg/kg). ..(< y: . ..,’;,,, ,, ,,,.,’.,,,,; ,,, ::..,,,,, .-, ,.”! $, SummaW ::,,: ,, ,,>, ., ?’),:.: “:!; ,; , ,,.., A short summa~ (Figure 19) of the investigated parameters residence time, cxackirrg severity, ,.,,. ‘- ,. ,,.,, ,,, ,;/ .,. ,’;’,.4 pressure and sleam/HC ratio shows different quantitative influences on pygas production: ,. ,,, ., ; ,>.<. , -, ,,, !.’,.,; . More influence on pygas + benzene yield => cracking severity and residence time ,. ., :’ :; ’:;j”;, ,:,,.. ,:, ● Less influence => pressure and S/HC ratio ,. ...’>.. , , ,,,+,(~-..l‘ ,; ., ,T,, $.,f.’ ,,<: ,. ..,...... , ,!.. 6 Summarv and Outlook for the Future ,; :):..-,\ ,. ‘,,;;: ,.,,>-: .. ,,,>, ~. ,- ,,,,,,: .,.;“.:;: ,::’:::: $., ,“’,: Pygas is one of the inevitable by-products of steam cracking. The main components of pygas are ,:l.j ),;,.,” ;\Y; aromatics, especially benzene and C5-C1 O olefirrs. Up to today the main portion of pygas has been )-1, :,::,”: fed to the motor gasoline pool. Due to tighter regulations for motor fuels in 2000 and 2005 ~European ‘‘ ,, ‘:.’!,p,, ., Auto Oil Program I and 11”)a surplus of BTX on the market is expected. . . ,,, , . <,,tf, ; . ‘, ..’~:,,+,“$.;, ... ,; -1; +,p:.,! ,, .:,:7 ,>,, ,’ , ,. ,.$’,., f .~,} ‘‘::,,/ .,, :, ,,,,, .,,,< ,,, ‘,;,; -. ,.,,! ,, ,., ,. .“.:;;,;: .’ :>’/ . ,’,1.~:, 121 .; ,. $}, .< :<“.” .- . ~, ;> .,.’,,1..,, . :, I ,,, . ,,, ,., ‘1 The fundamental parameters influencing steam cracker performance and therefore also pygas production are: . Type of cracker feedslock ● Origin of liquid feedstock . Pygas processing steps . Furnace operation parameters For existing crackers with tiquid feedstock the pygas/aromatics production can be sirrnificanlly influenced only by avoiding aromatics already contained in the feed. The furnace operation parameters within the usual design limits for cracking seventy, sleamfhydrocarbon ratio and TLE outlet pressure have only minor effects. However, in the case of grass root plants, new furnaces or furnace revamps, considerably higher ethylene and lower pygas production can be achieved by the installation of highly selective radiant section coils with shorter residence time. In addition, in the separation train, different processing steps can be integraled to reduce the pygas production and to extract aromatics. In the following, an attempt is made 10 estimate the effects of current developments on pygas production. Technical as well as economic aspects are considered. (Figure 20). # , 1. There is a general tendency 10 equip cracking furnaces with short residence coils. Tfris holds true . . for grass root plants as well as for revamp sewices. As previously shown, short residence time coils give a higher selectivity to ethylene and less by-products. This effed is more pronounced for liquid feedstocks than for ethane. On the other hand, stale-of-the-ad cracking at residence times below 0.2s (e.g. Linde PyroCrack l-l; see Figure 21) requires somewhat higher investment costs (due to a higher number of coils for a given feed flow rate) Ihan coils with residence times of approximately 0.5- 0.6s (e.g. Linde PyroCrack 4-2). 2. Long term developments in materials such as new alloys/ coatings for cracking coils and surface treatment for reduced coking enable cracking conditions at higher coil outlet temperatures and therefore facilitate eilher . more severe conditions (less pygas yield but higher benzene share in pygas), or . higher furnace load at constant severity (more pygas and benzene production due to higher feed rate) or . longer run length at constant feed flow rate (constant pygas and benzene production). 3. In general customers demand wide feedstock flexibility for modern crackers. Vvithin certain fimits (determined by the design of the furnaces and the separation train), ethane as well as hydrocracker residue and various intermediate cuts can be used as feedstock. An example of the resulting widely different cracked gas composition of these feedstocks was given above. Ethylene producers can manage to switch over in a short time between different feeds and adapt their production e.g. on a new feed supply or on other needs. As a result, the pygas production will change depending on the actual feed scenario. 4. A very important point, which influences furnace operation and therefore pygas production in the nexl years, is the forecast of ethylene and propylene demand. Wihin the next decade ethylene is expected to increase approx. 4-5?6 per year whereas the propylene prediction is for an annual growlh of 5-7%. These Figures translate into a . less severe cracking and/or . installation of additional propylene producing plants such as propane dehydrogenation, metathesis, deep calalylic cracking and others depending on site conditions and economical considerations. In the foregoing it was shown that mild cracking conditions result in increased pygas production. Despite a lower benzene yield at mild conditions, the total production of benzene will still increase as under mild conditions a considerably higher feed flow rate can be processed which overcompensates the reduced benzene yield at low severity condlions. 5. The predicted feed slate scenario for 2005 shows an increasing naphtha share (Figure 22). In addition, within the naphtha pool a tendency to heavier naphthas is predicted as a result of the 122 ;,,/,,;>, f..,,‘*J ., .,, ,, changing crude composition. This results from the depletion of light crude wells and these being replaced by crudes of higher specific gravity. Consequently, a higher specific gravity and final t boiling point for the naphtha fraction of these crudes will result in a corresponding shift in the cracked gas yield towards increased pygas. Summing up all trends and effects mentioned in this PaPer. a foreust of PY9as production in 9eneml ‘ is hardly possible and exlremely dependent the demand on ethylene, propylene and by-products, ,-. furthermore on economic considerations and local site conditions. -i, .,!, ,, ., ,. ,<,.”,,, .,. ,,” .,. . . 123 ., ;:,.- . ,%., ‘ .-,, . .-:. ,,, .+:-, ,.,‘ ,.. 7 Fiqures Fig, 1: Motor Gasoline Specificalbn 2000 Fig. 2: Example Pygas fraction from Naphtha ..#u52&!. Cracking Feedstcck Full Range Naphtha EU EU Conciliation Averago Crscldng S3verii. P/E. 0.55 ky%g Commission Parliament now 2nd reading Parameter yJ n-rise.Paraffins 10.6 S.lph.r [wt.-ppm] 200 lal 150 mono. Olefins 5,7 di- Olerins 18.6 Aromatics [vol..%] 45 35 45 Naphfhenes 3,2 Afomsb 61.4 Benzene [vol..%] 2,0 1.0 1,0 2.3 Oil 0.5 Olerins [,01,-%] 18 14 18 Total lW.O Fig. 3 Product W.tribulion of different Feedstocks Fig. 4: Ethylene Potential of Feedstock Components at Batte~ Limit .,As:&,. tm@%e- Elhane Propane Nwhtia AOO HCR n.Paraffins Products [W-%] 65% conv, 92% Conv. PC. 0s0 PI-E.0,50 PIE - 0,s0 6,0 2.0 3,0 2,1 2,0 lso-Parefflns %RY 5,8 270 !29 98 86 Eth$me 81.6 45.5 33,2 27,9 32.8 I Naphthenos m9$*n8 12 14,6 16,6 13,9 16,4 Pm.,”. 0:1. C4 hatirl” 36 46 loo 97 $!3 Olefhss [Pylulym 02soble 4.3 4.8 2t,2 20,0 17.1 I PYCIPI! Fuel Oi 0.3 1,3 2,1 16,5 1%,1 Aromatics Told 100.0 100,0 100.0 400.0 100.0 Oecreaaino Ethylene yield Fig. 9: Benzene Conlent in Crudes (NearlMiddle Fig. 10 Pygas Processing in Steam Crackers &;;I.. da. East) C5 Seprm.llm and H@roomallm + Recydlng as Feedslock C5 Ohm and Recwuy - Prockdm of Resins, t315fMlW4 Pestiddes BenzeneSepnrallm 6 FlOdJciim of Chcmrmls In Own-drmm UnRs Dealk)4aUm (HYckoalkWm) w PrOd.Mlm Or Benzene and Feed Resyde c9K1O Recovew e Pro&dim or RasIns c9/clo SCpareooa e Fuel In the Cmck.lno Furnaces Fig. 11: Ethylene Yield Fig. 12: Pygaa/Benzene Yield ,&S&e. LA%. as Function of Residence Time as FuncUon of Residence Time Feed$lcck Full Range Naphlhtu P/S cx furnace Feedsbdc Full Range Naphlha: PI’S = 0.60 kgkg ex furnace 22.8- -30 37.0...... ’”.“’’’’’’”’””””””””’””’: ~ / 36.0-— . ~ OvcfettPIE=o.40 j 22.6- -28 $ 25.0-— — — - ● % ~ 8 !224 - .26 $ ~ ‘“” / . . “ + ; 23.0- . . ‘“ “- Benzene 2 ,,. 22.2. -24 s 32.0-- /““ OveralkPiE :0.60 ! ,, i 31.0- I I 22.0 ~ 22 30.04 I 0.10 0.20 0.30 0.40 0.60 0.10 0.20 0,30 0.40 0,50 Residence mme [s] Residence Time [s] ,,. , . : .’ ... .,, . .’ ,, .,,,. :,. . .,...... , ...-,, ,, ,.:, ..,., ., .. . ., . ,... ,’ ,- .,: . ...’ ., -, , ...... , ..’.,’..,”. .\” ,,, . ‘...... - .:>’:. . ,,. ,.. . ,,, ..r:,- ‘.<.:. -, .. .. . . Fig 5. simplified ReactIon Scheme of Pyrolysis Fig, 6: European Feed Slate 1999 Reactions of Elhane vs. Naohtha ‘A::?!,?. &s&.. 2 ‘A Ethane ? % Propa~3 “k Butane \ ‘b Gas oil n % Others Elhtiene n = —. = Ace~me Pmpyne Aedflmt Coke ‘A Naphtha Ftg. 7: Global Feed Slate 1997 -. . Fig, 8: Benzene Content in Crudes ‘SAA ..L$&ti. 2 % Other 52 % Naphtha 3%B ., 8 % Propa 28 % , # Fig. 13 Relative overall Ethylene Yield Fig. 14: ‘Relatiie Pygaa/Benzene Yield &$&,. as Function of Cracldng Severity ‘43?fc- aa Function of Cracking Seventy Feedsloclc FIJI Range Naphlhu Rcsldence 71mti 0.18s Fecdd+ Full Range Naphlha; Residence mmti 0.18s 130- ...... ? 40 120- ...... Benzene b \ / 120- ~ \ 110- g ~ tlo - x E ~ \ / $ l@J - > 100- - g + FJ’gas a ! m 90 - S0 - \ 80 + + 10 801 0.30 0.40 0.50 0.s0 0.70 0.80 0.30 0.40 0.s0 0.60 0.70 0,80 P/Erallo (CXfurnace) [k@g] PiErallo (ex timace) IWO] Fig. 15: Relative Ethylene Yield Fig. 16: Relative Pygaa/Ben3ene Yield aa Function of TLE Outlet Pressure .&$&@. aa Function of TLE Outlet Prasaure &;2f.. FeedslacK Full Range N.sphlha; Residence 71me: 0.18x Feeds$xk: Full Ranoe Nnphlha: Residence Tlm.x 0.18 x PiE = 0.60 kghg cx fimnce PIE E 0.60 kgtag ex furnace 103- ...... -r 2s 102 ...... ~ ...... ~ 1 1! ~ 1o2- -25 ~ ..,’ .“ aenzene + . ...” 101- ,.. . -24 g i ~ ,W. — — . . . -23 ~ E ~ C@. ““..” -22 ~ 981 T 21 98 ! I I 4 1.30 1.50 1.70 1!90 2.10 1.30 1.50 1.70 1.90 2.10 Pfessure [bar a] Pressure [bar a] ,. ,.- .; .-’.... ~ -,., ,, .. ~ ,, . .. . -. . G m auazuq Se6Ad aUa@413IIW3A0 :Jala.uejed)0 M07 +- q6!H :UOemanuul txOJIoseamea [O@J owea~ 06’0 Oro 09”0 0s”0 Oro 0s”0 08’0 OL”O 09”0 0s”0 OP”O 0s’0 , L 86 -66 : ., . ... ,, .... ,,. . $ ..... - 00L ~ ... Ipa,w ...... ; !, .... >,BU,W~~ - 10I 921 ,,,1, ,,,,,J,,,,,.,...,,,,,,.l..,,,,,,,,,,..,..l,.,.,,,,,,,,,,,..l~o, : .,..,.. ,,,,,,,,, .. ,,.,,,,,,,,,,,,,...,..,,,.,,,,,., ,,,,,,,,.,,,,,,,.. 20, O>UUJIIJXO 8@+ 0s’0. ~d aa.dtq Xa Ex,loq 09.0.~d !S81”0:WJU03UWSOM!mnqd8N aOuoIJ Ilnj :w~aaj :s 81”0:WJu a3uaptsaM :BUl14d8Naouqj [[nj :~qspoo~ :,, -: ,. ,>,.,.’ s u-l In .,:. ..’. .“ ,,. , .,, , ‘, -— .’, ,.. ,( /-, ., ,’”. .,, :., ., .,, ...... ,,.,-,. .’,. . ,’. .’ -,’:,, .’. :.; .,. :, $..,,,..,.<.:;”.. ,,. , ,,, ,...... —) ,!, ,;, ,,tw, r’., ., ,f.,, .,.. .- .-,” ,-, ., .,., ... .,’.,,,,: ,., . .“}.. -.x. : ;, ,, ., .,, ,,L ,, ,., .:.,, .,, ,....~ 129 .....,..,,., :. ,~...-7... ,. -. ‘,, .;.-. , ,. .,. ,...,., , ,.~;,. . 130 ‘, ,, DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Eriangen 1999 !’ ,, J. Rault, P. Renard, F. Alario IFP, Rueil-Malmaison, France ,. ,,,,. Maximizing ParaxyIene Production with ParamaX .: Introduction .,. In spite of the recent reduced demand and low margins, demand for paraxylene is ~ . expected to recover by the year 2001. The average annual growth rate for paraxy- Iene is expected to be 6% for the new decade. This will require increased production “ capacity, the lion’s share coming from new, world-scale grassroots units. Today, with the availability of the ParamaXm aromatics technology suite, IFP is a leading BTX technology and service company. ParamaXm - The BTX Aromatics Technology Suite The processes incorporated into IFP’s aromatics potiolio provides customers with a complete suite of advanced technologies, unmatched in performance and cost effectiveness for all BTX produc- tion goals. Together, they make an attractive first choice when selecting the full complement of BTX technologies and services from a single source. The technolo- gies exclusively licensed by IFP in grassroots ParamaXpackages are: ● Aromizing - CCR reforming for aromatics production, ● Arofining - reformate purification for drastically reduced clay consumption, ● Su/fo/ane (Lyondell) - high purity benzene and toluene extraction, ● Motphy/ane (Krupp Uhde) - toluene and high purity benzene extraction, the latest technology to be integrated in the portfolio, ● MTDP-3 (Mobil) - toluene disproportionation to benzene and xylenes, 9 E/uxy/ - simulated countercurrent adsorption paraxylene separation, 9 Crystallizathn – enables the production of ultra-high purity paraxylene when com- bined with E/uxy/ in the hybrid version, 9 Octafining // (Engelhard) – CS aromatics (xylenes and ethylbenzene) isomeriza- tion, featuring the high selectivity Oparfs catalyst, ● Advanced MHA/ (Mobil) – xylenes isomerization with ethylbenzene dealkylation, using a new catalyst with improved activity and selectivity, ● TransP/us (Mobil) – toluene/Cg+CIO aromatics transalkylation, with the Proven ability to process high amounts of C9 and Cl Oaromatics. Different combinations of these technologies are employed to fulfill customer specific needs and objectives. A typical combination of ParamaX processes is illustrated prior to focusing on some key features of selected technologies. Typical Paraxylene Production Complex (Figure 1) A naphtha heart cut is first hydrotreated to remove sulfur and rutrogen prior to feeding an Aromizing unit yielding a full slate of aromatics compounds and hydrogen. Considerably more hydrogen is produced than consumed in the complex, making it a net expotter. .. The Arofining reactor, located upstream of the Aromizing effluent stabilization, hy- ‘,.’ drogenates undesirable olefin and diolefin compounds present in the high seventy ‘. reformate. This unit significantly prolongs clay (not shown) lifetime, resulting in lower ., ,,. .< ,.,., Figure 1 Typical Aromatics complex Raffmate k Hydrogen L r=- ~...–.-- RizF1’ + n A“”* ,’ . # ‘. c,+ C,o+ Heavy Aromatics ‘.6 1 + + Non-aromatic compounds are removed from the C7- fraction either in a Morphy/ane extractive distillation unit or in a Su/fo/ane liquid-liquid extraction process. The latter is preferred when very high purity toluene production is envisaged or when an exter- nal feed, such as hydrotreated pygas, is destined to go through BTX extraction. Morphy/ane is ideally used when toluene is internally transalkylated, and is offered ~h~esively by IFP for ParamaX packages through a licensing agreement with Krupp The deheptanizer bottoms are sent to the xylenes rerun column, together with the effluent from the TransP/us and the recycle from the isomerization unit. The C13frac- tion, taken overhead, feeds the Eluxyl unit, which produces high purity paraxylene from the mixed Ca stream at high recovery rates. The paraxylene-depleted stream then feeds the isomerization section for which EB-reforming type (Octafining //) or EB dealkylating type (Advanced MHAI) technologies may be selected. The effluent from the isomerization process, a close to equilibrium mixture of xylenes, is then re- cycled to the xylenes rerun and E/uxy/. The selection of isomerization technology is a function of the desired BTX production pattern and economical concerns. The xylenes rerun bottoms are sent to a heavy aromatics column where Cg and Clo aromatics are taken overhead for the TransP/us feed. TransP/us affords a significant increase in paraxylene production capacity by disproportionation / transalkylation of toluene and C9+ aromatics, yielding additional mixed xylenes and benzene. 132 ‘.’ ..,,, .. ,$. , .... ,,, ,, $:, .’ ,’ :, ,. ,., . .,,, Features of Selected PararnaXTechnologies The following provides some key features of selected ParamaX technologies. Aromizing is IFP’s state-of-the-art CCR reforming technology for aromatics produc- tion. The process employs the AR series of catalysts designed to maximize ,,, .’ ‘., aromatics yield and operates at low pressure and htgh seventy. The first Arcwniz”ng unit was started-up in 1977 and is still in use today. Arorn/zing is schematically represented in Figure 2. This technically advanced proc- ess for the production of aromatics from naphtha offers excellent product yields, low investment and operating costs and an exceptional on-stream factor. Here are some of its advantages: . The continuous catalyst regeneration system is fully automated, controlling all catalyst circulation and regeneration during start-up, shutdown and normal opera- tions. The latest generation CCR system, RegenC, features a fine control of moisture content in the various steps of the regeneration, a proof-burning zone and a fully independent control of each of the regeneration parameters. The safe and reliable RegenC operation affords improved regeneration quality, resulting in an ex- tended catalyst life and very stable catalyst performance. .. ,, .’-.., ● The AR405 catalyst is the third generation of Aromi~”ng catalyst featuring an im- proved selectivity towards aromatics owing to the uniform and ultra dispersion of active metals on the carrier. A harder support material provides excellent mechanical ,,,. ‘. ., ,.,;. ‘, resistance and assures low catalyst consumption rates. The catalyst also features a ,’ .,, ,;”~y,” ,’:, .--’, . , reduced platinum content, which has a direct impact on investment cost. ,>(,,:.8‘ ‘ ‘ ,,,; .; ● The side-by-side reactor arrangement allows simple reactor design, implementa- ‘ r,,.5.,,’,; ,,). ., , ~: .-.: :; tion and maintenance access and easy handling of thermal expansion issues, ,“,,.;Al,.,.,:,,,, resulting in cost-effective engineering and construction. ‘:.’,+:(:’‘;’., ,~,./,,’ ,,, ;:,;; Figure 2 Aronrizing process flowscheme “,,“.,, , >,,,, ,.,,:~1:..,,-,,.,,,., :,,. ,, Reactors and Heatera Regenerator Booster ,.. < Hydro~g~-Rich Compressor :. * ,;:,,’ ; ,. ..’, Separator ‘1 ., ,,’, ,.~“: Recovery System ,, ,,’ ?J+ * Recycle Aromizate to ;ompressor Stabilization ‘, ,,:. ,.. :,,., . .,, ,’, ,. ., f.’:. .,.,, .”$, . . .. . ~“ ,,.,,,. ‘,. ,, -,,... : ,’” ,,, 133 ;,$’::!<,:.. ,, .“.?.. .1: ,-,. ,,,, . ,., , .,,,, ...... ,., .’ < ..,7, >, ;..,,, ;,.: } .: .’;,. - ,., .. ..., .;:: .,”’ .,,. ; ! ,.>:,>> .,, The above features enable high aromatics and hydrogen yields with reduced utilities consumption. E/uxy/is based on the concept of simulated countercurrent adsorption. This tech- nique (Figure3) is based on the principle of liquid chromatography where a liquid, containing more or less strongly adsorbed molecules, flows through a fixed bed of molecular sieve. These molecules are then displaced by an even more strongly ad- sorbed desorbent, creating a concentration profile through the sieve bed. The sequential displacement, along the column, of feed and desorbent injection points and raffinate and extract withdrawal points provides for the separation of high purity paraxylene from the C8 aromatics feed. Figure 3 E/uxy/ simulated countercurrent adsorption . Paraxyiene o Ortho. and mela.xyienes, ehyl benzene Desorbenl time = O time = T 1 t Concentration 1 t Concentration — Ild -— — cle Px — — OX + MX “-kEB E/uxy/ includes a number of innovations that are incorporated in the world’s largest operating single-train paraxylene purification unit, affirming its leading edge position in paraxylene purification technology (Table 1). ● The microprocessor controlled ordoff valves system clearly distinguishes this technology, bringing a high degree of flexibility and a continuously optimized opera- tion. Dependence on large and expensive critical equipment such as a multiporf valve has been eliminated. A fully redundant architecture of the microprocessor system associated with a number of on/off valves leads to complete reliability and a high stream factor potential. ● On-line maintenance of an indhfidual valve during unit operation is achievable owing to the ability of the microprocessor to automatically detect valve failure. Con- sidering the stringent valve qualification procedure, this feature is an additional guaranty on the reliability of the system. ● On-line Raman spectroscopic analysis is applied to visualize concentration profiles along the E/uxy/ adsorber on a real time basis. This powerful technique al- lows a fine and continuous supervision of the adsorption section, enabling the control and optimization of the process. 134 Table 1 E/uxy/ Process Features f Highly Selective Molecular Sieve Advanced Dynamic Control . ● High purity (up to 99.9~o) Extraction is continuously optimized as feed ● Excellent mechanical resistance composition changes . ● Optimized particle size dktnbution Can manage valve failure or valve mainte- ● High productivity nance while maintaining production ,,,,., Individual On/Off Valve System Optimized Internals Design ... , ● Easy to operate with microprocessor control system ● Minimum dead volumes :-.;, . . ● ., ● Simple and reliable High blending snd distribution efficiency .:. . . ● Less expensiva than a large multiporl valve Ease of assembly . ,, , ● Piping length is minimized Segregation of clean and dirty streams ,. ,,, ,, ,.,,. , ● Well adapted for advanced control ,, ::.,’ ..,:. ,.> ,. ~ High Operation Flexibility ,,, ..,,,, A-., ,. ; ,,, , ● High or ultra-high purity operation ,., .1 I .-. / . Wide range of feed compositions “.,’. ‘.”. ,, .~.,,>,. ,. ..:,,: -,.,:,’,;.: ‘,, ,, ,. Eluxyl Stand-Alone and Hybrid Versions are in service today. The stand-alone .,,’ ,“,; version shown in Figure 4, employs two adsorption columns-; the hybrid version ? ‘, ’:,,, combines one adsorption column with a single stage crystallization unit enabling ,:,: >, , ;,, !$, .,,; flexible operations for new units or increased paraxylene production at existing ‘ : ,’:j; :,+ + , crystallization based complexes. ,. !,,( ,, ,.,,$ ., ,, Seven Eluxy/ licenses have been awarded for a total paraxylene production capac- ,, . ‘,,. ,,, :,; ity of over 2.5 million metric tons per year. There are four stand-alone Eluxyl units ., :.. varying in size from 180,000 to 600,000 metric tons per year and three Hybrid Eluxyl units. The 600,000-ton per year paraxylene production unit started up in December 1997 is the largest single train, stand-alone, paraxylene separation unit in the world. Figure 4 E/uxy/ unit flowscheme stand-alone version ,,.’ . . . Mixed Xylenes _ Heavy Ends Oesorbent ~;~qy:j A;~u~tio~n Extract Light Ends Raffinate Rerun Column Column Column Column .. Octafir?ingII is an ethyl benzene (EB) reforming type isomerization process, where ,’ ,,., ,, EB is converted to an equilibrium mixture of xylenes. As opposed to kinetically rapid xylenes isomerization reactions, EB isomerization proceeds via a CE naphthenic in- termediate equilibrium that limits the rate of reaction. Good control of the CfJ naphthenes content throughout the reaction section is therefore essential for opti- mizing process performance. A simplified flow diagram of Octafirring // is shown in Figure 5. Figure 5 Octafining II Reqcle Hydrcgen Off-Gas t t> Hydroger Yd--”I Make-up — Light Ends Eluxyl Ralhnate .’. . +, ., Octafining // has some very specific attractive features: . ● The light isomerate recycle column is a distinctive Octafining // feature that is enables precise and fast-acting control of the CE naphthenes recycle in the isomeri- zation unit. ● Oparis is a step-out isomerization catalyst representing a breakthrough in per- formance as compared to the previous generation of catalysts. ~ Superior performance and robustness of the Oparis catalyst were proven in semi-industrial, long-duration pilot tests employing industrial feeds. The robust- ness of the catalyst was evidenced by its response to induced severe process upsets. EthyIbenzene conversion in the range of 357. to 400/. iS achieved, while maintaining C6 ring loss per pass at less than 2?4..The paraxytene yield attains 93’Yo,expressed as the ratio of paraxylene produced to the amount of net mixed xyienes feed; this compares well the commonly achieved figure of 890A. z High activity. This excellent level of performance is achieved at a higher space velocity and under milder operating conditions providing savings from both capital and operating perspectives. Oparis is therefore ideal for catalyst replacement or revamp studies affording instant debottlenecking. Advanced MHAL The Mobil Technology Company has a strong worldwide market position in xylenes isomerization. At present there are 19 units using Mobil xylenes isomerization technology that account for over one-third of the world’s xylenes iso- merization capacity. Two additional units are scheduled to start-up in 1999-2000. The first of nine MHA/ units started up in Japan in 1990. 136 .,, ., ,, .:.-, :. ,.. ,,,.:; ,..’ ., ...,.t,:, ...... $,-~ ,’ . : ,’ ,. .,,:- ,: ,* Advanced MHA/ has only recently been offered for commercialization and is a con- . . . tinuation of this successful process with a significantly improved product yield. .. The process is distinguished by a unique dual-bed catalyst system (Figure 6). Ethyl- benzene dealkylation to benzene and non-aromatics cracking occur in the first bed, .... while the near-equilibrium isomerization of xylenes takes place in the second bed. ,,..,, Figure 6 Advanced MHAI process flowscheme Recyck+ Hydrcgen Olf-Gas n )’ ~— t I t+ ❑ EB & NA mnversion ,3 ,,,, Benzene Hydrogen .--.,. .- & Make.uP Toluene “1 ❑ Eluxyf Xylenes isomerizafion 2 $!LI Raffinate -1 Isomerateto Xyiene Recovery CMnnn m i5Y- ,, Advanced MHAI exhibits some key performance enhancements compared to MHAI. . ,,, ● High EB conversion per pass is achieved without detrimental effect on the cata- lyst cycle length. Conventional MHA/ typically achieve three-year first cycles. ,. ,,. Advanced MHAI is expected to reach even longer cycles, and owing to a higher ac- ., tivity, high EB deaikylation rates are obtained at milder temperatures. .. :-, ● Selectivity and activity in xylenes isomerization are also enhanced in Advanced MHAJ with reduced xylenes losses per pas:, and paraxylene approach to equilib- rium m excess of 100% throughout the operating cycle. ● High space velocity is one of the outstanding characteristics of MHAL Advanced MHA/ features a higher level of activity, leading to higher space velocity for an equivalent level of performance. ● Lower H2/HC is another key improvement in Advanced MHA/. Cost effective technology. Owing to the above features, Advanced MHAI is an ...-,,.,,, ~ttractive process both in terms of initial investment and operating costs. .’ ,, ,,, ,, ,“ .’ ,, ,, f, Trar?sP/us is Mobil Technology Company’s recently commercialized toluene/C9+ ,.,,.; ,,$l. :: ., .,., aromatics transalkylation technology co-developed with the Chinese Petroleum Cor- :J. ,!. :., ,, poration (CPC) of Taiwan. The first industrial unit, put on stream in June 1997, is in .’,,.,,,. , ‘ . ,1, CPC’S petrochemical plant in Lin-yuan where an existing unit has been switched to G ., .,. : TransP/us. This technology builds on extensive experience in toluene dispropor- ,, ‘. tionation that began in the mid-1 970’s. ,., . . ,,,.,.: ,’. 4< ,, ,. ,., . The TransP/us process utilizes a proprietary catalyst that has superior yield perform- ,,1 ‘.: ‘v ance. This is achieved by a catalyst design that maximizes desirable reactions such ,,. :., . as disproportionation, transalkylation and dealkylation and minimizes undesirable side reactions. In addition, TransP/us has the advantage of low capital and operating costs due to more favorable operating conditions relative to competitor processes. This results from the more robust nature of the catalyst. 137 .-— . .,..,,... .-, .’., The process flow scheme is typical for a vapor phase reaction in a fixed bed reactor. Toh.rene and C9-CIO aromatics are combined with hydrogen recycle gas and con- tacted with the catalyst bed where the disproportionation and transalkylation reactions occur to produce equilibrium mixed xylenes and benzene. By-product yields are small. The liquid separated from the reactor effluent is stripped to remove light ends and fractionated to recover benzene and mixed xylenes. The GEfraction is then sent to the E/uxy/ unit via the xylenes rerun column for paraxylene recovery. ● Feedstock flexibility TransP/us technology has the flexibility to process up to 60 wtO/~ of Cg+ aromatics in the fresh feed while maintaining long cycle lengths. In addi- tion, the robust nature of the 7iansP/us catalyst allows up to 25 wt’Yo Clo aromatics in the Cg+ feedstock enhancing the yields towards xylenes production as illustrated in Figure 7. Figure 7 Improved xylene yields via transalkylation of C~C1o aromatics 80 75 Xylene VieIds 713 on Fresh Feed, Wt% 65 60 55 50 0 10 20 30 40 50 60 70 80 90 100 Cg+ Aromatics in Fresh Feed, wt% ● Long cycle length: The catalyst stays on line longer than competitor processes. Typical performance of the fully regenerable TransP/us catalyst shows that cycles in excess of one year can be expected even when processing feedstocks having over 60 wtYo Ce+ aromatics. In addition, benzene product purity is better than 99.85 wt%. ● Low H2/HC mole ratio and higher weight hourly space velocity also charac- terize TransPlus technology and make it possible to build grassroots plants at lower capital costs relative to competitor technology. This also results in reduced operating cost. f. ,. Conclusion IFP’s ParamaX Technology Suite contains a complete set of technologies from a single source which can be configured to meet all BTX production needs. ParamaX is a continuously growing portfolio, incorporating the latest technological advances, such as Mor@ylane extractive dktillation, Advanced IW-IAI and Oparis, the latest Octafining // isomerization catalyst. 138 ,, DGMK-Conference ‘The Future Role of Aromatics in Refining and Petrochemist@’, Erlangen 1999 . ‘!, B. Lticke, A. Martin ., .,., Institute of Applied Chemistry Berlin-Adlershof, Berlin, Germany ,,. ,:,”,, ’.: ,, ,.! ,:;. ... ., Recent Advances in the Oxidation and Ammoxidation of Aromatics ., ,,.’,,., ,. ,, ,, “ ,/ -,. .,, ,,, 1) Introduction .’,.- ‘ .’, . . , .,..,,’::<.::.. , , ,.,. ...- ,., , Large-scale products, intermediates as well as fine chemicals can be produced .-,,.,’,.; .,, ., “,,’:, by selective oxidation of hydrocarbons. Thus, the catalytic oxidation of ,, ,,, ,.,,-...,,- hydrocarbons opens the way for the manufacture of different valuable products ...,-. ;,.,. from aromatic feedstocks. ,,,.-...... ,,:,,’ ...” ,:” :’,, ,,,, ,,?L, . . ..,. .; ,., ., - Because all the products of selective oxidations are intermediates kinetically ., ..., .,,. ,- ,, ,,, :,, controlled on mostly various parallel or consecutive ways to total oxidation the ~, . . -, choice of catalysts and of reaction conditions (particularly the control of .“. exothermy) is important to control selectivity and production rate. In the selective oxidation of hydrocarbons we can distinguish between three general reaction types: i) formation of oxygenates without splitting of ring-C-C-bondings (e. g. formation of phenoles or quinones) ,“. ., ,’ ii) formation of oxygenates under splitting of ring-C-C-bondings ?’ (e. g. formation of maleic or phthalic anhydride) iii) oxidation or ammoxidation of aromatic side-chains (e. g. formation of terephtalic acid, benzaldehyde or benzon.tile) , ,,, .,. Regarding the general mechanism of the catalytic activation of the oxygene we have to distinguish between the deChOphiliC oxidation on catalysts like V@S and ,’ .-. ., ., [email protected]@ (connected with splitting of C-C-bondings, e. g. in the formation of maleic anhydride from benzene) or nucleophilic oxidation on catalysts like Vz05- ... ,,- ,“ DGMK-Tagungsbericht 9903, lSEIN3-931850-59-5 139 TiOz (connected with the insertion of heteroatoms in the oxidized hydrocarbon, e. g. in the formation of aldehydes or nitriles by oxidation of side chains). 2) Oxidation of the aromatic nucleus without C-C-bond-splitting The direct oxidation of the aromatic ring without C-C bond splitting results in the formation of either phenols or quinones. The more important formation reaction is the direct hydroxylation to phenols. Particularly, the direct oxidation of benzene to phenol is a challenge of current interest. At present, phenole is produced by the cumene oxidation process resulting in equal amounts of phenol and acetone. The yield of phenol amounts to more them 90 %. The challenge is to replace this established process by the direct hydroxylation of benzene considering the fact that the demand for acetone will be decrease. There are numerous attempts to realize the direct hydroxylation of benzene either in gas- or in liquid phase using as oxidants: oxygene (also in mixture with reducing acceptors like hydrogen), hydrogene peroxide, nitrous oxide, and others, also anodic or photocatalytic oxidation. Examples characterizing the state of art are given in tab. 1. The selectivity of the phenolformation is frequently high with respect to benzene, whereas the benzene conversion, especially in reactions carried out with oxygen, is limited. The better results are given for special oxidants like N20 and H202. Particularly the application of nitrous oxide is of interesting in connection with the present manufacture of adipic acid. NZO is formed as waste byproduct in the oxidation of cyclohexanone with nitric acid to adipic acid and can be used at the same place to form additional phenol for the cyclohexanone production. A. K. Uriate, G. J. Panov et. al. pointed to the Monsanto situation with a production of more than 600 million pounds per year connected with the formation of nearly 200 million pounds of NZO. This amount could be equivalent to 400 million pounds of phenol [9]. Nevertheless, there are some problems for a successful commercialization, particularly concerning the catalyst selectivity with time on stream and the low phenol selectivity with respect to N20 [9]. 140 . . ,’- . ,, ., ,, :., The overall Process Performance as described by A. K. Uriate, G.-J. Panov et. al. on the base of pilot plant data is given in Table 2. Other direct hydroxylation processes using oxygene or peroxides are also investigated intensively. Vanadium containing catalysts - may be of more complicated structure (biomimetic catalyst system) - are the most promising one for further studies. The application of H20z for the hydroxylation of aromatics is more suitable for the ,., ,,’ formation of diphenols starting already with phenol. Thus, the Rhone-Poulenc ,,, process [1O] bases on the use of strong mineral acid as catalyst, (e. g. at 90 “C using phenol: Hz02: HCI04 = 20:1: 0,1). The disadvantage of acid homogeneous catalysts in the phenol hydroxylation is the product distribution with o: p – ratio of nearly 2. Zeolite catalysts give the possibility to overcome this problem. Enichem developed a process basing on a MFZ-type silicalite (TS-1) firstly in 1986 [11]. A comparison of the homogeneously catalyzed Rhone-Poulence- and the TS-l- catalyzed ENICHEM-Process is given in table 3 [cited from data of G. Belussi and C. Berego [12]]. .! ,. Also other aromatic compounds can be hydroxylated with HZOZ using TS-I or ,,.. , ~ similar catalyst systems. The reaction is limited to compounds having not electron withdrawing-or bulky-substituents. Furthermore side chains can be ... .,! , ., oxidized more easily than the aromatic nucleus resulting in the formation of preferently aldehydes. . ., 3) Oxidation of the aromatic nucleus connected with C-C-bond splitting This type of reaction marks a milestone in the heterogeneously catalyzed ,.. , oxidation technology. The first commercial production of phthalic anhydride ,, started in 1917 (vapour phase oxidation of naphthalene over vanadium oxide), .,: ,, ... Dufing the last fi~ years the raw material changes to O-Xylene. AIsO Maleic ,,. ,‘ ~~ ;.- anhydride can be produced by oxidative decomposition of the aromatic ring, the ,.. ,.. . 141 catalyst for the benzene oxidation is based on vanadium-molybdenum promoted mixed oxides. . The yield of maleic anhydride can reach 70 % of the theoretical amount. 6, ,. Nowadays, for the production of maleic anhydride benzene is replaced by butane ., as feedstock. There are three reasons for this replacement : butane as raw material is distinct cheaper, the formed maleic anhydride is free of quinoid traces and furthermore the process is more ecological with respect to the toxicity of benzene. Thus, benzene oxidation process are more of historical interest. Nevertheless, in oxidation processes of aromatic side chains (see next Chapter) the possibility of unselected pathways (electrophilic oxidation type) can lead to maleic anhydride as undesired by product. 4) Oxidation and Ammoxidation of aromatic side- chains Aromatic compounds having an easily abstractable H atom (forming a benzylic intermediate structure) can be oxidized or ammoxidized to aldehydes or nitriles, respectively o-dimethyl aromatic compounds (e. g. o-Xylene) can be transformed into acid anhydrides (e. g. phthalic anhydride). Oxidation of O-Xylene to phfhalic anhydride The oxidation is commonly carried out at 350 – 400 “C; the yield can reach 70 – 75 mol- %, main byproducts are CO/C02 and in only small amounts maleic anhydride (by direct oxidation of the aromatic nucleus) and tolualdehyde (by insufficient side chain oxidation). Commonly used catalysts are V~-oxides (particularly monolayer catalysts). Particularly in the preparation of the catalysts are some advantages are reported: influence on the surface molecular structure, dispersion of the active material, spreading of the active material on the surface [13]. 142 L,. ! ;,. ... ,. “,,..:} :.’~,,, ., ,“.:., ,,:, .,, , ...... ’” ,’ .- .’ ..,.,, .,, ,,., ,. -, Oxidation of foluenes to aldehydes Toluene and substituted toluenes can be transformed into aldehydes or nitrites depending on the composition of the reaction mixture containing ammonia or not. Essential steps of the catalytic reaction are . . - chemisorption of the aromatic reactant on the surface of the catalyst and H- abstraction forming a benzylic intermediate, - formation of a benzaldehyde like intermediate (also chemisorbed on the surface). After these steps the reaction pathway splits ott the benzaldehyde desorbes in the case of oxidation, and in the case of ammoxidation we have to consider further steps to discuss later. ,,, Nevertheless the oxidation reaction is not so selective than the ammoxidation ,.. reaction because ammonia can partly reduce the surface for an enhancement of ,.,,....,. selectivity [14] or may block sites for weakly bond oxygen, increasing the ,,-, selectivity of the partial oxidation compared with the one for total oxidation [15]. .,. The position, the size and the electronic effects of substituents significantly ,,’. influence the chemisorption step. The electronic interaction of substituted toluenes with structural welldefined vanadiumphosphates can be directly observed by in situ-ESR-spectroscopy the interaction of the substrates with the catalytic system is stronger in the case of toluenes with electron donating substituents (e. g. methoxy-toluenes) than with electron withdrawing substituents (e. g. halo-toluenes) [16]. Thus, the reactivity of toluenes with electron withdrawing substituents is significantly higher [171. The benzaldehyde-like species formed during the reaction on the catalyst surface is stronger chemisorbed in the oxidation than in the ammoxidation due to an additional hydrogen bridge bonding to surface Bransted-sites (which are blocked by ammonia during ammoxidation) [18]. ,,, ,’ ~.. . 143 This selectivity decreasing effect can be suppressed by adding of a non oxidable amine (e. g. pyridine) to the reaction mixture [18]. Nevertheless, there are also technological achievements in the oxidation of toluenes having electron donating groups. One example is the NIPPON- SHOKUBAI- and KAGAKU KOGYO-Process for the oxidation of 4- methoxytoluene to 4-methoxy-benzaldehyde on alkaline containing vanadium catalyst carried out at about 400 “C [19]. The selectivity mnserfing effect may be ensured by the controlled surface acidity by the alkaline. Examples for aromatic side chain oxidation are given in table 4. Ammoxidation of foluenes.to nitriles Ammoxidation refers to the formation of nitrites by oxidation of hydrocarbons in the presence of ammonia. The ammoxidation of lower hydrocarbons leads more to large scale products such as actylonitrile whereas the ammoxidaton particularly of substituted toluens may open the way for the synthesis of fine chemicals or of intermediates for tine chemicals syntheses. Vanadium-containing oxides (V/Ti; VISG; VIPIO and others) are preferently used as suppoiied, bulk or multicomponent catalysts. On vanadyl catalysts at least two closely adjacent VO groups are necessary for the chemisorption step and the subsequent running catalytic reaction on vanadyl units-containing catalysts. Additionally, high catalytic activity imperatively requires ,’- . effective exchange pathways for the electron transport. Therefore, active V/P/O- # , . . catalysts show that the altering electron density (at a discrete surface site) can be easily delocalized along coupled centres (detected by in situ-ESR [20]). The first intermediate species are already discussed in connection with the side chain oxidation. The ammonia to be insert into the benzaldehyde like species does not directly act from the gas phase. The activated N-species can react from imido species [21] or from NH4+-sites [22]. The temperature range for the ammoxidation of toluene mostly may reach from 320 “C up to 420 ‘C, increasing temperature leads to decreasing selectivity; thus, 144 ,,, ... , ,,,,. ‘$ optimum conditions for maximum yield must be adapted for the catalyst system ., ,,-, applied. Methyl substituted toluene’s can be ammoxidized to either mono- or .,. dinitriles [e. g. 23]. The. ammoxidation of o-xylene can give phthalimide or o- “, , phthalodinitrile depending on the reaction conditions. p-Tolunitrile or terephthalonitrile, respectively can be formed from p-xylene, also shape selective ammoxidation of o-or m-xylene is described recently [24]. Halosubstituted toluenes are easily to convert in nitrites due to the electron withdrawing effect of the substituent. The selectivity in the nitrile formation depends on the kind of the halogen substituent p-CI > p-Br >> p-I, whereas the reaction of dichlorosubstitutet toluenes is more influenced by the position of the substituent (2,6di-Cl c 2,5di-Cl < 2,3diCl c 2,4-di-Cl - 3,4-di-Cl) [25]. Nevertheless, 2,6dichloro-benzonitnl can be formed in 85 mol YOyield under fluid bed regime (separate introduction of substrates) on V/Mo-multicomponent catalysts [26]. P-Hydroxybenzonitrile can be formed from p-cresol on B~o oxide or related catalysts, nevertheless the catalysts are deactivated very rapidly (by coke-like deposits. After protection of the OH group the p-methoxytoluenes can be ammoxidized in 65 mol YOyield (V/Ti oxide [27]). Similarly, p-phenoxytoluene can be converted to the nitrile with 85 Y. conversion and 58 % selectivity on a multicomponent catalyst [28]. Joint influences of electronic and position effects hamper the conversion and ,.,, ,’ selectivity of higher substituted methoxy toluenes. Examples for aromatic side chain oxidation are given in table 5. 5) Further development Further development in the selective oxidation of aromatic compounds withe be ,’ highly influenced by the development of catalysts, considering detailed mechanistic knowledge particularly about the selectivity determining sites of the catalyst and steps of the reaction. The search for more selective catalysts may ... i’ 145 ,,. , ...==.. . .-J<...’,-.~.-.< ., .,:, ,.. .:,.-;,,- .. . . : :..->:.,,: ...... , include the search- for biomemetic catalysts (including more sophisticated host guest structures) and oxidation catalysts with defined disperse metal or metaloxide sites on the surfaces. ,“- . .4 ‘. . 146 [1] R. Raja, P. Ratnasamy, Progress in Zeolite and Microporous Materials, Stud. in Surf. Sci. and Catal. ~, 1997, Elsevier, Amsterdam, p. 1037. [2] T. Kitano, Y. Kuroda, M. Mori, S. Ito, K. Sasaki, J. Chem. Sot., Perkin Trans. z, 1993,981. [3] G. Panov, A. S. Kharitonov, V. 1.Sobolev, Appl. Catal. u, 1993, 1. [4] G. Panov, V. 1.Sobolev, K. A. Dubkov, A. S. Kharitonov, 1Iti International Congress on Catalysis - 40* Anniversary, Stud. in Surf. Sci. and Catal. ~, 1996, Elsevier, Amsterdam, p. 493. [5] M. Hafele, A. Reitzmann, D. Roppelt, G. Emig, Erdol und Erdgas, Kohle - Petrochemie ~, 1996,512. [6] P. T. Tanev, M. Chibwe, T. J. Pinnavaia, Nature ~, 1994,321. r] A. V. Ramaswany, S. Sivasanker, P. Ratnasamy, Microporous Mat. 2,1994,451. [8] S. K. Das, A. Kumar Jr., S. Nandrajog, A. Kumar, Tetrahedron Lett. 3, 1995,7909. [9] A. K. Uriate, M. A. Rodkin, M. J. Gross, A. S. Kharitonov, G. L Panov, 3 d World Congress on Oxidation Catalysis, Proc. (R. K. Grasselli, S. T. Oyama, A. M. Gaffney, 1.E. Lyons, eds.) 1997, Elsevier Amsterdam, p. 857. [10] F. Bourdin, M. Constantine, M. Jouffret, G. Latignan, Ger. Pat. 2064497,1971. [11] A. Esposito, M. Taramasso, C. Neri, F. Buonomo, US-Pat. 4396783,1983. A. Esposito, M. Taramasso, C. Neri, M. G. Clerici, Chim. Ind. 72,1990,610. [12] G. Bellussi, C. Perego, Handbook of Heterogeneous Catalysis (G. Ertl, H. Knozinger, J. Weitkamp, eds.), VCH, Vol 5,1997, p. 2329. [13] J, Haber, Handbook of Heterogeneous Catalysis, (G. Etil, H. Knozinger, J. Weitkamp, eds.), VCH, Vol.s, 1997, p. 2329. [14] A. Martin, A. Brtlckner, Y. Zhang, B. LiJcke, Stud. Surf. Sci. Catal. u, 1997,377. H. Bemdt, K. BOker, A. Martin, A. Bruckner, B. Lttcke, J. Chem. Sot., Faraday Trans. ~, 1995,725. ,. ,. 147 ,“ . ‘. [15] R. G. Rizaev, E. A. Mamedov, V. P. Viloskii, V. E. Sheinin, Appl. Catal. A ~, 1997,241. P. Cavalli, F. Cavani, 1.Manenti, F. Trifiro, M. E1-Sawi, Ind. Eng. Chem. Res. 26, 1987,804. [16] A. Bruckner, A. Martin, B. Lucke, F. K. Hannour, Stud. Surf. Sci. Catal. ~, 1997,919. [17] A. Martin, B. Lucke, Catal. Today 32, 1996,279. [18] A. Martin, U. Bentrup, B. Lucke, J. Chem. Sot., Chem. Commun. 1999, 1169. [19] B. Delmon, 3rdWorld Congress on Oxidation Catalysis, Proc. (R. K. Grasselli, S. T. Oyama, A. M. Gaffrey, J. E. Lyons, eds.), Elsevier, Amsterdam, 1997, p. 43. [20] A. Bruckner, A. Martin, N. Steinfeldt, G. U. Wolf, B. Lucke, J. Chem. Sot,, Faraday Tram. 92, 1996,425. [21] R. K. Grasselli, Handbook of Heterogeneous Catalysis (G. Erth, H. Knozinger, J. Weitkamp, eds.) VCH, Vol.s, 1997, p. 2302. [22] A. Martin, Y. Zhang, H.-W. Zanthoff, M. Meisel, M. Baerns, Appl. Catal. A ~, 1996, L 11. [23] M. C. Sze, A. P. Gelbein, Hydrocarbon processing 1996, 103. [24] K. Beschmann, L. Riekert, Chem.-lng.-Tech. 65, 1993, 1251. [25] A. Martin, B. Lucke, G.-U. Wolf, M. Meisel, Chem.-lng.-Tech. 66, 1994,948. [26] Y. Sasaki, H. Utsumi, A. Mori, K. Morishita, (Ger. Off. DE U, 176,88883, 1998). [27] A. Martin, J. French, H. Seeboth, B. Lucke, E. Fischer, J. prakt. Chem. w, 1990,551. [28] M. C. Sze, A. P. Gelbein, ,. Hydrocarbon Processing 1976, 103 +. [29] F. Cavani, G. Centi, F. Trifiro, Chim. and Ind. 74, 1992, 182. 148 ./. , ‘,, , ,,, ,., [30] E. V. Ivanov, L. A. Stepanova, E. M. Guseinov, B. V. Suvorov, ,.-, ,,,.: ,”::,., Neftekhimiya U, 1990,63. r., . ..., [31] Y. Sasaki, H. Utsumi, A. Morii, K. Morishita (Nitto Kag. Keg.), .-r ,, ,, .’ Ger. Offen., DE ~, 746,883, 1998. , .’ ,, -,, ,,, ,.. ; [32] H. Yamachika, K. Nakanishi, T. Nakaishi (Sumitomo) ,. JP 07,145,521,1993. [33] F. BrLihne, E. Wright in: Ullmann, 6* Edition 1998, Electronic Relaese (benzaldehyde entry) [34] H. Seko, Y. Tokuda, M. Matsuoka, Nippon Kagaku Kaishi 1979,558 [35] EP 0723949 (1996), (Hoechst AG) [36] A. Martin, B. Lucke, H.-J. Niclas, A. Forster, React. Kinet. Catal. Lett. ~, 1991,583. ,. ::.,, [37] R. Abele, J. Jovel, M. Shymanska, React. Kinet. Catal. Lett. m, 1993,69. ,,, .,, ,’, ,:, /. ,,” ‘;, .,,.., . .,,, 4 .. ,, . . . ..’ >’::’ ,.. ,“,,: ...,, .. . ,,, ,, ,, .“’, ?,,.,.~,-. , ; ,,! ~,,. .,,.,,.. , ,., ,,, ,,, $ ,., ,. ,., , ( ,,,,,:,. )4 ) ‘,,,,; .,,,. t.,.,., .i 149 -,, ,,., ,.’ , . ..., ., .,,, ,-. ~.,. .. ., . . Liquid phase (L) Oxidant Catalyst Selectivity Conversion Gas phase (G) (%) (%) Ref. (phenol) (benzene) F 02 Cu-pthtalocyanine 100 5,4 [1] (in Zeolites) G OZIHZ PdlCu >90 <1 [2] G NzO Fe-ZSM-5 98 27 [3] G NZO ZSM-5 95 8 [4] G N20 AIP04-5 100 4,5 [4] G N20 H-Ga-ZSM-5 71 28 [5] F Hz02 Ti-MCM-41 >98 68 [6] F H202 TS-I >95 31 [6] 88 F H202 TS-2 17 [7] 90 F H20z VS-2 7 [7] 100 F Hz02 sup. Vo 2+ 30 [8] Table 1: Hydroxylation of benzene to phenol: characteristic examples Reaction Temperature ~c] 400-450 Contact Time [see] 1-2 Selectivity: Benzene + Phenol [mot %] 97-98 Benzene + Cox [mol %] 0.2-0.3 N20 + Phenol mot 70] 85 Benzene + DioIs [mol %] 1 Productivity [mmol Phenol/g “:A,] 4 Table 2: Process Performance resulting from pilot plant data (Monsanto and Boreskov-lnstitute of Catalysis) [9] ., 150 Conversion Selectivities Catechol: Phenol [%] Diphenols a Diphenols b Tars hydroquinone Rh6ne-Poulenc 5 70 90 10 1.4-1.5 ENICHEM 25-30 84 90-94 6-10 1 awith respect to H20 .! bwith respect to phenol ,.’ Table 3: Performance of the Rh6ne-Poulenc- and the ENICHEM-Process for the Phenolhydroxylation Substrate AIdehyde Catalyst Conversion Selectivity Yield (%) Ref. Toluene benzaldehyde V-oxides 10-20 40-60 [33] (div. processes) ,., , p-methoxy p-methoxy VIPICu-oxide 70 [34] toluene benzaldehyde ., p-chloro- p-chloro- VICslFe-oxide 68 20 [35] ,. toluene benzaldehyde .,- ,,, , ,.,, , 4-methyl pyridine- V-phosphate 80 55 [36] pyridine 4-carbaldehyde 2,6-dimethyl- 2,6-pyridin- WMolOxide 95 65 [37] pyridine dicarbaldehyde 6-methyl- 2-pyridine- ,’ ,,:,,’, carbaldehyde ,,-, 1 .:>.,.. ‘ -. “, ,,,. ‘. :., .,: ,,, . ./. ,,,.O,,’,<,.:;. , ,,.,’.-. ‘ ., .,, !,- ~ ,,,,;- ,’,, ,,.,, ,.’, ., >-,.,,, ,.,..~ Substrate lAidehyde lCatalyst 2onversion lSelectivity field (%) Ref. p-toluene benzonitrile V containing 50-100 70-95 [28] mukicomp. cat. p-methoxy p-methoxy ViTi oxide 55 [29] toluene benzonitril VISblK oxide 50-75 40-60 [30] 2,6-dichloro 85 [31] toluene += p-phenoxy 85 58 [32] toluene I : Table 5: Examples for aromatic side chain ammoxidation to nitriles 152 ,, DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Eriangen 1999 ,. F.-J. Mais ,, Bayer AG, Leverkusen, Germany ,- ,, ,. .’ Recent Progress and Challenges in the Electrophilic Substitution of Aromatics .,. Introduction The industrial production of polysubstituted aromatics usually yields mixtures of isomers. For example, the electrophilic aromatic secondary substitution of chlorobenzene or toluene yields mixtures of mainly ortho and para isomers. 1; 1; .,>’,., . o “2/ o~ HN03 f .. ‘, ., cl .:’ 1; & &N02 b\ HNO, “2/ olp-chlorotoluene olmlp-Nitrotoluene I OH - CH3 olm/p- olp- l~OH Dichlorobenzene Chloronitrobenzene G OhS/D-cWSOi . . . .,, ,, ,.-’, Figure 1. Product family tree of basic chemicals for disubstituted aromatics ., ... .. $.,,’. However, there is otlen a preferential demand for only one specific isomer. This demand may ,, .. also fluctuate very widely, a situation which we at Bayer describe as the “isomer problem”. Balancing isomer ratios is a continuing challenge to the chemical industry. Apart from the various options for disposing of the surplus isomers that are not required - for example, the development of new downstream products, the marketing of these isomers for inexpensive ., ,,, , applications, isomerisation, breakdown into constituent components for reuse or, at worst, ‘, ,’ ., ,:, <., -, thermal recovery - the most elegant solution to the problem is to prevent the occurrence of ., :!’”, ,. ;’.’, surpluses altogether. Catalysis is an important approach here...... ‘..’. ,. ,..’.’ ,, 153 ., ,”,, ,, DGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 ,,, .,. .,, Example: chlorotolucnc In 1998, global demand for pure chlorotohrene isomers was around 35,000-40,000 tonnes. For many years, the isomer ratio has been tipped heavily towards para isomers because of the primary use of p-chlorotohrene in crop protection agents. The demand ratio for ortho and para products is between I :2 and I :2.5. Chlorotoluene is produced by the liquid-phase chlorination of toluene in bubble columns at 40-50 “C. The standard catalysts are either iron(III) chloride or iron(III) chloride with sulphur as the co-catalyst. (+1 Lit. 250 ppm FeCl, Off.1= 1:0,52 65% 2,0 y. 33,0 % 1 250 ppm FeCl, / OIP = 1:0,89 52 % 0,5 % 47,5 % 1 85 ppm S Figure 2. Nuclcar chlorination of tohrcnc Using stslphur as the co-catalyst makes it possible to improve the ortho : para ratio in the chlorotoluenc mixture to 1.1 : 1. However, it is by no means sufficient to balance the isomer ratio. hr the 1980s, chlorotohrene manufacturers made great efforts to solve the isomer problem. If, instead of sulphur, six-membered dibcrwocondenscd hetcrocyclic compounds of types such as thianthrcnc, phenoxathine or phenothiazine are used, it is possible to increase the para content even further. Pam content (%)in the chlorotoluene fraction Lit cl~:~’m 54 ‘% OXYCHEWUSA 1976 2) n c,~:)y 52-56 % IHAIVVJPN 1981 3) HODOGAY,WJPN 1982 4) n HOECHSTIGER 1986 6) ~.x ATOCHEMIFRA 1984 6) CQ3 “0’ lHARA/JPN 1985 7) I Figurc3. Co-catialysts forthcnuclcar cllIorination oftolucnc 154 I ,,,,, .’,,,,.. . ,,,.. , ,- ,.’,!, ,, .,.., ,,, :, :,, ; “,/ , ,, ;, ..,, , ,! ‘!,>+ ,., , ‘.., ,, ..; -,.,,:. ,.:, However, this group of co-catalysts has considerable disadvantages such as the large ,;, , ‘.! .,’, ., ,’ j ,;,: quantities required, the need to use SbCIJ instead of FeC13as the main catalyst and the ,... ,,!’ .,} :,,,~:’ ,,.j. < comparatively low reaction temperature (e.g. < 20°C). At Bayer, we have worked intensively ,, ., .“,,:, ,, ,,, ! ,, on the nucIcar chlorination of toluene and have found anew group of co-catalysts – ., ’,,,,.+ ,, ,’. heterocycIic seven-membered rings of the benzothiazepine type – which avoid these disadvantages. Max.para contont (%) Lir. Ho N 1; 57% a aj s o H 1> N 56,5% 8) d] s 60,0% 8) S-MO H,C N- 1; 64,5 % 8) s Tb ,, ,,, CH3 ‘, , ..’ Figure4. Benzocondenscd scvcn-mcmbcred S,N-hctcrocyclic compounds asco- catalysts for the chlorination of tolucne (50 ‘C, 170 ppm FcC13,40-60 ppm co-catalyst) With the aid of these co-catalysts, it is possible to yield a para content of up to about 60 V. in industrial production. The still available srrrphrs of o-chlorotohrene has a direct influence on the production of crcsols, for which Bayer uses alkaIine hydrolysis under pressure of chlorotoIuene mixtures. ,, ,. 155 .,: ~, :,’ ., ... , ,,J ,, ,, ‘,- :- . #. ,. ., AH Input (orlhodominant) 60% OIUIO 307. 30 % 1 “h meta 39 % para 20 % 20% z 30 % 50 % 20% Figure 5. CrcsoIs from the chlorotohscnc mixture - aryn mechanism On the basis of a 1998 global demand for single-isomer cresok of approx. 100,000 tonnes with an ortho : meta : para ratio of 1 : 1 : 1, it is evident that the isomer mixture is not matched to requirements. Above all the more valuable para isomer is produced in only small amounts. This situation yields a large number of tasks for researchers in the future. These are summarised in Fig. 6. Further optimisation of para selectivity in the chlorination of toltrene Position-selective hydrolysis of chlorotoluene (para as target) Alternatives for the selective manufacture of p-cresol New products based on o-chlorotoluene and o/m-cresol Figure 6. Challcngcs in chlorotolucnc/crcsol production Example dichlorobenzcnc . On an industrial scale, the chlorination of benzene is carried out in the liquid phase using FeCl] or FeC13/sulphur as the catalyst. In 1998, global demand for dichlorobenzenes was around 150,000 tonnes. The ortho/para ratio of 1:1.7 is still heavily weighted towards the para isomers. The main applications for para-dichlorobenzene are in the hygiene sector and in the production of PPS (polyphenylene srdphide). Ortho-dichlorobenzene is the far more valuable isomer because of its key application in the manufacture of active irrgredicrrts. [n future, the ratio is likely to shift towards ortho-dichlorobcnzene as a result of the decrease in demand for para-dichlorobenzene in the hygiene sector. 156 .’ ,! ,., . ., >-, ,, ., ‘., .,, ., , ~.,. ., ,. -,.,’ ‘, ., ,.: -.,>.’ ,., ,, ,?L ,, :, ::..., ,., .,. .,. , “.’-,..,. ,-,. ., ,’. ,L’ .’ .-’.-,, 1; 0 ~&’60-70” c b.,+ cl ~’.’, ‘,:, ,,4 500 ppm FeC13 o/p = 1:1,5 40% 1,0% 59,0 % .,” ‘; .,,’.,, 500 ppm FeC13/ Olp = 1:3,0 24% 0,3 % 75,7 % :, ’,,, 85 ppm S ,.,,,, ;,\, ,tj .,,,, ’-,., ,.,,/, Figure 7. Dichlorirmtion of benzene, technical process ., -’,-’ Lit. 9), 10), 11) Comparison of the isomer mixtures produced against actual requirements shows that we have reached the upper limits using the FeCIJ catalysts. The control principle of co-catalysis can ,, also be applied in the case of dichlorobenzenes. Fig. 8 shows an overview of documented options. These also reveal that, to date, only the process using FeCIJ and FeC1~sulphur catalysts is suitable for industrial applications. Para content (Y.) in the dlchlorobenzene I frsction LiL MnC12. 0H2 5070 PPG 1975 12) ,,!?. ,,.,,.;:,,,,..: ,:..! I ,,,.../ ,1,/-,,.. -,, FeCl, 59 v. - 10) +’.,..’> :’; I ,,~[,. ‘., > ,. .,,,. FeCl~S od. S2C12 76% N. KAYAKU 1961 11) ., ,. .:, ,., .,,. . . I OMNYCHEM 1962 ,,, ,, ye-x . .,. .“,, :. , ,..,-~,, ; 83% ATOCHEM 19S4 6) .,, . .,’. .,, “C’”an BAYER 1990 13) I X-Cl,-CF3 Zeolites: L-zeolNes >90 % KUREHA 1980 14) H-mordenIte IHARA 1964 15) I HODOGAYA 1985 16) Figure 8: Influence of the catalysts on the ratio of ortho- to para-dichloro- benzene ,. The use of phenothiazine-based co-catalysts is technically possible but has so far not been implemented. Given the expected decrease in demand for para-dichlorobenzene, an ortho- selective process for the manufacture of dichlorobenzenes is required. However, our work bas shown that the proposed process using MnClz ● HzO catalysis is not feasible as the chIorine loss is some magnitudes higher than when using Fe-based catalyst systems. 157 ~..’ ‘““ -. .-<-..>. ,.,,.~ .,j. ,:, ‘.> A further aspect of dicblorobcnzcnc production is meeting the demand for mcta- dicblorobcnzcnc. Only 1 % is present in the isomer mixture yielded by chlorination yet demand is already at 5 ‘7.and showing an upwards trend. This shortfall can be covered using a thermodynamically controlled isomerisation process. &d’+c,+ c1 c1 Lit Liquid phaae AlC13/H20(cat.) 8 40 52 17), 18) 150” C, [2 h] Gas phaae H-ZSM-5, HI 10 53 37 19) 30 bar, 350” C Thermodynamic 37 47 16 I equilibrium Figure 9. Isomerisation of dichlorobenzcncs However, this route is aLsolaborious and should be replaced with a selective direct synthesis as soon as possible. identification and development of an ortho- selective process suitable for industrial :. . application +,,. Alternatives for selective manufacture of ., m-dichlorobenzene Figure 10. Challenges in dichlorobcrsmne production Example: nitrotohscnc Total world production of nitrotoluene in 1998 was around 150,000 tonnes. Tbe main isomer here is ortho-nitrotolucnc with its applications in crop protections agents. Worldwide, the ortho/ para ratio is 1.6: 1. This ratio is almost the same as can be achieved by the classic nitration of toluenc20). 158 .. , ,. .’.,’: ?,, . ..,. . . . . ,.,’;. ,,, ,, ..,” ,, Olp No* H“”, I H=SO, 1.62:1 59% 4,5 % 36,5 % Industrial process H“”, I H,PO, 1.02:1 48 Y. 5,0 % 47,07. Ussble In prhrciple ., ~ Pr-O-NO# 0.05:1 5 % - % 95,070 Not usable -$ H-ZSM-’3 N02PFC / 2.34:1 69 V. 1,5 % 29,5 % Not usable /.i Me-NOz , .;.,,,,.’, ,, ,.. -,:,4,, ‘, Figure 11. Nitration of tolucnc - isomer control optionszo) .. :,,’,,, ,,.,,,:. \, i.{ $, :,.,,, ,.,.,,1, ,.. - This balance in the ratio applies only on a globaI scale and not necessarily to the individual manufacturers of nitrotoluene. At Bayer, the demand for pure nitrotohrene isomers is subject to very wide fluctuations. Unfortunately, it is not possible to control the isomers by means of co-catalysis. A number of other options have been proposed. However, none of these processes arc yet in industrial use, usually because of the high costs involved. Irs our view, the continuing lack of detailed knowledge of the reaction mechanisms involved is unsatisfactory. In the case of nitrotoIuene, a switch from sulphuric acid to phosphoric acid, for example, results in a significant shift towards the para isomer whereas the nitration of chlorotoluenc under the same conditions causes a clear shifl towards the ortho isome?). ,,, The task of producing single-isomer aromatic intormediatss to match demand has still not been solved satisfactorily. -,.,,, cially not in the case of simple reactions such as Q ~g~rinatio”,”itition, alkylation, etc.! The selective manufacture of the desirsd isomers in quantities .. . which meet market and downsbsam rsquirsments remains one “. > of the resin goals of industrial organic chemisw. @ Challe”gei”caWlysisandprocessengi”eeri”gl Also of great interest am the technically feasible and economically atixactive new processes for rscoveringlrscy c!ing isomers not raquirod. ... .‘, . . isomerisation, transhalogenation, transalkylation, ~ ~galoge”atio”,dea[kylatio” ,, ::, , ., ‘. .,,, ,, ,, ,, ,-! Figure 12. Remaining challcngea ; :..,,’!!,,; l,;::,% ... , ;1. ., .: ; ,,.,,. , . ,, r.!, ,,. ; ,, ., ,, -<,,,..,’ , ,, .,, 159 In conclusion it can be said that there is as yet no satisfactory solution to the problem of producing substituted aromatics in accordance with demand, especially in the case of such simple reactions as chlorination or nitration. The selective manufacture ofsingle-isomer aromatics remains an extremely important goal of enormous economic significance. Literature 1) U1lmann’sEncyclopedia of Industrial Chemistry, Vol. A 6,343 2) OXYCHEM, US 3989715 (1976), US 4024198 (1977), US 4031142 (1977), US 4031147 (1977), US 4069263 (1978), US 4069264 (1978), US 4250122 (1981) 3) IHARA, US 4289916 (1981) 4) HODOGAYA, EP 63384 (1982) 5) HOECHST AC, EP 173222 (1986) 6) ATOCHEM, EP 126669 (1984) 7) IHARA, JP 60/ 125251 (1985), JP 60 / 136576 (1985), JP 61 / 171476 (1986) 8) BAYER AC, US 4851596 (1989), US 4925994 (1990), US 4990707 (1991), US 5105036 (1992) 9) Ullmann’s Encyclopedia of Industrial chemistry, Vol. A 6,333 10) PPG, US 4235825 (1980), UNION CARBIDE, US 3226447 (1965) 1I) NIPPON KAYAHU, JP 64 / 3821 (1961), OXYCHEM, GB 988306 (1962) 12) PPG, US 4017551 (1975) 13) BAYER AG, EP 474074 (1990) 14) KUREHA, JP 57/ 077631 (1980) 15) IHARA, EP 118851 (1984), EP 154236 (1985), JP 61 / 172837 (1985), EP 248931 (1986) 16) HODOGAYA, JP 61 / 183235 (1985) JP 61 / 183236 (1985) 17) BASF AG, DE 1020323 18) UNION CARBIDE, US 2666085 19) BAYER AG, US 5466881 20) Ullmann’s Encyclopedia of Industrial Chemistry, Vol. A 17,421 21) G.F.P. Harris, ACS Symp. Ser. 22 (1976) 4, . . . . .? 160 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 H. P. Roger 1),K. P. Molter 2),W. Bohringer 2),C.T. O’Connor 2) l)s~d-chemie AG, Division Catalysts, Bruckmuehl, Germany 2,University of Cape Town, Catalysis Research Unit, Cape Town, South Africa Fundamental Relations between Modification of the External Surface of Zeolites and Catalytic Performance ,.,.,, ,.. Introduction ...... Zeolites are increasingly applied as catalysts in the refinew and petrochemical industw. They are crystalline alumino-silicates having micropores with uniform and characteristic dimensions in the range of Angstrram. One of the unique features of zeolites is their shape selectivity which is based on the fact that most of the active sites are located in these intracrystalline micropores. In customizing shape selective properties for individual process needs, the modification of external surface properties may be of increasing importance. Modification techniques, such as pre- mking [1], poisoning with voluminous strong bases [2], Silicalite coating [3], and silanisation of the external surface [4,5] are aimed at introducing a diffusion resistance between the intracystalline pore space and the extracrystalline fluid phase (e.g. pore mouth narrowing) and / or inertizing the non shape selective external surface. This paper illustrates fundamental relations between the modification of these external surface properties and the catalytic performance of HZSM-5 using a well selected set of catalysts and model reactions. Experimental approach The set of catalysts used comprised unmodified HZSM-5, a homologous sequence of 1 to 14 times silanised HZSM-5 samples with a progressive degree of external ;., ,,’ .. ,,,$ surface modification, HZSM-5 coated with a Silicalite 1 shell, and amorphous silica- ,,, .,. .. -,.. ;, : “, alumina. The catalyst powder was diluted with inert sand and loaded into a stainless ,.. , .’.., ,, >, .,,...,, ,, ,. steel, tubular, down-flow reactor. The silanised samples were prepared in-situ by . . . . : ?-, ,., ..,’,. !<- multiple treatments of unmodified HZSM-5 where each treatment consisted of low ,’, -, :, temperature, chemical vapour deposition using tetraethoxysilane (TEOS) and subsequent calcination. Detailed descriptions of the in-situ silanisation method [5] and the ex-situ Silicalite I coating [6] have been given elsewhere. Some properties of the catalyst samples are given in Table 1. Toluene, 1,2,4-trimethylbenzene (1,2,4-TMB), and 1,3,5-triisopropylbenzene (1,3,5- TiPB) were separately converted in the gas phase at the conditions given in Table 2 followed by regeneration after each run (flowing air, 500 “C, 5h). Selectivities were compared at similar conversion levels adapting the space velocity where necessary. . 161 ,, ?’ —. ...,, ,,. +..:.?’ . , . ... ,.., , ,.>,<:., ,~:”. ..L. ------. . . . I At5Lt 1: GatalyStS Catalyst Steric condition Physico-chemice woperties . SWI Particle Morphology size , unmodified HZSM-5 Micropores 5.5A 45 0.5:.5 spherulitic, , agglomerates 1 to 14 times silanised I a) Progressively decreasing 45 o) . spherulitic, HZSM-5 samples pore mouth size [5,6] = agglomerates b) Progressively decreasing external surface activity (z) Silicalite I coated HZSM-5 I a) Increased ~ffu5ion Zfj (1)/loo 1.5-3.0 spherulitic, resistance agglomerates b) Reduced external surlace +-- activity amorphous silica-alumina No constraints 10 0.5-1.0 agglomerates (1) Si/Al of the parent material (2) The external surface becomes almost inert after the fifth silanisation treatment TABLE 2: Reactions Reaction Steric demands Products Conditions in relation to HZSM-5 micropores Toluene Low Benzene T = 450”C disproportionation Xylenes x =2% pF_,~ = 130 mbar 1,2,4-TMB “) High Xylenes T = 450”C disproportionation TeMBsi2j x =7% Pr~,.= 35 mbar 1,2,4-TMB Highly restricted to the 1,2,3-TMB T = 450”C isomerisation external surface 1,3,5-TMB x =7% PFW,,.= 35 mbar 1,3,5-TiPB ‘3) External surface only Crack products T = 270°C cracking X =2-70% PF-.= 1.7 mbar (1) TMB =: trimethylbenzene (2) TeMB =: tetramethylbenzene (3) TiPB =: biisopropylbenzene Results and Discussion The effect of external surface and pore mouth modifications may generally be seen as a modification of the overall steric conditions of the zeolite c~s~al. In that context the conceptual approach to study the effects of external surface modifications on catalytic performance was based on varying both the steric conditions of the catalyst (Table 1) and the steric demands of the reaction (Table 2). Over unmodified HZSM-5 the spectrum of steric demands covered by the set of reactions in Table 2 ranges from almost no steric limitation by the micropores over intracrystalline mass transfer control to complete restriction to the external surface. 162 .,. . The external surface, and therefore the steric conditions of HZSM-5 were modified -,-’ ,,, by multiple silanisation treatments and Silicalite I coating. The repeated application of ‘, .- in-situ silanisation treatments lead to 14 silanised HZSM-5 samples where the amount of silica deposited onto the external surface of the zeolite crystal, the degree of external surface inertisation, and the degree of pore mouth narrowing continuously increased with increasing number of silanisation treatments. The intrinsic properties of the micropore space were not affected [5]. As a second external surface modification method HZSM-5 crystals were coated with a Silicalite I shell which ,- +;, reduced the external surface activity and introduced a diffusion resistance. In order ,, ,,:. to simulate the non shape selective external zeolite surface 1,2,4-TMB was also ,’-, ., ,’,,, converted over amorphous silica-alumina. ,,, i (. ., ., ...”<. , .. :.,,. Activify ,, .“.,,. .,. ‘. ,,, -,, ,, Figure 1 compares the effect of progressive external surface silanisation on the ,, .“,. , f ..,. ,. .,. activity of HZSM-5 during toluene disproportionation, 1,2,4-TMB disproportionation, .,,,...,’::,,, . . ,. (,,, ~,, ,, .,.-. 1,2,4-TMB isomerisation, and 1,3,5-TiPB cracking over HZSM-5. In order to obtain ,.. ,,. , .- ,., , .“. . comparable data the activities for each reaction were normalised by setting the “,,’:. ,,” activity of the unmodified parent HZSM-5 equal to 1. The original absolute activities “ ,, ,.. .,,’,, - were measured using the differential reaction rate of the respective reaction. To ,. .,- account for the integral conversion during 1,3,5-TiPB cracking the differential .,, .,.,.’. ,, ,, ...... reaction rate was determined assuming first order reaction kinetics. ,.. .$’ ...,;-, .:,: ‘.’.., ,,. <.,. ~ For all reactions the activity decreases more or less with increasing degree of ...... ,,, , . ,, .: silanisation. The almost complete inertisation of the external surface during the first 5 ,..,. ,, ,* silanisation treatments has the most direct effect on reactions which are restricted to .. ..-$’ the external surface of unmodified HZSM-5. In Figure 1 these reactions are represented by the isomerisation of 1,2,4-TMB and 1,3,5-TiPB cracking. Progressive pore mouth narrowing results in a mass transport resistance for molecules entering or leaving the micropores and obviously only affects reactions which take place ,, “;. ,,’ ..,, inside the micropores such as toluene and 1,2,4-TMB disproportionation. In ,:,, ,,. conclusion it is intrinsic to the nature of external surface silanisation to tend towards J ,,,., ,, :,.... ‘, ., !.’: decreasing the activity. ,.,,$, .; .,, , At a given degree of external surface silanisation the catalytic activity of HZSM-5 is reduced in the following order of reactions (7) 1,3,5-TiPB cracking and 1,2,4-TMB isomerisation, (2) 1,2,4-TMB disproportionation, (3) toluene disproportionation. This sequence is consistent with the findings for the Silicalite I coated HZSM-5. The normalised activities during 1,3,5-TiPB cracking and 1,2,4-TMB isomerisation are almost identical and may be considered as a lower boundary benchmark representing reactions which are limited to the external surface. In contrast the normalised activities during toluene disproportionation are an upper boundary benchmark representing reactions which easily take place in the micropore system with minimal diffusional constraints. , ,,, ,,: 1ss ‘,.,,.. .:. . ...... ,,’,’ f ., ,. , The normalised activities in the disproportionation of 1,2,4-TMB are lower than the normalised activities in the disproportionation of toluene, showing that the sterically more demanding disproportionation of 1,2,4-TMB is clearly more sensitive towards the inertisation of the external surface and pore mouth narrowing. On the other hand the disproportionation of 1,2,4-TMB shows higher normalised activities than the sterically more demanding 1,2,4-TMB isomerisation and 1,3,5-TiPB cracking. It may be concluded that the more sterically demanding the reaction the more sensitive its activity towards modifications of the pore mouth and /or the external surface. ■ ■ ■ ■ HZSM-5coated m= m m ■ mm ■ witha Silicalite m 1shell I I m Toluene-dis I * ~TMBdis TMB-is TiPB-crack L 0123456789 101112131415 Number of silanisation-calcination cycles FIGURE 4: Effect of repeated external surface silanisation and Siiicalite I coating on the activity of HZSM-5 in various model reactions with varying steric demands. Selectivity This section discusses the effects of external surface modifications on the selectivity by comparing the distributions of the xylene- and the TeMB-isomers of the respective fractions obtained over unmodified HZSM-5, the various modified HZSM-5 catalysts, and amorphous silica-alumina. During the disproportionation of toluene the proportion of p-xylene in the xylene fraction continuously increases after the 4“ silanisation treatment up to a p-selectivity of 9i’~0 after the 14ti silanisation as shown in Figure 2. Figure 2 also shows the proportion of p-xylene in the xylene fraction obtained during the disproportionation 1,2,4-TMB as a function of the number of silanisation treatments. As already observed during the disproportionation of toluene, the first 4 CVD cycles which inertize the external surface of HZSM-5 almost completely (see Figure 1, 1,3,5-TiPB cracking activity) did not affect the xylene distribution. The same was observed for the Siiicalite I coated sample. At this low degree of modification the 1s4 ,’ . #. ., /,., ,4, ,, ..,. ,. ’,. ‘. ,’,, . . ;.. ; .,.”646 ;.; .> ‘ # . . . .- . ,.. ‘.( ,. ;:.;.., ‘ :.’ :“”. ,, .$.~.;;: .’ ..:” -+. : f“. .’. _ xylene distribution tends towards the thermodynamic equilibrium. The 5h to the 8ti cycle increase the p-selectivity at the expense of the more bulky o-xylene and m- xylene due to progressive pore mouth narrowing. In contrast to the p-selectivity obtained during the disproportionation of toluene, the p-selectivity obtained during the disproportionation of 1,2,4-TMB deteriorates after the 8ti silanisation treatment and the xylene distribution approaches the by intrinsic kinetics controlled distribution which was observed over amorphous silica-alumina. It is concluded that the progressive pore mouth narrowing increasingly prevents 1,2,4-TMB from entering the micropores and restricts the reaction to the now modified, non-shape selective, external surface. The silanised external surface still shows a vev small remaining activity which corresponds to approximately 2% of the original external surface activity of the unmodified HZSM-5 as measured by 1,2,4-TMB isomerisation and 1,3,5-TiPB cracking (Figure 1). 1,0 ~ 0,8- ❑ ❑ L,.,”:,,: G 0 1,2,4-TMB dis over ❑ z ~ 0,6- HZSM-5wated witha % shell SiIiraliteI w + + ; 1’2’4Tm ‘k T ● M g 0,4- 1 ~ 4:~:g” & ❑ + = (3,2 ------Thermodynamic 1,2,4-TMB d~ over -n equilibrium amorphous silica-aIumina 0,0 , 01 2345678910 1112131415 ,, ... .A. , . . .,,, ‘d “., ...!,. ,.,2,, ,4 Number of silanisation-calcination cycles ~“”;l .<;,;,,, r “f.,,,-:: ,:,!:[J FIGURE 2 Effect of repeated external surface silanisation and Silicalite I coating .,, ,,.,.. , ;;,., ,<’.,.,”:.,... on the proportion of p-xylene in the xylene fraction during the sterically non , .,,!, .,1 demanding disproportionation of toluene and the sterically demanding ‘ ::,..).~ $;: .,, ,,,.;,..~,,. -:,$.; disproportionation of 1,2,4-TMB over HZSM-5. ,,;i,’, ‘,, Figures 3 (A) and (B) compare the disproportionation of toluene and the disproportionation of 1,2,4-TMB with respect to the compositions of the xylene fraction obtained with increasing degree of silanisation. The xylene composition obtained during the disproportionation of toluene approaches 100 YO p-xylene (Figure 3.A) at minimal activity loss (Figure 1). In contrast the xylene composition obtained during the disproportionation of 1,2,4-TMB follows a loop (Figure 3.B) accompanied by a decreasing activity (Figure 1). It can be concluded that pore mouth narrowing has not forced the disproportionation of toluene onto the external surface yet, even after 14 silanisation treatments. This is due to the disproportionation of toluene to benzene and p-xylene being the sterically least demanding reaction of the investigated set “of reactions. - REACTION WITH LOW STERIC DEMANDS REACTION WITH HIGH STERIC DEMANDS pxyiene (shape selectively favored) pxyiene (shape selectively favored) 1 ber (B) es Themmdynamic Equihbrium k ~ ~ -- m-xylene o-xyfene m-xylene o.xyiene ,. (thermodynamically favored) (inbinsically favored) hernmdynamically favoured) (Intrinsically favoured) I FIGURE 3: Effect of repeated external surface silanisation on the xylene distri- bution during the dispropodionation of (A) toluene and (B) 1,2,4-TMB over HZSM-5. Figure 4 shows the effect of progressive external surface silanisation on the molar product ratio [1 ,2,4,5-TeMB]/[l ,2,3,5-TeMB]. This ratio responded more sensitively, i.e. at lower degrees of modification, than the distribution of xylene isomers (Figure 2). The ratio [1,2,4,5-TeMB]/[1 ,2,3,5-TeMB] increased already during the first 4 silanisation treatments and over the Silicalite 1coated HZSM-5. In contrast the xylene distribution was not affected at these low degrees of modification. The high sensitivity of the ratio [1,2,4,5-TeMB]/[l ,2,3,5-TeMB] is most probably due to the larger molecular diameter of 1,2,3,5-TeMB relative to the xylenes and 1,2,4,5-TeMB. Due to the large molecular diameter the formation of 1,2,3,5-TeMB is almost restricted to the external surface. Therefore 1,2,3,5-TeMB would be the most sensitive product with respect to inhibition due to the inertisation of the external surface and would directly affect the above ratio. The enhanced shape selectivity towards 1,2,4,5-TeMB begins to deteriorate already after the 5* silanisation cycle. For comparison the shape selective formation of p- xylene began to deteriorate only after the 8’h silanisation cycle during 1,2,4-TMB disproportionation (Figure 2) and no deterioration was observed during toluene disproportionation (Figure 3.A). As expected, this shows that the progressive pore mouth narrowing restricts the sterically more demanding formation of TeMBs onto the modified external surface at lower degrees of silanisation than the formation of xylenes. lee ,,, ,,,’ ., E 4- 2 ...... ”...... — HZSM-5coatedwitha g + SilicaliteI shell .. Iil l-- ‘J.. 5’” ,, ~ I ‘1’hennodynbceq.ilibri.m :“0! .-. . ., , ., .:’ ‘. :,/;:.,,,. ., , 012 3456789101112 131415 ~.,.;+ ,- .. ,,. ,{ Number of silanisation-calcination cycles ,!,.~..{., ,,.,;4 ., ,, , ,,f, -,-’ ,,,, ,,.! :.; FIGURE 4 Effect of repeated external surface silanisation and Silicalite I coating ,’.. t: “;.,., +, ,’ ,1.:,.,.,.,, on the molar ratio of tetramethylbenzene isomers during the sterically demanding ,. :},’ .’, , ,;:,;:-.,, ,. .,,:.,,,: ,>,, ~ disproportionation of 1,2,4-TMB over HZSM-5. ,:j,;<,:$,.;j,. ;:,;.4 .,:,-,,../’”. ,.,.! ,’ ‘: .,;., :1 .- ,,’;‘:,//, ‘;! Generalised interpretation ,$’:,;{,,,,!;.,,~+,<~ ., -. !,,. ,.,. ,, ,.,, !,!, -,, .!:, , , .,.:.; ?J.*4 The studied set of reactions and catalysts yields several fundamental insights into ~-;, “. .. ‘. “$).,-- ,,:>+ the effect of external surface and pore mouth modifioations on the activity and shape >’: :,, ,“J ,,.~ ,’f,.’~ .,-, ,,,, ,;,;.: ‘;,, ‘: selectivity of zeolites as illustrated in Figure 5. ,. ,. “,.; ,,.,.,,$//4 ,., . ,,...>, J .;,! .,’.,.; . 1) It is intrinsic to the nature of external surface modifications such as silanisation ~,’.,,:., T : ,...,’:.,.“.,,,-...II$ ‘ ,,,::<. ,’,,.,,, and Silicalite I coating to tend towards decreasing the activity with progressive ,.,,.‘ .,,.<‘1 reduction of the external surface activity and pore mouth narrowing. ,, 2) The activity loss caused by such external surface modifications increases with the steric requirements of a reaction. Reactions with low diffusional requirements are only inhibited at high degrees of pore mouth narrowing (Figure 5, upper curve). Reactions with high micropore diffusion requirements are clearly more sensitive ,- towards inertisation of the external surface and pore mouth narrowing (middle cuwe). ~~~:, . These are reactions involving fast intrinsic kinetics and / or relatively large reactant ,,..) ,.-,; and /or product molecules. Reactions which are restricted to the external surface are :. ,”.: ,, of the category with the highest steric demands. They are obviously not affected by ,,.. C:* pore mouth narrowing but strongly suppressed by the inertisation of the external ‘ :‘- ~‘ ,; ,. ,. surface (lower curve). .’-. f, , :, -,, ,, .,. , 3) Progressive pore mouth narrowing increasingly inhibits reactions from taking place ,, ..’,. ‘. inside the micropores and restricts them to the modified external surface. The degree :’,,.,...<,,..’,,., .:;..,,. ,, ,-,”., of pore mouth narrowing necessay to restrict a reaction to the external surface , !.”:,..,-,. !., decreases with increasing diffusional requirements of a reaction. ‘, ’,. ;<,, ,. .,,,., ,-. , -, . ,’,,.,.”,, 167 , .....-.,, ‘,, ,’.. ,.,’, ., .< ,..... ;, .~....,,, ,’~’,,, 4) The consequences of point (2) and (3) for the course of selectivity are illustrated in Figure 5 (B). Over unmodified HZSM-5 the selectivity of the studied reactions tended towards partial thermodynamic equilibria due to rapid secondary reactions (i.e. xylene isomerisation subsequent to the disproportionation of toluene or 1,2,4-TMB). At low degrees of modification the inertisation of the external surface and / or the reduction of the pore mouth size shift the shape selectivity towards sterically less demanding reactions to which the micropore space is still accessible. With progressive modification the micropore space accessibility continues to decrease thus increasingly restricting reactions to the non shape selective external surface. This effect is counteractive to the initially positive effect on shape selectivity, and causes a maximum after which the shape selectivity decreases until the reaction is totally restricted to the external surface conversion. The final selectivity and activity are then determined by the intrinsic kinetics given by the modified external surface. (B) Selectivity (A) Activity towards the sterically least demanding . reaction paths Non-diffusion controlled reaction, if over Enhancedshapeselectivitydue to unmodified HZSM-5 a) reduced micropore space accessibility I 1 b) inertisation of the external surface I Selectivity controlled by /- in!rinsic kinetics since [hc IN ! \ micropore space has \ — Diffusion controlled reaction, already \ I become inaccessible and n1, over unmodified HZSM-5 I \ \ the reaction is now \ / .- / \ restricted to the \ Reaction restricted 10the extema \ & \ external surface \ surface \ \ ‘. \ \ Selectivity over unmodified HZ;M~5- ---- (tends towards thermodynamic equilibrium) c Number of silanisation-calcination-treatments Number of silanisation-calcination-trcatcncnts FIGURE 5: Generalised interpretation of effects of progressive modifications which affect both the external surface and the pore mouth on (A) the activity and (B) the selectivity of zeolites. “0” silanisation treatments as reflected by the origin mark the unmodified parent sample the activity of which has been normalised to “1”. 1. Low, C.C.D.; Lawson, R.J.: Kuchar, P.J.; Gray, G. L.; US Paterrt 5,321,184 (1994). 2. Namba, S.; Nakanishi, S.; Yashima, T.; J. Catal., 88,505 (1964). 3. Rollmann, L. D.; US Patent, 4,203,869(1980). 4. Bhat, Y.S; Das, J.; Halgeri, A. B.; Appl. CataL, 115, 257 (1994); Niwa, M.; Masato, E.; Murakami, Y.; Res. Cfrerrr. hrtermecf., 21 (2), 127 (1995); Wang, l.; Ay, C. L.; Appl. Catal., 54, 257 (1 989). 5. Roger, H. P.; Kramer, M.; Moller, K. P.; OConnor, C.T.; Microporous and Mesoporous Materials, 21 (4-6), 607 (1998); Roger, H. P.; PhD thesis, University of Cape Town (1 998). 6. Roger, H. P.; Molter, K. P.; OConnor, C. T.; Microporous Materials, 8, 151 (1 997). 16a DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 :,, E. Dietzsch, D. Honicke Chemnitz University of Technology, Chemnitz, Germany ,,, Gas Phase Hydrogenation of Benzene with Unusual Cyclohexene Formation ,!, , ., , ,’ -.,, Abstract ,,’ ,: The partial gas phase hydrogenation of benzene towards cyclohexene on ruthenium catalysts is ,,’ described. To yield the desired cyclohexene the presence of a reaction modifier in the gaseous feed is .,. ,4.’: :1 ,?, necessary. Methanol was found to be an effective modifier. By its influence the formation of . , .,,-,, ,1 ,,.,.: cyclohexene with high selectivities was observed. Different types of ruthenium containing catalysts, viz. coated Ru/A120@ and sol-gel derived alumina supported Ru catalysts, were investigated. With ,, ‘,, .“> ,+,,:: ,,-,. those, the formation of cyclohexene with selectivities of more than 50 % at conversions of about 5% ,,, ! J ,,., was obtained, which corresponds to a yield of z.s~o. By addition of metallic promoters e.g. zinc to ,., :.,, ,,, ..! ‘,, ,:.; ruthenium, cyclohexene yield was increased up to 8Y0. .!> ,, ,:,,., ,,,?,’ ,, : :,{ “,, ,,, ,. . ,, ,. ,, ,’ ‘, I .,’ ,, ;’, I 1 Introduction ,-, , ‘.1 i’:.. ,., f ;] Benzene is one of the main aromatics produced as by products in steam reforming ‘.; ,f .,:, .’ :> .1 processes. An excess of benzene is expected for the next years due to its substitution in .’, ! (, .,, ,,$, / 4 gasoline by other non-polluting components. Therefore, the surplus of benzene is available ..,, .,..1 ,1 ,, ..:.’ as feedstock for chemical utilization. Until today, one of the most important conversion of the ,::,.,.,,,,., . aromatic benzene is its catalytic full hydrogenation to cyclohexane. About 20% of the ,, ,, worldwide produced benzene was converted to this hydrocarbon in the last years. The benzene hydrogenation process is mainly carried out in the liquid phase and to a minor extent in the gas phase [1]. The formed cyclohexane may be converted into more valuable products, e.g. cyclohexanol and cyclohexanone. A more efticient route for the formation of such derivatives may be starting from cyclohexene which is produced by the partial hydrogenation of benzene. Recently, a corresponding liquid phase process for the production of cyclohexene is introduced in industry [2]. Benzene is batch processing ,. hydrogenated to cyclohexene in an aqueous solution of an alkali hydroxide in the presence of a Ru catalyst with product yields up to 60Y0.Drawbacks of this liquid phase hydrogenation are the non-continuous operation of the reactor, the diftlcult and costly separation of the -, ... , , ., ..:,: catalyst from the reaction products, high hydrogen pressures which are necessary to convert ,,.’:,,: .,, : benzene successfully and the short lifetime of the ruthenium catalyst [3]. [n the present study . ,., ,’;,,, , ... ,; j the formation of cyclohexene by a continuous catalytic gas phase hydrogenation of benzene .,. ,, ~,4’,+., “$! at atmospheric pressure was investigated, with the aim to offer an alternative way to conduct ,,4;,’,’,. ,’,,,.,,,I .?, :.’, ,:, the reaction with the benefits of a gas phase process. .. ;.. . ., .. ,.,,.. .-“ ‘.”...,!., ‘,,,).,:,,,., ,., k., ,,; ,. :,1 DGMK-TagungsberichI 9903, ISBN3-931850-59-5 169 2 Particularities of the reaction and aim of the study The formation of cyclohexene in the gas phase hydrogenation of benzene is more diftlcult, because of the thermodynamic and kinetic preference of cyclohexane in the reaction between benzene and hydrogen on catalyst surfaces (Figure 1). From thermodynamics, the formation of cyclohexane over cyclohexene is favored by about 75 kJ mol-’. The reactivity of the C=C bond in cyclohexene is much higher than the reactivity of benzene and furthermore, the control of the reaction by competitive adsorption is not possible because the values of the adsorption enthalpies of both benzene and cyclohexene on catalyst surfaces are of the same magnitude [4, 5]. These difficulties to overcome require a thoroughly tuning of the catalyst properties and the reaction conditions. Ruthenium (Ru) as one of the hydrogenation +,,. active component showed the greatest potential for the partial hydrogenation of benzene [6- 9]. The usage of a so-called reaction modifier, which interfere in the adsorption and mass ., transfer, is necessary to obtain at all the desired cyclohexene [6,9]. It was assumed, that the reaction path adsorption enthalpy / au. Figure 1: Scheme of the thermodynamic and kinetic particularities of the benzene hydrogenation to cyclohexene reaction modifier forms an adsorbate layer on the catalyst surface. This layer repels the intermediate cyclohexene from the catalyst surface and hinders its readsorption, and thus the further hydrogenation to cyclohexane [7]. This complicated mass transfer influenced coherence is illustrated in Figure 2. Benzene and hydrogen are adsorbed on the modifier covered ruthenium surface where they react firstly to cyclohexene. Part of the formed cyclohexene desorbs to the gas phase and because of the presence of the reaction modifier the readsorption is suppressed. The remaining adsorbed cyclohexene is further hydrogenated to cyclohexane, because of the very high rate of this reaction step [9]. However, the cyclohexene selectivities attained in the partial gas phase hydrogenation of benzene on powdered, supported and sol-gel derived Ru catalysts using water vapor as reaction modifier within the feed are substantial lower than those reached in the liquid phase reaction as described in the literature [2,3,5,6,8,9]. The search for a more efficient reaction modifier as well as for a novel designed ruthenium catalyst would be a suitable contribution to discover the hidden reserves of that reaction. From the literature about the liquid phase 170 ,-. =====p A H, Uu :1 o - . . .-~ ~.,,.- ~ ~ . .--,.’.. .’.-:. . >:: .;J. ! ...3 .,..,.,. Figure 2: Schemeof the reactioncourse in the partial gas phase hydrogenation of benzene on a .,,, ruthenium cakalystin the presenm of a reaction modirier. ., hydrogenation it is known the positive effect of the addition of metal salts to the liquid phase ., ,’”’ . on the cyclohexene selectivity. Especially due to the adsorption of Zn ions on the Ru surface ,,.’:.’1 > the formation of cyclohexene was enhanced ~,8,10]. Therefore, the effect of the addition of a second metal to ruthenium in the catalysts for the gas phase hydrogenation on the cyclohexene formation is particularly of interest. In the present study results of the partial gas phase hydrogenation of benzene are presented, which were achieved on specifically coated and on sol-gel derived ruthenium catalysts. Because of the low intrinsic residence times for the reactants, coated catalysts are predestined for the formation of unstable intermediate products like cyclohexene. Recently published results of the benzene hydrogenation on nickel containing coated catalysts confirm this assertion [11]. Using coated Ru catalysts several reaction modifiers were examined to find tirst the most effective one. Then the influence of the texture of the porous support oxide of these catalysts on the cyclohexene selectivity was examined. Eventually, by the / application of the sol-gel method the effect of the addition of second metals to ruthenium on the cyclohexene formation was proofed. , .. .,,, 3 Catalysts and experimental conditions The coated Ru/A120~Al catalysts (CRUC) were prepared as follows: Aluminum wire were anodicelly oxidized in an aqueous acid solution forming an oxide layer having cylindrical, regular and non-branched pores [12]. The resulting material with a thin porous oxidic layer on a metallic substrate is an excellent catalyst support for preparing coated catalysts. Ruthenium was immobilized by impregnation of the prepared support in a toluene solution containing ruthenium acetylacetonate. For sol-gel derived alumina supported Ru catalysts (SRUC) the same precursor Ru(acac), and, if necessa~ Zn(acec), or Cr(acec),, were dissolved together with aluminumtriisopropylate in alcohol. The hydrolysis of the alcoholate in the presence of water and acetic acid led to alcogels containing the catalytic active component [13]. Both types of catalysts were finished by a calcination in air at 573 K and a reduction in a H~N2 stream at a mild temperature of 473 K. A reduction at a higher temperature would lead to catalysts with an abated cyclohexene formation [13]. Some ., .,., 171 ,: ,~.,, !, characterization data of the resulting coated and sol-gel derived catalysts are summarized in Table 1. The coated catalysts CRUCI and CRUC2, formed by anodic oxidation and subsequent impregnation show great differences in their properties Ru content, pore diameter dP, and BET surface area. Note, that the Ru content and the specific surface area of the sol-gel derived Ru catalysts are higher than that of the coated catalysts, except the sol- gel catalyst SRUC2. However, the pore diameters of the sol-gel derived catalysts are much smaller than that of the coated catalysts. Table 1: Characterization dala of the mated Ru/Al,C ~satalysts (CRUC) and the sol-gel derived supported RuIAIZ03cataiy (SRUC) notation CRUCI CRUC2 SRUC1 SRUC2 SRUC3 catalyat RuIAI,031AI RuIA1203 Ru-ZnlA1203 Ru-CrlA1203 :- . electrolyte .5% (COOH)Z 1% H,PO, t.,. anodlc temperature I K 298 293 oxidation voltage I V 50 150 time I h 5 Id pm 34 24 porous Ru content’1 0.14 %3) 0.05 %3) 2.0 Y. 2.7 Y. 2.2 % metal content$) 1.6 % 1.6 ‘% texture dP I nm 2, 36 150 7 10 10 BETI m’lg’) 6.53) 1.53) 123 5 a4 V,J cm’1~ 2] 0.073) 0.043) 0.23 0.03 0.33 determined by 11WOX 2)N2adsorption 3)related to the oxide layer Hydrogenation experiments were performed in a conventional continuous flow apparatus consisting of devices for the adjustment, mixing and conditioning of the reactants, the hydrogenation reactor and traps to condense the organic products as well as analytic devices. Benzene and the liquid reaction modifier were vaporized with the help of two nitrogen bubbled saturators, afterwards mixed with hydrogen and passed over the catalyst. The composition of the feed and the reaction products were on-line analyzed using a gas chromatography. The benzene partial pressure and the molar modifier/benzene ratio in the experiments was kept constant at 0.11 kPa and 3, respectively. The hydrogen partial pressure was varied between 22 and 44 kPa. The total gas flow rate was adjusted to 6 I/h, balanced by nitrogen. 4 Results and discussion Benzene hydrogenation on coated Ru/A120/Al catalysts A variety of reaction modifiers, likely able to initiate the cyclohexene formation, were tested. The modifiers carbon monoxide, carbon dioxide, and formaldehyde gave no or only a slight amount of cyclohexene in the benzene hydrogenation, but experiments using water, methanol, ethanol and tetrahydrofuran were more successful. The benzene conversion and cyclohexene selectivity as well as the calculated yield on the coated Ru/A120~Al catalyst CRUC1 in the presence of the last mentioned modifiers are shown in Figure 3. The modifiers contain oxygen with free electron pairs but in different chemical environment. Note, that much higher water pressures were needed to initiate the cyclohexene formation than for the 172 . ... -,.- ,“,..,.:, .,, ,, ,. ,\ .-i’ ,. . -m ,., ,-, “3; .,,: -’- ,,. : pU=PE=p~ 0.33 kpa ethanOl (E),,_,H20~H, ~ ~ ,.. .’ ,,& ‘- ~.-: ,,, ~ tetrahydro: ; ‘H$~H2” ., --., .’.,,4- ‘ furan ‘(7) ,, ‘2C’6GH? ~ .,, .; ,. ,’ 0.05 0.1 0.15-’0.2.0.25 0.3 0.35 0.4 . . ““”’m”.. ‘conversion/ selectivity/ yield ...... __-...... Figure 3: eeruene conversions, aydohexene seleclivities and yields in the partial gas phase hydrogenation of benzene on the mated Ru catalyst CRUCI in the presence of various reaction modifiers T=353 K, P=IIO IrPa, p~=0.11 kPa, pm=44 kPa, WIP~= 19000 gh/mok organic modifiers. Whereas in the presence of water a high benzene conversion and a low .,,. cyclohexene selectivity was found, cyclohexene selectivities were much higher using the organica as modifiers. The selectivity to cyclohexene increased in the order of .-.,, ethanol , ,. ,, o~ -i. o 200 5 ‘ 10 Is time on streaml h Figure 4 Time on slream behavior of the benzene sonveraions, specific activities and cysfohexene seleclivitieaIn the partial gas phase hydrogenation of benzene on the coated Ru catalysts CRuCland CRUC2 T=353 K, P=I 10 kPa, p,= 0.11 kPa, PM=0.33 kPa,pw=14 kPa, WIPs= 19000 ghrmob 173 ><, ,.- .—_.- -L, ., . ‘..,. > !.. , ., ~=, .! state of the coatedcatalyst . . q . . ‘Z 80 - E a.: ~, A aftar 2nd rageneratlon ‘ m.-; ....>>..,> ,, . .. ,. ., n., ” . .-’. -o ..,’ 5 !. ,,. :. .,, < ‘‘ time on streaml h“’ ~ Figure 5: Time on stream behavior Specific activities and cyslohexene selectivities in the partial gas phase hydrogenation of benzene on the coaled Ru catalyata CRUCZ T=353 K, p=l 10 kPa, PB=0.11 kPa, PM=0.33 kPa,pm=44 kPa, WIP,= 19000 gtimob deactivation kinetic of second order [15]. The deactivation rate of the catalyst CRUCI is higher than that of the catalyst CRUC2. In addition, the latter catalyst showed the higher activity value at steady state. However, the higher benzene conversions was reached on the ,- . catalyst CRUC1, because of its higher metal content. The observed cyclohexene selectivities $, ,. increased initially, followed by a period of nearly constant value and reaches 38 ‘%0 in the case of catalyst CRUCI and 51 ‘Yo in the case of catalyst CRUC2. These results suggest, that ,, higher activities and selectivifies may be obtained on coated catalysts with large pores and low surface areas. Besides the high cyclohexene selectivity the possibility of regeneration of the coated Ru/A120~Al catalysts is another benefit. This is illustrated in Figure 5 where the time on stream behavior of the coated catalyst CRUC2 is depicted. The specific catalyst activities and selectivities were observed during three catalytic runs. First a fresh catalyst sample was investigated in the partial benzene hydrogenation, afterwards a regeneration procedure including reoxidation at 573 Kin air and subsequent reduction at 473 Kin a H~Nz stream was done. Then a second hydrogenation experiment was started with the regenerated catalysts. It can be seen, that in the hydrogenation after the first regeneration step the specific activity of the catalyst was slightly higher, whereas the cyclohexene selectivity was lower. In the third catalytic run, after the second regeneration the specific activities and cyclohexene selectivities were identically with the data of the foregoing hydrogenation experiment. Benzene hydrogenation on sol-gel derived sr.rpporfed Ru/A1203 catalysfs Besides a Ru/Alz03 catalyst, several others were prepared by the sol-gel method described in chapter 3, where a second metal was added to the ruthenium. This was done in order to investigate the influence of the second metal within the hydrogenation catalysts on the cyclohexene formation. It results, that second metals, e.g. Fe, Cu and Ni enhance the cyclohexene selectivity, but best results were obtained by the addition of Zn or Cr. In Figure 6 the benzene conversion, specitic activity and the cyclohexene selectivity are shown as a function of time on stream for the Zn modified catalyst SRUC2 and the Zn free catalyst SRUCI, which contains only ruthenium. The course of the conversions and selectivifies was similar to that described for the coated ruthenium catalysts (see Fig. 4). The specific activity 174 . ..., ., .7 ., .6 ; q. s .5 % ~ ‘o. .4 % 0 ., ,’, ~.., .3 ; ,., 0. o .. 25 0 .0. .1 s ,. ?... .. ’.,, ‘-. ‘, .:, time’on e$eam/ h Figure 6 lime on stresm behsvior of the benzene mnveraions, specific activities and cyclohexene ,, selectivities in the parlial gas phese hydrogenation of benzene on sol-gel derived catalysts SRUCI and :,’ SRUC2. (T=353 K, p=l10 kPa, p.= 0.11 kPa, PM=0.33 kPa,pm=44 kPa, W/p,= 1670h4mol) ‘,-, ., ’-’.,, .;”: and conversion of the Zn modified catalyst SRUC2 were slightly lower than those of the :..<:, catalyst SRUCI, but the maximum cyclohexene selectivity was enhanced from 2070 to more ,, .,:...,,. than 40% due to the availability of Zn. In comparison to the coated Ru catalysts without metal ,! ,,.’ .’ . ,’. promoters, the sol-gel derived catalysts showed lower specific activities and cyclohexene , . . . . However, high cyclohexene yields was obtained using the sol-gel derived Ru/A1203 catalysts in the partial benzene hydrogenation. In Figure 7 is depicted the benzene conversion, cyclohexene selectivity and the calculated yield of cyclohexene as function of the time on .,, ,. ,<.,,.,, stream in the benzene hydrogenation on the catalyst SRUC3, in which Cr was added as .--, ,., -., second metal to Ru. In the first period of 18 hours at a hydrogen partial pressure of 44 kPa ., .,., ,, the observed benzene conversion was sz~o and the cyclohexene selectivity 157. ;...;..; “’,.’.,’, corresponding to a yield of about 8Y0.After diminution of the hydrogen partial pressure to 22 ,.... ‘:; ,, . .,- :.,. :,, ..,. -:, ., ! , J’ . ■ wnverslon Aselestkity and O yield of svclohexene ., .,,., ..- .- .,,7 ,. ., -! 0.8 ,’.’ ,; 0.6 :“:,,,, . . ,: ,,. , 0.5 ,’” .:, .,’.., ,,.:. , f :,.;,, .,, :, ., .~,,;f. ‘. i ,,,.“’%,, : 1 .. time on etream/ h Figure 2 Time on stream behavior of the benzene conversions, syclohexene selectivities and yields in the parilsl gas phase hydrogenation of benzene on se-gel derived satslyst SRUC3. (7=353 K, P=I1O kPa, p,= 0.11 kPa, p.= 0.33 kPa,p@41ZZ kPa, WIP~= 1670 gtdmol) ..- .,, ,., J., ., ~<,:,..,, ,, kPa in a second period of 12 hours, the benzene conversion decreased to 13’10 and the cyclohexene selectivity increased to sq~.. This resulted in a yield of cyclohexene between 4 to 5 Y.. 5 Conclusion The cyclohexene formation in the gas phase hydrogenation of benzene on coated Ru/AlzO~Al catalysts and sol-gel derived Ru/AlzO~ catalysts was investigated. By the application of several reaction modifiers in the gaseous reaction mixture which can act as a kind of a diffusion barrier, methanol was selected as the most effective modifier to yield cyclohexene with high selectivities at moderate benzene conversions. Coated Ru/A120~Al catalysts with various porous texture, which was adjusted by the preparation conditions gave different cyclohexene selectivities in the hydrogenation reaction. From that, it may concluded, that large pores and low surface areas of the coated catalysts improve the cyclohexene formation. The regeneration of the coated catalysts were performed by a treatment in air. Sol-gel derived supported Ru/A1202 catalysts are also suitable for the cyclohexene formation in the presence of methanol, especially if a second metal beside ruthenium was present. On such catalysts containing zinc or chromium as a second metal, cyclohexene yields up to 8% were obtained. The investigations of coated ruthenium catalysts which contain zinc as the second metal, together with the determination of the effect of the second metal on the catalyst properties establish the focus of further interest. The authors express their gratitude to the “Max-Buchner-Forschungsstiftung” and the “Fends der Chemischen Industrie” for financial support of this work. [1] Campbell, M.L. in Ullmann’s Encyclopedia of Industrial Chemistry, 5rncd., Vol. A8 (Verlag Chemie, Weinheim, 1987) p. 209. [2] Nagahara, H., One, M., Konishi, M., Fukuoka, Y., Appt. Surf. Sci. 121/122 (1997) p. 448. [3] Odenbrand, C.U.I., Lundin, S.T., J. Chem. Techn. Biotechnol. 30 (1980) p. 677. [4] Smith, H.A., Meriwelher, H.T., J. Am. Chem. Sot. 71 (1949) p. 413. [5] Mizukami, F., Niwa, S., Ohkawa, S., Katayama, A., Stud. Surf. Sci. Catal. VOI.78,eds. Guisnet, M., Barbier, J., Barrault, J., Bouchoule, C., Duprez, D., Perot, G., Montassier, C., Elsevier Amsterdam (1993) p. 337. [6] Don, J.A., Scholten, J.J.F., Faraday Discuss. Chem. Sot. 72 (1982) p. 145. r] a) Struyk, J., Scholten, J.J.F., Appl. Catal. 62 (1990) p. 151. b) Struyk, J., d’Angremond, M., Lucas de Regt, W.J.M., Scholten, J.J.F., Appl. Catal. 63 (1992) p. 263. [8] Dbbert, F., Gaube, J., Chem. Eng. Sci. 51 (1996) 11, p. 2873. [9] Patzlaff, J., Gaube, J., Chem.-lng.-Tech. 69 (1997) 10, p. 1462. [10] Soede, M., van de Sandl, E.J.A.X., Makkee, M., Scholten, J.J.F., Stud. Surf. Sci. Catal. Vol. 78, eds. Guisnet, M., Barbier, J., Barrault, J., Bouchoule, C., Duprez, D., Perot, G., Montassier, C., Elsevier Amslerdam (1993) p. ?45. [11] Dietzach, E., Rymsa, U., Hbnicke, D., Chem. Eng. Technol. 22 (1999) 2, p. 130. [12] Keller, F., Hunter, F., Robinson, D.L., J. Eleclrochem. Sot. 100 (1953) p. 411. [13] Gonzalez, R.D., Lopez, T., Gomez, R., Calal. Today 35 (1997) p. 293. [14] Zhanabaev, B.Z., Zhanozina, P.P., Utelbaev, B.T., Kinet. Catal. 32 (1991) 1, P.191. [15] Germain, J.E., Maurel, R.. Compt. Rend. 247 (1958) p. 1854. 176 ., DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erfangen 1999 L. Simon 1),J. G. van Ommen 1),P. J. Kooyman 2),A. Jentys 3),J. A. Lercher 3, 1)Catalfic processes and Materials, Faculty of Chemical Technology, Enschede, .. The Netherlands ,, ’-,,, z)National Centre for HREM, Delft University of Technology, Delft, The Netherlands ,,,.! 3)Technische (.jniversitatMOnchen, Lehrstuhl fur Technische Chemie ii, ,..., Garching, Germany ,., ,/, Sulfur Tolerance of Alkali Exchanged Zeolites for Benzene Hydrogenation Absfract . ,-, ,:, .,.’, Pt suppo~ed on Na+ containing mordenites and K+ containing L zeolites was applied for benzene :;, .: ,. hydrogenation in presence of thiophene. The effects of pressure, temperature, and proton / alkali ,,,,<,..:’, : metal cation molar ratio upon the catalytic properties were studied. Catrdysts were characterized ... ,, ,, by XRF, BET, ammonia TPD, hydrogen chemisorption, HREM, and EXAPS. Increasing the ‘, ,,,, ,. hydrogen partial pressure or the total pressure significantly increased the benzene conversion in ,. -,, ,,, presence of 50 ppm thiophene. Increasing the temperature also decreased the impact of 50 ppm .,. :. I .,”. thiophene upon benzene hydrogenation. Arr optimrd proton / alkali metal cation molar ratio seems ’.,.;,. ::,. .,...... ,,. ~ to exist for an optimat sulfur resistance of the catalyst. EXAPS and HREM did not indicate ,,, ., : ., .,, .,, .:,-,,,, -,, - ~ changes in the platinum particle size aller two hours of benzene hydrogenation in presence of 50 ,t - ppm thiophene. ., “ ,, ,. .“;-. ~ ,“ -, ,“. f-:, .,, .,. s Introduction .:...... : . . ., .- ‘.,.:, , ‘) One of the major drawbacks of the noble metal based catalysts in hydrogenation reactions ,. ,, f. :, “,, ,, .-: -:.,!.:.-. is their sensitivity towards sulfirr. In the case of biflmctional reactions, it has been reported that “.,,..;, ,. -<. catalysts based on Pt/Pd supported amorphous silica alumina and deahrminated zeolites can ,, ,.’ .. . tolerate up to 1000 ppm sulfur in the feed [1]. Most of the sulfur tolerant noble metal catalysts .-$’ reported are supported on acidic silica-aluminas or zeolites. Tire effect of acidity on the sultirr tolerance has been explained via influencing the electronic properties of the noble metat [2,3]. In the case of monofinctionrd reactions such as the noble metal catalyzed dehydrocyclisation of n- .-.,, - .,,,... ; -1, , heptane to benzene, Pt supported on basic supports such as on K+ and Ca2+ exchanged zeolite LTL are reported to be very sensitive to sulfur and may be poisoned by the presence of a few ppm of sulfur in the feed [4]. In both cases, pronounced sintering and migration of particles out of the zcolite pores in the presence of nucleophilic sulfur compounds were identified as the main cause of catalyst poisoning. In order to better understand the sensitivity of noble metal based catalysts towards sulfur during monofunctional reaction, the influence of exchange level of alkali cations in MOR and LTL type zeolite during benzene hydrogenation in the presence and absence of thiophene was studied as model reaction. Experimental The H form of the mordenite was prepared by ion exchanging three times 20 g of NaMOR (TOSOH, Si/Al = 18.6, ref. HSZ-641NAA, lot T960402 ) with 300 ml of a 0.1 M solution of NfiN03 (MERCK) at room temperature. After each step, the zeolite was filtered and washed with 50 ml of deionized water. Calcination in dry air with a flow of 100 ml.min-l per gram zeolite was performed at 723 K for 3 hours with a heating rate of 5 K.min-l. Platinum supported catalysts were prepared by liquid phase ion exchange at room temperature using tetraammine platinum hydroxide (Pt (NH3)4 (OH)2.X H20, 59 %Pt, STREM) as precursor, identified as the best method for platinum incorporation [5]. The required volume of the precursor solution in order to obtain a platinum loading of 1 wt% Pt was added dropwise at a rate of 12 ml.min-’ to the zeolite suspended in deionized water. The suspension containing the zeolite and the precursor was stirred overnight. The sample was then filtered, washed with deionized water, dried for 24 hours in air at room temperature, calcined in a flow of air of 40 ml.min-] per gram of catalyst at 493 K for 2 hours (heating rate of 0.5 K.min-t) and finally reduced with pure hydrogen in a flow of 35 ml.min”l per gram catalyst at 623 K for 1 hour (heating rate 0.5 K.min-l). For comparison, Pt supported on L zeolite catalysts (LTL, TOSOH, Si/Al=12, Ref. HSZ500KOA, Lot 1041) were prepared via the same method. In the case of Pt-Na/MOR and Pt-K/L samples, back exchange of the protons present in the zeolite after reduction was earned out at 323K by stirring the Na+and K+ form of the catalyst with a solution of 0,033 M of NaN03 and KN03, respectively. After filtering, the catalysts were washed with deionized water and dried in air. Chemical compositions of the catalysts were determined by X-Ray fluorescence spectroscopy [6]. The concentration of accessible Pt was determined by hydrogen chemisorption. Between 0.5 to 1 g of the calcined and pre-reduced catalyst was loaded in the sample chamber of an all glass volumetric system. The catalyst was reduced in flowing (50 ml.min-]) hydrogen for at least one hour at 623 K. The system was then closed and evacuated (p < 1.105 mbar) at the reduction temperature for one hour. After cooling the system to 293 K, hydrogen (about 500 mbar) was admitted into the sample chamber and was allowed to equilibrate with the catalyst overnight. To obtain a sorption isotherm, the hydrogen pressure was lowered in equilibrating time steps of one hour per point. HREM pictures were obtained using a Phillips CM30 T electron microscope with a LaB6 filament as the source of electrons operated at 300 kV. Samples were mounted on a microgrid carbon polymer supported on a copper grid by placing a few droplets of a suspension of ground sample in ethanol on the grid, followed by drying at room temperature. 178 EXAFS spectra were obtained on beamline B-2 at CHESS, Ithaca, NY, USA and on beamline ROMOII at Hasylab, DESY, Hamburg, Germany. The storage ring operated with an electron energy of 4.5 GeV and a current of 100 mA. The Si (111) double crystal monochromator (CHESS) and the Si (311) double crystal monochromator (DESY) used were detuned to 80% of the maximum intensity to minimize the presence of higher harmonics in the X-ray beam. The weight of the sample was chosen to obtain a totaI absorption of 2.5. The pre-reduced samples were reduced in-situ at 623 K for 1 hour prior to benzene hydrogenation in the presence/ absence of thiophene (623 K, 1 bar, H~C61&=39.1). Data were collected at the Pt Llll edge (11.564 eV) and analyzed with the WINXAS97 software [7]. For the kinetic experiments, the catalysts were tested in a stainless steal tubular reactor (5 mm diameter) using 90 mg of catalyst mixed with at least the same amount of quartz. The catalysts were calcined in-situ at 673 K in a flow of 55 ml.min-l of air foIIowed by an in-situ reduction in a flow of 200 ml.min-l of pure hydrogen at 623 K for 1 hour. A solution of 50 ppm thiophene in benzene (ALDRICH) was mixed with a flow of pure hydrogen using a high-pressure syringe pump (ISCO) and passed through the reactor. The reaction was carried out for two hours. Rmction products were analyzed using an on-line gas chromatographywith equipped with a DB-1 column and a FID. Results Characterization Some resuks of the characterization of the supported platinum catalysts are listed in Table 1. The XRF analysis showed that the alkali back exchange of the catalysts after reduction of platinum increased the amount of alkali cations exchanged in the zeolite by 23% for Pt-Na/MOR and 14% for Pt-K/L. Acid site concentrations calculated by integration of the peak areas obtained by ammonia temperature program resorption showed a decrease in the acidity of the catalysts with increasing alkali content. Table 1 Characterization of the platinum supported catalysts. Fraction of [~] Pt loading BET surf. area Acid site concentration (mol %) (Wt%) (m’/g) (mmollg) Pt-HIMOR 97.9 1.13 250 1.30E-03 Pt-NaH/MOR 29.6 1.14 283 2.57E-04 Pt-Na/MOR 22.9 0.94 n.d. 1.IIE-05 Pt-KH/L 15.2 1.09 n.d. 4.4~E.04 Pt-K/L 3.3 1.06 n.d. n.d. HREM pictures of reduced Pt-NaI-I/MOR, after 2 hours time on stream in benzene hydrogenation in presence and absence of 50 ppm thiophene (1 bar, 623 K, WHSV= 0.05 h-l, HZ/CcHcmoku ratio = 7.6) are shown on Figure 1. An average 4 nm platinum particle size was determined by this method. The platinum particles are mainly located inside the zeolite pores and only a small fraction is present on the surface of the zeolite particles. Changes of the platinum particle size . . . . . ‘., ,, ,, 179 !, Figure 1 HREM pictures of P(-NWfIdOR(a)reduced.(b)after120minu[csof bmzenchydrogcmnion,(c) after 120 minutes of bcnzcnc hydrogenation in presence of 50 ppm thiophcne. were not detected after two hours time on stream during benzene hydrogenation. The mme average platinum particle size as for Pt-NaH/MOR was also obtained by HREM in the case of Pt- WMOR. Also in this case, a significant increase in the particle size was not observed after 2 hours time on stream. The BET surface area showed a decrease of 11 and 13% for Pt-WMOR and Pt-NaH/MOR after platinum incorporation compared to the parent materials. This indicates that minor pore blocking may have occured in the presence of the 4 nm platinum particles inside the mordenite framework. XRD patterns taken after platinum ion exchange showed a decrease in the zeolitc diffraction peak intensities and a slight amorphisation of the support indicating the generation of structural defects in the zeolite framework around the platinum p~iticles. This decrease in cristallinity was not observed on the XRD pattern of the Pt-K/L indicating Icss dcforrrmtion of the zeolite framework. in-sim EXAFS was performed on the catalysts. The Fourier transform magnitudes of kz-weighted (J) —..Rp.a Iwsn. bcfoltmwlam —-- PIFoJ .. .. . P&N.JMOk bcfom n-.ctmn I — Pc.m. .8klxo Iluml.w ----- P.K13-WII,Ihb”l,lm II,, — PvNLJMORwflmti Ihwn!ux ----- l%N.IhiORwth Ownhx Figure 2 Fourm Irmsform of the PIIIIedge l?-weighted EXAFS oscilkmon, (a) PI-K/L and (b) PI-NtiMOR. I&l .,. ., , EXAFS oscillations in Figure 2 show a very small peak of the first Pt-Pt coordination shell for all the catalysts compared to the Platinum foil. This peak was found to be larger in the case of the acidic support compared to the more basic ones. After reaction in presence or absence of thiophene, the intensity of the Pt-Pt first coordination shell was smaller than for the fresh catalyst. Pt-Pt coordination numbers and interatomic distances calculated from the EXAFS analysis are compiled in Table 2. Significant changes were not observed in the Pt-Pt coordination numbers after 60 minutes time on stream benzene hydrogenation in presence/ absence of thiophene for the alkali exchange samples. In the case of Pt-H/MOR an increase of the coordination number after one hour ben~ene hydrogenation in presence of thiophene was observed. Table 2 Results of the EXAFS analysis of the catalysts reduced and after benzene hydrogenation. Nn.R rn.fi (~) Auz ~.n *10_3(~z) Pt-H/MOR Prior to reaction 9.3 2.77 1.3 After reaction without thiophene 9.1 2.76 3.1 After reaction with thiophene 10.7 2.76 2.3 Pt-NaH/MOR Prior to reaction 9.0 2.75 4.0 After reaction with thiophene 10.5 2.75 4.4 After reaction without thiophene 10.7 2.75 4.6 Pt-Na/MOR Prior to reaction 8.8 2.74 3.5 After reaction without thiophene 8.1 2.73 3.6 After reaction with thiophene 8.1 2.73 5.9 Pt-KIWL Prior to reaction 5.8 2.79 7.6 After reaction without thiophene 7.8 2.77 1.2 After reaction with thiophene 7.0 2.79 3.5 Pt-K/L Prior to reaction 7.2 2.76 4.4 After reaction without thiophene 6.6 2.74 6.0 .. After reaction with thiophene 5.3 2.76 4.4 ,,.,. ‘ Average particle sizes between 1.5 and 4 nm, containing 200 to 1000 atoms can be estimated from the Pt-Pt coordination numbers between 8 and 10 which is in agreement with the average ‘. ’...-4 ,2 ~.,,:(;,’ $ estimated by HREM [8]. The coordination numbers in the case of Pt supported on LTL zeolite are : <;;,, .’:’? smaller than for the other support in line with earlier observations [9]. EXAFS results support ., : )$ ,,1..,:./,,., ,,’:; those obtained by HREM showing that the dispersion of the platinum inside the zeolite ,, J,,.!#.,,. ., ,;, >,J framework does not depend on the concentration of Na+ or K+cations present. .,,.!:,,, . ,“+; ,-,,.,,\.,..,’:,.. ‘1 ‘~.’,f::,:,-:,. ,-.,:< Kinetic resrdtsfor benzene hydrogenation The effects of temperature and the proton/ alkali metal cation ratio on the sulfur tolerance during benzene hydrogenation are compiled in Figure 3. In the case of Pt supported on MOR the increase of the temperature from 523 K to 623 K increased the rate of benzene conversion in the presence of 50 ppm thiophene significantly. The acidic form and the 77.1 % exchanged Na+ or the 96.7 Yoexchanged K+ catalysts showed a much lower conversion than the less exchanged >’ alkali samples indicating that an optimum in acidity / basicity may exist with respect to the sulfur ,., resistance of the catalysts. In the case of Pt supported on LTL, the rate of benzene conversion in ,,. . ,.. ,1. 181 —.- . 2m 4 rwi?vcn 1.4E02 n R-NIH’MU A h-h- 1.2602 X60! I.CGO?. / Lsrm MG03 rKw13 Itlw 4&03 ● 5MX35 ‘MEo.? -——n WX3Co acE3co 5034-+2msowmlrim6Y36WY31.ms$omswao(colxo I Taqwuc(K) Tarpmm(K) Figure 3 Effect of tcmpcraturc on the rate of benzene hydrogenation in presence of 50 ppm thiophcnc isf[cr 60 mlmncs time on strum (25 btw. WHSV=O.05 ht. H2/C6H6=7.6). prcscncc of thiophene was much higher than in the case of Pt supported on MOR. However, also Pt-KH/L showed a markedly higher rate of benzene hydrogenation in presence of thiophene compared 10 Pt-K/L. This high benzene hydrogenation rate results in a almost independency of temperature, possibly caused by apprcmchitrg [he equilibrium (80% benzene conversion) or diffusion limitations. The effects of pressure on the sulfur tolerance during benzene hydrogenation are shown in Figure 4. All catalysts showed a strong initial deactivation during the first 20 minutes followed by stable catalytic activity. Quantitatively, this amounted to 5%, 9?Zoand79Yorelative Iossofactivity after 60 minutes time on stream in the presence of 50 ppm [hiophene for Pt-NaH/MOR, Pt-KH/L and Pt-H/MOR, respectively. Pt-NaH/MOR and P[-KH/L were significantly more active and stable than Pt-H/MOR. To study [he influence of pressure upon the catalytic properties the total pressure of the system was varied between 1 and 10 barkecpingthc WHSV of 0.05 h-t and the HJCGH6 molar mtio of 7.6 constant. The TOF increased for all catalysts with increasing total pressure. Presence of thiophene in the feed increased the initial deactivation of the catalysts, but the catalysts were all stable after20 minutes [imcon strram. When the partial pressure of hydrogen was increased from 10 to 25 bar(WHSV =0.05 h“’, H2/C6H6=20.6), 5 to 6 times higher TOFs were observed for all catalysts in presence of 50 ppm thiophene. P~rtially exchanged platinum suppoficd catalysts exhibitcd:igain thchighest TOF. Thcdccrease of the initial activity was much lower at high pressures than in the case of lower pressures. Discussion Kinetic results show that under all experimental conditions tested the catalysts reach stable catalytic activity after an initial strong deactivation in presence of thiophcne. The sulfur compounds are expected to deactivate the catalysts by two mechanisms, i.e., sintering of the platinum particles and direct poisoning of the platinum surhce. The Pt particle size was stable in the case of alkali cxchangcd zcolitcs and increased in the case of the Pt supported on acidic sample. This suggests that the presence of alkali cations induces stabilintion of the platinum 182 .. . . ,., , 0.5 %nu n ❑ ❑ nun :., o L •4.+~—*a-.-*-* o o 50 Ico 50 100 Tw on strwrr(nin) Twon slrum(rrin) Figure 4 Benzene hydrogenation in presence of 50 Pr-WMOR ppm thiophene, 623 K at + 1, Q 10 mrd A 25 bar. particles inside thezeolite pores. Due to the 2 large Pt p~rticle size, we expect those i particles to be trapped inside the structure. 1.5 Thesize and electrostatic interactions of the 1 alkali cations favored this entrapment. The relatively low sulfur sensitivity of Pt- 0.5 ~AAAA A N@MOR and Pt-KH/L compared to Pt- AAA AA H/MOR agrees well with the stabilization of o Lr_s- the Pt particles inside the zeolite by the alkali g o 50 100 metal cations. Even in presence of sulfur Trmon slrmrrr(rrir) ., which supposedly decreases the interaction ,!.,. between ‘metal ~articles and sumIort and favors their mirm~tion [101.-. the me~~luarticles remain immobile. ConsecwentlY,.- dmctivation by sintering does not occur in Pt-NaH/MOR and Pt-KH/L. This is concIuded from the results of in-- situ EXAFS after benzene hydrogenation in presence or absence of thiophene. Although 4 nm particles trapped inside the zeolite will certainly block the pores (4 times the average pore diameter), the local destruction of the lattice detected byXRD could still allow the flow of .,” reactants/products to reachlleave the surface of the metal. Increasing the hydrogen partial pressure markedly enhanced the conversion of benzene. We assume that the increase of the partial pressure of hydrogen induces an increase of the coverage with hydrogen which increases the rate of benzene hydrogenation and markedly reduces the coverage of thiophene on the platinum surface. The presence of the alk~li in the zeolite framework increases the electron density of the platinum particles. As shown by Mallmann et al. [11] in the case of zeolite X the electron density of the ;.? platinum particles increases in the presence of alkali cations in the order Li < Na e K c Cs c Rb. . . Applied to the presently studied catalysts, the presence of the Na+ and K+ cations should make Pt ‘. ,.,,. .; somewhat more electron rich compared to H+. The fact that theplatinum containing 77.1%Na+ ,,,, . . on MORand96.7% K+on Lzeolite aremore sulfur sensitive than platinum containing 70.4% .,., Na+ on MOR and 84.8% K+ on L suggests that an optimal proton/ alkali metal ratio exists for a ,., , ,,,’ 183 maximum sulfur rcsistonce of the catalyst during benzene h y(frogcnation. Note that as the influence of [he cations is not the same for the different zeolitcs. Pt supported on LTL which was shown to be very active in reforming reactions [12], but very sensitive to sulfur [13], showed the highest activity for benzene hydrogenation in presence of 50 ppm thiophene. This suggests that the high sulfur sensitivity is related to the specific hydrocyclizution and the specific reaction conditions. It should be noted that typical reaction conditions for reforming over Pt-BaK/LTL catalysts are 673 K, atmospheric pressure, under hydrogen flow [10]. Under these more severe conditions sintering has been shown to be importmrt, while it did not occur under the reaction conditions used in the present study. Conclllsiolls Pt supported on alkali exchange MOR and L zeolites has been shown to be more sulfur tolerant for benzene hydrogenation than Pt supported on the purely acidic one. The presence of the alkali mctul cations seems to stabilize the platinum particles inside the zcolite pores during the benzene hydrogenation in presence of thiophenc. The proton / alkali metal ratio for these two types of cinalysts has been shown to be an important parameter for the increase of the sulfur tolerance of [he platinum containing alkali metal zeolitcs during benzene hydrogenation reaction. An optimum alkali cxchangc degree seems to exist. This work wos supported by STW/NOW, The Netherlands under the project number 349-3787 and has been performed under the auspices of NIOK and PIT. [1] J.K. Mindcrhoud and J.P. Lucien, Ew-opmmpakwr No. 303332 (1989). [2] R.A. Dana Bcttu and M. Boudwt, Proc. S’” Meru. Congr. CafaL, Palm Beach 2, 1329 ( 1972). [3] P. Gallezot, J. Datka, J. Massardicr, M. Primet and B. Imelik, Proc. 6’”hfern. Congr. Carol., London, 696 (1977). [4] G.B. Vickcr, J.L. Kao, J.J. Zlcmiak, W.E. Gates, J.L. Robbins, M.M.J. Treaty, S.B. Rice, T.H. Vwrderspurt, V.R. Cross, and A.K. Ghosh., J. Card. 139,48 (1993). [5] S. Feast, M.English, A. Jcntys, and J.A. Lercher, Appl. Card. A 174, 155 (1998). [6] M.H.J. Bckkers and H.A. van Sprang, X-Ray Spectroscopy, 26, 122 (1997). [7] T. Ressler, J. Physique W, 7, C2-269 (1997). [s] M.S. Tsou, B.K. Tee, and W.M.H. Sachtler, Lmgmuir, 2,773, (1986). [9] M. Vaarkamp, F.S. Modica, J.T. Miller, and D.C. Konningsberger, J. Cafal. 144,611 (1993). [10] M. Vaarkamp, J.T. Miller, F.S. Modica, G.S. Lane and D.C. Koningsberger, J. Caful. 138, 675 ( 1992). [11] A. de Mallmann and D. Barthomcuf, J. Chiol. Phys., 87,535 (1990). [1~] J.R. Bernard and P. J. Nury, US pa[eiir No. 4,104,320 (1978). [13] T.R. Hughes, W.C. Buss, P.W. Tamm and R. L. Jacobson, in Y. Mumkami, A. Iijima and J. W. Words (Eds), New Developtnenfs in Zcolite Science ond Technology (Proc. 7’”In[ Zco/ife Conf, Tokyo, Au:Hsf 17-22, 1986) Kodansha/Elsevier, Tokyo/Amsterdam, p 725 (1986). 184 ,, ., .“ DGMK-Conference ‘The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 ,. ,. ,.. ,. M. W. Haenel 1),St. Oevers 1),St. Hillebrand 1),W. C. Kaska’) >. 1)Max-planck-lnstitut fi,ir Kohlenforschung, Mulheim/Ruhr, Germany z)University of California Santa Barbara, Department of Chemistry, Santa Barbara, USA Polycyclic Arenes and Heteroarenes as Backbones of Diphosphine/amine Ligands for Thermostable Homogeneous Catalysts Dedicated to Dr. Gerd Collin on the occasion of his 6P birthday 1. Introduction Phosphine and amine Iigands are important tools to stabilize transition metal atoms as complexes in organometallic chemistry and homogeneous catalysis and to modify the reactivity of such metal complexes within a broad range. We have been pursuing the concept of using polycyclic arenes and heteroarenes as rigid backbones for the synthesis of new diphosphine/amine Iigands for some time. Examples are the di- phosphines 1-6 (R = phenyl, alkyi) derived from anthracene, acridine, dibenzofuran, dibenzothiophene and 9,9-dimethylxanthene [1-8]. fq ~,~,, ~ @ 2 2 s Pt+P PPh2 l: X=CH 3X. O 2:X=N 4X. S 5 6 We report hereon the syntheses of various anthraphos and acriphos Iigands 1 and 2 which serve as tridentate PCP or PNP Iigands to form highly stable transition metal complexes of the type 7 or 8, respectively. ‘1’; ‘1” \ / \ y“ ‘ R2F;—PR2w R2P—~—w PR2 n 7“ 8 Some of these complexes 7 and 8 are expected to be useful as highly thermostable homogeneous catalysts. First examples of catalytic reactions in carbon-monoxide chemistry, in C-C coupling reactions and in the functionalization of alkanes (C-H activation) are presented. 2. Results and Discussion 2.1. Syntheses of the Ligands Anthraphos (la, R = Ph) and acriphos (2a, R = Ph) can be prepared from 1,8- difluoroanthracene (9) and 4,5-difluoroacridine (10) by nucleophilic substitution with potassium diphenylphosphide. We have developed practicable synthetic routes for the preparation of the previously unknown difluoro compounds 9 and 10 ~,8]. Similar to 1 and 2, 1,8-bis(dimethylamino) anthracene (11) and 4,5-bis(dimethylamino) acri- DGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 1s5 dine (12) were obtained by nucleophilic substitution of 9 and 10 with lithium dimethyl- amide [8,9]. Since in this case it was observed that the substitution of the second fluoro substituent of 9 and 10 is considerably retarded for electronic reasons, two subsequent substitution reactions by first applying amides and then phosphides are exDected to lead also to anthracenes and acridines which have one amino and one phosphino group each [9]. Ph2P bPh2 # F Me.$ ~Me2 la X=CH 9: X=CH ll:X=CH 2aX=N 10:X=N 12X=N The substituents on phosphorus and nitrogen can be easily varied by using corre- spondingly substituted phosphides and amides in the substitution reactions of 9 and 10. Even chiral anthraphos and acriphos Iigands should be available by our synthetic route [1O]. We have also prepared anthraphos-tbu (1b) and anthraphos-cy (1c) in high yields from the reactions of 9 and potassium di~butylphosphide or lithium dicyc- Iohexylphosphide, respectively [9]. qq_/q L_$jlq ‘Bu2P P1Bu2 cy2P Pcy2 lb 9 lC 2.2. Metal Complexes Treatment of anthraphos la with the dichlorides of the group 10 metals nickel, palla- dium and platinum in refluxing 2-methoxyethanol eliminated hydrogen chloride to form the cyclometallated square-planar chelate complexes 13, 14 and 15 [1,9]. Because of the cyclometallation at the anthracene C-9 atom, anthraphos la is acting as a tridentate PCP Iigand in these complexes. Similarly, the reaction of 1a and rhodium trichloride in refluxing 85 % aqueous i-propanol resulted in the cyclometalla- ted octahedral dichloro-rhodium complex 16 in which one coordination site is occu- pied by a solvent molecule. The extreme stability of the cyclometalled anthraphos MC12, - HCI RhC13, -HCI :1” :1” :1” //~ //~ // M = Ni, Pd, Pt ‘PrOH / H20 w w Ph2Pw‘M ~ PPh2 Ph2P PPh2 ‘h*:~ ~hz~ 13: M= NI cl la 14 M=Pd lPrOH 15: M=Pt 16 ‘1” ‘1” \// \ // \‘1” // Ph2P—~—w PPh2 1Bu2Pw— ~d—P1Bu2 Cy2P—w?Pd— PCY2 CN cl cl 17 18 19 166 .,,, . complexes is shown by treatment of 13 with aqueous potassium cyanide, which did not remove the nickel(n) ion, but formed the cyano-nickel complex 17 by chloride/cyanide exchange (complex stability of [Ni(CN)d]2-: K = 1031;compare also [1l]). Similarly to la, anthraphos~bu (lb) and anthraphos-cy (It) reacted with palla- dium dichloride with elimination of hydrogen chloride to form the cyclometallated chloro-palladium complexes 18 and 19. However, unlike la anthraphosJbu (1b) and rhodium or iridium trichlonde in refluxiing 85 Y. aqueous Fpropanol yielded the chloro- hydrido-metal complexes 20 or 21, respectively. Their formation is explained by reduction of the metal tnchlorides to metal monochlorides by the solvent (MCIS + ‘PrOH + MCI + 2 HCI + acetone) followed by oxidative addition of the anthracene C-H bond [11-13]. In agreement with this mechanism 20 also was formed as the major product from the reaction of 1b and bis(cyclooctene) chloro-rhodium(l) dimer in refkrxing benzene. In this case, however, in addition to 20 two further rhodium compounds were formed as side products, one of which was identified by NMR as the alkylidene-chloro-rhodium complex 22. When a tokrene solution which contained 20 as a major component and 22 as a minor component was heated, the ratio changed in favor of 22 which indicated a possible rearrangement of 20 into 22. H, ,H 20+ ,Bp—M IL PtBu2 - ,Byq”2-‘Bu2Pw— Rh—ptBu2 M = h, Rh lb I c1 cl 20: M = Rh 22 21: M=lr In analogy to the preparation of the dihydrido-iridium complex 34 of the PCP Iigand 2,6-bis(di-’butylphosphinomethyl)benzene (see below) [13], the red chlorohydrido- iridium complex 21 was reduced in pentane solution by LiBHEta under 1 atm of hydrogen to yield the yellow tetrahydndo-iridium complex 23. Upon applying high vacuum at room temp. 23 is converted with elimation of molecular hydrogen into the red dihydrido-indium complex 24. ., ., - H2 .,’,,, :1;; LiBHE~ ‘l’; .—* ‘1” \ 0 \ /0 ,, H2 + H2 ‘(VQ4 w w ‘Bu2P — h-ptBu2 ‘B”’j7ir– “B”’ lB”’~<~~[’B”’ H/ %. cl H 21 23 24 ,, As anticipated acriphos (2a) serves as a tridentate PNP Iigand in its metal com- plexes. Treatment of 2a and 2,5-norbomadiene-molybdenum tetracarbonyl in tolu- ~: ene at room temp. gave the dark green octahedral acriphos-molybdenum tricarbonyl complex (25). Treatment of 2a with the dichlorides of the group 10 metals nickel, ,,, palladium and platinum in dichloromethane or n-butanol in the case of nickel yielded the square-planar acriphos-metal dichloride complexes 26, 27 and 28. The ionic , . ; : .,-, structure having one chloride coordinated to the metal corresponds with the ,,, . . conductivity and with the mass spectra (electrospray ionization) showing the parent ,,. . ions of the complex cation [acriphos + M + Cl]+ ~]. ,. ,’. :.,,,-, ” ,, 187 MC12 + ‘1;’, \ :1;’ N’ M = Ni, Pd, Pt N’ Ph2Pw PPh2 -WTqPh2P—+~ PPh2 c1 - cl 26 M.Ni J 25 2a 27: M = Pd 2& M.Pt We also attem~ted to me~are metal comtiexes of 1.8-bis(dimethvlamino) anthracene (11) and 4,5-b~s(dime~hylamino)acridine ~12). Treatment if 11 with bis(benzonitrile)- palladium dichloride in refluxing 2-methoxyethanol led to a black precipitation of palladium. Obviously the reduction potential of the extremely electron-rich aromatic diamine is high enough to reduce the noble metal salt. On the other hand, from the reaction of 12 and bis(benzonitrile) palladium dichloride in acetone at room temp. the complex 30 was obtained. m~w.w+ Me2N— Pd— NMe2 Me2N NMe2 Me2N—jd— NMe2 Cl - I cl ll:X=CH c1 29 12X=N [130 All new transition metal complexes of the diphosphine/amine Iigands derived from anthracene or acridine were characterized by mass spectra (El and/or ESI), NMR spectra (lH, 31P and mostly 13C) and in many cases by elemental analyses. Single cystal X-ray analyses could be performed for complexes 14 [1], 16 [14] and 25 [7]. 2.3. Catalytic Chemical Reactions In the past numerous monodentate and bidentate phosphine or amine Iigands were synthesized and tested for applications in homogeneous catalysis. Compared to this, tridentate and polydentate Iigands have found much less use in catalysis, possibly because they generally are considered to limit the availibity of open sites at the metal center by the non-dissociative nature of multiple chelating bonds [15]. To our know- ledge only very few reports have appeared which describe catalytic activities for metal complexes containing Iigands structurally related to anthraphos (1) or acriphos (2). Examples are the metal complexes 31 of the PNP Iigand 2,6-bis(diphenylphos- phinomethyl)pyridine and 32-35 of the PCP Iigand 2,6-bis(dialkylphosphinomethyl)- + 1: 1’/ 1’/ y (’R Ph2P—~i— PPh2 R2P— ~d— PR2 (’9 [m-cl - R2P ~ k,~PR2 cl — &OCCF3 H 31 32: R = ‘EIu 34 R = %U 33: R = ‘Pr 35: R = ‘Pr benzene. The nickel complex 31 is reported to possess some activity to catalyze the water gas shift reaction (WGSR: CO + HZO = C02 + Hz) [16], the palladium com- 108 plexes 32 and 33 can be used as highly active catalysts for the Heck reaction (C-C coupling reaction) [17], and very recently the dihydrido-iridium compounds 34 and 35 ~’: attract special interest for the dehydrogenation of alkanes (C-H activation) [18-20]. . ., Compared to the benzylic PNP and PCP Iigands in 31-35, which in some cases are ,, still flexible enough to allow tridentate as well as bidentate coordination [16,21,22], anthraphos 1 and acriphos 2 have the advantage of the much more rigid polycyclic aromatic skeleton and the lack of reactive benzylic hydrogen atoms. Hence by ex- changing the benzylic Iigands in 31-35 with corresponding acriphos and anthraphos Iigands the development of highly thermostable homogeneous catalysts is feasible. This is of high interest with respect to the catalysis of endothermic reactions such as the alkane dehydrogenation. Here we present our first results on the catalysis by anthraphos and acriphos metal complexes 7 and 8. Acriphos-palladium dichloride (27) was found to catalyze the WGSR at comparably ., low temperatures under neutral conditions, i.e. without adding any base or acid [23]. ,,,.: Using a solution of 27 in n-butanol/water 9:1 and 30 bar starting pressure of carbon monoxide, at 130°C within 15 h a 20’% conversion was observed, corresponding to a ~:-, turn over number (TON) of 300 and a turn over frequency (TOF) of 20 h-l. Raising ~ : the temperature to 180”C did not increase TON and TOF, though the catalyst seemed to remain stable. The corresponding acriphos complexes 26 and 28 of ., nickel and platinum, the acriphos-molybdenum tricarbonyl (25) and the anthraphos- chloro-palladium (14) did not show activity under the conditions where 27 is active. ,’. The WGSR activity of 27 demonstrates that carbon monoxide coordinates to palla- dium and then can be attacked already by a weak nucleophile such as water. This observation might be useful for other catalytic processes of carbon monoxide. The ,’ catalytic cycle proposed by scheme 1 is different from the cycle which was sug- gested in case the nickel complex 31. There it was claimed from spectroscopic data ~~. that some nickel intermediates are coordinated only bidentately to the benzylic PNP Iigand derived from pyridine [16]. In our opinion this is not possible for acriphos ~.. complexes, and hence different cycles might explain, why the nickel complexes of ‘.,’ : ,, the two PNP Iigands behave differently in their ability for CO activation. .- ,,... ,., 27 ... . ,,. .’ .,, ,, ,.,,.., ,,, .-. , .’, . . ... ,:, .,, . .:,. . , ,, “ Schemel: Catalytic cycle proposed for WGSR catalzed by acriphos PdC122 .,, .,..,,, . .. . . : ,., !,.”. ,’ . ,.,,,<,, ,,.,.,. 189 ,,~,: .,, ... The high activity reported for the palladium trifluoroacetate complexes 32 and 33 as catalyts in Heck reactions [17] prompted us to investigate the catalytic properties of related anthraphos-palladium compounds. Surprisingly the anthraphos-chloro-palla- dium complexes 14, 18 and 19 turned out to be active catalysts for the Heck reaction, and it was not necessary to use palladium compounds having less strongly coordinating anions such as trifluoroacetate or acetate. Table 1 summarizes selected results obtained for the C-C coupling reaction of bromobenzene (36) and methyl or n-butyl acrylate (37a or 37b) in N-methylpyrrolidone (NMP) at 140”C. With 36 and 37a as the educts and 14 as the catalyst, a complete conversion of 36 was obtained yielding a 70:30 mixture of methyl E-cinnamate (38a) and methyl 2-phenylcinnamate (39a). The unusually high amount of 39a, the product of the twofold Heck reaction, was decreased almost to zero by using rr-butyl acrylate (37b) instead of the methyl ester. Using sodium carbonate or sodium acetate as the base, had no significant effect. The much lower catalytic activity of 18 as compared to 14 can be apparently attributed to steric effects caused by the bulky t-butyl substituents of the phosphorus atoms. For 19 a similar catalytic activity was observed as for 14, but the selectivity towards the monosubstituted product 38b apeared to be somewhat lower than with 14. Although in 14, 18 and 19 the chloride Iigand attached to the palladium did not prevent the catalytic reactions of bromobenzene, the catalysts proved to be inactive for the conversion of chlorobenzene. On the other hand, no decomposition of the catalysts was obsetved on raising the temperature up to 180°C, and even after complete conversions the reaction solution maintained the characteristic yellow color of the anthraphos-palladium complexes. /1 /\ catalyst — — — base Br+- .% + \/ o- COOR NMP COOR – COOR 140°C, 64 h 36 37a: R = Me 38a R = Me 3\ 1 39a: R=Me 37b: R = ‘Bu 38b R = ‘Bu 39b: R = ‘Bu Table 1: Results of the Heck reaction catalyzed by anthraphos-chloro-palladium(ll) complexes 14, 18 and 19. Conversions and turn over numbers (TON) are based on the consumption of 36. run I 36 37alb catalyst base conv. 38:39 TON mmol mmol mmoi.104 6 mmol % 1 I 5.5 a: 6.9 14:8.7 NazCOs 100 70:30 6350 2 4.9 b: 6.4 14:4.7 Na2C03 58 99:1 5900 3 5.0 b: 6.2 14:7.0 NaOAc 100 96:4 6400 4 5.0 b: 5.7 18:7.1 NazCOs 11 100:0 750 5 5.0 b: 6.1 19:7.0 NaOAc 99 87:13 7050 According to the generally accepted mechanism of the Heck reaction a palladium(0) phosphine complex is the catalytically active species and the catalysis proceeds by going round a Pd(0)/Pd(ll) cycle. The possibility of an alternative Pd(ll)/Pd(lV) cycle is currently under debate [17,24]. In the case of the anthraphos-chloro-palladium(ll) complexes 14, 18 and 19, as well as in the case of 32 and 33, one hardly can imagine that reduction to a palladium(0) species is possible. Without having done 190 ,“ detailed mechanistic studies yet, we propose two possible intermediates. inter- mediate 40 is the ionic octahedral palladium complex which is formed from 14 by “ oxidative addition of bromobenzene and coordination of the olefin [25]. in inter- mediate 41 the addition of the phenyl group to the anthracene C-9 atom generates ,. an arenium cation, but avoids the formation of the debated palladium species. Precedents of structurally related cationic arenenium systems are known. However, they differ insofar that they are formed by adding an alkyl cation rather than an ~ , unfavorable phenyl cation [26,27]. ,. + :1;; /.\-.. ,. \ :+:/ .,., cl ,,. c1 \,. Ph. ,,, Ph2P—Pd(lV)- PPh2 Ph2P$9?- Pd(ll)-PP~ Br - 9?? Br- ~h< . i .,, : COOR_ L COOR_ J ,., 40 41 Furthermore, the utility of our new Iigands as components of highly therrnostable homogeneous catalysts is impressively demonstrated by (anthrphos~bu)dihydrido- iridium (24). Whereas the dihydndo-iridium complexes 34 and 35, which presently are intensively studied as homogeneous catalysts for the alkane dehydrogenation [18-20], instantaneously decompose at temperatures beyond 200°C, 24 did not decompose when the red solution in cyclooctane was heated to 250°C for 30 min. Since the temperature was limited by the used oil bath, this might not be the end of the thermostability of 24. When a solution of 24 in cyclooctane (42) was heated to 230”C for 3 h, the formation of cyclooctene (43) and molecular hydrogen was detected by lH NMR and MS, respectively. Experiments to precisely determine the ., thermostability of 24 and to study the potentials of 24 and related complexes as ,,.,. , catalysts for the alkane dehydrogenation are currently under way. .’ ,’” ,Bv ‘tableat2500c! —lr— PtBu2 / “,. H 24 + H2’11 230”C, 3 h “o o42 43 3. Conclusion Using anthracene and acridine as examples, we have shown that polycyclic arenes and heteroarenes are useful as rigid and stable backbones of diphosphine/amine li- ., gands. Anthraphos (1), acriphos (2) and corresponding phosphine/amine ligands can be prepared from 1,8-difluoroanthracene (9) and 4,5-difluoroacridine (10) by . ~ ,,.,, nucleophilic substitution with alkali metal phosphides or amides. The PCP and PNP ,, Iigands 1 and 2 bind transition metal atoms in a ‘T-shaped” planar geometry (rner- “.“,:. ., ,.$. configuration) to form highly stable complexes 7 or 8, respectively. Some of these ,,, complexes have been found to be useful as highly therrnostable homogeneous .,, 191 catalysts: Acriphos-palladium dichloride (27) is able to catalyze the WGSR, the anthraphos-chloro-palladium com lexes 14, 18 and 19 are active catalysts for the f’. Heck reaction and (anthraphos- bu)dlhydrido-iridium (24) catalyzes the dehydro- genation of cyclooctane to cyclooctene and possesses the highest thermostability yet reported for such homogeneous catalysts. The development of other thennostable homogeneous catalysts derived from 1,2 or nitrogen analogues is feasible. 4. References 1. M. W. Haenel, D. Jakubik, C. Kruger, P. Betz, Chern. Ber. 124, 333(1991). 2. M. W. Haenel, D. Jakubik, E. Rothenberger, G. Schroth, Chem. Ber. 124, 1705(1991). 3. M. W. Haenel, H. Fieseler, D. Jakubik, B. Gabor, R Goddard, C. Kruger, Tetrahedron Leti. 34, 2107(1993). 4. S. Hillebrand, J. Bruckmann, C. Kruger, M. W. Haenel, Tetrahedron Leti. 36,75 (1995). 5. E. M. Vogl, J. Bruckmann, C. Kruger, M. W. Haenel, J. Organomet. Chem. 520, 249 (1996). 6. E. M. Vogl, J. Bruckmann, M. Kessler, C. Kruger, M. W. Haenel, Chem. Ber. 130, 1315 (1997). 7. S. Hillebrand, B. Bartkowsk3 J. Bruckmann, C. Kruger; M. W. Haenel, Tetrahedron Lett. 39,813 (1 998). 8. M. W. Haenel, S. Oevers, J. Bruckmann, J. Kuhnigk, C. Kriiger, Syrr/ett, 301 (1998). 9. S. Oevers, M. W. Haenel, unpublished; S. Oevers, doctoral thesis in preparation. 10. T. Morimoto, N. Ando, K. Achiwa, Syn/ett., 1211 (1996). 11. C. J. Moultonr B. L. Shaw, J. Chem. Sot. Da/ton, 1020(1 976). 12. S. Nemeh, C. Jensen, E. Binamira-Soriaga, W. C. Kaska, Organomefa//ics 2, 1442 (1983). 13. M. Gupta, C. Hagen, W. C. Kaska, R. E. Cramer, C. M. Jensen, J. Am. Chem. Sot. 119, 840 (1997). 14. C. W. Lehmann, S. Oevers, W. C. Kaska, M. W. Haenel, to be published. 15. H. A. Mayer, W. C. Kaska, Chem. Rev. 94, 1239(1994). 16. P. Giannoccaro, G. Vasapollo, A. Sacco, J. Chem. Sot., Chem. Commun., 1136 (1980). 17. M. Ohff, A. Ohff, M. E. van der Boom, D. Milstein, J. Am. Chem. Sot. 119, 11687 (1997). 18. W.-w. Xu, G. P. Rosini, M. Gupta, C. M. Jensen, W. C. Kaska, K. Krogh-Jespersen, A. S. Goldman, Chem. Commun., 2273(1997). 19.F. Liu, E. B. Pak, B. Singh, C. M. Jensen, A. S. Goldman, J. Am. Chem. Sot. 121, 4086 (1999). 20. F. Liu, A. S. Goldman, Chem. Commun., 655 (1999). 21. P. Dani, T. Karlen, R. A. Gossage, W. J. J. Smeets, A. L. Spek, G. van Koten, J. Am. Chem. Sot. 119, 11317 (1997). 22. J. A. M. Brandts, R. A. Gossage, J. Boersma, A. L. Spek, G. van Koten, Organo- metahs 18, 2642 (1999). 23. Review on the homogeneous catalysis of WGSR: R. M. Laine, E. J. Crawford, J. MoL Cat. 44,357 (1988). 24. B. L. Shaw, S. D. Perera, E. A. Staley, Chem. Commun., 1361 (1998). 25. A. J. Canty, J. L. Hoare, J. Patel, M. Pfeffer, B. W. Skelton, A. H. Whiter Organo- mefa//ics 18, 2660 (1999). 26. P. Steenwinkel, R. A. Gossage, G. van Koten, Chem. Eur.J.4,759(1998). 27. B. Rybtchinski, D. Milstein, Angew. Chem. 111,918 (1999); Angew. Chem. /rrt. Ed. EngL 38,870 (1999). Financial support of the Max-Planck-Society (MPG), the Fends of the German Chemical Industry (FCI), the German Academic Exchange Service (DAAD) and the German Research Association (DFG) is gratefully acknowledged. 192 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 !.-, .,, ..: -,-1 V.A. Zazhigalov 1),A.i. Kharlamov 2),L. Depero ‘), A. Marino 3), I.V. Bacherikova ‘j, J. Stoch 4),J. Haber 4, 1)Ukrainian-polish Laborato~ of CatalySi5: Institute of Sorption and Problems of Endoecology, National Academy of Sciences of Ukraine, Kyiv, Ukraine ,,. z)Institute for Materials Science Problems, National Academy of Sciences of ,;.....,,,,:,,:;.,, ,.. Ukraine, Kyiv., Ukraine ,,, : ‘) Llniversi~ of Brescia, Structural Chemistty Laboratory, Brescia, Italy :!, <.,:’,-* d)LJkrainian.PofiShLaboratory of Catalysis: Institute of Catalysis and sUhCe . ‘::: . Chemistry, Krakow, Poland .,, , .,,f- ., .’, ... ..- ,.. .’,:..’~. -.x. ;. Low Temperature o-XyIene Oxidation to Phthalic Anhydride on V-Ti-O :,,, . ,, :$, Catalysts Prepared by Mechanochemistry ...... ,-. ...’ ‘.,,.;,-:- ,,, ., ,...... , . .. “ ‘ ,:,:7,....$ r, ,,’/’’:.7:,;,; 1 ,, !.,, ,),,j The V-Ti-O oxide system is a catalyst of many processes, including industrial $, , ,:. , ,.: o-xylene oxidation and NO. reduction [1,2]. They are usually prepared by widely ,,J-, ! ;,.,, f ,.’,’ used methods: i) impregnation of Ti02 with different vanadium compounds, ii) .,.” ,,,.’.., vanadium chemical bonding with functional groups on TiOz, and iii) V and Ti co- ,.:’,1 ,..,,,,., ::r, precipitation from compounds of Me(OR). type [2,3]. The catalyst was also obtained ,,, :,( ,’, from the powder mixture of vanadia and titania oxides by stirring them in an agate t .“’,)>,,, ‘, , mortar [4]. ,,.’,,,,,.: ::;i In this paper a possibility of the catalyst synthesis by a mechanochemical . ;,. ,..,*.J ,,,; .. ‘.’,.:< , treatment of the V20S and Tioz powder mixture is considered. ,’!:“::,.,,.;j”:,,,,,,,;;’,’7, j :::-,.f,,.:~J Experimental ‘, ‘1,,.,\.., ,’‘,. / ,,, ,,.:,,:.’,’./,:,y Initial reagents were TiOz powder (Aldrich, p.s.), V@5 - p (powder, REACHIM, ,, ‘ ,,:;.’.’rjr.... p.s.) and V@s - m melted in air and then crushed. Powders were precisely mixed in ,,.,.:;,! :,,~.,,,$,,,,,J “~ agate mortar and by different periods in a planetary mill (3000-rpm). As dispersing !!’:,,,;$, ,~,,, .,., , .! ,,:.,’ medium ethanol or water were used. Mass of mixed oxides was 25 g, mass of balls t ,,: ,.’! -,... . :;,f.: 450 g. After the treatment was completed, catalysts were filtered and dried at 393 K. , ,,,;..,.; ,.”If:... ,. ,“/ Specific surface area was determined by thermal resorption of argon ‘- ~ , :,.:/;:, ,,,..+ (GASCHROM-1). A phase composition was studied with a Philips MPD 1880 ~ ./”,;, .,,.,: ,.C,,.’: ,’ .,,,”,..:, ,..,,,,,4 ,1 diffractometer (CuKa radiation, graphite monochromator). Relative line intensity and $, ,. ,,,,.;,,.,,~, ,.,.,. width (FWHM) was determined with Automated Powder Diffraction Program (version . ,,, ,.,.:..,,’, ““ 3.5, Philips). Micro Raman spectra were obtained with Dilor Labram spectrograph ... . ,,,,,:,!,,.,.- ,.: .,;.,; ,, ,:,,.; .,’..<, (HeNe laser - 632.8 rim). The spectrograph was linked to a microscope for ! ,, .’,, ‘,. ,/<;:: ~; selection of proper analysis place (resolution better than 1 pm). Surface composition ,.., ., “.:; was determined with XPS using VG ESCA-3 spectrometer. The spectra were ,.-. ,. .:,~. ,,, , calibrated using Cls line at 284.8 eV the spectra handling were described ,,, ” .. . ,!, elsewhere [5]. ,, ,“. , ,’ ,’,’, ,,~..,.-’.. , ., ,,, $,,. ,s, . ; ,’-., . “, ,, :,!, ,., :. ,,, ,: ;.. ,,,, ,“), ” ;,.. ?,,,.,, 193 ,, ,, .,.,. OGMK-Tagungsbencht 9903, ISBN 3-931850-59-5 ,, ”:, -, .,:.: ‘t Catalytic properties of the samples (fraction 0.25 – 0.50 mm) were studied in a flow system with steel microreactor (IKFCP PAN). The reactor was loaded with 0.5 cm3 of catalyst. The contact time was modified by controlling of a speed of feed mixture in the range 12-65 cm3/min. Reaction of o-xylene oxidation (0.9 vol. ‘XO in air) for the testing was selected. Reacted gases were analyzed in the on-line system of two chromatography (Chrom-5) equipped with: a) 2.5 m column filled with 4.3% F- 50 on Chromosorb G (temp. program 373-598 K) for the analysis (DTP) of products in partial oxidation of o-xylene, b) 2 m column filled with silicagel KSK-2.5 (temp. program 323-393 K) for the analysis of C02 and hydrocarbons, and c) 2 m column filled with molecular sieves CaA (temp. 273 K) for an analysis of 02 and CO. Chromatograms were recorded and then analyzed on a PC 486 DX 100 instrument. Results and discussion In V@5 – p and VZ05 - m XRD revealed the most intense reflex at 263= 20.3° corresponding to the (01O) plane. However, while in the first sample the reflex intensity ratio (010)/(1 10) was near 3/1, in the second one it was about 10 times higher. This is evidence that V@5 – m was in the form of large crystals developed along the (01O)plane. The XRD spectmm of TiOz was typical for anatase with strongest reflex at 20 = 25.3° corresponding to the (101) plane. In the case of VzOdTiOz mixtures after the mechanochemical treatment the XRD spectra were a composition of superposed diffractograms of simple oxides (Fig.1 ). Simultaneously this treatment changed intensity of all reflexes, initially typical for simple oxide, remaining their FWHM unchanged. For V20s-m/TiOz an increase in time of milling did not change the reflex intensity ratio (101 )/(01 O) though it was remarkable smaller comparing to v205- plTiOZ. This is connected with differences in intensity of the (01O) reflexes in initial forms of vanadium oxide. In the mixture of V@SpfiOz, after 10 min of the treatment in alcohol, the intensity ratio of (101)/(01 O) reflexes (informing about WV ratio) decreased from about 12 to 6 (Table 2). However, when time of the treatment increased from 10 to 30 rein, the ratio rose from 5.9 to 8.3. [n water this ratio already after 10 min of the treatment decreased to 10 and reminded almost constant under further exposition. Somewhat lower intensity of the (01O) reflex after the treatment in water can be caused by partial dissolution of V205 earlier already observed during its ,’- . mechanochemical treatment [6]. c. ,. The Raman spectrum (Fig. 2) from TiOz presented typical set of bands at 145 (very strong - v.s.), 400, 520 and 645 (v.s.) cm-’. Bands at 155 (v.s.), 205, 295 .. (v.s.), 307,405,487,540,705 and 1000 (v.s.) cm-l were characteristic of both V205 samples. The spectra from the VZO#l_i02 mixture and treated samples were the result of superposition of that of initial oxides. However, micro-Raman spectra taken for two different points (A and B in Fig.3) of the near-surface region showed the following. [n the case of the V20s-p/Ti02 after 10 (Fig. 3 a, points A and B) and 20 min of the treatment in water, differences in intensity of absorption bands 145/155 and 645/1 000 cm-l were reflecting heterogeneity of the sample. After 30 min (Fig 3 b, points A and B) almost equal ratios of the band intensity in these points were observed, proving homogeneity of the V205-p/TiOZ system. On the treatment in alcohol, already after 20 min the system was nearly homogeneous, while after 30 194 ., min it was fully homogeneous. The treatment of the VzOs-m/TiOz system in ethanol after 20 min also gave a uniform component distribution. AS it follows from XPS data (Table 1), binding energies of electrons in the VzOS/TiOz system was not much influenced by the mechanochemical treatment being close to that already published [7]. This shows that the oxidation state of elements did not change under the treatment. On other hand the atomic ratio of surface elements was substantially changed. Treatment in ethanol is increasing the WV ratio, more significant in VzOs-m, while in water this ratio decreased (Table 1). The data obtained, including those concerning mechanochemical modification ,, of V205 [8] allowed us to assume the following picture of changes accompanying the .. treatment of the V20#l_iOz mixture. As it was already established, the ,.,.,,.’ mechanochemical treatment of VQ05 in ethanol produces anisotropic crystal deformation [8] resulted in an increase of the (01O) plane content. !, Such change in the plane contribution was observed in the sample V205-p : ~, ~~ ,..;,,;- ,,, after 10 min of the treatment (Table 2) when it increased from 3.1 to 5.3. But further ,, exposition unexpectedly brought small reversion of the process, i.e. decreasing in the vanadium XRD (010)/(110) ratio (Table 2). Simultaneously the XRD (101 )/(010) ‘.,- ratio initially decreased by twice, mostly as the result of strong development of the .’ (010) vanadium plane. On further treatment this plane intensity ratio was consecutively increasing from 5.9 to 8.3, partly reflecting decreasing of the relative vanadium (010) plane content and partly by lost of the vanadium visibility by XRC) in strongly dispersed oxide. Changes in XRD data (Table 2) suggest that comparing ,’ TiOZ and V205 oxides, the last one is more tractable to destruction and easily ,. dispersing. In conclusion, our results show that prolonged mechanochemical ,.. r treatment of the V205-p/Ti02 composition in ethanol produces a mixture of larger TiOz crystals and much smaller V205 particles. Since the XPS Ti/V ratio remained ; , high, there is the reason to expect a coating of Ti02 with V205 or agglomeration of ,,., ~ dispersed vanadia should be expected. ..’ Different picture is observed in VzOs-m. Very high XRD (010)/(110) ratio in the : :“ , , ~, starting powder (Table 2) points to strong preference for the (01O) plane. During the ,.. , treatment this big, flat crystals initially are crushed vertically to this plane, lowers its ,- relative content, but after some time, a gliding along the (01O) plane begins to be dominant. Considering the treatment of the V@5-p~02 mixture in alcohol, the most striking feature is breaking the vanadium (01O) plane evolution. In this system, with exactly repeated handling, the only reason for the effect was the presence of titanium atoms. On the other hand each inorganic compound dissolves little in every : , solvent. These trace amounts of both substances are available during the treatment. :,” Table 1 shows that preferential dispersion of the vanadium phase, braking upon ‘:) “ ; time, is accompanied by increasing the XPS Ti/V ratio. These both facts univocally ;,< ~ bring us to the conclusion that single titanium atoms or oxygen-titanium polyhedra .,,.. <:!. ,. .,,, ,, were incorporated to the Vz05-p lattice, demrating edges or staking-faults and in ,,,‘,,’, ~ consequence braking the gliding of vanadia (01O) planes. The surface enrichment in ,. titanium can be detected by XPS and could not be reflected by any XRD data .” “u,~ ~, , ; because of XRD insensitivity in such systems. ‘,, When V205-m was used, the anisotropic deformation was less extended for its :;::: ~.,j. ‘”’. higher hardness. This is evidenced by absence of changes in the XPS and XRD . .,,., . , ‘.’ ‘,, , !,-, ,,,,;-,, $ composition during the treatment (Table 1 and 2). .,: ,,.,,,,,’ - ,- .-, ,, .,“(-.,., ,, . .,..,,.; .:.;::.. ,., 195 .“ ‘.. , ;! , :’ .,,., ,,,,. , ... ,: ..,, ,., ,,. In the presence of water a chaotic destruction of crystals [8] and partial dissolution [6] proceed in VZ05. As the result, the highest values of the (101)/(01 O) reflex ratio and the Ti/V ratio at the surface have been observed. Mechanochemical modification of the V-Ti-O composition leads to the change of its catalytic properties (Table 2). The treatment in ethanol influences much o- xylene conversion which significantly increased. After the treatment in w=ter conversion of hydrocarbon also increased, but less than in the case of ethanol treatment. The activation energy o-xylene oxidation remains constant after mechanochemical treatment and equal to 30 kcal/mole. The selectivity towards phthalic anhydride increased independently on the environment of treatment, while that towards maleic anhydride decreased or not changed when water was used This study showed that catalytic properties can be significantly influenced by compounds accompanying the mechanochemical treatment. It seems that the mechanism do not rely on direct chemical action, but the chemical action controls the physical process which produces the more effective product. Concluding, the mechanochemical synthesis of V-Ti-O compositions up-rises their catalytic ability to selective oxidation of o-xylene. References 1. Nikolov V., Klissurski D., Anastasov A., Catal. Rev.-Sci.Eng., 33, 319(1991). 2. Dias C.R., Portela M.F., Catal. Rev.-Sci.Eng., 39, 169 (1997). 3. Grzybowska-Swierkosz B., Appi. Catal., A., 157,263 (1997). 4. Centi G., Giamello E., Pinelli D., Trifiro F., J.Cabal., 130, 220(1991). 5. Zazhigalov V.A., Haber J., Stoch J., Pyatnitskaya A. I., Komashko G.A., Be!ousov V.M., Appl. Catal., A, 96 135(1993). 6. Zazhigalov V.A., Haber J., Stoch J., Kharlamov A. l., Bogutskaya L.V., Bacherikova I.V., Kowal A., Solid State Ionics., 101/1 03, 1257(1997). 7. Reddy B.M., Chowdhury B., Ganesh l., Reddy E.P., Rojas T.C., Fernandez A., J.Phys. Chem., B., 102, 10176 (1998). 8. Zazhigalov V.A., Kharlamov A. I., Bacherikova I.V., Komashko G. A., Khalamejda S. V., Bogutzkaya L.V., Byl O.G., Stoch J., Kowal A., Teoret.Experim. Khim., 34, 180 (1998). 196 TABLE 1 Surface properties of the V-Ti-O composition afler mechanochemical treatment Treatment Binding energy, eV Atomic ratios Medium Time, Ti 2p v 2p o Is TiN ON Clo min 459.2 517.7 530.8 1.31 4.78 0.42 Ethanol 10 459.1 517.6 531.3 2.02 6.20 0.61 Ethanol 20 459.1 517.6 531.0 1.83 5.80 0.60 .. Ethanol 30 459.3 517.6 530.9 2.06 6.69 0.60 ,).,.’ Ethanol * 10 459.1 517.4 531.0 2.61 7.63 0.72 Ethanol * 20 459.2 517.6 530.9 2.32 8.35 0.71 ,,:. ., Water 10 459.2 517.6 531.0 0.90 4.37 o.4~ Water 20 459.4 517.9 531.4 0.91 4.19 041 Water 30 459.3 517.8 531.1 1.09 4.56 0.42 ,.,,,..‘.+ ,.,j,. !, * V20s-m ,. 4,,,1,,,..,;., ....<.! , .; .:?: ~~,’,..... J: ~., $:: ,,, ,,,..r .,:,.;,,,.;:,.,.:‘,~..,;;!”,.,1, ., ..!. .,~j,, ,.1//. .:,,., ,,~,. ,;>:,,,:. l ;,:,,:P,,,,,.. :}-’f.,~ ‘: ,’;,<.,,,’,<,,..’ { TABLE 2. ,., . ,:,, ‘,, ,-,,; ~; . f,.$, ,1,1 /!, :.-., < ,;(.,,.!,,:;,; Catalytic properties of the V-Ti-O composition after the mechanochemical treatment i. :., ,,,,,‘-, .,< :,,.,, ,., , ,.,., Treatment XRD ratio o-Xylene oxidation*, T = 533 K .:, Medium Time TiN V205 SW, .,,, Conversion, sPhA, ,’. (101)/(010) (01 0)/(1 1 o) 0/0 mol. ‘%0 mol. ~$ 12.09 3.07 38 12 48 Ethanol 10 5.90 5.33 73 7 61 Ethanol 20 6.25 4.66 72 8 70 Ethanol 30 8.33 4.50 83 9 79 Ethanol** 20 3.45 25.00 89 3 87 Water 30 ,, 11.10 3.20 69 14 56 ., * SMAand SPH– selectivity in maleic and phthalic anhydride formation, respectively ,’, ,., . . ~ V205-m !’, , .: ,’.. ,, , ,. ~,” ., . -, ., .“ 500( Ice.,, ts 450( 400, 3501 300( 250( 200( Isoc 100< — F 50( E O.c A_ IL,,T —, , t 3S00 3000 2s00 2000 1s00 1000 SOo 0.0 Sooo 4000 3000 2000 1000 A 0.0 A 4’0 1.;01 Fig. 1. XRD spectra Ti02 (A), V20EiJTiOzmixture after mechanochemical treatment in ethanol 10 min (B), 20 min (C) and 30 min (D), V205-P (E) and Vzosm (F). 198 2ooo- $, ,,,,. A 1500- -,.> 1ooo- ., ,, 2500. 1500- 500’ ,,.!’,.,,,’,,. ,, ,. ,., . — I 200 400 600 800 1000 v, cm-’ Fig. 2. Micro Raman spectra TiOz (A), VZOSP (B) and VZOs-m (C). a 6, ,. lo- ., 10 . g. - 20 20- ~- 30 j 30- 40- Ii 30 40 50 ; 10 30 40 50 Lcnglh X (pm) Length X (pm) ) .— -J -— -.. — —-1 A ‘i I I I I 1. . . . ,..._—.——--.— 200 400 600 800 1000 260 400 6iI0 860 lb \ \\pwcnumbcr (cm-1) Wavenmnber (cm.]) . . .. —.— —. I .- —-.I i I I I I i-l‘* .—-—— L 200 .$no 600‘—”—;00 1000 200 dOO 600 800 1000 . I \Vm cnumbcr (cm-1 ) Wmwumbe, (cm-l) –> Fig. 3. Micro Raman spectra V@Sp~Oz after mechanochemical treatment in water 10 min (a) and 30 min (b). 200 DGMK-Conference ‘The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 A. 1.Lutsyk, E. S. Rudakov, V. N. Mochalin L.M. Litvinenko Institute of Physikal-Organic and Coal Chemistry, National Academy of Sciences of Ukraine, Ukraine A New Calculation Method for Arenes Solubilities in the Whole Composition Range of Water-Sulfuric Acid System Data on volubility of arenes and their derivatives in water - sulfuric acid system are of great interest because of wide use of this system in researches and industrial technologies. Such data for aniline and nitroarenes has allowed M.1.Vinnik to establish the causes of decreasing effective rate of arenes derivatives nitration in a ., .,,’ sulfuric acid with the increase of the medium acidity at [H2S04]s 90 % weight [1]. . . However the experimental determination of the arenes solubilities is often ,: difficult or even impossible due to high viscosity and aggressiveness of a sulfuric acid .( and the associated problems of mass transfer and solutes stability. The data .,,’,, ,: available are not numerous and usually do not cover the area of high sulfuric acid concentrations. The effects of specific solute-solvent interactions may also distort the findings. The new volubility data for non-reacting non-electrolytes (NE) (helium, hydrogen, saturated hydrocarbons Cl - Ca of various structure) in the whole range of H,2SC)4concentrations (frOr_nOup to 100 %f.H2S04) at 298 K obtained from our recent studies [2-5] have given new opportunities of calculating solubilities in such complex systems. The experimental data were obtained as limiting distribution coefficients (a) of ,.. , NE between a gas phase and solution a = lim([NE]~,~[NE],Ol) at [NE]g,, + O. The ,- value a-l is equal to volubility of the solute at its fixed concentration in a gas phase .,, ,: and coincides with a widely used form of representation of gas volubility - Ostvald ,,)-;: coefficient. The correlations of a value with the other ways of representing the ,- volubility of gas in liquid described in literature are the following: “..’,,~,., ,, a = 273,15/B.T = p2V1. (1-x2)/(xzR.T) = H.V2/(R.T) (l), where X2- NE mole fraction; H - Henry constant; B - Bunsen coefficient pz - partial ., pressure of NE; R - universal gas constant; T - absolute temperature; VI and V2 - molar volumes of pure solvent and solute, accordingly. ! . .,, .,; ... Having analyzed the experimental and l~e}ature data we succeeded in ,,, .,,.,,-,,’,,,:;,,: ., : .P.? :. establishing the following rules of solvent composition and solute structure influence ,/, b = (0,3 f 0,1) .10-2 - (2,5 * 0,2) .IO-2.XA (3). Il. For non-electrolyte solutions in the range of acid concentrations from 84 up to 100 %(o(here and further the weight percentage is used) the following equation is applied [4]: . lg(a./aW)= C.[HSO~] (4), which is analogous to the Sechenov’s equation if to consider 84- 100 YO HzS04 as the solution of electrolyte HSOiHSO+ in pure sulfuric acid being the solvent. As the sulfuric acid is diluted by water the concentration of electrolyte HS04-HSO+increases, this resulting in salting-out of NE. Ill. In the whole range of compositions of the water - sulfuric acid system the following equation is applied [5]: (lga)’ = K.p = LWE/V, (5), which eqUaliZeS the excess value of Iga ((lga)’ = iga - Xwlgaw - xAlgaA) and iOII fOrCe (p), and relat$ne compression of solution due to of interactions between its components V IV = (V - XWVW - XAvA)A/ (VEAf are derived from the density of sulfuric acid solutions [6]). xw and xA are mole fractions of water and acid, K H L are ratio coefficients for given solute. L value is linearly related to NE molar volume: L = - (4,1 * 0,4) - (4,9 f 0,5) .10-2.~fiE (6). These rules allow to calculate the volubility of non-electrolytes including arenes and their derivatives in water - sulfuric acid system in the whole range of it compositions. The essential advantage of the approach presented is that it allows to predict the volubility in the cases when direct measurements are impossible because of the difficulties mentioned above. Let’s consider two examples of application of the rules formulated for benzene and nitrobenzene solutions in sulfuric acid - water system. BENZENE The volubility of benzene has been measured in the range of O-90 ~. H2S04. [7-1O]. The data were obtained by applying various methods: the spectrophotometric analysis, the kinetic distributive method, the method of an equilibrium vapor analysis and the direct NMR-measurements of the solute concentration. The literature data are summarized in Fig. 1. At higher concentration of the acid the volubility measurements are impossible because of a high rate of the sulfonation process. We have calculated the a values of benzene in the whole range of compositions of H20-H2S04 system by the equation (5) (see the line in Fig. 1). The necessary values of benzene volubility in 100 ~0 H2S04 and L parameter were obtained by equation (2) ((XA= 0,014) and equation (6) (L = -8) using the partial molar 202 ,’, , ,, volume of benzene in water ~~E = 83,2 cm3/mol from [11]. A good agreement between the experimental and calculated data (Fig. 1) in the range of 0-80 ‘%.sulfuric acid ,,,, has led us to the assumption that the ,. calculated values for a more concentrated acid are also correct. It .“. ,> should be mentioned once more that ,:. . thbre are no experimental volubility data for benzene at [H2S0.J s 90 %!.. ,!’ 0.0 I 1 I I I NITROBENZENE 0:0 0.2 0:4 0.6 0.8 1.0 One of the causes of a sharp ‘H2S04 growth of nitroarenes solubilities in the H@HZSOd system with the increase Fig. 1. Benzene volubility in system water of concentration of the acid is - sulfuric acid. The line shows the results considered to be the protonation of . . calculated by the equation (5). The points solute [12] and formation of show literature data: ● [7]; A [8] (303K); compounds RN02-H2S04 [13]. In such + [9]; + [10]. case the total process could be represented by the scheme (7, 8): RN02 (gas) ~ RNO* (sol) (7), RN02 (sol) + A ~ A. RN02, (8), where A = H+ (protonation) or H2S04 (complexation). According to the scheme the measured value %,P must be related to the value of a determined in the absence of specific interactions as: : ,’:,J ~,P = cd(l + K.[A]) (9), - ‘ where [A] = ho in case of protonation or [A] = [H2S04]in case of interaction with the acid, K - equilibrium constant for step (8). If the process of nitroarenes dissolution is actually accompanied by proton ation or complexation with the non-dissociated acid, as represented in the scheme (7, 8), then in a case of protonation the value Ig{(dGxp) - I} will be linearly related to the media acidity function HO(Ho = -Ig(ho)), and in case of interaction with non-dissociated ti2S04 - to the logarithm of mOlar concentration of the acid. These assumptions could be verified for nitrobenzene, its volubility in water - sulfuric acid system being measured in detail by Hammett and Chapman in the range [H2S04] = O -79,3 YO at (298,15 * 0,02) K [14] 1. The experimental SOhJbility ,. ,,.,., data of these authors (in mol/1) we evaluated in a values (see Table 1) using the ,., - .,, 1 In ref. [15] the data on volubility of PhNOz in the range 0-95,9 Y. H$30A are given .,’,, #, . . ., .. pressure of saturated vapor of nitrobenzene which is equal to 0,284 in_Hg [17]. Then, using the experimental values of nitrobenzene partial molar volume in water (~~E = 96,1 cm3/mol [18]) and its volubility in water (a~ = 9,7.104, see Table 1) the volubility of nitrobenzene in pure sulfuric acid was calculated by the equation (2) (a*= 2,2.10-5), and parameter L of equation (5) was calculated by the equation (6) (L = -9). The results obtained were used for calculating the presumable volubility of nitrobenzene in the absence of interaction with the media components in the whole range of sulfuric acid concentrations (see Table 1). Table 1. The comr3arative data of experimentally measured and calculated ;olubilities of nitrobenzene in water - sulfuric acid system-at 298 K. [H2S0,] a [H,SO,] a 0/0 Literature Calculated on 70 Literature Calculated on weight data [14] Eqn. (5) weight data [14] Eqn. (5) o 9,7.104 9,7.104 52,8 5,6.104 2,8.10-3 19,8 1,4.103 1,7.103 58,2 4,1.104 2,8.103 35,4 2,6.103 2,4.10-3 64,3 2,9.104 2,7.10-3 35,8 1,2.103 2,4.103 70,4 2,0.104 2,4.10-3 44,5 8,5.104 2,6.10”3 74,5 1,3.104 2,1.10-3 48,9 7,2.104 2,7.103 79,3 6,0.105 1,6.10-3 The comparison of calculated and experimental data shows that the presumable volubility is lower (and a value is higher) than the experimentally measured ones and the difference increases with the growth of the acid concentration. In concentrated H2S04 solutions the difference amounts up to hvo orders. The result presented points to the fact that nitrobenzene actually interacts with the components of the water - sulfuric acid system in the process of dissolution. The value lg{(aalJ~,P) - 1} is linearly related lg[H2S04] to lg[HzSOA] (Fig. 2), this confirming the earlier 02 04 06 08 10 12 I I I I I assumption concerning complexation process. 15- ~ Thus the equations (2), (4) and (5) ● prove to be effective for the analysis of 0 -1 : 10 0 volubility of non-electrolytes including arenes c and their derivatives in water - sulfuric acid : system. They could be recommended for wide !?~ 05 ● m use in investigating volubility and the mechanisms of reactions of non-electrolytes in ~2 00 _. -— acid media, and for the development and optimization of technological processes. -0.5 : ●x 0 I I I I 1 1 I 1 -8 -7=5-5-4-3-2-1 Ho F[g.2. The correlation of lg{(~lc/ a+,. ) - 1} with Hammett acidity function [19] ( O ) and with lg[H2S04] ( ● ) for ntrobenzene solutions in water - sulfuric acid system. 204 ,.. ,,,,. ,; , ,, (,-:<, :~...... - ,,, :,,, :.’ ,.. ., ’.,:,,-, y ., ,.’ .,,.., LITERATURE 1. Vinnik M.I., Grabovskaya G.A., AEamaskova L.N. Zhurnal tizicheskoi khimii. ~, 1102 [1967] (in Russian). 2. Rudakov E.S., Lutsyk Al. Russ. J. Phys. Chem. ~, 731 [1979]. 3. Rudakov E.S., Lutsyk Al. Intermolecular interactions and reactivity of organic compounds. Ed. E.V.Titov. Naukova Dumka, Kiev. P.3 [1983] (in Russian). ! 4. Lutsyk A. l.; Suikov S. Yu, Rudakov E. S. Doklady Akademii Nauk Ukrainy. ~ 40 [1986] (in Russian). . . 5. Rudakov E.S.; Lutsyk A. I.; Suikov S.YU. Russ. J. Phys. Chem. U, 601 [1987]. . . 6. Perri J.H. Chemical Engineerk Handbook, 41hEd. McGraw-Hill Book Company, ,., New York, 1, [1963]. ,, :,, 7. Cerfontairr H., Telder A. Rec.trav.chim. 84,545 [1965]. !& 8. Hanson C., Ismail H.A.M. J,App/.Chem.Biofec. 25,319 [1975]. ,, 9. Deno N.C., Perizollo C. J. Amer Chem. Sot. U, 1345 [1957]. 10. Lobachev V. L., Rudakov E.S., Lutsyk Al. Doklady Akademii Nauk LJkrainy. Q 51 ., [1980] (in Russian). ,, 11. Masterton W.L. J. Chem. Phys. 22, 1830 [1954]. ,. 12. Hammett L.P. Physical Organic Chemisky (Reaction Rates, Equilibria and ~, Mechanisms), 2“d Ed. McGraw-Hill Book Company, New York [1970]. .,,,,, 13. Strahan A. N., Field J.P., Fleming K.A. /SEC 80:/n-t solvent extr. conf - Liege. ... . 1980. P.6211-6216. 14. Hammett L.P., Chapman R.L. J. Amer. Chem. Sot. 56, 1282 [1934]. ,, 15.Arnett E.M., Wu C.Y., Anderson J.N., Bishick R.D. J. Amer. Chem. Sot. 84, 1674 . . . . [1962]. 16. Rilkis C.G., Visotskaya M.P. Izvestiya Kievskogo Politechnitcheskogo Irrstituta. .. . 29, 143 [1960] (in Russian). ,, ,’ 17. Hine J., Mookerjee P.K. J. Org. Chern. 40,292 [1975]. 18. Shahidi F. J. Chem. Sot. Faraday Trans. /. 77, 1511 [1981]. 19.Vinnik Ml. Uspekhi khimii. 35, 1922 [1966] (in Russian). .. ,., , ,, ,. ./. 205 .’. . 6, ,. ., DGMK-Conference ‘The Future Role of Aromatica in Refining and Petrochemishy”, Erlangen 1999 D. L. Hoang, A. Trunschke, A. BrOckner, J. Radnik, H. Lieske ,. Institute of Applied Chemist~ Berlin-Adlershof, Berlin, Germany A New Oxidic Catalyst System for Selective C6+Paraffine Aromatization ‘. ,,, ,. Introduction Reforming of CS+parar%nes is an important source of aromatics as raw materials for “.’ chemical syntheses. For catalyst R&D, it is an attractive task to make catalysts available to the petrochemical industry, which are able to convert appropriate .,, . paraffined selectively into alkyl aromatics, without undesirable side reactions as ‘. ~ ,. cracking or different kinds of isomerization. ., ,, Recently, we reported on attempts to develop such types of catalysts. Composites of microcrystalline lVb oxides and carbon [1,2] as well as group Ivb oxides, which had been modified in their bulk and on their surface by further oxidic components [3], proved to be able to convert n-hexane and n-octane with high selectivity into corresponding aromatics, e.g. n-octane almost exclusively into ethylbenzene and o-xylene. An interesting system among the latter type of catalysts is CrO,/LazOJZrOz. In this paper, we would like to represent catalytic properties of this catalyst system and, moreover, we will report on our present knowledge of the general nature of this catalyst system, of state and function of catalyst components, ., especially of chromia [4-7], as well as of the nature of deactivation of this catalyst ,,,,.’ ‘- system, which is, of course, an essential problem with respect to an industrial ‘. application. ,, As an catalytic test reaction, we applied the conversion of n-octane. The catalyst has “., ~! been characterized by TP methods, in situ FTIR, XPS and in situ ESR. ,- Experimental ..’ Cata/yst preparation. Catalyst samples were prepared by impregnating amorphous zirconium hydroxide doped with 7.0 wt’%0LazOs (MEL, UK) with an aqueous solution of (NHd)zCrOq. The support was immersed in a solution mntaining the appropriate amount of (NHq)zCrOdto yield a catalyst .4CIZ loaded with 4.0 wtVO Cr. By ammonia addition, the solution pH was kept at 10. Under stirring, the excess water was slowly ,., ! evaporated at 50-60°C. The products obtained were calcined in air at 600”C for 4h. Test reaction. Aromatization of n-octane was carried out in a home-made ;.”..,,‘ ,’ characterization apparatus, consisting of a fixed-bed quartz reactor and a gas ,, -. analysis system under normal pressure at 550”C. As feed, a hydrogen or nitrogen ,”.:- flow, respectively, (flow rate=20 ml rein-’) was saturated with n-octane (7.7 kPa). The .’ . ,’, .. ,,. , reaction products were analyzed by an on-line gas chromatography, equipped with a 50 m PONA capillary column. :,,.,! ,,..-. ,, ..,,.,/,, TP methods. Temperature-programmed reduction as well as ammonia temperature- programmed resorption experiments have been carried out in the same reactor as the catalytic tests. TPR was carried out in an 5.17 ‘Yohydrogen in argon flow at a heating rate of 10 Wmin and a gas flow rate of 15 ml/min. The first TPR run (TPRI ) was followed by re-oxidation by a 20 YO oxygen in helium flow at 600”C for 1h. Afterwards, a second TPR run, TPR2, was carried out as described above. TPD of ammonia was carried out in a helium flow of 15 ml/min at the above heating rate. in-situ FT/R. For transmission FTIR studies of adsorbed CO and NO, self-supporting discs of the catalyst were placed into an IR cell allowing thermal treatments in vacuum or in controlled atmospheres. All spectra have been normalized with respect to the weight per mm2. For DRIFTS measurements, a diffuse reflectance attachment (Harrick, USA) have been used. The spectra were collected on a spectrometer FTS-60 A (Bio-Rad), coadding 256 scans at a resolution of 2 cm-l. XPS. The photoelectron spectra were recorded by a VG ESCALAB 220 iXL spectrometer (VG Instruments) with a MgKa source operating at 20 mA and 13 kV. in-situ ESR. ESR spectra were recorded by a C.W.spectrometer ELEXSYS 500-10/12 (Bruker) in X-band. In-situ investigations under catalyhc reaction conditions were -- performed using a home-made flow reactor, which was placed into the cavity of the spectrometer [8]. Results and Discussion n-Octane aromatization on the CrO#LazO#7rOz catalyst ,,4CLZ,,. In Fig. 1, the selectivity pattern of the catalyst 4CLZ is compared with corresponding patterns obtained on the known bifunctional catalysts CrzOJA1203, Pt/A1203and Pt/Sn/AlzOs. :Iectivitv. “h 100 80 Fig.1. 60 Selectivity patterns of n-octane conversion in Hz atmosphere on 40 4CLZ, Cr203/A1203, Pt/A1203as well as on Pt/Sn/A1203at 550”C. 20 0 4CLZ 4Cr/A1203 PtlA1203 PtSn/A1203 4CLZ exhibits a very hiah total aromatization selectivity and only low amounts of side reactions. The di~tribution among the aromati~ is characterized by the practically exclusive occurence of the alkyl aromatics ethylbenzene and o-xylen. 208 ‘ .,,,,,. This product spectrum can only be explained by a C6 ring closure of the n-Cc paraffinic chain. Products of acid catalyzed side reaction like cracking, isomerization, ,, dealkylation and transalkylation are essentially missing. This means that 4CLZ does ,’ not behave like USUd bifunctional Cddpk like pt/At203 and CrzOdAlzOs, see the corresponding product spectra in Fig.1. s “ gso - Z Fig.2. +’60- Conversion of n-octane at : 550°C in Hz atmosphere .-g40- (a), as well as selectivities E al for aromatics (b) and ;20- c octenes (c). 0 D . ------*. -..--* ---- 04 ...... 0 60 120 1s0 240 300 reaction tima, min In Fig.2, the n-octane conversion as a measure of the catalytic activity is plotted ,., .. . . versus the reaction time. Avery high initial activity, but also a significant deactivation of the catalyst is observed. The decline of the activity is accompanied by a certain loss of aromatization selectivity, which is essentially due to the formation of n-octenes as by-products. Cafa/yst acidify. Tab.1 contains the result of ammonia TPD on the 4CLZ catalyst. The data are compared with those on alumina and on the acid zeolite H-ZSM-5. Tab.1. Ammonia TPD results on 4CLZ, y-AlzOSand H-ZSM-5 I I Iow-temoerature desorotion I hiah-temr)erature desomtion I samnle------I l-’ Tm, “c acid sites, mmol/g T- acid sites, mmol/g 4CLZ 220 0.05 indistinct very low ~-AIZOs 200 0.52 indistinct low H-ZSM-5 235 ca. 0.8 480 ca. 0.6 -.. ,, . With 4CLZ, the ammonia main resorption takes place already at low temperatures and its intensity is low, compared with alumina and especially H-ZSM-5. Hence, 4CLZ is of very low acidity, both with respect to strength and number of acid sites. :. This finding explains the selectivity pattern of the catalyst. The catalyst is only able to produce aromatics, but, because of its low acidity, not able to catalyze the - ‘ undesirable side reactions cracking, isomerization, dealkylation and transalkylation, ,’ which take place on the acid sites of bifunctional catalysts. Hence, 4CLZ is not a bifunctional catalyst, but a monofunctional one. ,, ... The low acidity also gives a certain hint to the reaction mechanism: 4CLZ neither should be able to ring closure reactions of paraffined or mono-olefines, which demand acid or metallic sites. That means that on 4CLZ the necessa~ ring closure within the reaction step sequence of aromatization should occur only in a late stage, i.e. by the easy ring closure of a .triene. [4] alkane-+alkene+.alkadiene+ (cis-)alkatriene+cyclohexadiene+aromatic. This is in line with a monofunctional catalyst, the essential ability of which is to dehydrogenate. Finally, the low acidity of 4CLZ is relevant to reflections on the mechanism of deactivation, see the corresponding chapter. A mechanism of coke formation with an involvement of acid sites should be excluded. Oxidation state and dispersion of chromium. We characterized oxidation state and chromia dispersion by TPR, XPS, ESR and FTIRS of adsorbed probe molecules. The results show, that during calcination in air at 600”C, an inhomogeneous CrO. overlayer is growing on the catalyst surface, containing chromates and CrzOs as well as chromium in intermediate oxidation states. XPS (Tab.2) clearly shows the presence of Cr& besides a second peak, which is attributed to Cfl (3sn~5). Tab. 2. XPS results on 4CLZ XPS signal binding energy, eV I surface atom YO I assignment calcination in air at 600”C Cr 2P= 579.2 2.3 Cre+ 578.2 2.4 Cr’3* With ESR spectroscopy, non-interacting Crfi and Cfi ions and Cr20~ clusters were observed (Fig.3). The higher valent chromia species are reduced to three valent chromium by hydrogen at 550”C. XPS (Tab.2), ESR (Fig.3) and FTIRS of adsorbed CO (Fig.4) reveal that the valence state of chromium in the reduced catalyst is +3. Fig.3. [n situ ESR of 4CLZ during reduction in hydrogen at increasing temperature. 210 The single peak in the IR spectrum of adsorbed CO at 2185 cm-l due to C~(CO) indicates that other oxidation states like C~ are not present on the surface in detectable amounts. TPR of 4CLZ (Fig.5) shows that CrO, can repeatedly be reduced and deoxidized, indicating a high chromium dispersion. This observation is consistent with the results of FTIRS of adsorbed NO. The very intense signals at 1852 cm-l and 1710 cm-l due to Cfi(NO)2 complexes, which have been observed on 4Clif (Fig.6a), but not i.e. on the Ianthana-free 4CZ catalyst (Fig.6b), are in accordance with a high number of surface C~ sites on the Ianthana-modified 4CLZ catalyst. In situ ESR spectroscopy (Fig.3) confirms the presence of isolated surface C? species besides CrzOs clusters. The beneficial influence of Ianthana on the dispersion of chromia is probably due to the increased basicity of LazOs-ZrOz in comparison with ZrOZ and a specific interaction between chromia and LazOs, which leads to a more efficient anchoring of the surface chromia species. ? .. 2- A b .15 s o r Fig.4. .,.. b “1- a Adsorption of 0.1-10 mbar CO n at 40”C after reduction of 4CLZ c .05- at 550”C. e ,!, ,. o- } I “,j, 2W0 2250 2200 2150 2100 2050 2CO0 1s50 ,, :.,, Wavenumber 1 m“’ .. . /“’, Fig.5. TPR Profile of 4CLZ and 4CZ reduction in H#4r ,. (TPRI); reoxidation in Oz/He, reduction in HJAr (TPR2). ~ 2rxr Temperature, ‘C ,. -,.,’;.;. ,,; ., l,”; , ‘, ,,., 211 ./,,, ,! .,> ‘.,, , ,.! .’ r .’,” ,.:; 1710 A 1.5 b s o Fig.6. 1852 : ‘- Adsorption of 1 mbar NO a after reduction of a) 4CLZ n c and b) 4CZ at 550”C. e s 0, ~ ,’- . #, Wavenumber I cm-’ Catalyst deactivation. In the chapter “Catalyst acidity”, we anticipated that coke formation on acid sites cannot be a probable deactivation mechanism of 4CLZ. The question arose, if coke formation at all is an essential source of deactivation. Fig.7. n-Octane conversion at 550”C in H2 (a), in N2 (b), as well as percentage of coke deposited in Hz (c) and in Nz (d) atmosphere. ~iJ ~ 0 y–.--r.—. —~— —— 0 60 I20 180 240 300 time on stream. min In Fig.7, n-octane conversion and carbon formed on the catalyst are depicted as function of the reaction time. Curves a and b describe the conversion in a n-octanelhydrogen and a n-octanelnitrogen mixture, respectively, and curves c and d depict coke deposition in these two reaction mixtures. The results allow to conclude: i) The deactivation is less significant in hydrogen atmosphere by far. This is typical for deactivation by coke formation, which is mostly retarded by presence of hydrogen. ii) The deactivation is always accompanied by coke formation. In case of stronger deactivation, stronger coke formation is observed. This is a further 212 ,’ ..,,,, . ‘,<.,..’ ., .-,. ,,, ,,. , ;’.,.,,.. ., ,.:.,,.. -. :.,,: argument for deactivating coke. iii) Burning off the coke in oxygen containing :.; ~,’: :,,j.~., atmosphere practically completely restores the catalytic activity, see the arrows at ‘ ~.~’~~~.-.~”.;: ~ the respective curves. This can be taken as strong evidence for deactivating coke. ‘-$ ,, .,,.: .,,. Nevertheless, a certain contribution of changes of the surface chemistty or of the ‘ ~ ; ‘..’ -’ , -, dispersion of surface species to the deactivation cannot yet be excluded. 40 30 1 c 20 .-0 . . . 0 fn , , , IDo y 50 , z z 40 - Fig. 8. 30 Acidity changes during the deactivation process at 550”C ‘. 20 - in Nz and in Hz atmosphere. 10 5 rnin o t t , 1 100 200 300 400 500 600 ,., , temperature, “C .,,. ., ., .,-.,. Fig.8 gives information on acidity changes of 4CLZ, determined by ammonia TPD, during the deactivation process. The acidity measured hardly changes with ~” ,’- , deactivation. This result is, on the one hand, unexpected in some way, because hardly a blocking of the catalyst surface by coke can be verified by ammonia TPD. ‘ ‘ Nevertheless, the finding is in line with the assumption, that the deactivating coke ‘ : has not been formed on acid sites. Obviously, similar mechanisms of coke formation as on metal catalysts must be taken into consideration for CrO./LazOJZrOzcatalysts. ,. Conclusions ,. ● CrOJLa20JZr02 is a selective catalyst for n-C6+ paraffine aromatization. The catalyst system is non-acid and, therefore, monofunctional. Concerning the ~ reaction mechanism, C6 ring closure of a triene is assumed: alkane+alkene+alkadiene+ (cis-)alkatriene+cyclohexadiene+aromatic. .. . ● The active catalyst contains Cfi species in two agglomeration Skh.?S, ,. isolated ,, .,,.: ,,, ,.,. , Cfi ions and CrzOs clusters, the catalytic importance of which is not yet clear. ~~.,. ,,.: ,.;,., There are no hints on existence and importance of other Cr oxidation states. “;’J .. ,. .,,, ,,.: ,:,>::.,”,! ~ ,-..~ ,,,..., ,’-,, .,,,,:,,, ,, { . ,::,; ,,r .,, . 213 ,. ,, . , ,,,’.:: .-’: -~-, ~,,’$ . Lanthana in CrO,/LazOJZrOZ stabilizes the existence of reoxidable highly dispersed chromium species and isolated coordinatively unsaturated C~ ions. ● Catalyst deactivation of CrO./LazO4ZrOz is predominantly due to carbonaceous deposites, which must be formed without involvement of acid sites. References [1] H. Lieske, D. L. Hoang, 1-1.Preiss, German Patent DE-OS 19516318. [2] D. L. Hoang, H. Preiss, B. Parlitz, F. Krumeich, H. Lieske, Appt. Catal. A General, 182, 385(1999). [3] H. Lieske, D. L. Hoang, German patent DE-OS 19612000. [4] B. M. Weckhuysen, R. A. Schoonheydt, J.-M. Jehng, 1.E. Wachs, S. J. Cho, .’- - R. Rye, S. Kijistra, E. PoeIs, J. Chem. Sot. Faraday Trans. 9f, 3245(1995). +.,. [5] J. R. Sohn, S. G. Ryu, Langmuir 9,126 (1993). [6] K. Arata, M. Hino, H. Matsuhashi, Appl. Catal. 100,19 (1993). [7] V. Indovina, Catal. Today 41,95(1998). [8] A. Bruckner, B. Kubias, B. Lucke, R. Sto13er, Colloids and Surfaces f 15, 179(1996). [9] Z. Paal, Adv. Catal. 29,273 (1983). Acknowledgement This work was supported by the Berlin Senat Department of Science, Research and Culture, by the Federal Ministry of Education and Research of the Federal Republic of Germany as well as by Deutsche Forschungsgemeinschaft. Moreover, the authors are grateful to BASF AG, Ludwigshafen, Germany, and MEL Chemicals, Manchester, UK, for supporting parts of this work. 214 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 L. V. Bogutskaya, S.V. Khalamejda, V.A. Zazhigalov Ukrainian-Polish Laboratory of Catalysis, Institute of Physical Chemistry, National Academy of Sciences of Ukraine, Kyiv, Ukraine .? :“, . d’ ,>,... ‘} -,, , Benzene Oxidation to Phenol by Molecular Oxygen on Medicated Mo03 ,. , :4 .-.. : ,. ,,.’“ \ ,’, ~ Direct oxygenation of aromatics into their hydroxylated derivatives is one of the most ,> ,<,,’:?:;,:” ..:.-’. ,., : , challenging topics in organic synthesis and has also been expected in industrial chemistry, for 3 .. ,, ..,-, :’,,- example, benzene oxidation to phenol as an alternative method to the cumene process. Some ,, ,. success in this direction was achieved when using NzOas an oxidizing agent [11.Any attempt ,. .’ ,,-, to realize benzene oxidation with molecular oxygen leads to mainly complete oxidation products (COJ and a very few per cent conversion into the partial products. The present work was aimed to modification of the molybdenum oxide catalytic properties by means of its mechanochemical treatment in different conditions. The catalytic anisotropy of MoO] was shown in partial oxidation of propene, alcohols and Cq hydrocarbons [2-3]. But while V-MO oxide composition is a basis of commercially used catalysts for benzene selective oxidation into maleic anhydride, individually Mo03 is very low active in this reaction and even at 500-550 ‘C benzene conversion reaches a few per sent, maximum value of the selectivity (S) to phenol is about 3.6 ‘A. The treatment of MOOSin a planetary mill (3,000 rpm) with balls of 10 mm diameter in the presence of water leads to significant growth in specific surface area (from 0.6 to 5.1 m2/g) and increase (OkO) crystallographic planes exposure. In so doing, some amount of maIeic anhydride forms (S = 7.0 0/0), but the selectivity to phenol drops to 0.3-2.0 O/O.However, an appearance of the reduced phase h’b02.8 in the catalyst impmves the catalytic @ivitY - the l,’ reaction starts at the lower by 100 ‘C temperature. The similar results are observed at the treatment in the presence of ethanol. Further increase in the reIative intensity of (OkO)planes can be reached by continuation of the treatment (up to 50 rein) or replacing the dispersing medium on water-benzene mixture. But in such a case phenol forms in traces amount and the selectivity to maleic anhydride riches very high value -40 ‘%. at 54 ‘Y. benzene conversion. For all the treated samples R spectra show some weakening of the lattice Mo-O and terminal Mo=O bonds. Under long mechanochemical treatment with smaller balls (6 mm diameter) in the presence of aIcohoI the led up mechanical energy is spending mosdy not to oxide particles crushing but its structural and chemical transformations. The specific surface area increases up .,! . ,> to only 2.4 m2/g (50 rein) but all the (OkO)crystallographic planes lose their intensity and (021) .+, plane becomes the most intense one. Atler XRD study the crystallographic shear structure of MoXOZJ composition is found as a new partially reduced phase produced from initial molybdenum oxide during such treatment. The obtained sample is proved to be the most active one in the reaction of benzene oxidation by molecular oxygen and the selectivity to phenol is ,’ about 7.5 0/0 at 335 ‘C. References 1. G.L Panov et al. Appl.Catal.A General. 82 ( 1992) 31. 2. K. Bruckman et al. J. Catal. 104 (1987) 71. 3. M. Abon et al. J. Catal. 134 (1992) 542. 215 DGMK-Tagungsbericht 9903, ISBN3-931850-59-5 ,’- < . ,. .! 216 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangan 1999 M. Walter, E. Schwab, M.G. Koch, P. Trtibenbach, S. Dining BASF AG, Ludwigshafen, Germany Selective Hydrogenation of light Reformate for Production of high-purity Benzene ., ., .<’, 1. Introduction Catalytic reforming is one of the most important processes to produce high grade motor gasoline. Feedstock for catalytic reformers is mainly naphtha from crude oil distillation. By catalytic reforming the octane number of these gasoline components is increased from 40-60 RON to 95-100 RON. Besides isomerization and dehydrocyclization reactions, mainly formation of aromatics by dehydrogenation of naphthenes occurs. Thus, catalytic reformers within refineries are an important j,-,, source of BTX aromatics (benzene, toluene, xylenes). ,,.,, ,.-’, Frequently, high purity aromatics are recovered from these streams using modern ,.. liquid extraction or extractive distillation processes, e.g. the Krupp-Uhde -,.,1. .,,’ Morphylane” process. Aromatics product specifications, notably the so-called Acid- Wash-Co/or according to ASTM D-848 which is a visual appreciation of the achieved purity, have obliged producers to use clay treatment to remove traces of impurities .,, such as diolefins (mainly methylcyclopentadiene, MCPD) and olefins in order to meet the required AcWVash-Co/or of max. 1. The conventional clay treatment is a multiple vessel process which periodically requires disposal of the spent clay in a suitable environmental manner. The disadvantages of the clay treatment are obvious: s Frequept exchange of used clay against fresh clay is necessary. ● Disposal of the spent clay is costly due to strict environmental regulations. ● Yield losses of up to 5% aromatics have to be taken into account. . Installation causes high investment costs. BASF has developed a front-end continuous Selective Catalytic Hydrogenation Process (SCHP) in order to achieve high product-purity as an alternative to clay treatment. This new process is very efficient, cost effective and environmentally compatible and is already verified in front of Krupp-Uhde Morphylane” Extractive Distillation Units. ... ! ,,..,,.;; .-, ,,-, , ,,,, .,. ., “.”,..,.- -,; ,. ,...... , DGMK-Tagungsbericht 9903, ISBN3-93185@59-5 217 2. Process Scheme The SCH process of reformer streams can be easily integrated into the downstream section of a catalytic reformer unit. It is located between the catalytic reformer itself and the extractive distillation unit (Figure 1). Clay Treatment (CT) Selective Hydrogenation (.SCHP) Catalytic Reformer Catalytic Reformer Section Section v SCHP (BASF) GX I v v Extractive Distillation Extractive Distillation Separation Separation I Clay Treatment Separation II I Pure Benzene b (95-97 % yield) (100 % yield) Figure 1: Selective hydrogenation (SCHP) versus clay treatment (C7J. The process flow scheme of the SCHP unit is illustrated in Figure 2. The unit consists of a feed pump, a feed pre-heater and a fixed-bed reactor and is processed as a one stage process. The reactor consists of a very efficient distribution plate, one catalyst bed, an integrated separator and an off-gas cooler to minimize hydrocarbon losses. Optionally a recycle unit for the product stream can be installed in order to 218 ,., ;“: ,, ,, :’ dilute the feed and therefore minimize the temperature increase across the catalytic , .. bed and guarantee a homogeneous coverage of the catalyst. :,. ,, .’ Core of the process is a highly efficient nickel hydrogenation catalyst. ~., . . .- ,. :. M PE1’+ Feed Pre-heafer B B+ Feed ~Pump 1 I Reactor 1 1 r 1 LP-Steam 3 I P ~ am Off-gas Cooler Recycle ; Unit ~ ,,~-,.. ------~m ., .,:: ., :’ Figure 2: Process flow scheme for the SCH process. 3. Experimental ,:,,:. ,.,::,I , ;:. ,:-, ,, . The experiments have been performed in a 200 ml trickle bed reactor tilled with . ~:..,.,,-,.,,, . ,, ,,,>,.:. 50 ml catalyst, using the following reaction conditions: ,,~,, ,) ,f?,.A!1 ..<,,.,,.;,, ., .!.. ~:, ,-.. ~$,,,-’f . . . .,, ,;7, . pressure: 6-14 bar ,’ ,. ?,,.-:, ,;. : ;.;V::,’, ,$,”; :,, , ,,, . temperature: 20-70”C ,,... ,, .. ’,, . ,7 ! ,. .; ● space velocity (whsv): 6.0 kg/(l*h) ,., , :,:”,’;.. ,p{ ,. ..’, .:,~.,:,.-‘f .,J.T;.; ,,. . ,’, ., ,.l..:. - !, The feed used for the experiments consisted of reformate with a benzene content of ,> f ,.,’,?j -, -,,...... ,.,..-. 55-65 !4. and a methylcyclopentadiene (MCPD) content of 800-1400 ppm. Further :,, ,, ,.,.,.’, ‘.O$: ,. ,-:’~ components of the reformate feed were Cs to CTalkanes and alkylated aromatics like , ,’,:?2,:.;;,,:/ toluene and xylenes. .~.,,,.,,.;,, y 1 ,,..>,,.,’, ,:..:, Pure hydrogen (99.9 vol.-%) was led through the reactor in oncethrough operation. ..j >.,, ;{ ,;.’ ., The liquid product was then analyzed by gas chromatography (GC) in order to .. .<-.., ,’,,,,’:,; , ,, ,..~, 1 determine the concentrations of benzene and MCPD. The Acid-VVash-Co/or (AWC) ,,:,.,<.,;, ! ;:,! ; ,-,., ,,:,,< of the product was measured by adding concentrated sulfuric acid and comparison ,,,.’,,’.”,‘ ,;,, j: ,:# with reference samples according to ASTM D-848.1 ,, ‘ .,,.-,< , :“(;, :.;’‘. In order to evaluate the quality of the final product coming out of the EDU, the ,, ,1, selectively hydrogenated reformate was submitted to a distillation procedure for ,...,,!,:, .W‘l ... ..<:+,.,.,’, achieving a benzene purity of >95 wt.-’%o prior to AWC determination. :.:>‘1 ,,,.-s.’;.,, :,,.~,::: ., !, ;-,,.:,,, . . . ,;,.;.,,$:,~‘ ?1.’< .,.. ,.. , y., , ,. ,,f,, ,: The aim of the selective hydrogenation process is high diolefin conversion and low benzene loss. In order to meet those requirements even at mild operating conditions a new nickel-based low acidic catalyst has been developed at BASF. The suitability of this new material in a low messure trickle-bed rxocess is outlined in Table 1. where the obtained results are ‘compared with the performance of a state-of-the-a~ palladium-based catalyst. Table 1: Performance test of a Ni versus a Pd catalyst (6 bar, r.t.). Ni Catalyst Pd Catalyst (HI-89RED) start of run (SOR) MCPD [ppm] <10 170 benzene loss [Y.] < cJ.3a <0.1 after i ~A h “,.”, .“ . ,,, MCPD [ppm] < 1(I I ~90 (14 bar, 70”C) benzene loss [70] < O.la <0.1 a Benzene hydrogenation is slightly higher at SOR due to a higher initial Hz concentration at {he surface of the freshly reduced catalyst. The decisive criterion for these catalytic hydrogenation experiments was to lower the content of methylcyclopentadiene (MCPD) in the feedstock to less than 10 ppm, minimizing the saturation of benzene to cyclohexane (< 0.3 ‘Yo). The results show that the nickel based BASF catalyst HI -89 RED is superior to a t)alladium based catalvst which is frequently used for- selective hydrogenation of diolefins. Using the nickel catalyst no MCPD could be detected (< 10 ppm) in the reaction product, while the palladium catalyst showed significant amount of MCPD, even after increasing reaction temperature and pressure. The effectiveness of the nickel catalyzed selective hydrogenation process for MCPD reduction can also be demonstrated by monitoring the Acid-Wash-Co/or (AWC) of the reformate before and after selective hydrogenation (SCHP) and after distillation. These results are listed in Table 2. Tab/e 2: Influence of the SCH process on the A WC value of benzene after the extractive distillation unit (EDU). Reformate-AWC AWC after SCHP AWC after EDU Conventional process >14 4-5 (EDU followed by clay treatment) SCH process >14 3-4 <1 (BASF) The table demonstrates that using the conventional process without SCHP the AWC value for reformate (AWC s 14) can only be reduced to AWC 4-5 after the extractive distillation unit (EDU). Insertion of the BASF selective hydrogenation process however produces AWC <1 afler EDU and therefore allows to meet the specification for benzene (AWC = max. 1) without additional clay treatment. 220 .,. - 5. Environmental aspects ,,, : The conventional clay treatment used for removal of diolefins and oletins from , ., ,, reformate streams is environmentally critical since it requires a frequent exchange of ; !, the used clay. These by-products produced in huge amounts during this process ‘, ,“ cause severe disposal problems since they cannot be recycled. Furthermore a ,. significant amount of benzene remains on the used clay and leads to yield losses of U; tO s~o. The SCH process, however, achieves a selective removal of diolefins without yield losses and without by-products. This demonstrates nicely the beneficial effect of ca- ... - talytic reactions for obtaining environmentally friendly chemical processes according to important industrial goals like Sustainable Development and Responsible Care. ,, 6. Summary -,., ,,. , The SCH process, developed by BASF, is a modern and innovative solution to purify .,. reformer streams. The catalytic process is characterized by low investment and ,. ,, operation costs, high flexibility concerning feedstock composition, mild operating conditions and a highly active and selective nickel-type catalyst removing all .: impurities and avoiding formation of any by-products. The advantages of the SCI-I process in comparison to the conventional clay treatment are summarized in Table 3. Table 3: Purification alternatives for reformate streams. Selective Catalytic Hydrogenation Clay Treatment ● selective removal of diolefins and ● oligomerization due to acid centers olefins ● no by-products ● formation of green oils ● long cycle and lifetime ● short lifetime . overall more economical ● no flexibility on purification depths Until today three SCH units according to this process-design have been licensed in front of Krupp-Uhde Morphylane@ Extractive Distillation Units. The first plant at Tonen in Kawasaki/Japan has been started up successfully in 11/99and fultils the required specifications. Further plants located in Germany and the Middle East are scheduled to come on stream in 111/99...... References 1. H.O. Folkins, in Ullmann’s Encyclopedia of Industrial Chemistry, Vol. A3, p. 475-505, VCH,Weinheim(1985). 2. P. Polanek, H.M. Hooper, J. Mi31er, M. Walter, Purification of Reformer Streams by Catalytic Hydrogenation, NPRA annual meeting, San Antonio, Texas, USA (1996). 3. BASF, patent application, PCT/EP97/00960 (1996). ., .. ‘ 221 ‘ ;:. , ..,’ —,---- . . ,, :: ,..,,, .. <%.,’, .,. – ...... ‘, .:::>,. ., ~ .- ,- ,;~,; ,.!’ 6, ,. 222 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Ertangen 1999 E. Dietzsch 1),P. Claus 2),D. Honicke’) 1,Chemnitz University of Technology, Chemnifz, Germany, 2)Institute for Applied Chemistry Adlershof, Berlin, Germany ,,. . ,,. . : Supported Ru-catalysts for the partial Gas Phase Hydrogenation of Benzene ,. .,, The objective of the present study was to determine the influence of the support material of Ru catalysts on the cyclohexene formation in the gas phase hydrogenation of benzene. Supported catalysts were prepared by impregnation of some common support materials with a ruthenium precursor or by the application of the sol-gel method. The cyclohexene formation was investigated i, during the hydrogenation of benzene over these catalysts in the presence of methanol as reaction .,,. modifier. It was found, that several support materials are suitable as catalysts supports, but there was a dependence of the activity and selectivity on the catalyst properties caused by the support material. ,. By an increase of the hydrogen partial pressure it was possible to enhance the yield of cyclohexene. A ,,, .-. ,, comparison between different preparation methods of the catalysts is given. ., ,,, ~,> 1 Introduction For the liquid phase hydrogenation of benzene to cyclohexene many different ruthenium catalysts were suggested in the literature. Besides support-free powdered ruthenium [1,2], also supported ruthenium catalysts such as Ru/y-Alz03 [3], Ru/SiOz [4] and especially Ru/LaZOS [5,6] were successfully used for the formation of cyclohexene. However, the influence of the support material of ruthenium catalysts on the cyclohexene selectivity in the gas phase hydrogenation has not been thoroughly investigated yet. Gonzalez et al. found that the support material of Ru catalysts influenced the deactivation rates of these catalysts in the gas phase hydrogenation of benzene to cyclohexane ~. A low surface area supported -:, ,”,.,, ,:$: I Ru/a-A1203 catalyst was described for the cyclohexene formation by van der Steen and ; Scholten [8]. [n a more recent study the reaction was performed on supported Ru/La20~ and Ru/Si02 catalysts [9]. However, cyclohexene selectivities using these supported catalysts were lower than on a non-supported Ru catalyst [10]. In all cases the presence of water vapor as reaction modifier was necessary for the formation of the desired cyclohexene. It was assumed, that the ruthenium catalyst is surrounded with the reaction modifier (Figure 1) ,, ?,.. which repels the formed cyclohexene from the catalyst surface into the continuous gas phase and suppresses its readsorption and further hydrogenation to cyclohexane [11]. The rate of cyclohexene hydrogenation to cyclohexane is three orders of magnitude higher than that of 223 =.-—-.., . ..,.. .. :-.. ,$ ,,, <. ,, .-,. s .:. L ,, ,’ Continuous Dhase #,. . .-w--- F’1 A ~ ;,,%Q.~H2 reaction .; modifier 47 O& ’co supported Ru-catalyst Figure 1: Scheme of the reaction course of the partial gas phase hydrogenation of benzene on a supported catalyst surrounded with a reaction modifier. benzene to cyclohexene [9]. Therefore, satisfactory selectivities to cyclohexene may only be obtained at low conversions. In the present study several supported Ru catalysts prepared by different techniques were investigated in the partial gas phase hydrogenation of benzene under equal reaction conditions. As previously reported, methanol was selected as a more eficient reaction modifier than water [12]. 2 Experimental A few supported Ru catalysts were prepared by a liquid phase impregnation of the common support materials SiOz, y-A120~, TiOz (P25) (Degussa), ZrOz (Aldrich), and C (activated carbon, Takeda Shirasagi C) using solutions containing RuCl~ or Ru(acac)~ (Aldrich) as a precursor. Several characterization methods were applied and results are summarized in Table 1. It can be seen, that the ruthenium content of the catalysts determined by ICP-AES is about 1 weight70 for all samples. The specific surface areas of the catalysts measured by Nz adsorption varied over a wide range between 3 and 1400 m2/g. The dispersion expressed as the ratio of H/Ru differs probably as a consequence of the unequal surface properties of the support materials. Prior to each hydrogenation experiment 0.2 g of the supported catalyst mixed with quartz was first oxidized for 1 h at 573 K and then reduced in a hydrogen/nitrogen stream for 2 h at 473 Kin order to activate the catalyst. 224 :,’, ., ~. .’ ...: ,’,,,...... ,:...... ,. , ~’. .’ ,. Table 1: Characterization data of the supported Ru catalysts prepared by impregnation technique. catalyst RuLSiO. Ru/y-A1203 RufTi02 RufZr02 RuIC ,, .,, precursor RuC13 Ru(acac)~ Ru(acec)~ Ru(acack Ru(acac)j Ru content’] 1.00% 0.9570 1.04 Y. 1.28 % 1.00 y. :, WRU 2) 0.03 0.04 0.09 0.02 ...,,. , ij,, ‘, dPI nm’) 20 4 14 4 <1 :,’ BETI m’lg 3) 17~ 110 33 3 1436 V& cmalg 31 0.90 1.01 0.17 0.02 0.74 determined by:’) ICP-AES ‘] Hz chemisorptiow 31Nz adsorptioru In addition, selected supported ruthenium catalysts were prepared by application of the sol- gel method. A solution of a support corresponding alkoxide (Al(i-C3HTO)3 (Merck), Ti(i- C3HTO)4 (Aldrich), M!3(C2H50)2(Hfils), Fe(C2H50)J in ethanol was mixed with Ru(acach (Merck) and heated undar reflux. Afler the addition of water to the mixture, the hydrolysis of ,., ;,f the alkoxide occurred. The resulting Ru containing gel were first dried at room temperature ,., ,.. - .’, .”,,, ,< and then at 373 K, calcined in air at 573 K for 6 h and then reduced in a H~N2 stream for 2 h .,, ., “. . ,: ....’.. ,.: at 473 K. With the help of the described procedure supported RU/A1203, RU/T102, Ru/MgO, -, .- ‘,; ..:’> ,.: and Ru/Fe20a catalysts were prepared. Some characteristics of these catalysts are ,, ,:’:..~,,:, ..,.~y,v,, ,~, >,. .,. .: summarized in Table 2. The ruthenium contents of the catalysts were measured by X-ray ,;-’ ,,. ,. ,-,,,- ... ,, ,.’ microanalysis and were between 0.5 and 2 weightYO. Note, that the support oxides are .-. ::-> amorphous materials. The specific surface areas of the sol-gel catalysts (BET/ Nz- adsorption) varied between 38 and 360 m2/g. ‘, L,,, ,. ,j ,, The apparatus used for the benzene hydrogenation experiments at 110 kPa and 353 K was ,. ‘;. .’ ., ,. previously described [12]. The benzene partial pressure and the molar methanol/benzene :.:: :,4 .,, ..., ratio in the present study was kept constant at 0.11 kPa and 1, respectively. The hydrogen $. .,, ,! ,. ,,>, , partial pressure was varied in the range between 11 and 66 kPa, while the total gas flow rate was adjusted to 6 I/h (STP) balanced by nitrogen. ,, Table 2 Characterization date of the supported Ru catalysts prepared by the sol-gel-method. t, catalyst Ru/A1203 Ru/Ti02 Ru/MgO Ru/FezO, Ru content 11 2.0 y. 1.0% 0.5 % 1.070 H/Ru ‘1 0.11 dpI nm’) 6 6 20 50 BET/ m’lg’) 360 304 164 38 V+ cm’/g ‘i 0.64 0.56 1.10 0.20 determined by’) WDX 2)Hz chemisorptiorx 31Nzadsorption: 225 ,. ., ,, . 3 Results and Discussion In the investigation ofall supported Ru catalysts an interesting timeon stream behavior of the conversion and selectivity was observed. An example of this behavior is shown in Figure 2. The benzene conversion decreased rapidly and achieved a nearly constant value after 5 hours. This deactivation corresponds toa kinetic ofsecond order, already described for Ru catalysts [7]. Thecyclohexene selectivi~ increases initially, followed byaperiod ofa nearly constant value. The selectivity to cyclohexane conversely decreased with the identical time behavior. The results of the hydrogenation experiments using the supported Ru catalysts prepared by impregnation are summarized in Figures 3 and 4 at steady state conditions. The calculated specific activities (Fig.3) as well as the selectivities and yields of cyclohexene ■ benzene Ocyclohexane Acyclohexene o. 0. 0.3 .- 0.2 0.1 ; 0 5 10 15 ,. time on stream/h #, . . Figure 2: Time on stream behevior in the gas phase hydrogenation of benzene on the supported impregnated catalyst Ru/ y-AlzO~(Table 1) in the presence of methanol; T= 353 K, p= 110 kPa, p,= 0.11 kPa, pm= 0.11 kPa, pH,= 22 kPa, bal. N,, W/FO~=747 gh/mol; (Fig.4) are depicted as a function of the hydrogen partial pressure at the given reaction conditions. The specific activities of the catalysts increased with increasing hydrogen partial pressure but decreased in the order of RurTiOzzRu/ZrOz> Ru/SiOz>Ru/y-AlzOa> Ru/C. Note that on the alumina supported catalyst activities similar to that on SiOz, Zr02 or TiOz supported Ru catalysts at higher hydrogen partial pressures were reached. The increase of the specific activitiy of the carbon supported Ru catalyst was very weak compared to the other supported Ru catalysts. In contrast to the increase of the specific activities of the catalysts, the selectivities to cyclohexene decreased with increasing hydrogen partial pressure. . . 226 ,. ... <., ,. , ::“ ~ ,’,. ,.,~ ,. .,:. ,. i,’.:. ,...,” : ;, ., :,..”, ,.,, ...... ’. ,, ’$’ <,’. ‘,’,,;. ,:, ., ,.., .,, , ,:, .,,. .. . ,, A Ru/ SiOz :0 .20 +RUITIN203 , ‘ -“ r= “HRu/Ti02 w •R~zro, .,: .’” ~ :15 ‘VR~c ,, g - “-.”..’ . ,- 1 ; 10 -, ...... = ,- .-> . . -G (U5 ,, , . . . ., / ., .,. ,’ , i, u 10 20 30 40 50 60 70 ,- k,pH2/kpa .. . . Figure * Specific activities versus hydrogen partial pressure over impregnated Ru catalysts (Table 1) in the gas phase hydrogenation of benzene in the presence of methanol; T= 353 K, p= 110 kPa, p~= 0.11 kPa, PM=0.11 kPa, bal. N2, w/PB= 747 ghlmol; i, The selectivity decrease was more stronger for the SiOZ, ZrOz and Ti02 than for the y-A1203 . ,,, ... . and carbon supported Ru catalysts. The highest selectivity to cyclohexene of nearly 38 YO was attained on Ihesuppotied Rui1102 catalyst. The yields to cyclohexene increased with .::.,\, ... . hydrogen partial pressure and decreased in the order of Ru/TiOPRu/ZrCIPRu/y- ,, ‘ 0.51 ■ Rti Ti02 “ A Rul Si02 ,, ❑ Ru/,Zr02 ‘ ● Rul@tL03 v Ru, ~ -0 z “s 0.03 0.02 \ 0.01 L g, ,- oLA—w_al 0 10 20 30 40 50 60 10 20 30 40 50 80% . -, ..’ Figure 4 Cyclohexene selectivities and yields versus hydrogen partial pressure over impregnated Ru catalysts (Table 1) in the gaa phase hydrogenation of benzene in the presence of methanol; T= 353 K, p= 110 kPa, p~= 0.11 kPa, PM=0.11 kPa, bal. Nz, W/P~= 747 gh/moh .,, . 227 .’ - #, ,. .! A -A I 1 ‘“20 30 40 50 60 70 .. .: p$ I kPa”” Figure 5: Specific activities versus hydrogen partial pressure of sol-gel prepared Ru catalysts (Table 2) in the gas phase hydrogenation of benzene in the presence of methanot T= 353 K, p= 110 kPa, P8= 0.11 kPa, PM=0.11 kPa, bal. N2, W/Foe= 747 gtr/mol; A1203>Ru/Si02>> Ru/C. In sum, Ti02 and ZrOz supported Ru catalysts showed the best catalytic performance followed by the AlzO~ and Si02 supported Ru catalysts. The results of the hydrogenation experiments using the supported RU catalysts prepared by the described sol-gel method are summarized in Figures 5 and 6. The calculated specific .<, -., .~. >,-, , .,; , + ,. ,,* ;;’7. ,..:,. 0.5 ) ,, .,,> .’ , ., V .Ftu) A@3 O Rr/ MgO ‘-: ,, : “” ❑ W Ti02 A Ru/ Fe203 “L Q .: a3 \ \ “: 0.02 ,. c 2’ :“ <’ ~ 0.01 J g I I I , t 1 0 o 10 20 30 40 50 60 10 20 30 40 50 60 ,!-- },, PH2 I kPa . Figure 6 Cyclohexene selectivifies and yields versus hydrogen partial pressure of sol-gel prepared Ru catalysts (Table 2) in the gas phase hydrogenation of benzene in the presence of methanol; T= 353 K p= 110 kPa,, P8= 0.11 kPa, PM=0.11 kPa, bal. N,, W/F08= 747 gh/mol; 228 ,,,., .... activities (Fig.5) as well as the selectivities and yields of cyclohexene (Fig.6) are shown as a function of the hydrogen partial pressure at steady state conditions. The specific activities of these supported catalysts increased with increasing hydrogen partial pressure and decreased in the order of RulTi02>Ru/Al,0a >Ru/MgO>Ru/Fe*OS. In comparison with the -,.,, impregnation catalysts, the activity of the alumina supported catalysts was nearly in the same range, but the activity of the titania supported impregnation catalyst was about five times higher than those of the corresponding sol-gel catalyst. The reason for this may be ,, presumably the different cristallinity of the titania phases. The titania of the sol-gel derived catalyst is mainly amorphous, whereas that of the impregnation catalyst is crystalline and contains rutile and anatase in a ratio of 41. The highest cyclohexene selectivity (Fig. 6) of @Yo was observed on the sol-gel prepared Ru/MgO catalyst. The selectivities decreased in ,. :,,. ,:, the order of Ru/MgO>RulliOZ> Ru/Fe*OS>Ru/Alz03. The ‘attained cyclohexene yields ,,, . (.: . . increased with increasing hydrogen partial pressure and decreased in the order of Ru/A120S= .;, ..<:: Ru/TiOz>Ru/MgO>> Ru/Fe20~. The cyclohexene yields on the alumina supported ruthenium ., :.,,. ,..,,,., ‘. .,... ,,, catalysts were independent from the preparation method in the same range, but quite ,. .,., ,, ..-,, different on the titania supported catalysts according to their different activities. ., .:.. ,,: \’ ,”,- ‘,, . :,$: ;, ,,f.:~,,.” . .,, .,.. 4 Conclusion !, ,,,,-I .. ,,.,,’,,< ,... The cyclohexene formation in the gas phase hydrogenation of benzene on several suppofled ,.-”.- ,,, “<- Ru catalysts prepared by impregnation and sol-gel method was investigated in the presence .,, , ,,. ., , .,... ,, .,.-. .,- of methanol as reaction modifier. Many common support materials, except microporous .,.: .; :.,,- ;! activated carbon are suitable supports for the preparation of Ru catalysts by impregnation. .,...... ,,..,->.,. ,..., .. . For these catalysts the order of suitability of the support materials was TiO~Zr02>SiOzW- . .,,.,,.,,., ,,, ,.’ A120~ relating to the catalyst activity. Without any optimization cyclohexene selectivities ,.-.,.. . between 25 and 38 %’. were observed. The enhancement of the hydrogen partial pressure was proved to be an effective tool to increase the yield of cyclohexene. This was possible because the increase of the specific activity of the catalysts with increasing hydrogen partial pressure was stronger than the corresponding decrease of the cyclohexene selectivity. Some supported Ru catalysts were prepared by the sol-gel method. It resulted in very appropriate catalysts which gave cyclohexene selectivities up to As~o. But their activities were lower in ., ,. /,!,,” ;;” comparison with that of the impregnation catalysts. While sol-gel prepared Ru catalysts with ,’, , ..:,.-.. ; ~,{1 . . ,,.“:::’;,:,’,’,,,,...! TiOz, A120~and MgO as supports were effective ones, Fe*03 was an unsuitable support. l,,:’! ,.<,:,,’ ,, ?,.:,:, . ,,,.j “. ‘! ../:,,’, ,. .;’,., ,,,,,-,.>,..,,., ., ..,l Acknowledgement ,,:t ‘:,$‘* :, 1, “:i?.,, ,-’, ,,,.--:: ,,,,1 The authors express their gratitude to the .Max-Buchner-ForschungssIiftung” and the .Fonds der ‘:~~,~.4, !! ,, :,,,;,-,- ,:, ~ Chemischen Industrie” for financial support of this work. ~,,-I ‘,,,,>:,, ,,...... ( ,. :$,.;: ;,,.!.?; ,, ,;:, .: . /; ,,v,;<~i’,. .: ;>,.,,.!{” ,, ,,. :J - .. ,,,: ~.,,1 ,, “. .,,.,, ,,,,>,; 229 ,, . ,’. l,:.,, , . ,,,,,$.L,. .,,, :{ :,,,,,,, ...,,, ~ [1] Odenbrand, C. U. I., Lundin, S.T, J. Chem. Techn. Biotechnol. 30 (1980) p. 677. [2] Nagahara, H., One, M., Konishi, M., Fukuoka, Y., Appl. Surf. Sci. 121/122 (1997) p. 448. [3] Zhanabaev, B.Z., Zhanozina, P. P., Utelbaev, B.T., Kinet. Calal, 32 (1991) 1, p.191. [4] Niwa, S., Mlzukami, F., Isoyama, S., Tsuchiya, T., Shimizu, K., Imai, S., Imamura, K., J. Chem. Techn. Biotechnol. 36 (1986) p. 236. [5] Dobert, F., Gaube, J., Chem. Eng. Sci. 51 (1996) 11, p. 2873. [6] Reisinger, M., Doctorate Thesis, RWTH Aachen (Germany) 1995. [71 Viniegra, M., Gomez, R., Gonzalez, R.D., J. Catal. 111 (1988) 429. [8] van der Steen, P.J., Scholten, J.J.F., Appl. Catal. 58 (1990) 281. [9] Patzlaff, J., Gaube, J., Chem.-lng.-Tech. 69 (1997) 10, p. 1462. [10] Don, J.A., Schollen, J.J.F., Faraday Discuss. Chem. Sot. 72 (1982) p. 145. [11] Struyk, J., Scholten, J. J.F., AppI. Catal. 62 (1990) p. 151. [12] Dietzsch, E., Honicke D., Proc. 7h DGMK Conf. 1999 Oct. 13-15, Erlangen/Germany, to be submitted. 230 .> . . DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 .. .. ,~,,,. ,. ., ,, .-, , .,’,,” R. Fodisch 1),W. Reschetilowski 2),D. Honicke 1) 1)Chemnih I_fniverSi$rof Technology, Chemnitz, Germany z)Dresden University of Technology, Dresden, Germany ,. Heterogeneously catalyzed liquid-phase Hydrogenation of nitro-aromatics using Microchannel Reactors .,.. Abstract ,;..’ ,,,, -, ,., The objective of the present study was to develop a procedure and a microchannel reactor for ,. ,. the continuous hydrogenation of nitro aromatics in the liquid phase. As model compound .’,, n-nitrotoluene and as catalysts commercial pd Wpes were chosen. ~fferent waYs of edu~ ... ~osage were tested and th=ir applicability for feeding the microchannel reactor. The results obtained when using the microchannel reactor were compared to those when using a conventional f~ed bed reactor. Best results for both reactor types were achieved when they operated as loop reactors. f Introduction .’ --’ .’,’’.:!,.. “.-l Nitro aromatics are important chemicals among aromatics. They are major ..,.’: “: : ,,..,’. intermediates for the production of aniline and its derivatives. These compounds are ,, ..’,:, ,., ,,. , ,,’~ . . of great importance for the production of fine chemicals such as pharmaceuticals. ,., ,.’ . .. ,: ., There is a wide variety of catalyzed hydrogenation processes for nitro aromatics. The ,,. ,, ,., :-, ‘; simple compounds nitrobenzene and nitrotoluenes mainly are continuously .’, ,. -. >..,,.::/.:;. ‘.. ,.. ,,-..,,”... hydrogenated in the gas phase [1-3]. More complex compounds, e.g. higher alkylated ,.,., .,:. . ,ff,,’. and functionalized nitro aromatics are hydrogenated in liquid phase batch processes. This can be done using molecular hydrogen [4-7] or hydrogen donors [8,9] which can be e.g. complex hydrides [10],triethylsilane[11] and 2-propanol [12]. A major difficulty of these hydrogenation processes is the high reaction enthalpy of about 550 kJ per mol hydrogenated nitro group [13]. Therefore, the aim of this study was the development of both a microchannel reactor and a procedure for the mntinuous liquid ., -, phase hydrogenation of nitro aromatics. The major advantages of microchannel reactors for this exothermic reaction are the excellent heat and mass transfer 231 t . . —. . ,., ;, :..,- j: ;,:, ., .;;-. :.,. !,. . .,:;:,:;, - ,,, \ , .,.,--.,- ,, properties. For the performance of such a reaction hydrogenation of p-nitrotoluene with both molecular hydrogen and hydrogen donors, e.g. cyclohexene has been considered, The utilization of molecular hydrogen usually requires working under high pressure. Security considerations relating to this procedure can be avoided using hydrogen donors in the so called hydrogen transfer catalysis. One major drawback of the latter is the costly separation of the desired product from the co-product formed from the donor molecules and its expenditure. In contrast, using this procedure a high selectivity towards the desired product can be achieved even if the educt molecule comprises other unsaturated groups. 2 Experimental The catalysts used for the hydrogenation were commercially available. For hydrogenation with molecular hydrogen o,s~o Pdly-A1203 and with hydrogen donors Pal/coal and Pd black were used. As solvents in this study 2-propanol and ethanol were used for the experiments using hydrogen and donor molecules, respectively. The corresponding temperatures of hydrogenation were 363 and 351 K. The apparatus used for the hydrogenation experiments with molecular hydrogen was a standard flow type suitable to pressures up to 50 bar. As reactors both a fixed bed and a microchannel reactor were used. The catalyst fixed bed had a diameter of 1 cm and a height of 7 cm. The microchannel reactor contained a stack of microstructure aiuminium wafers having a total of 140 parallel channels (0,3 x 0,3 x 70 mm3). The surface of these channels was anodically oxidized and impregnated with 0,5% Pd. The liquid educt solution was fed with a HPLC pump and hydrogen with a flow controller. The reaction products consisting of non-converted hydrogen and the liquid product mixture were released through another flow controller. The pressure inside the apparatus was kept constant by controlling the product flow. Characterization of the reaction products was done off-line by gas chromatography using a 5890 Series II plus chromatography from Hewlett Packard and a 30m DB-wax column, isothermal at 180”C. High time resolution analysis was done by on line mass spectrometry with a mass spectrometer DSMS from Hiden. The hydrogenation with hydrogen donors was done batch-wise in a flask under reflux at normal pressure. 4. ,. 232 .::.. .. . !, ;..’. ,., 3 Results and discussion ,,, ,, As a model educt p-nitrotoluene was chosen in order to avoid any selectivity problems whereby the development of the procedure for the continuous hydrogenation in the ,., ,. liquid phase was not complicated. From that, the expected selectivity towards the ‘?, . desired product p-toluidine should be above 99% provided that no C-C-bond breaking >., ” takes place. The first experiments with gaseous hydrogen were conducted in the fixed bed reactor which operated first as a trickle bed reactor with the liquid educt solution flowing from the top to the bottom and the hydrogen flowing from the bottom to the top and afterwards as a plug flow reactor with both of the reactants flowing from the bottom to the top. In the latter case a thoroughly mixing within the liquid phase was assisted by the gas flow. The results of the hydrogenation of both experiments are ., shown in figure 1. The yield to p-toluidine using the plug flow configuration is much ,$,,., ‘, ,, ,, .! .“. .<, .,’,,,, t 1 1 I 1 1 I I 1 I 1 0 102OYJ4O5JJ6O 705JJ9JIC0 Vield p-toluidine ~A] Figure 1: Yield of p-toluidine by varying the fixed bed reactor configuration. p=40 ban LHSV=I O h-l; threefold .. hydrogen amount related to the stoichiometric need i. >,, higher than when using the trickle bed configuration. Therefore, the following ,. . hydrogenation experiments in the fixed bed reactor were done using plug flow .,,.,. , , ,, configuration. The selectivity towards p-toluidine was indeed above 9$1,S% so that the given yields correspond to the conversion degrees. Hydrogen supply for these ., experiments was three times of the stoichiometric requirements. The volumetric ratio ,., ,.~..; ,-,,’,>::.,’, of the liquid flow to the gas flow under process conditions was 1:2. Using the ,..“>,.:,.$.l;.,,,! ,, ., ... ,, ,/..., ,,,.l ,,. “ ,.; ‘,’.’; ‘,, , ,,’ ,,:2;; f .!;-,, gas ~quid “ gas — ,, :,, , ,.; ,., ,, ,, l-a .,’ , , ,.,. ,. ,,,: .,, :,’, ,, ,. ,,,j,.,,:< 4j,- ,,4 ..,; ..4 .,,’ ;., ,.,. .,,.,..... :..) Figure 2: Scheme of the desired liquid and gas flow in a ,,, .- ,,. ..::. ,,’.,”,’ microchannel of the microchannel reactor when applying ,.,, : -,’.;,:.~,,,~.,-;:,$, ,,? ,,,,,,,,,J,,,>; ,,.’ , ‘. ‘J’”1 three times of the stoichiometric hydrogen supply at 40 bar ,,,:,,+<,’,,’.} ; , ,.’. .. . ‘“‘ , .JJ.;I.I., , ,! ,. (:,:,...1, ,,:,.< , :..”’, ,,,,’:{p~ !;;t;? !, . . . , .,. ~,.,,~.,,.r..!.,...:,,.”,;(J,,,,4 233 :{ , ,,}-?~:,, , ;;,j ~.,.-’,.,.,<.,., .,,’ . ,,, :,, : ; ,,:!’,#Iz f;; ,;!;;! ... . ‘lylj,;,,,’; ‘c<:&r ..>, “,.,,.,~.,{: ,:.~~,!, ,!. ,:’, .:,,>$,, ,-:4,..,.., ,,..~.-. T - :- liquid. As a consequence, an insufficient amount of hydrogen will be available to the chemical reaction. Therefore, in the following experiments only one third of the hydrogen used in the former hydrogenation was applied corresponding to the stoichiometric amount. The gaseous and the liquid phase were delivered simultaneously. The hydrogenation results using the fixed bed reactor are depicted in I— ‘ # / 50 / 1 45 20 25 w 35 40 45 Pressure [bar] Figure 3: Yield of p-toluidine as a function of pressure. LHSV=I Oh-l; stoichiometric amount of hydrogen figure 3 and show an increase of the yield with increasing pressure up to 30 bar and then reach an upper yield limit of about 64%. From that we concluded that the major problem here was the mass transport from the gaseous to the liquid phase. The volumetric ratio of the liquid flow to the gas flow under process conditions was 1:1.3 at 20 bar and 1:0.65 at 40 bar. By the use of the microchannel reactor the distribution of ‘as—~=j =’s JS-- Figure 4: Scheme of the liquid and gas flow in a micro- channel of the microchannel reactor when applying stoichiometric hydrogen supply and 30 bar 234 :.4, . : ,,. the fluidic phases would be as illustrated in figure 4. Under this conditions it is much less likely that part of the microchannels are completely filled with gas and thus, ,. realizing this procedure in the microchannel reactor is probably successful feasible. .,.,~- ., Thus, the hydrogenation was carried out in the microchannel and the fixed bed reactor . .,,\ under the descriped and identical conditions. The results are depicted in figure 5. As .. ,,! .’, 1; 20- ,, ..-. murochannel reactor ... 28- ,, ~ . ~ 26- .G .5‘a ~ 24 & ,Q > I 20- A ‘ 4) ,. 18- 20 25 30 35 Pressure [bar] Figure 5: Yield of p-toluidine as a function of the total .. ,. pressure in both the microchannel reactor and the fixed bed ,., reactor used in plug-flow-mode. LHSV=60 h-l; ,,, -, stoichiometric amount of hydrogen shown, the yield obtained in the microchannel reactor was somewhat higher than the yield obtained in the fixed bed reactor. This means the microchannel reactor is suitable for the liquid phase hydrogenation of nitro aromatics under the applied :,,, conditions. ;, To avoid the described mass transport limitations (fig. 3) we did additionally two experiments using different homogeneously solved hydrogen donors instead of molecular hydrogen. The results conducted batch-wise in a flask are shown in table 1. Two donor molecules and &vo different catalysts were tested. The advantage of this procedure was the absence of a gas-liquid phase boundary which means that it would ,P,’.{ ~ ,: .,, >.:., ..,, ,,; , be easily to perform these experiments using a microchannel reactor. However, the ., ... J ,, j.,;,. ,, ..,’ ,,.; 7.,,”-, ,:,....,,~ ‘ “j; major drawback was the much higher reaction time of several hours compared to some . ‘”. ,- %:4”) -, ,. r.! ; ,, .: ‘,’.~.’. few minutes for the experiments with molecular hydrogen. That is why no further ., , ..c<.,eg,.:,?,c.. -: Fey,+,,..,....<;_. “, ,< .,-:, ,,. ,,, . . . :, +J.t.+G~,a ..’,.:: ‘,. , -! f,, .;: 7; \ .::.:<,h- ;: ~~,.;;- I : , , <{;;;; 0::[{ : ‘.<: L-,., : .“” ‘“’, !:-.~”:$-~.”... ,,::, .> =V , .>, < ,., , ‘, : ,;’: ~., .’.:,..: .>.. ,, . ,.;;y:. !:l+ *.: ,,, ...... , ..-:..:..’”,~z...... , .-..>. ~...... * ~~ ‘.s hydrazine Pal/coal 2 27 cyclohexene Pd black 17 14 Finally, we tried to increase the mass transpofl from the gas phase to the liquid phase in the fixed bed reactor by recycling the liquid phase in a loop. The reactor worked as a plug flow reactor with both of the reactants flowing from the top to the bottom. For the folIowing experiment six seventh of the liquid outlet of the reactor were recycled to the input of the reactor. The reactor throughput was kept constant at LHSV=I O h 1. Applying a pressure of 10 bar and stoichiometric hydrogen supply a yield of 46?L0was achieved. This is almost the same yield as obtained in the fixed bed reactor experiments without recycling using twice the pressure as depicted in figure 3. The volumetric ratio of the liquid flow to the gas flow under process conditions within the reactor was 1:0.37. Again, by the use of the microchannel reactor the channels would be filled with gaseous hydrogen only to a minor extend as illustrated in figure 6. At the .- Figure 6: Scheme of desired liquid and gas flow in a micro- channel of a microchannel reactor when recycling the liquid phase in a loop outlet of the reactor would be a continuous liquid flow, provided that hydrogen is entirely consumed due to the chemical reaction. Therefore, applying this procedure to a microchannel reactor is probably successful feasible. Corresponding experiments using the microchannei reactor showed indeed almost the same yield as obtained when using the fixed bed reactor. 4 Conclusions The liquid phase hydrogenation of nitro aromatics to the corresponding aniline derivatives is performed industrially mainly as a batch process. In our study we tried to develop a procedure to perform this reaction continuously. Therefore, we used several 236 . ..,,..’’!.. :’,’ ,..i..:....,”., ,,. procedures. The use of a fixed bed reactor as a plug flow reactor and applying $, ,. .,. -’ ,.-, hydrogen in excess was very suitable. Almost complete conversion could be achieved in a short reaction time of only a few minutes. Disadvantageous was the high pressure needed and it would be difficult to perform these experiments using a microchannel reactor because of the stight volumetric ratio of the liquid flow to the gas flow under process conditions. The afterwards performed experiments with stoichiometric amounts of hydrogen showed that due to the difficulties of mass transport from the gas J phase to the liquid phase it was not possible to achieve high yields. On the other hand :!. . under similiar conditions using a microchannel reactor the same or even somewhat “, higher yield could be obtained than using the fixed bed reactor. To avoid the mass :, :, transport limitations the use of hydrogen donors was taken into consideration. This ‘ ?,’ had some advantages, viz. the reaction mixture was homogeneous and working under ,, normal pressure was possible. By use of a flask as a batch reactor the long reaction ,., time of several hours was the major drawback. By operating the fixed bed reactor as a loop reactor the pressure could be decreased without loss of yield. This is of .i importance because of security considerations. Furthermore, this procedure could be ,,, ., successfully applied to the microchannel reactor. The yield obtained by using this .“ .! reactor was similiar to the yield obtained by using the fixed bed reactor. Major -.., ,., , advantages of microchannel reactors for this reaction are their excellent heat and .,. mass transport properties. The long reaction time of typical industrial hydrogenation ,. ,, processes is caused by the very high reaction enthalpy of about 550 kJ per mol hydrogenated nitro group. An uncontrolled increase of the temperature during the , ~ reaction often results in a loss of selectivity. Application of microchannel reactors effectively increases the heat removal whereby reducing temperature gradients was achieved. This allows a pronounced diminution of the reaction time down to some few ,. minutes. 4 Acknowledgements . .. :., We thank the Sachsisches Staatsministerium for Wissenschaft und Kunst and the ,, Fends der Chemischen Industrie for financial support. ,. .,.’ ., 237 [1] R. Langer, H.-J. Buysch, P. Wagner, U. Pentling, EP 0696573, 1996 [2] R. Langer, H.-J. Buysch, U. Pentling, EP 0748789, 1996 ,“ . [3] R. Langer, H.-J. Buysch, U. Pentling, EP 0748790, 1996 6,. . [4] L. CervenK 1.Paseka; V. StuchlK V. Ruzicka; Collection Czechoslovak Chem. Commun. 1982, 47(3), S. 853-857 [5] B. Yang; Z. K. Yu; Y. Xu; S. J. Liaq D. R. Yu, Chinese Chemical Letters 1996, 7(7), S. 663-664 [6] V. M. Belousov; T. A. Palchevskaya; L. V. Bogutskaya; L. A. Zyuzya, J. Mol. Cata/. 1990, 60(2), S. 165-172 [7] M. A. A. Lopidana; V. B. Bolos; C. J. Sanchidrian; J. M. M. Rubio; F. R. Luque, Bull. Chem. Sot. Jpn. 1987, 60(9), S. 3415-3419 [8] J. W. Chen; C. Q. Qin, Reactive and Functional Polymers 1992, 16(3), S. 287-295 [9] M. Burli; T. Buhlmann; B. Troxler, Speciality Chemicals 1993,13(6), S. 346-348 [10] S. F. Pan; P. D. Ren; T. W. Dong, Chinese Chemical Letters 1996, 7(11), S. 981-983 [11] H. R. Brinkman; W. H. Miles; M. D. Hilborn; M. C. Smith, Synth. Commun. 1996, 26(5), S. 973-980 [12] T. T. Upadhya; S. P. Katdare; D. P. Sabde; V. Ramaswam~ A. Sudalai, Chem. Commun. 1997, S. 1119-1120 [13] T. E. Daubert, R. P. Danner, Physical and Thermodynamical Properties of Pure Chemicals, Hemisphere Publishing Corporation 1989 238 DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erfangen 1999 M. Weber Phenolchemie GmbH, Gladbeck, Germany Safety of large Reactors for the Cumene Oxidation Introduction Worldwide most of the phenol is produced by oxidation of cumerrc to curnenehydroperoxide (CHP), which is in a second step decomposed under acidic conditions to phenol and acetone. The cumene oxidation is a very slow exothennic reaction so that usuaIIy kmgebubble columns are used as reactors in which cumene is oxidwd by oxygen horn air. Typical operating conditions include pressures rang-brgfrom atmospheric to about 700 kPa and tempemtures in the 80- 120°C range, the fmaI CHP-concentrations range from 20 to 40 ‘Y. /11. The liquid phase in the bubble column is backmixed so that the temperature can be very well controlled by intcmal or external heat exchangers/2/. In case of a failure of the compressors for areation the liquid phase in the reactors is no longer mixed. The CHP decomposes thermally with heat generation and it is important to make sure that there is sufficient heat removal to decrease the temperature and so to avoid a runaway reaction. Some form of auxiliary nitrogen should be available in case of such a fdure of the air compressors to improve the liquid mixiig for a better heat removal by for example external heat exchangers /2/. It is important to know to which temperature a reactor has to be cooled down to avoid a subsequent heating up. At th~ temperature the heat which is generated by thermal decomposition of CHP is equal to the heat which is removed by heat losses tlom the reactor. This equilibrium femperafure depends of course on weather conditions and especially on the intensity of tiee thermal convection in the non-mixed reactor which has to be taken into account in every safety concept for cumene oxidizers. In addition to our practical experience we made some calculations with Compufafional Fluid Dynamics (CFD) to describe the free thermal convection in our reactors for cumene oxidation. Experimental and Calculations The thermal decomposition of CHP is a reaction of first order d CHP I dt = - kcllP . CHP (1) We determined the temperature dependency of kCIWvery accurately for temPcmtures above 70”C. OGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 239 “~ogether with the heat evaluation of the reaction (270 KJ/Mel) onc can calculate for example the adiabatic temperature arise of a certain CHP-solution as it is done in Fig. 1. “~hiscurve shows that in a Cl IP-in-cumenc solution with 23 ‘%0 CHP having a starting tcmpcraturc of 92.9°C a runaway reaction will occur allcr about 20 hours if there is no heat removal. Actually there arc heat losses caused by tbc heat transfer through the cylindrical wall of the reactor, sce Fig.2. The overall heat transfer coeflicicnt k depends on the outer heat transfer cocflicient ~ the heat conduction through the wall I.w / SWand the inner heat transfer cocl?icicnt a,, which is here of special intrcst as a, dcscribcs the free convection in the non- mixcd reactor with heat sources llk=lla, +slk~ + Ilao. (2) During a shut down we Iookcd at the tempcratorc dcvclopmcnt of one of our reactors (diameter 4.6 m) having a starting temperature of 92.9°C and a CHP-concentration of 23 %, so the same starting conditions as for tbc calculation shown in Fig.1. The reactor wasn’t mixed neither by air or nitrogen nor by circulation pumps. Tbc outer tcmpcraturc To was 25”C, it was slightly windy. Ifig. 3 shows the temperature development over time. T, is the mcm temperature of the CHP-sohrtion, calculated Ilom seven measuring points distributed in the reactor. T, was dccreascd of about 0.8 “C in 180 min. With this cooling rate onc can calculate tbc overall heat transfer coeflicicnt k with tbc following equation: k= Q/( A. AT~) . (3) A is the surface of the wetted cylindrical wall (275 mz) and AT~ is the mean temperature difference, which is in this case #,. . AT~ = (T,,~~. + T,,Cd)/ 2 -T. = 67.5 “C . (4) Q is the rate ofhcat tmnsfcr, which is calculated by Q = m.co. AT, /At + Q, = 125000 W (5) with QK as the heat generated by the cxothcrmic decomposition of Cl 11’(23 O/.) at this tcmpcraturc Icvcl. Eq.(3) finally gives k = 6.7 W /(m ‘.K) 240 The outer heat transfer cocl%cicnt at a wind velocity of 4 rds can be estimated as /3/ u+=1O W/(m Z-K) The heat conduction through the wall doesn’t determine the overall heat transfer coefficient ,. ”-’:., , because I.w/ sw is much more higher thank. So eq.(2) gives .“ ai =20 W/(m2. K) which is the average heat transfer coefficient for the tlee thermal convection in this type of reactor under the above mentioned conditions. The calculations with Computational Fluid Dynamics for th~ case gives ai = 18 W /(m’. K), which is in a very good agreement to the experimental result. To calculate now equilibrium temperatures for different CHP-concentrations the following ‘1, ’,” assumptions are made 3 u is in case of calm weather in minimum 7 W / (m’. K), /3/, 2 the outer temperature is 25°C (summertime in Europe), ~, 2 a, is for the given reactor geometry 20 W/ (m 2. K), ,’.,,, so .,. , .,,-,,, k=5 W/(m*. K). Whh this assumptions for example the following equilibrium /empera/ure.s are calculated: ,, 96°C for 20 ‘7. CHP-solution and 87°C for 35% CHP-solution. Conclusion ,’,:,:,,, The fiec thermal convection at elevated temperatures in a non-mixed cumene oxidizer of .’, ,, ,’,{ ,, production size was measured as an average heat transfer coefficient ai. The results are fitting ,, -,,,,,-i, !. .,’,, very well to CFD calculations. By taking into account the weather conditions (a,) the overall ,’ :,:,:,:.,,,:j, heat transfer coeilicient k and so depending on CHP-concentrations the equilibrium ~ .‘!,,, , <, q ., , .,’ fcmpm-afurcof a “stillstanding” cumene oxidizer can be calculated. , ., , .,j ,,. , ,, 241 .. :.;,-- -,- >.-,,- :“’i-.: ,.’-: <’,’ ,..$. .’, . ,.- ,... ‘. Acknowledgcmcnts The author gratefully acknowledges support by Dr. Hcntschel of Degussa- HulsJVcrfahrcnstcchnik, who made the experiments for determining the kinetics of thcrrmd CHP-decomposition and by Dr. Birtigh of Degussa-Htilflcrfahrenstechnik, who made the CFD-calculations on FLUENT 5.0. # ‘. Refercnccs Ill Encyclopedia of Chemical Processing and Design, Volume 35, p. 373-391 (Editor J. J. McKetta) 121 Safety in Phenol-Cumene Process Ilydrocarbon Processing, Jan. 1976, p. 185-196 131 W. F. Cammerer: Wiirrne- und KWeschutz Springer-Verlag, 5. Auflage 1995 . 242 ,. ‘s . ,,,-, u) N ,,, :’.,.. ,. .,’;.., :, .,, .,. ., ,,,. ..- ,. “.... ., o .,, ,,’ ol : ,, In o ,., ;,! ,,.,, ,- ,.-,.-. ., .,,,’, ,.,,‘,,’ .,. , .,... ., ...... ,,,;,:.,,:~ “.’~,... f, .’.., :, .,.,,’ ?, ,., ,. ., . In -, ~ ‘,:;,,,.’,,,, ,,, ,,. y,, ., ,, :,,. .,,,.,,,, .,.,.,.,‘.?+ , ,,,.,+- , ,’,~,,,,., , t , o n w x II + + w w , 244 r ,.< .,,. ,,, ,. .,,’ !, ~., ~;. , , -,‘; ,,’ ,, , ,,. >::” ;! ; .. :’ ,,, ,,,. ., . ~,,-,’ : J ‘ ‘. ,, ‘- ., .,’ ., .,,’ :,, ., -, . ; ,, ,. !,., ■ ■ ., ... ,. ,, -’ ■ ,.. ,,. , ,,,, ,. 9 ,“,’ :,’...,, ,., ,: ., .. .,,. {’, .:.,,,’ ,,, .. ,.. ,: .“-, , ... - ,,. . ,,. ,,$,,’,,:,.:;, ., ,:, . . ‘. ,,,..:,\, “,;.. ,,:, ,., ,’, ., .,. ,, -,--- .,-.,, .,’ .,” ...... ’. :. ,.: .’, ..,, (%) U ameladma~ ,, , :’,: ;:’ ~ 245 ..::.:’ ;:; :,, ,, ,’ ... ’’~,,.’, ?,,. ., ~-, :. 246 4 . . ,-.,e.~ . ..” ‘...... -. . -, . . .. ,“. - . . . . , DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Etlangen 1999 ,,,- .. J. Ackerrnann, E. Klemm, G. Emig , !, Friedrich-Alexander-University of Erlangen-Ntknberg, Institute of Technical Chemistry 1,Erlangen, Germany , ,- ,,, {,:: ,..,,, . , ,- Base-catalyzed Synthesis of Ethylbenzene by Dehydroisomerization of 4- ;. ,, , .,-. ,,. , Vinylcyclohexene , ,,,.’/ , ,, . ,. .,’ !;: ,“, ~: ,,. , ,., ,, ,, ,,. .,, , ,.,, ‘,, .,,:,, 1. INTRODUCTION ,,: ,...‘:’.’:”, -,;:.,,. . The base-catalyzed vapour phase dehydroisomerization of 4-vinylcyclohexene (4-VCH) is ..- .,- one important step in the synthesis of ethylbenzene (EB) from 1,3-butadiene. This new route .,’, .1 ., ., ,, , could probably be a useful way for the economic production of ethylbenzene, especially by ,; :,- : considering the growing butadiene surplus [1]. The parent compound 4-vinylcyclohexene can ,. ..,,,’ : ~. “!,,,,.,,>;; be produced in high yield from 1,3-butadiene by cyclodimerization using e.g. homogeneous ., ,, ,. catalysts [2] or CU(I)-Y zeolite in the liquid phase [3]. .,o, ,, f.! ‘,, The aromatization of 4-vinylcyclohexene runs via an isomerization with subsequent ,.. 4, ‘., ,; .4 dehydrogenation of the intermediately formed isomers and is known to be catalyzed by solid :.. ,,. ‘. ; . . ,, ..,..; f bases like e.g. supported pure alkali metals [4]. Since such catalysts are very sensitive to the ,. ,’, , ,.1 contact with water and oxygen, we focused on the use of the following other well-known ,., , basic solid materials . MgO supported on Y-A1203; ● Mg-Al mixed oxides obtained by thermal decomposition of hydrotalcites ● FAU-type zeolites modified by alkali metal ion-exchange or impregnation with alkali metal oxides. ... .’ Alumina-supported magnesium oxide was chosen because it is a typical representative for ‘., . a catalyst having a certain degree of surface area and providing basic properties. The active sites are represented by structure defects of MgO resulting in low coordinated 02’ in the lattice. Additionally, its acid-base characteristics are well studied over a broad range of varied MgO/AIZO~-ratios [5]. Materials with the same chemical composition and also possessing basic properties can be obtained in a totally different way by calcination of synthetic hydrotalcite-type anionic clays. Hydrotalcites are layered mixed Mg-Al hydroxycarbonates with the general formula [Mg,JJW(OH)2](C0,)WmmH20 (with Os ws 0,4 and Os m S 5). The structure consists of : ,,, ,, positively charged brucite-like (Mg(OH)z) layers, with A13+substituting some Mg2’, separated ?>” ‘,., by intertayer carbonate anions for charge compensation and water molecules. Thermal !,:. . .. treatment transforms these materials to interdisperaed high surface area Mg-Al mixed oxides : >“! with the structure of MgO. Their unique basic properties induced by OH, Mg-O pairs and 02- .:,’.! can be modified e.g. by varying the Mg/Al-ratio and the activation conditions [61. ., ‘, ‘.. . .: -,:’,:,‘:,,, %; Beside the wid~spfiad ap~cation-of zeolites as Bwmsted or Lewis acidic-catalysts, they ,,“,.. ,,. , , ,.,,.,!$., , -:.‘ also can act as solid bases ~]. Especially FAU-type zeolites exhibit framework oxygen sites ; ! ..;.,,,, ! ,, ,,,, ,. ,,! with Lewis-basic properties. Generally, this so-called intrinsic (structural) Lewis basicity .,,, ,f, ‘, ... , .! .:;,,j~. ,., ~ increases with decreasing Si/Al-ratio of the framework. Thus, X-type are more basic than Y- ...’., type zeolites. It is important to mention that basic sites always exist in connection with acid q~: : ,:’{ .::~ sites which consist of exchangeable countercations for charge compensation. These Lewis- .t..c.l .,, , , ,., L.,,,, ‘,,!, acid/base characteristics can be controlled by variation of nature and number of extra- ,. .,,:! ,,, ., . OGMK-Tagungsbericht 9903, ISBN 3-931S50-59-5 247 framework cations. The basic strength of the zeolite increases as the electronegativity of the cation decreases. Therefore, the Cs-form offers stronger basicity than the Na-form. Furthermore, intensification of the basic character can be achieved by encapsulation of alkali metal oxides [8]. The aim of our work was to get more knowledge about the structure-reactivity relations of the above described fairly different basic catalysts particularly on condition that both hydrotalcites and zeolites were not described for this reaction in literature as we know so far. Additionally, we tried to use computer simulations in the sense of molecular modeling to improve our knowledge on structure-reactivity relations in zeolites, especially with regard to sorption phenomena. 2. EXPERIMENTAL The alumina-supported MgO catalysts were prepared by wetness impregnation of Y-A1203 (from Akzo) with magnesium nitrate followed by evaporation of the suspension and sub- sequent calcination at 500 “C for 24 h in air [9]. Three samples with a MgO content of 50, 75 and 95 mol-% were obtained (named MgO-50, MgO-75, MgO-95). Mg-Al hydrotalcites were prepared by coprecipitation of magnesium and akrminium nitrates with a highly basic carbonate solution [10]. The first 1 M solution contained Mg and Al dissolved in destined water with nominal molar Mg/Al-ratios of 0,5 / 1,2 / 3,0 corresponding . to a MgO content in the final catalysts of 50/70/ 85,7 mol-% (named HT-50, HT-70, HT-85,7). The second solution with equal volume contained KOH and KZC03 in the following ratios: C032-/(A13*+Mg2+)= 0,48 and OH/(A13’+Mg2+) = 1,62. These two solutions were slowly mixed (feed rate: 1 ml/min) under vigorous stirring. Afterwards the mixture was aged at 45 “C for 15 h under slow stirring. The precipitate was washed several times until the pH of the rinsing water was neutral. After drying at 110 “C over night, calcination was performed at temperatures between 400 and 550 “C for 1 h in ftowing nitrogen. The thermal decom- position of Mg-Al hydrotalcites during calcination was studied by TG and DSC running with a heating rate of 10 OC/min. Commercial FAU-type zeolites with different Si/Al-ratios were used in their sodium form as starting material (NaY. Si/Al = 2,69 from Grac~ NaX Si/Al = 1,26 from Fh.rka). Cesium ion- exchange was carried out according to the corresponding ion-exchange isotherm [1 1]. Following the method of HATHAWAYANDDAVIS [12], NaX was threefold exchanged with 0,1 N solution of cesium chloride at 75 “C for 3 h. The sample (named CsNaX) was filtered and washed chlorine free afler each exchange. Finally, the catalyst was dried at 100 “C in air. According to MiPIWs ET AL. [13], alkali metal oxide added zeolites NaX were prepared by impregnation with an aqueous solution of sodium acetate or cesium hydroxide of appropriate concentration and subsequent calcination at 550 “C (heating rate: 1 OC/min) for 1 h in air. Different loadings of n = 4-5 additional sodium or cesium atoms per super cage (corresponding to 8m additional sodium or cesium atoms per unit cell) were obtained by this method (named NaX+4CsOH, NaX+5NaOAc). Catalytic measurements were carried out using an isothermal fixed-bed reactor (inner diameter 20 mm, length: 340 mm) operated at atmospheric pressure. The catalyst screening was performed under the following standard reaction conditions: T = 400 “C, W/F(Ot = 800 (g.min)/mol, m=, = 2 g, pn,tiwen:pd.vcH= 9:1, dwn,d, = 1,1-2 mm. Analysis of organics was done by on-line gas chromatography equipped with a FID-detector. Prior to use, catalysts were activated in flowing nitrogen at 450 “C for 1 h. The textural data (BET surface area and micropore volume) of the catalysts were measured by N2-physisorption. For determining the chemical composition XRF or ICP-AES were used. The compositions of impregnated zeolites and alumina-supported MgO were calculated from their synthesis mixture. Crystal structures were determined by XRD. Computer simulations were performed on the basis of molecular modeting techniques with the commercial software Cerius2 (version 3.5, developed by MSI). 248 ., ’,,,,, >,., ,. :-.. ,. .. . .-,’ ,- -, 3. RESULTS AND DISCUSSION ,: , ,, 3.1. Characteristics of Catalysts .,,.,,,,,.. . i .’. ,,;:, ,’ ,. .; The nomenclature as well as some physico-chemical characteristics of the different obtained mixed metal oxides are given in table 1. Table 1: Characteristics of the tested mixed metal oxide catalysts: nomenclature, textural ;.. data and chemical composition ‘, ..: surface area nominal molar nominal molar measured molar measured molar ., Mg/Al-ratio Mg/At-ratio A1/(Al+Mg)-ratio ,’ cataly5t MgO content ~1,’, [:yg] [mol-Y.] [-1 [-1 [-1 MqO-50 125 50 0,5 0,5 0,67 MqO-75 80 75 1,5 1,5 0,40 MgO-95 48 95 9,5 9,5 0,10 HT-50 329 50 0,5 0,4 0,71 HT-70 242 70 1,2 1,2 0,45 :. - HT-85,7 250 85,7 3,0 3,0 0,25 !, , ,,,.. Alumina-supported MgO catalysts show a decline of surface area with growing MgO .. content. The same behaviour applies to the calcined hydrotafcites the more the Mg/Al-ratio ,. decreases, the more high surface area mixed oxides are formed by the thermal Mg-Al ... ,. /, decomposition of hydrotalcite precursors. During the heat treatment two weight losses are ,. ,, :-,,, .’. ; obsewed for all samples. The TG/DSC curves of the stoichiome~”c hydrotalcite structure with ,, a Mg/Al-ratio of 3 (whose formula can be written as: Mg~Z(OH)1GCOx4HzO) are exemPlarY ,,, :, .,. ,,, ,,, shown in figure 1. Firs~ below 300 “C ,,,,.,,, ,“. , ,. .,, adsorbed interlayer water and crystal water are released without collapse of the structure. Then the dehydroxylation of the lattice and the decomposition of the carbonate anions accompanied by C02 evolution takes place. These two endothermic processes lead finally to the destruction of the layered structure. The XRD patterns of uncalcined samples showed the characteristic 50 lao 350 350 450 temperature [“C] reflections observed in typical hydro- Figure 1: TG (!) and DSC (2) curves of hydro- talcites and the reflection intensity decreased with rising Mg/Al-ratio. Heat talcite sample with M&Al= 3,0 .,, ,, treatment destroyed.,-—--—clearlv...—the Iavered ,<:~,.,,, . , structure and only slightly crystalline MgO phases were built, characterized by diffuse’XRD patterns. As a result of this calcination process an amorphous mixed oxide with a high degree of dislocations and different basic sites is created. The determined chemical compositions of calcined hydrotalcites are almost equivalent to the nominal values. For the ,, stoichiometric hydrotalcite structure the agreement is optimal. ,,, The characteristics of the zeolites are shown in table 2. The ion-exchanged sample shows the expected reduction in micropore volume due to the partial replacement of Na+ by the larger Cs+. The decrease in micropore volume with growth in loading of additional alkali metal cations reflects the successful formation of intrazeolitic metal oxide species. It is important to mention that the two impregnated samples have different loadings but are nearly equal in micropore volume. XRD measurements revealed that the framework structure was not destroyed by the post-modifications. 249 ~, —..-, . ,,,..: .,, ;~.>., - 1 :’.,: ,,.-’,, # ,. Table 2: Characteristics of the tested zeoiites: nomenclature, textural data and chemical composition micropore volume surface area Na content Cs content loading catalyst v~e n~n~ n~n~ mpvmmrlmz~ ,W [cm’/9l [:79] [-1 [-1 [wt.-%] NaY 0,32 660 1 0 0 NaX 0,25 526 1 0 0 CsNaX 0,19 399 0,38 0,62 0 NaX+4GOH 0,10 202 1 0,38 35,8 NaX+5NaOAc 0,09 191 1,47 0 24,5 3.2. Mixed Metal Oxides Catalytic Behaviour of Mg-Al Mixed Oxidas Prelimina~ experiments showed that the calcination temperature has a strong influence on the catalytic behaviour of the Mg-Al mixed oxides derived from hydrotalcites. An optimum relating to the selectivity to ethylbenzene was found at a temperature of 450 “C. Further- more, pure MgO (synthesized by thermal decomposition of Mg(OH)2) was found to be nearly inactive in this reaction. This might be due to the relatively low activation temperature of 550 ‘C. Higher activation temperatures above 600 “C where MgO is described to be activated for several base-catalyzed reactions [14] were not used because such high temperatures are not very interesting for an industrial application. The catalytic activities and selectivities of the Mg-Al mixed oxides which were obtained in the described different ways are illustrated in figure 2. All samples show no deactivation with- in the observed period of 6 h time-on-stream. Alumina supported magnesium oxides as well as calcined Mg-Al hydrotalcites show high activities; the conversion of 4-vinylcyclohexene increases with rising MgO content and reaches values of 100 ‘A at the highest Mg/Al-ratio. Concerning the selectivity to ethylbenzene both catalyst systems show the same trend. With growing proportion of MgO in the catalyst the selectivity to ethylbenzene rises from around 25 ‘A up to more than 80’70. On the other hand the selectivity to the main side- product p-xylene falls from initially 20 % down to only slightly detectable traces as the acidity offered by A120a decreases. The deficit in selectivity can be ascribed to the isomers of 100 Mgoconversion of yield of .- Catalwt Icontent 4-VCH EB I [mol-O/O]I [%] [“/0] ~ 0 ~ 40 50 60 70 80 90 100 nominalMgO content[mol-%] Figure 2: Catalytic activity and selectivity of different Mg-Al mixed oxide catalysts in the conversion of 4-vinylcyc/ohexene (after 200 min time-on-stream) as a function of the nominal molar MgO content: tilled symbols: MgO supporied on FA120,, open symbols: calcined Mg-Al hydrotalcites N O se/eclivity to ethy/berrzene, O 0 selectivity to p-xylene 250 4-vinylcyclohexene and its partially hydrogenated products. In comparison with the supported catalysts, calcined hydrotalcites offer the substantially better catalytic performance. By using them, comparatively higher yields ot ethylbenzene accompanied by a lower formation of p-xylene can be obtained with increasing Mg/Al-ratio. This might be explained by their higher surface area and the combined availability of more and probably different basic sites. Nevertheless, it seems that there exists a synergism between Al and Mg in hydrotalcite derived catalysts as well as in alumina-supported MgO. One remaining interesting question deals with the mechanism of the formation of p-xylene in this actually base-catalyzed reaction. Possibly the answer can be found in the mechanism ,,, , of the isomerization of ethylbenzene to p-xylene in the presence of hydrogen on bifunctional :,, ’-, ,, ... . . ,,, catalysts containing Pt and H+. For this, one can find the below mentioned reaction scheme r .,. , in literature (scheme 1) [15]. ... ,’: ,. .. ,. ? :: .,.., ,,. ,,, .. ... ~ ,. 0“ D ~,,-’,, ., ,- ,’,, ,,, :,, ..’.-, ,, Pt Pt ,. .,,, ‘: .,.. 11 II , ;,. ~:,’,,.-,,,”...... - . .:,.:,,,,.,...... ,. . ,P, ,,. .,,,,, , ,,.,.,. .,} “- ,.’ Scheme 1: Simplir7edscheme of the bifunctional (PHI-I’)isomerization of ethylbenzene to ., ,,,’, p-xylene ., :’ .,~.,,.. . .,.. . , ...,. ,,-..: ~ On Pt/y-A120s catalysts which provide weak acidity resulting from the Y-A1203 the ..... , ‘::/:. :. isomerization process starts with the Pt-catalyzed formation of partially hydrogenated ,,’. . ,-, .- .,, ;,-, intermediates, namely ethylcyclohexene. In the presence of Bnzmsted acid sites this one ., ,. \,& reacts via a ring contraction and expansion mechanism to p-xylene. Since in our reaction .--’ ; “4 ethylcyclohexene probably is produced by hydrogenation of 4-vinylcyclohexene, the formation of p-xylene can occur on accessible acid sites provided by A1203 following the described mechansim. The needed hydrogen results from the main reaction (viz. the 4 ,,,; I dehydroisomerization of 4-vinylcyclohexene). The existence of a precious metal component -,,.:...... ! /, .’..’ for dehydrogenation of the intermediately formed p-dimethytcyclohexene seems not to be .,,,, necessary as this reaction step can also proceed base-catalyzed. 4,. ,,, .,, ,,., .,. , ,, .:,.,,? ~ ,., ,,, .,, ,.’ ..; 3.3. Basic Zeolites Computer Simulation of the Sorption in Basic Zeolites Computer simulations on the basis of molecular modeling techniques were employed to study the sorption of reactants inside the intrazeofitic micropores at reaction conditions. First, a simulation method was developed considering: (1) Lowenstein’s rule, (2) energy mini- mization of the structure using the augmented CVFF force field [16], (3) calculation of the ?,’, charges on each atom of the sorbate molecules using the semiempirical method AM1 [17], (4) assuming both sorbate molecules and zeolitic framework to be rigid, (5) calculation of van-der-Waals energy using Demontis force tield [18], (6) calculation of long-range Coulomb interactions by an Ewald summation, (7) a Monte-Carlo method working at fixed pressure (grand canonical ensemble). The simulation was tested on the well studied problem of the adsorption of benzene in NaX. Because of the good agreement with experimental data like e.g. the isosteric heat of adsorption published by RUTHVENANDGODDARD [19], the simulation method was applied to the sorption of 4-viny lcyclohexene and ethylbenzene in NaX and CsNaX. !-,,., 251 The calculated values of the isosteric heats of adsorption AH,~~of reactants (which is a measure for the strength of adsorption) are shown in table 3. Obviously, on both samples ethylbenzene is stronger adsorbed than 4-vinylcyclohexene. Comparing NaX with CsNaX a weaker adsorption of 4-vinylcyclohexene and a stronger adsorption of ethylbenzene is found on the more basic sample. The latter can be explained by the stronger interaction between the aromatic x-electron system and the more electropositive extra-framework cesium cations [20]. Additionally, the value for AHA of ethylbenzene on NaX is nearly the same as the one measured by AUST ETAL. [21], showing also the good qualitiy of the simulation. Table 3: Lsosteric heats of adsorption AHad~of reactants in different zeolites calculated by molecular modeling (in kJ/mof) catalyst 4-vinylcyclohexene ethyl benzene NaX 65 73 ~NaX 59 81 Cata/@”c Behaviour of Basic Zeo/ites The catalytic activity and selectivity as well as the deactivation behaviour (observed within 6 h time-on-stream and given as relative loss of initial activity) of the tested basic zeolites are presented in table 4. On unmodified zeolites conversion as well as selectivity to ethylbenzene increases with decreasing Si/Al-ratio due to the growth of the intrinsic basicity of the framework. The main side-product p-xylene as well as all other side-products (the isomers of 4-viny lcyclohexene and traces of o-xylene) diminish with decreasing Si/Al-ratio. On all these samples deactivation occurs and gets stronger with decreasing basicity of the framework and the entailed increasing number of extra-framework cations which represent the sorption sites for ., aromatic systems. Compared to the unmodified sample, the Cs-exchanged X-type zeolite shows a reduction in conversion. This might be caused by the reduced micropore volume which affects diffusional problems and the lower aftinity for the sorption of the reactant 4-vinylcyclohexene as it is derived from the computer simulations (see table 3). Although CsNaX has actually an enhanced basicity of the framework, the selectivity to ethylbenzene is reduced while the formation of side-products is favoured (esp. for p-xylene). The reason for the formation of p-xylene is probably the same as described before. Bifunctional zeolites (e.g. Pt/HNaY) are also able to transform ethylbenzene via ethylcyclohexene to p-xylene [22]. The required acidity might result from small amounts of residual Brsmsted acid sites which are probably generated during preparation. Table 4: Catalytic activity and selectivity of various (unmodified and modified) FAU-type zeolites in the conversion of 4-vinylcyciohexene (after 200 min time-on-stream) conversion of selectivi~ to selectiviPf to yield of deactivation catalyst 4-VCH p-xylene (xt=6h-xl=H)/x(= 2m,n [%] [%] ;:] ;:] Na’i 57,9 4,6 30,8 17,8 -41,3 NaX 81,2 0,4 74,1 60,2 -22,9 ~NaX 46,0 7,5 63,6 29,3 -33,8 NaX+4CsOH 99,9 none 96,8 96,7 none NaX+5NaOAc 99,9 none 94,6 94,5 none 252 ! . ,., !.’ Figure3 shows the time-on-stream behaviour of NaX in comparison with CsNaX at different reaction temperatures. While on NaX conversion of 4-vinylcyclohexene remains constant on the high level of approx. 100 YO at 450 “C, significant deactivation occurs at ,,. , 400 “C and lower temperatures. The identical behaviour is observed on zeolite CSNSX. Comparing these two zeolites at the same temperature of 400 “C, the Cs-exchanged sample shows a lower initial activity and deactivation runs faster. Additional measurements on the same initial activity level (adjusted by varying residence time and keeping temperature constant) lead to the same result namely a stronger deactivation on CsNaX. As already ,, described, the computer simulation studies of the sorption indicate that a strong adsorption of ‘>., “.. , the product ethylbenzene might be responsible for the deactivation (see table 3) in the sense of product inhibition. To check this assumption, we performed some special experiments at .:,/ -,,, fixed 4-viny lcyclohexene concentration (10 vol.-%) ethylbenzene was added in different ,, .,.,,’ ratios to the feed as part of the inert (P.WCH:PEB = 2:1 / 1:1 / 1:2). The results in figure 4 show, ,. ~.’ that the more ethylbenzene is present, the stronger is the deactivation of the catalysts. The ., .-, loss of initial activity is ultimately higher on the Cs+ containing sample. .-, .,,., 100 Figure 3: Time-on-stream behaviour in %’ 2 ::: ...... the conversion of +vinvlcwb ::: , hexene on NsX and C;N~ x :;: :, w :: + NaX (400 “C, W/F= 800 gmin/mol) c 0 NaX (450 “C, W/F= 800gminlmol) P :.. El NaX (400 “C, W/F= 200 gmin/mol) g ;:: ;: c 20 - ...... !...... +...... ’.’’”:“’”’””+-’ ““’”” ● QNaX (400 “C, W/F= 800gmin/mol) ::, :, @NaX (450 “C, W/F= 800g.min/mol) ,:, ‘.. -, .,. ,, : o 50 100 150 200 350 300 350 ,f’,,. time-on-straam [mini ,, ,. <,!. . ,; .-. -’ ,,, ::: , ... .: ~ ,,..,,, . . -. ,: Figure 4: Influence of adding ethyl- ~ ., .,.,, benzene in different ratios ~ ...... ;...... :. ::! ;:. ,. .’,, ‘. to the feed forzeolite N& ~ :> ..,;,,:,,,;.’ , .,, ,,., . . ‘, and CsNaX . .. . ., 1? -~.. g ., t “ ..”.’ +0 P(4-VCH):P(E8)= 21 ,., ,’ ■ o p(4-vcH}p(EB) = U1 ,, ..,- ● ,..,’,:- 0 p(4-vcH):p(EB) = 1:2 O-!i:i:i: ,, o 50 100 150 200 Zso 300 350 .,... .,. . ; tima-on-3traam[rein] ,,7’:..,.,,,. ,., .,-..,.> ;.:.,:-,:! , [n comparison to unmodified and ion-exchanged samples, alkali metal oxide loaded X- :’. ,,. ! ...... ,“,~’, 8 . . . type zeolites show outstanding results (see table 4). They distinguish theirselves by high .:, .,;, -,. conversions of 4-vinylcyclohexene (= 100 ‘A) and simultaneous maximum selectivities to .’. :.’ ethylbenzene of295 ‘A, both existing over a broad range of varied reaction conditions. The formation of p-xylene is completely suppressed. In addition, no deactivation can be obsenred over a period of 6 h time-on-stream. The more basic cesium oxide containing sample shows a higher yield of ethylbenzene despite the lower loading of additional extra-framework basic oxygen sites. A remarkable effect is observed exclusively on this zeolite: during an initial period, toluene and cumene are produced in a 1:1 ratio (with a considerable quantity of initially up to 12 YO yield of each toluene and cumene) at the expense of ethylbenzene. Both the variation of residence time and the utilization of ethylbenzene as feed indicate the unusual base-catalyzed side-chain alkylation of ethylbenzene by itself. The reason for the suppression of this reaction with progressing time-on-stream remains still unclear. . . . . 253 Obviously, in contrast to Mg-Al mixed oxides impregnated zeolites show a significantly better catalytic performance. By using these catalysts and reducing the residence time to 100 (g.min)/mol, a space time yield of more than 5,5 kgd(kgw h) at still total conversion over 6 h time-on-stream can be obtained. Thus, they are possibly attractive for industrial application. 4. CONCLUSIONS In our investigations concerning the vapour phase synthesis of ethylbenzene from 4-vinyl- cyclohexene using basic solid catalysts which were not described in literature so far the following essential results were found: ● calcined Mg-Al hydrotalcites show a better catalytic performance than the corre- sponding alumina-supported MgO catalysts which agree with the state-of-the-art; on principle, growing Mg/Al-ratios give rise to beneficial effects on selectivity to ethylbenzene; ● faujasitic zeolites, especially those having a low Si/Al-ratio and being post-synthetically modified by impregnation with alkali metal salts, represent the most promising catalysts; ● applying computer simulations on the basis of molecular modeling techniques to study the sorption in zeolites at reaction conditions was (in combina~on with appropriate experiments) very helpful to demonstrate that deactivation on ion-exchanged samples results from product inhibition. REFERENCES [1] M.L. Morgaw DGMK Tagungsbericht 9705 (1997),9 [2] H.A.M. Duisters, J.G.D. Haenen; US-Patent No. 5545789 (1995) [3] LE. Maxwell, R.S. Downing, S.A.J. van Langen; J. Catal. 61 (1980), 485 R.S. Dlxit, C.B. Murchison; Proc. Ethylene Prod. Conf. 3 (1994), 118 [4] G. Ruckelshau13, K. Kosswig; Chem.-Z. 101 (1977)2, 103 [5] H. Noller, J.A. Lercher, H. vine~ Mat. Chem. Phys. 18 (1988), 577 [6] F. Cavani, F. Trifiro, A. Vaccari; Catal. Today 11 (1991), 173 [7] D. Barthomeu~ Catal. Rev. -Sci. Eng. 28 (1996), 521 [8] H. Tsuji, F. Yagi, H. Hattor~ Chem. Lett. (1991), 1881 [9] J.A. Lercher, C. Colombier, H. Vinek, H. Nolle~ Stud. Surf. Sci. Catal. 20 (1985),25 [10] Y.Z. Chen, C.M. Hwang, C.W. Liaw Appl. Catal. A General 169 (1998), 207 [11] H.S. Sherry; J. Phys. Chem. 70 (1966)4, 1158 [12] P.E. Hathaway, M.E. Davis; J. Catal. 116(1989), 263 [13] M. Lasperas, H. Cambon, D. Brunei, 1.Rodriguez, P. Genestq Micropor. Mater. 1 (1993), 343 ‘., [14] H. Hattori; Mat. Chem. Phys. 18 (1988), 533 [15] K.-H. Robschlager, E.G. Christoffel; Ind. Eng. Chem. Prod. Res. Dev. 18 (1979), 347 [16] P. Dauber-Osguthorpe, V.A. Roberts, D.J. Osguthorpe, J. Wolff, M. Genest, A.T. Haglec Proteins: Structure, Function and Genetics 4 (1988), 31 [17] M.J.S. Dewar, E.G. Zoebisch, E.F. Healey, J.J.P. Stewart; J. Am. Chem. Sot. 107 (1985), 3902 [18] P. Demontis, S. Yashonath, M.L. Klein; J. Phys. Chem. 93(1989), 5016 [19] D.M. Ruthven, M. Goddard; Zeolites 6 (1986), 275 [20] S.M. Auerbach, L.M. Bull, N.J. Henson, HJ. Metiu, A.K. Cheetham; J. Phys. Chem. 100 (1996), 5923 [21] E. Aust, W. Hilgert, G. Emig; Stud. Surf. Sci. Catal. 46 (1989), 495 [22] M. Nitta, P.A. Jacobq Stud. Surf. Sci. Catal. 5 (1980), 251 254 . .s . ,.. . ,- ,, .1 .,-1 !’ ..-, ,,, ,; ,, ’.,,, , ,;. - f.:.:!,’,,1 , ., .. ,. DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 ,; .;. ,/:Ji .,“,,,. , ,4 ,,, . G.!,,;! ,,, , ,:, .,. .“{,, ,;. ... ,,, A. Reitzmannl), G. Konigl), F.M. PetraV), E. Klemm’), G. Emig’) ,., . . OFriedrich-Alexander Lfniversi~ of Erlangen.Numberg, hlstitute OfTechnical ,., ,,; .: ‘;. ,, ,: ., ‘,., ,,.. ,, Chemistry 1,Erlangen, Germany .,,’:, ,, z)HU[slnfra~or GmbH, prUfinstitute fUr Analfik, Biologie und Toxikologie, ,,-+ ,’, ,,, ‘. ,h:, ,’., Marl, Germany ,1,.,,,.,,.’ .: /.!’”; ,, f’ ,,, ,., . Catalytic Behaviour of modified ZSM-5 Type Zeolites in the Hydroxylation of ~,. Benzene using nitrous Oxide .. ,- 1. Introduction Since the beginning of the 80ties the direct oxidation of benzene to phenol has attracted attention of several research groups. It has been found that the MFI type zeolites show considerable activity for selective hydroxylation of benzene to phenol using NZO - even relevant for industrial application [1]. It is believed that a monooxygen species with special characteristics generated by first step of NZO decomposition is able to oxidise benzene to Dhenol (Fig. 1) [1, 2]. The question about the nature of the active site(s) responsible for this phenomenon has not been answered, so far extrafrsmework iron clusters [1, .. ~H 6], extraframework aluminum species . ,!. ,’ [3] or Iewis acidic framework defects ‘‘ ,, N20 [4] and Bransted acid sites [5] are ,: :., assumed to be relevant. Remarkably, ,(, .,’ $, in almost all cases these special active sites have been achieved by cal- cination at higher temperatures [1, 3, 4, ,: 7]. But it is still under discussion how the properties of the zeolite influence its catalytic performance. N2 A further interesting and important - ~ )( question in benzene hydroxylation is onlv scarclv discussed in literature the Fig. 1 Principle of hydroxylation of benzene to phenol with NZO int&action - of the involved arvmatic hydrocarbons with the zeolite [8,9, 11]. -’ In this paper we try to demonstrate, firstly, how du~ation of ~Jcination at an oPtimum of ; temperature affects important physiochemical and catalytic properties of the ZSM5 zeolite. ,.,.,’ ,,, Secondly, it will be shown what influence the sorption processes can have on the behaviour ., .. -.’ of the process. For this reason NZO and benzene feed is introduced separatly which clarifias ,,. ., the important reaction steps on the catalyst surface. -,.,,,T, ,, 2. Experimental The parent catalyst sample used for the catalytic investigations was the NHd-fo~ of a ~.,’ zeolite of ZSM5 type (SM55, ALSI-PENTA). It possesses a modulus of 45, 0.05wFX0 of FezOa and a Na content lower than 0.03 wtOA Na*O (determined by XRF and in good correspondanca with company’s product specification). Protonic form was formed by calcination at 550”C for 3 hours. All additional samples were get by calcination in air at 900”C for different periods of time. All calcination procedures were performed in a tubular , 4, ~ ,,, :, . oven under flowing air. The heating rate was 5 KJmin. ,- ‘.,,.;, ,. ’:’, . OGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 255 :,< ,,.. (’, ,,, ,- ,.~,, ,, ,“, .<..~ —---- , .. ”,.,. -. :,, ,;,. : ,-. ,,:,’, -. w ,, .’. ,. -y ,.:’:: , .:.<,/:7 , .- The influence of calcination procedure on the catalysts was characterised by the following physiochemical methods % Microporevolume by Nz-physisorption. % Aluminum distribution by 27ALMAS NMR. % Volume of the unit cell by XRD. % Number of Bransted acidic sites by TPD of NH, (heating rate: 10tVmin) Hydroxylation of benzene was performed in a fully automated laborato~ setup containing a plug-flow reactor [8]. A micro gaschromatograph (Chrompack MicroGC 2002) was added to the conventional gaschromatograph to get information about the amount of coke calculated from the amount of COZ in the regeneration procedure. This method is justified because no carbon monoxide has bee found and treatment with air is immediately started afler the hydroxylation experiment to avoid loosing physisorbed hydrocarbons. Additionally, the Micro GC exhibits very short analysis times (COZ -25 s, NZO -30 s) and makes it possible to follow regeneration with high time resolution. Generally, two kinds of experiments were conducted: > Continuous flow experiments with .simultanous feed of nitrous oxide and benzene > Continuous flow experiments with alternating feed of nitrous oxide and benzene (Experimental procedure is described in Section 3.2) 3. Results and Discussion 3.1 Optimisation of the catalyst by calcination pretreatment The chosen parent ZSM5 sample (SM55, ALSI-PENTA) used for the following in- vestigation exhibits optimum characteristics: octahedral Al content < 2% (*’Al-NMOR), a very low content of FezO, (0.05wW by XRF), high crystallinity and no phase impurities (XRD). The catalytic behaviour of the zeolite after “activation” by calcination in air at higher temperatures (700-1 OOO”C for 2h) recently showed maximum phenol activity at 900”C [7]. So, the influence of calcination time on the activity was investigated using the same calcination procedure. Catalyst characterisation During the calcination procedure the zeolite significantly changes its properties, Table 1 gives an overview what influence calcination procedure has. Table 7: /nfkrence of the calcination procedure on the zeolite characteristics Calcination SiAl Framework Si/Al Framework Microporevolume Volume of unit procedure (“AI NMR) (TPD NH,) t-plot method cell (XRD) [-1 [-1 [m!Jgl [A’] +/-3 A’ Parent 22 24 0.14 5394 900”C 2h 57 44 0.12 5383 900”C 4h 65 58 0.12 5380 900°C 8h 71 0.12 5379 900”C 12h 76 (1:4) 0.10 5371 Firstly, all samples showed no significant loss of crystallinity and phase impurities in X- Ray diffraction pattern, even not after 12 hours calcinafion at 900”C. The first 2 hours of calcination at 900”C causes the most structural changes: framework aluminium, amount of adsorbed ammonia, micropore volume and volume of the unit cell significantly decrease (see Table 1). In the DRIFT spectra IR bands of the bridged OH bands (3600 cm-t) disappear afler 2 hours of calcination at 900”C. Longer calcination (up to 8 h) leads to further dealumination, but it leads to no significant contraction of the unit cell or micropore volume. The Si/Al ratio calculated from TPD corresponds well with the values from 27AI-NMR. ,,;,<;”,:: ..; .+, ., ,, ,,.:.,.; J-,, :- . .’1$ ,,, . ,..’..J,,.),. ,~,< ., ,, .,,, ~;:.”;-, ‘.:..,; L, ,r, ; .-. , ,,, ” ; Ar-on-si .’ —parent A Si-OH ,’,,.‘,, ,; ,,, . ,,, ?,,- ,, , “f ,, /., ,,, :, ‘, :,,,’.,: ‘,,: J. ,,,. ,,;;,.~ ,, ,-, :!. :“ , ,., ‘,’:,,;!’ , ‘ ‘:’..:. ,,,,:; ,-. 1,-- I .. : “’,,.,,,, ,, : 3400 3600 3800 0 0.2 0.4 0.6 0.8 1’ ,,,; ~, ... ,-,: Wavenumber [cm-’] Relative Pressure [p/pO] ,: -, ,, ,<,,,... ,,, :’? Figure 2: Catafyst characterisation: Ieff: DRIFT spectra of the OH-Bands; right: N2-Sorptionisotherms .,, , However, the 12 hours calcined sample exhibits new characteristics: [R bands of the bridged OH-groups totally disappear, but the non-bridged SiOH groups seem to increase . (Fig. 2 left). This effect could be due to silanol nests generated by dealumination [12], but the precise location of these kinds of OH-bands is still not clear [13]. In this sample also a big difference between both calculated Si/Al ratios in table 1 is obvious. The reason is probably due of the experimental error using NH~ TPD when almost no Bransted sites are Iefl. A ,., : ‘ closer look to the nitrogen adsorption isotherms (Fig. 2) is necessa~ to realise a structure change which is just invisible calculating the micropore volumes (T’able 1). The isotherm ‘ ,, ~- ‘ shows a step at p/pO of about 0.2 though at higher relative pressures adsorbed nitrogen ,, volume is very similar to the 2 hours calcined sample. A partial generation of secondary, wider pores is probably visualized which leads to no structure change determined in XRD patterns. This behaviour should occur from the extended dealumination of the framework [14]. Catalytic behaviour Selected results of the catalytic investigations at time-on-stream of 15min are shown in Figure 3. The high N20 surplus is chosen as standard condition because highest phenol yield is reached and beginning of total oxidation becomes visible. Parent sample exhibits total oxidation above a reaction temperature of 400”C. This is connected with a pronounced .- increase of benzene conversion, a total consumption of NZO and temperature runaway of . ‘},., ,’ 150”C. At higher temperatures total oxidation becomes dominant (yield of C02 up to 38%). So Bransted acid sites seem to be responsible for total oxidation, not for hydroxylation (s. ,.: ‘.,’,.. (, : Table 3). In contrast to that, all calcined sample show no total oxidation up to 450”C even at ~~,.’ this high N20 surplus. Having DRIFT spectra (Fig. 2 left) in mind total oxidation is prevented when Bransted hydroxyl bands are absent. This is in contrast to findings in literature [5] which support the acidity in the zeolite as essential factor for phenol production. With increasing calcination, from 2 hours to 8 hours, no significant change in phenol yield is reached below 375”C, while higher reaction temperatures lead to a continuous increase. At the 12 hours calcined sample phenol yield decreases again below 40WC but reaching the optimum phenol yield (45WC zzt%’0)of all samples at higher temperatures. (Fig. 3 right). Although characterisation showed the dealumination of the zeolite framework by contraction of the unit cell volume, 27AI-NMR (Table 1) and a loss of Bransted acidity (DRIFTS Fig. 2 left, Si/Al ratio by TPD tablel ) no clear correlation could be found with octahedral extra framework aluminum species (like [3]). The exclusive function of iron clusters as active sites !.,-,....- “i I ,,..,, for the hydroxylation also seems to be doubtable [1]. The formation of Lewis acid sites in the framework bv dehvdroxviation as active sites could also be possible [4]. But it seems to be clear that Br~nste~ acid sites are not necessary for the reaction. Mo~eover, those “’ detrimental for the phenol formation because of deactivation and total oxidation. 70 n -6-parent ~ 60 . : 50 -e-900”C 2h /7 ; 40 m -A-900”C 4h ~ 30 .— $ 20 -e-900”C 8h ; 10 ++900”C 12h o 325 375 425 475 325 375 425 475 Temperature [“C] Figure 3: Influence of calcfnation time at 900”C on catalyst activity versus temperature: Conversion of Benzene (left), Yield of Phenol (right), TOS: 15 rein, N,OI C,H6=6, W/F= 93 gminlmol Selectivity at TOS of 15 min is surprisingly low (maximum at 425”C: -70%). Even considering selectivities of the side products cannot fulfill the carbon balance (Table 3). Table 3: Selectivities to main and side products at 400”C Catalyst Selectivity Selectivity Selectivity Selectivity Selectivity Amount of Sample Phenol” Benzoquinone” Diphenols- Dibenzofuran- CO,” coke after [%] [%] 100 min p%l] [%] [%] [mmolig] Traces 6.7 1.81 -0.: Traces 0.7 3.1 -0.2 Traces o 3.1 -0.1 Traces o 3,7 Tre.-.. Traces o 2,7 EEEBEIL● calculated as amount of product related to amount of converted benzena Selectivity to phenol and benzoquinone increases during calcination procedure and reaches a plateau at 4 hours at 900”C. In the same way the selectivity to carbon dioxide, and the diphenols hydroquinone and catechol decrease. No resorcinol can be detected like in [8]. Missing hydroxylated hydrocarbons seem to remain on the catalyst surface to form coke. The amount of coke which is determined from COZ formation during regeneration shows a good correlation with the phenol yield at 400”C (compare Table3 (last column) and Fig. 3 right). This supports the opinion that coke formation is predominately caused by phenol as precursor. Dibenzofuran, a product of phenol condensation, is identified as preliminary stage of coke formation in phenol alkylation on USY zeolites [10]. In the case of ZSM5 it is not clear if the formation occurs at the outer or inner suriace. But if one has a closer look to time-on-stream behaviour of activity and selectivity during the first 100 min (e.g. 12 hours 900”C sample see Fig. 4) it is obvious that selectivity to 256 ,, ,, .,.: ,.’, $ phenol increases over time-on-stream. Moreover, at temperatures above 40WC maximum .,. selectivity are reached during 100 min while at 350”C a plateau is still not obtained. The “. ,,. reason for that behaviour could be the strong interaction of phenol with the zeolite. Higher temperatures favour the resorption of the product and accelerate the adjustment of sorption equilibrium, but is overlapped by coking. These sorption effects were also discussed in some previous studies on a Ga-ZSM5 [8]. Investigating the reaction behaviour dependend on the benzene concentration in the feed the same conclusions were drawn: A surplus of benzene forced the resorption of phenol resulting in high selectivities. Also findings of Vereshchagin et al. [11] concerning exceptional kinetic behaviour of benzene in contrast to linear alkanes at hydroxylation with N20 support this result. l+YPhenol 350”C +-Y Phenol 400”C -o-Y Phenol 450”C I- G .S Phenol 350”C -0 -S Phenol 400”C -0 -S Phenol 45WC ,>,. ,’- 25 Figure4: ,, g 2(3 ., :’. . Time-on-stream behaviour of ,,$, -,:(;,, ‘,’, zeolite (12 hours 900°C): ,,.);, ,’.: ,- ..,, .< .; ~ 15 N20/ C6H,= 6 !,,:!. - > W/F= 93 gminfmol : 10 ..’ 25 n o O 20 40 60 80 100 120 ,“ ,, ., .!, Time on Stream [rein] ,. ~..’ .,, 3.2 Separation of N20 and Benzene feed ,,~, ,. ,.. TAP- Experiments-overview .,,.,, , Recent studies described in [2] using continuous flow TAP-exrzeriments for the nitrous ‘, .. oxide decomposition showed that a 0.15 chemisorbed surface oxygen spe- + N20, N2, Ne (430”C) cies is formed upon N20 decom- ,, position on an H-ZSM5 without any ~ 0.10 - 02(500”C) special iron introduction. The gene- .g Lz ration of this species occured below In 30~C. Above this temperature sur- @ 02(458”C) face oxygen desorbs as molecular = 0.05 oxygen into the gas phase resulting ~z 02 (420”C) in a mean life time on the surface ti between 0.2s at 500 “C and 2s at 0.00 420 “C (Fig. 5). At higher tempera- 1 2 3 4 0 tures stoichiometric nitrous oxide Time [s] decomposition takes place. Figure 5: TAP raactor: single puke experiments [2] So even if the sites for the formation of surface oxygen are not exactly located, the for mation itself seems to be out of doubt. The results are in good agreement with Panovs tindings in his experiments [6]. 259 Investigations in presence of benzene and phenol seemed to be more problematic: Dual pulse experiments with benzene and nitrous oxide showed that no phenol formation was found though 99 % of the surface oxygen species reacted with benzene, Multipulse experiments lead to phenol production after 30-40 pulses. Therefore we concluded that phenol accumulates on the catalyst sutiace until a sorption equilibrium is reached. After this short review of the microreactor experiments in vacuum it is interesting if this effect is also visible in the continuous flow experiments under atmospheric pressure: B 1s it possible that the zeolite can store oxygen long enough to separate oxidation and reduction cycle? % How fast is phenol resorption if a separation of oxidation and reduction is possible? For these experiments we used the 12 hours at 900”C calcined SM55 (s.3.1). Alternation of nitrous oxide and benzene feed The first type of experiments consisted of the following steps (Fig 6, 8): @ pretreatment with 25% N,O/N, (1 hour, T=325- 400”C W/F=93 (g min)/mol) S purge with N2 to remove physisorbed and gas-phase N20 0 feed of 107. benzene/N, (T=325- 40rYC, W/F=93 (g min)/mol) G purge N, and heat up to regeneration temperature (T=500”C) 0 isothermal purge with N, (3 h, T=500”C) Measurements at a reaction temperature of 351Y’C and 400”C are exemplary shown in Figure 6 and discussed below. Generally, it is obvious that only very small amounts of phenol and low conversions of benzene can be reached. So it is difficult to get quantitative results. In Figure 6 all components are presented with their reactor outlet concentrations. 3.OE-03 1.OE-06 5.0E-070N Q z 2.5E-07j n. O.OE+OO O.OE+OO o 10000 20000 30000 0 10000 20000 30000 40000 Time [s] time [s] Figure 6: Alternated feed of N20 and benzene: Ieff 350”C, right 400eC After starting benzene feed it takes more than one hour until phenol can be detected. Even carbon dioxide appears time delayed. These effects depend upon temperature (i, phenol: 350°C: 100min, 400”C: 71 rein; ii, CO; 350”C: -36min, 400”C: -O min (Fig. 6, @). Additionally, after the benzene separation and further purging with nitrogen phenol appears still in the product spectra. Here, it becomes obvious how strong the interaction of phenol with the zeolite framework is. Thus, the idea of this process mode as a new reactor concept seems to fail. Section 0 of Figure 6 shows that carbon dioxide formation increases though no oxidant is in the gas phase. This is an hint that carbon dioxide can be formed by thermal decomposition of oxygen containing hydrocarbon species. Another possibility could be a recombination of remained, free monooxgen species to molecular oxygen during increase of temperature which is able to oxidise adsorbed hydrocarbon species completely. :. 260 # ‘. . . . Unfortunately, at a reaction temperature of 350”C the amount of phenol evolved does not decease during benzene purge (Fig. 6 left). So the experiment at 401YCis prolonged (Fig. 6 right). But also in this experiment the amount of evolved phenol does not decrease. However, total phenol amount is increasing (0.7E-07 + 1.3E-07 mollmin) as expected for higher temperatures. At 40WC carbon oxide formation is simultaneously detectable with the start of benzene feed and its amount is significantly higher (2.5E-07 ~ 5. OE-07 molimin). After the nitrogen purge at regeneration temperature of 500°C some amount of hydrocarbon is still on the catalyst surface. This fact is obvious in the last step of the experiment, a regeneration with air at 500”C (Figure 7). 0.04 A )25°C 375°C ~ 0.03 E : 350”C 400’ e 0.02 &. 8N 0.01 0 u L 0 25000 50000 325 350 375 400 Temperature ~C] time [s] Figure7: Regeneration with air at 500”Creation temperature:325”C-400”C left: CO1formation; righti Area of COZ peaks + amount of carbon C02 formation “peaks” during regeneration increase with increasing temperature until 375”C. In the same manner calculated carbon amount increases. However, in the long experiment at 400”C the C02 peak and amount of carbon is decreasing again. Up to 37SC the resorption of phenol seems to be slow enough that oligomers can be formed by condensation reactions. At 400”C resorption of phenol is faster. So coke precursors do not remain on the surface long enough and the prelimina~ stage for formation of bulkier oligomers is inhibited. These effects were also visible in the time-on-stream behaviour of the steady state experiments (see Fig. 4). In a next investigation the reverse experiment is done pretreatment with benzene 1.OE-02 0.1 ~ N20 (0), followed by a nitrous ~ 8.OE-03 oxide purge (0). o.075~ E ~ 6.OE-03 :0 ~ 0.05 g ~ 4.OE-03 L -. ~ 2.OE-03 0.025fj z Figure 8: O.OE+OO — 0 Alternated feed of banzene and N20at 375”C: 0 10000 20000 3000C ) time [s] Results at 37&C again demonstrate the importance of active oxygen formation (Fig. 8). When after benzene pretreatment nitrous oxide runs through the catalyst bed no phenol or 261 —— . other products than C02 are formed. The zeolite surface and, especially, the sites for active oxygen formation are blocked by sorption of benzene. Nitrous oxide itsself burns off the adsorbed benzene. Afler heating up and isothermal purge in nitrogen and even after addition of air no carbon dioxide is formed anymore. This effect is found independent on the temperature. The amount of carbon dioxide during NZO purge (Fig. 8 0) is higher than in the corresponding regeneration peak of the reverse experiments because no hydroxylated hydrocarbons leave the catalyst surface in this period (0.5 mmol compare to Fig. 7). 4. Conclusions After calcination of a ZSM5 zeolite at 900”C for different periods of time Bransted acid sites can be excluded to play a positive role in the hydroxylation of benzene. Coke formation can be inhibited and total oxidation can be prevented destroying these acid sites. Dealumination of the zeolite leads to an optimum catalytic activity and selectivity at reaction temperatures higher than 400”C. But a clear correlation was found neither with a special extraframework aluminum species nor lattice defects. Additionally, any iron species cannot be exclusively responsible for hydroxylation reaction. Moreover, we focussed on the interaction of the aromatic hydrocarbons with the zeolite. Experiments with an alternated feed of nitrous oxide and benzene supported the results got in the TAP reactor concerning the lifetime of the active oxygen species, on the one hand. On the other hand it was demonstrated what effect results from the patilcular strong interaction of phenol with the zeolite. Formation of phenol and coke as well as formation of carbon dioxide are strongly influenced by sorption processes. In literature this phenomenon has been scarcly considered, so far, but it will play an important role for industrial application of . one step phenol synthesis. Acknowledgements The authors are grateful for the financial support from the Deutsche Forschungs- gemeinschaff (DFG). We thank also Dr. Ti131erof ALSI-PENTA for offering the zeolites. References [1] A.K. Uriarte, M.A. Rodldn, M.J. Gross, A.S. Kharitonov, G.1. Panov, Stud. Surf. Sci. Catal. 110 (1997), p. 857. [2] A. Reitzmann, E. Klemm, G. Emig, S.A. Buchholz, H.W. Zanthoff, Proc. of DGMK Conf. 1998 Hamburg, p. 163, ibd. Chem. Ing. Tech. 70 (1998) p. 1017. [3] J.L. Motz, H. Heinichen, W.F. Holderich, J. Mol. Catal. A 136 (1998) p. 175. [4] L.M. Kustov, V.I. Bogdan, V.B. Kazansky, European Patent EP 0889018 Al, (1999). [5] Burch, Hewitt, J. Catal. Appl. Catal. A: 103 (1993) p. 135. [6] G.1. Panov, V.I. Sobolev, A.S. Kharitonov, J.Mol. CataI. 61 (1990) p.85. [7] S. Kowalak, K. Nowinska, M Swiecicka, M. Sopa, A. Jankowska, G. Emig, E. Klemm, A. Reitzmann, Proc. 12rnInt. Zeolite Conf., Mater. Res. Sot., 1999, p. 2847. [8] M. Hafele, A. Reitzmann, E. Klemm, G. Emig, Stud. Surf. Sci. Catal. 110 (1997) p.844. [9] J. Wang, E. Kiemm, G. Emig, Microp. Mesop. Mat. 26 (1998) p.11. [10] M. Guisnet, 1. Neves, F. Ribeiro, C. Canaff, P. Magnoux, G. Perot, Stud. Surf. Sci. Catal. 68 (1991) p. 735. [11] S.N. Vereshchagin, N.P. Kirik, N.N. Shishkina, A.G. Anshits, Catal. Lett. 56 (1998) p.145. [12] H. G. Karge, in: H. Robson, Microp. Mesop. Mat. 22 (1998) p. 547 [13] R. Salzer, U. Finster, F. Roessner, K.-H. Steinberg, P. Klaeboe, Analyst. 117 (1992) p.351. [14] F. Rouquerol, J. Rouquerol, K. Sing, Adsorption by Powders & Porous Solids, Academic Press, London, 1998 262 ,. .,,, ,.,,,.. .,. /.,, , ., DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 R. Monnig, W. Schwieger ,., Friedrich-Alexander University of Erlangen-Nurnberg, Institute of Technical ‘, , ,. .,..,, Chemistry 1,Erlangen, Germany ... ,,- ,, .,,. - ., Effect of ZSM-5 Zeolite Synthesis Following Different Routes on the Isomerization of Xylenes ...-.~-,, Abstract The objective of our work was a comparative study of HZSM-5 catalysts synthesized following different routes. Special attention was to be paid to the role of the template component in the reaction mixture with respect to its influence on physico-chemical properties of pentasiles. For this, ZSM-5 zeolites with different Al-contents were prepared with TPABr as well as without template. The Si/Al ratio of the samples from the template-free syntheses varied in the range from 12 to 25, whereas that from the TPABr-supported syntheses is between 15 and 45. The isomerization of xylenes was used for studying the catalytic behavior of ZSM-5 zeolites in a well-mixed batch reactor. In each case, we have found a good correlation between the Al-content of the HZSM-5 catalyst and its catalytic behavior. Additionally, it can be shown that the product synthesized without template but with boron as well as the product with the :, lowest Al-content synthesized with TPABr do not show the correlation mentioned ‘,,’, above, but exhibit enhanced activities. Some explanations for this are given in the paper. 1. Introduction HZSM-5, a pentasile-type zeolite, has provided unique possibilities of shape selectivity [1]. This has led to the development of several novel processes for producing the valuable p-isomer in the petrochemical industry [2]. The isomerization of xylenes is an acid-catalyzed reaction [3]. Only the strongest .,, sites of the Brrzmsted-type in HZSM-5 are involved in this reaction [4]. For xylene ;,.,., isomerization, previous measurements indicated that the monomolecular methyl ,, :-,, shift mechanism prevails in HZSM-5 catalysts [5-7]. The dimensions of the channels in HZSM-5 zeolites show diameters very close to those of the benzene molecule. The p-selectivity of ZSM-5 in various reactions with Cs-aromatics is the result of an interplay of the catalytic reaction with intracrystalline mass transfer. The generalized coupled network for the isomerization of xylenes involves six rate constants [8]. In general, it is assumed that the p-selectivity is caused by product-selectivity. The mobility of the p-xylene isomer is higher than that of the other xylene isomers in the intracrystalline void of ZSM-5 zeolites. However, there have been controversial discussions about this topic up today [7, 9, 10-14]. According to MIRTH et al. [1O], molecular modeling has shown a difference in size for the transition states in the meta-to-ortho intermediate as DGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 263 well as in the meta-to-para intermediate. Thus, there is a constraint for the formation of the transition-state in o-xylene formation. Consequently, an unknown contribution to the p-selectivity may result from transition-state selectivity. Furthermore, it is probable that the isomerization kinetic is adsorption-disguised. The number of adsorption and catalytic active sites themselves can therefore influence the p-selectivity. It was reported earlier that under template-supported crystallization conditions Al-gradients in larger ZSM-5 crystals have been observed [15-17]. The location of an acid site either in the zeolitic pore structure or at the external surface area as well as the site density itself can influence the acid strength and the adsorption behavior, respectively. In addition, some researchers underlined the importance of a carefully done zeolite crystallization and some difficulties resulting from subtle differences in the manufacture of ZSM-5 even from charge to charge: “The difficulty ... likely arose from differences in the processes vendors use to manufacture zeolites. Two vendors may offer zeolite products with the same structure and the same silicon-to-aluminum ratio, but their catalytic properties are /ike/y to be different” (SACHTLER in [18]). The relationship between synthesis conditions and physico-chemical properties of ZSM-5 zeolites has rarely been studied [19, 20]. [n the present work we aim at elucidating the influence of the zeolite Si/Al ratio in HZSM-5 catalysts on their isomerization rate under carefully selected reaction conditions, Furthermore, we will discuss the influences of the ZSM-5 zeolite synthesis method on catalytic properties. 2. Experimental 2.1 ZSM-5 Synthesis and Characterization Series of ZSM-5 were synthesized through both the template-supported and the template-free crystallization method described previously [20]. Colloidal silica, sodium aluminate, and NaOH solution (series B) as well as the template tetrapropyl ammonium bromide, TPABr, (series A) served as the starting raw materials. The crystallization of the ZSM-5 materials took place under hydrothermal conditions in permanently mixed stainless steel autoclaves (2000 cm3). The elimination of the template from the ZSM-5 products of the series A was achieved through a 5 hrs calcination in air at 873 K. The synthesis of the sample B4 was performed by substituting boron (as H3B03) for 40 mole YO of the aluminum in the reaction mixture. To obtain the protonic form, the ion exchange with 0.1 molar hydrochloric acid was carried out twice at room temperature. The compositions of the samples were determined through emission spectroscopy (Plasma 400, Perkin Elmer) after the HF-microwave decomposition. The X-ray powder diffraction patterns were taken with the help of the goniometer URD 63 from the Freiberger Prazisionstechnik GmbH, using Cu&radiation and a nickel .- filter. The crystailinities of the samples were calculated by making use of the external standard CZ-A1203according to the procedure described elsewhere [21]. The N2 isotherms were recorded at 77 K, using the volumetric apparatus Sorptomatic 1900 Turbo (Carlo Erba Instruments). 2e4 2.2. Xylene Isomerization Kinetics and product distribution in xylene isomerization were observed in a well- mixed batch reactor system according to RIEKERT and coworkers [22, 23]. This apparatus allows determining the gas phase composition by 9as chromatography (Varian Star 3400) as a function of the contact time between catalyst and xylenes. The main components in this reactor system are a mixing turbine (VM = 7000 cm3) and a quartz-reactor ( VR = 280 cm3) connected with the mixing turbine by a diaphragm pump ( V IOOP = ZOO Crn3d. The comPacted catalyst (500-800pm) was placed on a quakz-fritt in the heated reactor zone (Tf? = 573 K). Before every catalytic run was carried out, the catalyst had been preconditioned in synthetic air at 823 K for at least 4 hrs and then cooled down on the desired reaction temperature by N~flushing. Atler evaporating the educt (o-, m-, p-xylene) in the mixing turbine at TM = 453 K (P.YI = 5 mbar) in ni~rogen, the kinetic measurement of the reaction was started at t = O by connecting the mixing turbine w“th the reactor. The kinetic model employed accounts for the fact that an overall description of the xylene isomerization over ZSM-5 zeolites needs a coupled reaction network containing six first-order rate coefficients [8, 24]. This triangular scheme is given in FIG. 1. The estimations of the six rate coefficients were based on a numaric differentiation algorithm combined with an optimizing method (Newtonian approximation). Q ‘ CH3q\\ ~mP ,$0k’,m\\ 1.*. cH~ CH3 CH3 kmo ~ 0 k o o om d CH3 FIG.1.ReactionNetwork Scheme for the Xylene Isomerizationover ZSM-5 Zeolites 3. Results and Discussion 3.1 Characterization of the Samples The X-ray powder diffraction patterns of the samples are shown in FIG. 2. Obviously, the spectra of both series characterize the samPles as MFI structures [25] w“thout any additional peaks. The intensities of the structural reflexes are comparable for all samples. This is reflected in almost idantical crystallinities, given in Table 1 (row 3). These crystallinities are expressed as QAI values [21]. Numbers about one indicate crystallinities of almost 100 %. 265 :,, !& Two Theta Cqree T’.m Theta Cegme FIG. 2. XRD patternsof the HZSM-5 Samples: (a) Template-free Synthesis, (b) TPABr- supported Synthesis Some features of the catalysts are summarized in TABLE 1.The molar Si/Al ratio in series B varies in the range from 12 to 25. In Series A this ratio could be extended to 45. [n crystallization without any template, the Si/Al ratio of ZSM-5 products is limited to a maximum of about 30 [20]. In TABLE 1 it is to be seen that the Al contents of the corresponding samples are comparable in both series. Consequently, the potential number of active sites in these catalysts must also be almost identical. In addition to the XRD investigations, the results obtained from adsorption measurements confirm that all samples are pure ZSM-5 zeolites, which is indicated through almost theoretical micropore values. Table 1 ,’- , Designationand Some Physictihemical Characteristicsof the HZSM-5 Samples 6 , . . Sample Bulk Crystallinity3 Micropore Votume Surface Area Ctystal Size Si/Al (XRD) (NzAdsorption) (DUBININ’S Model) (SE Image) .. ~m3.g-11 [mz.g-l~ Ratio [pm] BI 12 0.90 0.15 413 <0.3 B2 19 1.30 0.14 385 8 B3 25 1.25 0.15 419 8 B4 25 —.....———1.15 —0.14______388 —-.. — .---—..-.9 Al 15 0.92 0.18 496 <0.1 A2 21 0.99 0.17 481 <0.1 A3 27 0.98 0.18 507 <0.1 A4 45 0.93 0.17 474 2 “)QNvalues, see Iexl The crystal sizes given in TABLE 1 were taken from SE images. FIG. 3a pictures the typical morphology of template-free crystallized ZSM-5 zeolites (sample B2). Here typical prismatic ZSM-5 single c~stals with a length of about 8 pm are to be seen. Their crystal surfaces are almost smooth and partially occupied by very small ZSM-5 particles. 266 In the case of the sample A2, which was synthesized with TPABr, there are aggregated particles consisting of very small crystallite (see FIG. 3b). ,., ., .. ~..-,, ,,, ., , ...... ,, , .,, , ,, ... ..- ,. .. .,- ...; ,,,- ,, ,,, ,,, : ::<:;’$ ..;,-’.,!. ‘: ,:,,, .,.,,.: -.,- ‘. ,:, (b) (a). . 1 Fl~. 3. SE Images of the ZSM-5 Samples: (a) B2, (b) A2 :.. . 3.2 Kinetics ,. FIG. 4 depicts the product distribution of the xylene isomerization over HZSM-5 ‘!, catalysts from series A and B as a function of the modified contact time. This is reaction time multiplied by catalyst weight. All three xylene isomers were used as feed in separate catalytic experiments. As expected, the rn-xylene isomerization yielded both o-xylene and p-xylene until equilibrium was reached (24.1 ?to o-xylene, 54.4 % m-xylene and 21.5 % p-xylene). The o-xylene isomerization also yielded two products simultaneously, m-xylene and p-xylene. ,. ~ This fact can be observed in the p-xylene isomerization as well. -This phenomenon confirms the usage of the triangular reaction scheme for descnbmg ,,,,,, the observations (see FIG. 1)and is a consequence of the p-selective behavior of both catalysts. The equilibrium was reached at different modified contact times. Consequently, ,> the activities of both catalysts must be different under identical reaction . . conditions. FIG. 4 does not only show the experimental points. Here the computed reaction paths are also given. The six first-order rate coefficients for .,. all catalysts are summarized in TABLE 2. . 267 [b) ,, m-Xy16me0 ..xylene I. px “1.”. 4, . . 0.8 . 0,8] OS ~ ‘., ❑ E : ““:: ““”- . . . . . 02 = 02 0.1 . Elliilli10 X301 020X, O %34 FIG. 4. Product Distribution of the Xylene Isomerization over HZSM-5 Catalysts: (a) B4, (b) A3 Table 2 Rate Coefllcients for the Xvlene lsomerization over HZSM-5 Catdvsts at 573 K Sample [Si/All k% ., ., kw km k pm k OP k em ~3 .t~.t cm, ., s.! ~ml .1~.1 ~ma .1 *.! ~, ., ~., (Cmgs B1 [121 23.8 16.3 10.2 58.4 14.8 31.6 B2 [19] 9.7 5.9 3.5 25.8 7.7 8.1 B3 [251 8.0 3.8 2.9 211 5.0 6.5 —B4~@-. 10.6__-__---_.——.4.9 3.3 27.5 5.2 ----—6.7 Al rla 19.6 3.9 3.2 39.6 2.5 8.5 A2 Elj 11.5 6.6 4.7 32.1 7.7 10.3 AZ ~?l 6.1 4.6 1,7 19.7 5.0 3.7 A4 [451 7.1 2.8 3.1 20.2 4.1 6.3 .) ~OmOmho@ sublduled wlh boron(43 mole%) As expected, in both ZSM-5 series the activities (rate coefficients) depend on the Si/Al ratios of the catalysts. Samples with the highest Al contents (i. e., lowest Si/Al ratios) are the most active catalysts. In particular, this is evident for the fastest reaction path, the p-xylene to m-xylene reaction. [t can also be observed that the rate coefficients decrease with increasing Si/Al ratio. However, the activity of the catalyst B4 is higher than that of the sample B3, although the Si/Al ratio is 25 in both samples. This is an unexpected result, since we assumed that due to the boron substitution either the number of active sites or the acid strength would decrease and, therefore, a slightly smaller activity would result. CHANG et al. [26] investigated the role of boron in pentasiles in different test reactions, e. g., in xylene isomerization. They concluded that framework boron has a negligible impact on Bransted acidity in ZSM-5 and on the catalytic activity in acid-catalyzed reactions by it. However, the ZSM-5 zeolites in their study was synthesized template-supported. We observed a slight increase in activity through boron substitution, even though the sample B4 is boron-free after the ion exchange with hydrochloric acid and calcination. Further studies should be dedicated to the clarification of this behavior. [n the case of an Si/Al ratio below 30, the HZSM-5 catalysts from both series except the sample B4 (synthesis with boron) possess almost the same isomerization activities relating to their Al-content. Besides this, if the catalyst utilization for all reactions of the isomerization network shown in Fig. 1 is nearly 268 ,, ,’.’ .,. , .,,,’ ,’ -’ identical, it is conceivable to conclude that the aluminum distribution in crystals from both series is homogeneous. In series A (syntheses with TPABr) the sample A4 (Si/Al = 45) shows a slightly ,“ ,: ;- enhanced activity relating to its Si/Al ratio. Unfortunately, we can not explain this ( . result at present. There is strong occasion to believe that the unique role of TPA+ cations in ZSM-5 crystallization process leads to an inhomogeneous Al distribution in larger crystals preferably. We assume that the enhanced activity of the sample A4 is caused by the Al-rich outer rims of the crystals. However, there is no further evidence for our explanation. , ,. Besides the isomerization activity of ZSM-5 zeolites, their p-selectivity is also an ,, ‘., -,, important subject of investigation. Our results show that a pronounced p- !, ..,, :,, -, selectivity can be observed over all HZSM-5 catalysts. ,,, ,.. , ,. ..- , ,..”,., ,, 4. Conclusions :,, , i. ZSM-5 products obtained from the TPA-supported as well as the template- ,,, , free synthesis route are very similar with respect to their crystallinities and adsorptive behaviors. However, different crystal morphologies and differences in the extent of the external surface in relation to the overall surface area have been observed. We believe that there is an important ,b,’., influence of the TPA+ cation on the mechanism of crystal growth leading to .-, ‘, ,.,., ,’ . ,’ differences in crystal morphologies during ZSM-5 crystallization processes. ,..,. , f .,.’:,, ,’” : ii. It has been found that the first-order rate coefficients correlate with the Al- ... $, contents of the HZSM-5 catalysts on condition that the Si/Al ratio is limited :..,.”’.,. ,~.; to 30. ZSM-5 zeolites from both synthesis routes show nearly identical :,:,.,:.,,., ,. ,. ,, xylene conversion rates relating to their Al-contents. ..: .,.,...... Ill. Since the raw materials and the crystallization conditions used in the ,:, ,”,-, ~.,, .-., template-free method are comparable to those of the TPA-supported ,,, . ,,. crystallization, we suggest the employment of the template-free synthesis ..>..,.,$,:,..,..,..5,.’- . “,:, ,.,---- method for ZSM-5 zeolites with low Si/Al ratios. Further studies should be ?. ,:., .’,< : \: ,,,,.,- ,’, . dedicated to investigations of the p-selective and deactivation behaviors of .’ ,’,.,;<, different synthesized ZSM-5 zeolites. ,.,. , >., -,. ,<. . .,, .,, .,,. ,..,~.,., ,., .,,’. , :.; ,,,<. ,,/,. ~,,“-’ ‘.:, References , .,.’,..-, ,., ,, :. :, .,,. ,, :... .,. ,:L 1. Kaeding, W. W.; Chu, C.; Young, L. B.; Weistein, B.; Butter, S. A.: ./. Catal. ~, 159 (1981). .-: .-, 2. Chen, N. Y.: Gamvood, W. E.; Dwyer, F. G; .Shape Selective Catalysis in Industrial Applications”, Decc.er, New York (1 989). 3. Ward, J. W.; Hansford, R. C.: J. Cata/.U, 154(1969). 4. Babu, G. P; Hedge, S. G.; Kulkami, S. B.; Ratnasamy, P.: ./. Cafa/. ~, 471 (1983). 5. Corms, A.; Sastre, E.: J. Chem. Sot. Chem. Commrm., 594 (1991). 6. Young, L. B.; Butter, S. A.; Kaeding, W. W.: J. Cafa/. ~, 159 (1981). 7. Olson, D. H.: Haag, W. O.: ACS$ymp. SeL N, 275 (1984). 8. Collins, D. J.; Medina, R. J.; Davis, B. H.: Can. J. Chem. Eng. ~, 29 (1963). 9. Chen, N. Y.: J. Catal.~,17(1988). ? . . 269 . ..’ 10. Mhth, G.; Cejka, J.; Lercher, J. A.: J. Catal.139,24 (1993). I]. Klemm, E.; Wang, J.; Emig, G.: Chem. Eng. SGi. = (18), 3173 (1997). 12. Kumtrner, U.: Jerschkewitz, H. G.; Schreier, E.; Volter, J.: APP/. CafaL 5L 167 (1990). 13. WU, p.: Debebe, A.; Ma, Y. U.: Zeo/ites a, 118 (1983). 14. Ruthven, D. M.: Eic, M.; Richard, E.: Zeolifes fit 647 (1991). 15. v. Ballmoos, R.; Meier, W. M.: Nature ~, 782 (1981). 16. Hughes, A. E.; Wilshier, K. G.; Sexton, B. A.; SmaIt, P; J. Catal. QQ,221 (1983). 17. Derouane, E. G.; Gilson, J. P; Gabelica, Z.; Mousty-Desbuquoit, C.; VerbSt, J.: J Cafa/. U, 447 (1981). 18. SACHTLER,W. M. H.: Chem. Errg. News 75,8 (1997). 19. Schw”eger, W.; Bergk, K.-H.; Alsdorf, E.; Fichhrer-Schmittler, H.; Loffler, E.: Lohse, U.; Parlifz, B.: Z. Phys. Chem. ~ 243 (1990). 20.Schwieger, W.; Bergk,K.-H; Freude, D.; Hunger, M.; Pfeifer, H.: ACS SymP. Ser. 39S, 274 (eds. Occelli, M. L.; Robson, H. E.), Washington, DC (1989). 21. Be@k, K.-H.; Schwieger, W.: Ab.str. hrtemaf. Corrt on Zeolites, Portoroze, 1984. 22. Riekert, L.: Z. .Hektrochemie ~, 201 (1962). 23. Prinz, D.: thesis, Kartsruhe, 1966. 24. Beschmann, K.; Riekert, L: J. Cafal. ~, 548 (1993). 25. Meier, W. M.; Olson, D. H.: Atlas of Zeolite Sfrucfure Types, Polycrystal Book Sewice, Pittsburg (1978). 26. Chu, c. T. W.; Kuchl, G. H.; Lago, R. M.; Chang, C. D: ./. Cata/. ~, 451 (1985). 4, 270 DGMK-Conference ‘he Future Role of Aromatics in Refining and Petrochemistry”, Erlangen 1999 R. Glaser, J. Weitkamp Institute of Chemical Technology 1,University of Stuttgart, Stuttgart, Germany ,., j t Alkylation of Napthalene on a Zeolite Catalyst in Supercritical and Gaseous ,., .,,, ,- :,.-,..., ,, Reaction Phases ,.. . . ., ‘., . . .,,. ,.... .,, .,. . .. .’, . \ ,, ., .;,, ,, :.,,>< :.. , ,., ...,,, ,,, .,,, ,.- ABSTRACT :,.,,.,;, ,’ ,,, .. --,.-, The alkylation of naphthalene with methanol and isopropanol over zeolite LaNaY-73 was investigated in supercritical carbon dioxide and in the gas phase. Monoalkyl- and dialkylnaphthalenes are the major products, and their yield ratio was dependent mainly on ,,. the almholhaphthalene ratio in the feed, regardless of the reaction conditions, time-on- stream, and the degree of catalyst deactivation. With both methanol and isopropanol, the deactivation of the catalyst at supercriticsl conditions was strongly reduced as compared to the one observed in the gas phase. Depending on the reaction pressure and temperature, a transition from thermodynamically to kinetically controlled product formation was observed which was different for either of the two alcohols. The distribution of the dialkylnaphthalene isomers follows qualitatively that for the monosubstituted products and is mainly governed by semndary alkylation rather than by isomerization. INTRODUCTION ,, Supercritical fluids have recently attracted increasing interest as media for heterogeneously catalyzed chemical conversions [1,2]. Many important physical properties of supercritical fluids, such as density and viscosity, are intermediate ,, between those of liquids and gases. These properties can be continuously changed over a wide range by small changes in pressure and/or temperature in the region near the critical point. Therefore, the mass transfer properties of reactants and products to and from the catalyst surface into the reaction medium is strongly dependent on reaction conditions which opens up possibilities to direct the selectivity in heterogeneous catalysis [3]. Another benefit of using supercritical fluids . for heterogeneously catalyzed conversions is the increased volubility of higher :.’, .4, . molecular weight compounds which, under gas phase reaction conditions, may form ;, unsoluble deposits on the catalyst surface that eventually lead to complete catalyst deactivation due to coking [4,5]. For the izomerization of 1-hexene over a .,. macroporous catalyst at a temperature close to Tc, an optimum density has been ,,” shown to exist for which the interplay behveen the volubility of coke precursors in the supercritical reaction phase, which by itself increases with density, and the mass transfer rate of reactants and products within the catalyst pores, which by itself is slowed down by increasing density, leads to an optimum in catalyst performance [6]. ,.’ -,,,’, :.-, !,,, J., ,., , DGMK-Tagungsbericht 9903, ISBN 3-931850-59-5 271 , .. ,...... , :f.i .“ ,,:,,:. ,,, ‘“’ ,“ .. ,:,’,:,.- “- ~ ... Y,:., - :, - ;:,’ EXPERIMENTAL SECTION Zeolite LaNaY-73 was prepared from NaY (Union Carbide, Tarrytown, N.Y., USA) by a two-step ion exchange with La(NO~)Sat 80 “C and a CSdCiIMtiOn Step at 150 “C for 3 days in bebween the two steps. For the catalytic experiments at supercritical conditions a high-pressure flow- type apparatus was used designed for operation at up to p = 520 bar and T = 250 “C. The feed mixture was prepared by passing a supercritical mixture of carbon dioxide and the corresponding alcohol over a saturator which contained a fixed bed of solid naphthalene (mixed with an inert solid carrier and metal beads). At the constant temperature of the saturator, Np dissolves in the supercritical phase according to the phase equilibrium; thus the desired concentration of naphthalene in the reactant mixture is obtained. This alcohol/Np/COz mixture is then fed to the reactor in which the zeolite catalyst as a fixed bed is maintained at the reaction temperature. With the catalyst in the reactor at 250 “C and without the alcohol in the supercritical mixture it took ca. 4 h until the adsorption of naphthalene was completed. Samples were taken periodically from the condensed product stream . after depressurization. Prior to the catalytic experiment, the zeolite powder was pressed without a binder, crushed and sieved to obtain the particle fraction of d = 0.2- 0.3 mm. The catalyst was activated in a purge of nitrogen (30 cm3/min) by heating with 2.5 IVmin from room temperature to 250 “C and keeping this 272 temperature for 6 h at ambient pressure. Subsequently, the reactor was flushed and pressurized with pure C02 to the desired reaction pressure. A similar flow-type apparatus was used for the experiments in the gas phase at ambient pressure. The carrier gas (nitrogen or carbon dioxide) was directed through two separately thermostatted saturators containing liquid naphthalene and the alcohol, respectively. The gas streams from these saturators were combined and sent to the reactor. Product samples were taken periodically on-line from the reactor effluent and analyzed by glc. RESULTS AND DISCUSSION Alkylation of naphthalene at supercritical conditions The main products of the alkylation of Np with either MeOH or IPOH are the mono- (R-NP) and dialkylated (DR-NP) products. Higher alkylated products (tri- and tetraalkylnaphthalenes) were only formed with yields c 0.1 %. The conversion of naphthalene in the methylation reaction at 250 “C decreases with time-on-stream from ca. 12 Y. at 3 h to ca. 9 Y. at 10 h (Figure 1). This activity decay, which is due to coking of the zeolite catalyst, is comparable to the one observed with IPOH as the alkylating agent [7]. In the isopropylation, however, a considerably higher conversion level (ea. 60 Y.) than in the methylation is reached, because of the higher stability of the isopropyl versus the methyl cation which act as the alkylating species. This is in line with the observation that more dialkylated products are formed in the conversion of Np with IPOH (YIP.Np/YDIP.Np= 4,7) than with MeOH (YM.N+YDMLNP= 6,4), if compared at the same conversion. For this purpose, conversions at different temperatures and at different degrees of catalyst deactivation have to be compared. Since it has been shown for the isopropylation reaction that the yield ratio YIP.Np/YDIP.NP is determined ● xNp 20 ❑ ‘M.NP O YDM-NP .0. Uj 15 t i 5 - “000000000 000 , 0 0 2 4 6 8 10 TIME - ON - STREAM I h Figure 1. Conversion of Np with MeOH over zeolite LaNaY-73 at supercritical conditions, TR = 250 “C, yNP = 2.8 Y., tk?OHtflNP = 0.8, WIFNp = 275 9“hlm01, P = 200 bar. 273 ,, !,’,, -. :i: ;,, ... :.,; “-. ,1:, by conversion only [7], this comparison is still valid. Influence of the nhIeOr@j - ratio and of the reaction pressure The conversion of naphthalene with methanol at 250 “C, but at higher nt.!aOH/nt@- ratio and at higher pressure, is shown in Figure 2 (left-hand part). At 201)bar a twofold kICW2aSe Of the alcohol COnh?nt h the feed frOIll nM@H/nNp = 0.8 tO 1.6 results in a decrease of catalyst activity also of roughly a factor of two. The activity is reduced from the beginning of the experiment. Upon a twofold increase of the reaction pressure to 400 bar the initial conversion is higher (ea. 25 %), and after 3-4 h it remains at a level comparable to the one reached at 200 bar and the lower nM~H/n~~- ratio. This may be caused by the higher density of the reaction phase at increased pressure which appears to efficiently slow down catalyst deactivation by an increased volubility of higher molecular weight coke precursors in the supercritical phase. This is corroborated by the color of the product condensate which was dark brown in the experiment at 400 bar and only slightly brown in the experiment at 200 bar. The yield ratio of methyl- and dimethylnaphthalenes is nearly constant if compared at the same conversion, regardless of the pressure, nMeOH/nNp - ratio or time-on-stream (not shown). A constant yield ratio of mono- and dialkylated product has also been found in the conversion of Np with IPOH for different reaction temperatures [7]. The yield ratio of 1- and 2-methylnaphthalene at different n~@@Np - ratios and at different reaction pressures are depicted in the right-hand part of Figure 2 as a function of time-on-stream. As also observed in the isopropylation of naphthalene at supercritical conditions, the fraction of the 1-substituted isomer increases during I , i ,’ . m 1, ,. x= z 20 - . 0 ~nno m cc w c1 > ~onQnoDo 00 z 10 - O%d G o 0 0 000 0 , , , 0 2 4 6 8 0 2 4 6 6 TIME - ON - STREAM / h TIME - ON - STREAM / h Figure 2. Conversion of Np (left-hand part) and distribution of methylnaphthalene isomers (yield ratio yI.M-NfiZWNp; right-hand part) in the conversion of Np with MeOH over zeolite LaNaY-73 in dependence of time-on-stream, at different n~tinN~ - ratios and reaction pressures; reaction conditions as in Figure 1. 274 .. :, ,-” ‘,. ,.,, ,,, .’-, ;. ,.:)’. .:.l :...,,’.’ ,, the experiment. From an initial value close to the one expected for thermodynamic “~ : ~; : equilibrium (Y1w.~$Y2.~.~P= 0.61 [12]) the yield ratio increases corresponding to a :,, ,:, , shift from thermodynamically to kinetically controlled product formation. Interestingly, ;:’: .. .’ ,; the change of the isomer distribution with time-on-stream is not affected at all when ~,’ ‘-;,’ ,, the fl!JoOtI/nt@- ratiO is doubled, although the CatalySt activity (and, thus, the yield of ,1,,..~ .- ; ‘ ,., $’ the methylnaphthalenes) is considerably lower. Therefore, the properties of the :;’, ...,:,:.,;,. ,,:.’, reaction medium surrounding the catalyst paticles appear to determine the resul~n9 :: :’ ‘T.,‘.” ‘:, isomer distribution rather than the nature or density of active sites on the catalyst ~, “:, ‘~”’ ~‘ . surface. It can also be seen from Figure 2 (right-hand part) that, at the increased ‘‘”.: ~ ,. ,- pressure of 400 bar, the distribution of the methylnaphthalene isomers is not strongly .,,,, ..’, ,’ different from the one at the lower reaction pressure. Two competing effects might . . ., be responsible for this The increased volubility of the products in the supercritical . .,,.,’ .,,.. reaction phase leads to a faster resorption of the initially formed 1-M-NP before .,..’ ... .,,? ,, ,, ,.$,, !.:. .’,’ subsequent isomerization to the 2-isomer can take place. A similar shift to kinetically .-, , . ,!.. . . controlled product formation with increasing pressure at supercritical reaction ‘$, - ;,,.”’; .,, conditions has been reported by 77/tscher et a/. [3] for the cis/trans isomerization of -, ,.’ ,. :.!’ 2-hexene. On the other hand, the removal of coke precursors from the catalyst surface becomes more effective at increased pressure leaving more active sites for re-adsorption and isomerization of 1- to 2-methylnaphthalene. Further experiments are, however, needed to clarify this issue. Distribution of fhe dime fhyh?aphfhalene isomers The distribution of the dimethylnaphthalenes in the conversion at 200 bar and n~m”/n~P= 0.8 is shown in Figure 3. Similar to the distribution of the methylnaphthalenes (Figure 2) the mole fractions of the dimethylnaphthalenes change the most during the first 4 h of the experiment. Like for the dialkylated products in the isopropylation of naphthalene (v], Figure 3), the fractions of the 2,6- and 2,7-disubstituted isomers decrease, while those of the 1,4- and the 1,5-isomer 1, increase. The 1,2- and 2,3-isomers, which are not formed with IPOH as the ~, . ,,, .,. .,, ,> .$ ,, :.,, o~ o~ o 2 4 6 8 10 0 2 4 6 8 10 ...... > TIME - ON - STREAM / h TIME - ON - STREAM I h Figure 3. Distribution of dimethylnaphthalene isomers in the conversion of Np with MeOH over zeolite LaNaY-73 at supercritical conditions; reaction conditions as in Fig. 1. 275 ,. alkylating agent due to steric hindrance of two neighboring isopropyl groups, do occur in the dimethylnaphthalenes, the 1,2-isomer being even the most abundant. Furthermore, the mole fraction increase in the early stages of the experiment is most pronounced for 1,2- and 1,4-dimethylnaphthalene, which are both secondary 6, alkylation products and are preferentially formed under kinetic control from the . . methylnaphthalenes. The decrease of the thermodynamically favored 2,3-, 2,6- and ., 2,7-DM-Np is accompanied by a decrease of the fraction of the 2-isomer within the monoalkyiated products. Due to the concomitant decrease of 1,6- and 1,7-DM-Np, which result from isomerization of the 2,6- and 2,7-isomer, respectively, and due to the fact that 1,2-DM-Np cannot be isomerized into any other dimethylnaphthalene [13], it can be assumed that direct alkylation of the methylnaphthalenes predominates over isomerization in determining the distribution of the dimethylnaphthalene isomers. A/ky/ation of naphthaiene in the gas phase In the gas phase reaction with methanol, the conversion of naphthalene drops rapidly from an initial value above 30 ‘A to a value below 2 ‘%0 within 6 h time-on- stream, and after 8 h no conversion at all can be observed any more (Figure 4). These results are in accordance with an earlier investigation by Fraenke/ et a/. [8] who observed a rapid deactivation of faujasitic zeolite catalysts in the conversion of naphthalene with methanol at temperatures around 400 “C. The n~m~/n~P- ratio in the feed has only a minor influence on catalyst activity. A higher methanol concentration in the reactant mixture results in a slightly higher initial activity, but also in a slightly more rapid deactivation of the catalyst. Also, there is only a small effect of the methanol concentration on the distribution of the methylnaphthalene isomers (insert in Fig. 4). I , , , I I 0 , 0 2 4 6 8 10 TIME - ON - STREAM / h Figure 4. Influence of the nM~n~P - ratio on conversion of Np and distribution of monoalkylated products (yield ratio YlW~fi2W~P; insert) in the conversion of Np with MeOH over zeolite LaNaY-73 in the gas phase, T~ = 250 “C, y~P= 3.0 ‘XO, W/F~P = 285 g.h/mol, p = 1 bar. 276 ., .,. , :, !,. The yield ratio yl.M.Np/y24J.Np remains nearly constant during the experiment, although ,, .,,, the conversion changes over a wide range. The inflUenCe Of the nMaOHhNp- ratio On ,. !’. ’,$, the yield ratio YI.M.Np/YZ-M.NP was also found to be negligible at supercritical conditions (cf. Fig. 2, right-hand part). In the latter case, however, the activity of the catalyst was considerably changed with the methanol content in the feed. Again, this is best ~’ understood in terms of an increased volubility of coke precursors in the supercritical ;- phase, maintaining the catalyst activity at a higher level for a longer time-on-stream. ‘, .,. Since coke precursors are not soluble in the gas phase to any appreciable extent, the transition from thermodynamically to kinetically controlled product formation cannot occur and, as a result, the yield ratio of the methylnaphthalenes YI-M.Npfyz-M.Np stays close to the value expected for thermodynamic equilibrium. ‘,,. The influence of the alcohol concentration in the reactant mixture is similar, :., ... but more pronounced, if IPOH is used as the alkylating agent (Figure 5). At variance to the conversion of Np with MeOH, however, there is an increase in the yield ratio ~j.k’,, of the monoalkylated products with time-on-stream, which also depends on the ., -,’ : alcohol concentration. Since water is formed as a by-product with both MeOH and ,.. , IPOH, and since more water is expected to be fomed at higher alcohol concentration . .”..’,,., in the reactant mixture, it will most probably not be responsible for the observed :,. .$ differences in the isomer distribution of the monoalkylated products. It might be considered that there is an interrelation behveen the shift from thermodynamic to kinetic control of product formation and the degree of catalyst deactivation, which is dependent on both the alkylating agent and the reaction conditions. For both methylation and isopropylation of naphthalene the deactivation of a zeolite catalysts observed in the supercritical reaction phase is appreciably reduced as compared to the conversions in the gas phase. This is best accounted 50 - 40 - 30 - ‘024681[ TIME - ON -STREAM I h 20 - ,’, ,,,, 10 - \ K — G “o 2 4 6 6 10 TIME - ON - STREAM / h Figure 5 Influence of the nlpo~nN~- ratio on mnversion of Np and distribution of monoalkylated products (yield ratio Yt.IP.N@’z.JP-Np: insert) in tie conve~ion of Np with IPOH over zeolite LaNaY-73 in the gas phase, TR = 250 “C, 4, .,,,. , .;,,. . , YNP = 1.3 Y., W/FNP = 300 g.hlmol, p = 1 bar. ‘:, ,., ... ,. ,, i for by an increased volubility of coke precursors in the supercritical phase. But even at supercritical reaction conditions deactivation cannot be completely supressed. CONCLUSIONS Supercritical fluids offer interesting opportunities as reaction media for zeolite catalyzed conversions. The deactivation due to coking can be strongly reduced at supercritical conditions as compared to the one for gas phase conversions. Also, the selectivity for products with similar physical properties, such as alkyl- and dialkylnaphthalene isomers, can be changed by adjusting the reaction conditions. Above ail, pressure may serve as an additional variable to control the outcome of a heterogeneously catalyzed reaction. Therefore, the application of supercritical fluids might be beneficial for many other catalytic conversions of aromatics like, for example, isomerization, transalkylation, or hydrogenation I dehydrogenation reactions. ACKNOWLEDGEMENTS The authors gratefully acknowledge financial support by the Deutsche Forschungsgemeinscha ft, Max-Buchner-Forschungsstiftung and Fends der Chemischen Industrie. REFERENCES [1] P.E. Savage, in: “Handbook of Heterogeneous Catalysis”, G. Ertl, H. Knoezinger and J. Weitkamp, eds., Vol. 3, pp. 1339-1347, Vol. 3, Wiley-VCH, Weinheim (1997). ,’ . [2] A. Baiker, Chem. Rev. 99, 453-473(1999). #.,. [3] H. Tiltscher, J. Schelchshorn, F. Westphal and K. Dialer, Chem.-/nTechech. 56,42-44 (1 984). .. [4] H. Tiltscher and H. Hofmann, Chem. Eng. Sci. 42, 959-977(1987). [5] G. Manes and H. Hofmann, Chem. Eng. Techno/. 14, 73-78(1991). [6] S. Baptist-Nguyen and B. Subramaniam, A/ChE ./. 38, 1027-1037(1992). [7] R. Glaser and J. Weitkamp, in: “Proceedings of the 12’h International Zeolite Conference” Baltimore, MD, USA, July 5-10, 1998, M.M.J. Treaty, B.K. Marcus, M.E. Bisher and J.B. Higgins, eds., Vol. 11,pp. 1447-1454, Materials Research Society, Warrendaie, PA (1998). [8] D. Fraenkel, M. Cherniavsky, B. Ittah and M. Levy, J. Cata/. 101, 273-283 (1986). [9] S.-J. Chu and Y.-W. Chen, /rid. Eng. Chem. Res. 33, 31 12-3117(1994). [10] M.J. Kamlet, J.-L.M. Abboud, M.H. Abraham and R.W. Taft, J. Org. Chem. 48,2877-2887 (1983). [11] J. F. Brennecke and C.A. Eckert, A/ChE ./. 35, 1409-1427(1989). [12] M. Neuber, Ph.D. thesis, University of Karlsruhe (1988). [13] G. Suld and A.P. Stuart, J. Org. Chem. 29, 2939-2946(1964). 278 ;’ DGMK-Conference “The Future Role of Aromatics in Refining and Petrochemisby”, Eriangen 1999 ~,’ ... ,, ;“’,.‘, ,. .-,:,’ ,.,;:, , , ,,, .,. ~f. ;.~,, M. Rep 1),A. E. Palomares 2),G. Eder-Mirth 1),J. G. van Ommen 1),J. A. Lercher 3) ...~, o Universiw of Twente, Faculty of Chemical Technology, Enschedet . .: .,,,; The Netherlands 3 lJniversidad Politecnica Valencia, ValenCja, Spain 3) Technische Universitat Mtinchen, Garching, Germany On Selectivity Aspects of the Alkylation of Toluene with Methanol over Zeolites Abstract The (co)-adsorption of methanol and toluene on and the methylation of toluene over H- ZSM5, Na-XandCs-X has been studied using i.xspectroscopy and therrnogravimetry. Adsorption resuks arc correlated to kinetic results indicating that successful side chain alkylation requires a ‘,,,-’. substantial interaction between toIuene and methanol with the cation and the basic lattice oxygens (resulting in methyl proton abstraction at elevated temperatures) and a balanced sorption stoichiometry. The polarity of the framework oxygens of CS-X induces decomposition of ,,;, ,,,” .1 intermediately formed formaldehyde to Hz and CO limiting the selective use of methanol as .,, ,. . ,/ .,,,./ methylating agent for side chain alkykrtion. In contrast, methylation of tohrene over H-ZSM5 and .,, ,, :,. ,,, Na-X yields ring alkylation products, resulting from a high preference of the catalyst for methanol . . :/ . .,,. ,S+ ,.,.,?,, adsorption. ,,,., 1. Introduction ,.. . The production of alkylbenzenes, such as xylene or styrene, has received significant ,, attention over the last decade. Styrene is commonly produced by dehydrogenation of ethylbenzene over potassium promoted iron-oxide catalysts requiring high stetam concentrations and high ,, temperatures to re~ch a satisfying yield [1]. While xylenes have been produced with liquid ,, ‘ ,,’ hydrofluoric or sulfuric acid catalysts [2], transition to solid acid catalysts is now being largely completed using molecular sieves such as modified MFI catalysts [3]. ,., ,. ~ , ,.’, ;, .,. ,-, ,.,:..,, ~, ,.. . ,’i ,. ,, ,,. !,., , .’”,,I :,,,-,,, . .,.,’, , .:,. ..,,,. .1 ,, J! ,“,~ ( ,,:.,,, .,:> 279 ISBN 3-931850-59-5 -, . .. . . DGMK-Tagungsbericht 9903, ., J- -—— . ,.-,... ,. ..,.: ,“,.. ,, , . Alkylation of tohrcnc.with methanol over acidic and alktdi cxchmrgcd zcolitcs leading to xylenes and styrcne respective y has been cxtcnsi vety studied in our laboratory. It has been shown that styrene yields are the highest over Rb+ and Cs+ exchmrgcd X zcolites [4], the most basic zeolites studied (following the Electrone@ivity EqualizzztionMethod of Mortier [5]). Alkylation of tohrcne with methanol over H-Z.SM5 Imds only to ring alkylation [6], with the enhancement .“ . of the selectivity to p-xylene by modi~ying the diffusiomrl resistance of the zeolites with 6. ,. postsynthcsis trciitment, such as the impregnation with metal oxides [2]. .. The present contribution focuses on the interaction between methanol and/or toluene with the zeolites CS-X, and Na-X and H-ZSM5 with the attempt to relate the sorption characteristics to [he selectivity in the alkylation reaction. 2. Experimental 2.1 Materials ZSM5 with a silicon aluminum ratio of 36 was synthesized according to a Mobil patent [7] and subsequent y four times ion exchanged at 368 K with a 0.1M NHJCI. For a dctai led characterization see Derewinski et al. [8]. Prior to the experiments the sample was activated at 600 K for 3 hrs (tempemture increment 10 K/rein). Alkali cations (Na+ and Cs+) exchanged X zcolites were prepared from a commercial Na-X zeolite (Fhrka; Lotnr. 69856; Si/Al rxtio of 1.3) by exchanging with 0.025 M aqeous solutions of the appropriate acetate (solidliquid ratio = 15 g.1”’)at 353 K, as described in a previous paper [4] ~~ctivation procedure: 10 K/rein up to 773 K for 1hr). For a detailed characterintion see P~lomares et al. [4]. For an easier compwison of the differences in zeolite composition the unit cell compositions as calculated from the Si/Al ratios are compiled in Table 1.Methanol and tohrene were obtained from Merck (p.is.)and used without further purification. Table 1: Chcmicd composition, Si/Al ratio, averaged charge (-5.) on the zeolite lattice oxygens and BET specific surt%ccarea of the samples investigated. Zcolite Unit cell composition Si/Al -60 BET (m’/g) H-ZSM5 Hz~3N%~~ALtiSig@lg~ 35 0.154 382 Na-X Na~.lAl~.lSilO@J~ 1.28 0.317 591 CsNa-X Csw lNa~ sAlw~Si104.8°384 1.~5 0.384 394 2.2 Infrared Spectroscopy The sorption and co-adsorption experiments of tohrcne and methanol were monitored bya BRUKER1FS-88 i.r. spectrometer (optical resolution 4cm-')equipped with a vacuum cell (base pressure below 10+mbar). The spectra were recorded in the transmission absorption mode. The zcolite powder was pressed into a self-supporting wsfer and placed in a cell equipped with IRtrmrsparent windows, where it wasanalyzed insitu during all treatments.Thew~fcrwas contacted with 10%mb~rofmethanol anWortoluenc vapor at308Kthrough ag~s dosing valve, until adsorption-resorption equilibrium was achieved (monitored by time-resolved i.r. spectroscopy). All i.r. spectra presented in this communication are difference spectra, ie., the spectrum of the activated zcolite is subtracted from the spectrum of the zeolite with the adsorbed molecules. hrthiskind ofpresentation all i.r. b~ndspointing upwards increased inintcnsity and 280 all pointing downwards decreasedin intensity, upon adsorption of methanol compared to the activated zeolite. 2.3 Reaclion studies For the iu situ reaction studies,an i.r.cellwith thecharacteristicsof a well stirred continuouslyoperatingtank reactor(volume = 1.5cm3) was used. To characterize the species sorbed on the material during the mtalytic reaction, i.r. spectra were recorded time resolved at different reaction temperatures (from 373 to 723 K) as the activated materia[ was contacted with the reactant stream. The feed gas composition had a partial pressure ratio of toluene/methanol= 45/15 (expressed in mbar) and He was added to 1 bar. The overall flow was 3.5 ml/min. This corresponds to a W/F ratio of 1.8104 g of catd ys.s/mol of toluene+methanol. Simultaneously, samples of the effluent gas stream were collected in the sample loops of a multi-port va[ve and subsequently analyzed by gas chromatography (a HP5890 II equipped with a capillary column DBWAX 30m and FID was used He was used as carrier gas). 2.4 Tbenuogravimetrie aualysis The therrnogravimetric measurements were performed on a modified SETARAM TG- DSC111 apparatus. The adsorption capacity fortoluene and methanol was determined at 103-100 mbar equilibrium pressure at 323 K. The equilibration was followed by means of weight changes (for Cs+ and Na+ exchanged X, 1 hour, respectively, several hours at low methanol partial pressures). For atypical experiment, approximately 15 mg of the samples was activated in situ by heating in vacuum (pc104 mbar) according to the procedure described above. 3. Results and Discussion Significant differences in the i.r. absorption maxima, the shape and the relative intensities of the CH and OH vibrational bands can be observed for methanol, when adsorbed on H-ZSM5 andalkali exchangedX (seeTable 1andFig. 1). Methanol adsorption on H-Z.SM5 at ambient temperature shows a broad OH band at 3545 cm-’ (seeFig. 1 and Table 1). The width of the band indicates strong hydrogen bonding of the zeolitc hydroxy group at 3610 cm-’ (broad structured bands at 2900, 2450 and 1690 cm.l) and weaker bonding of the methanol hydroxy group (3545 cm-’). The bands can be explained by hydrogen bonded methanol without the transfer of the zeolite proton to the methanol ,; -,,,. molecule (see Blaszkowski and van Santen [9] and H&se and Sauer [10]). In addition two sharp bands at 2958 and 2856 cm”] (asymmetric and symmetric C-H stretching vibration respectively) are observed. The 1 1 3s00 3000 2500 2000 1500 bands are similar to those observed for Wavmumbe,s [m-l] methanol sorbed on silica, where the methyl group is regarded as being free from lateral Fig.1: Adsoqxiou oj methauol ou H-ZSh45, interactions and to freely rotatein respect to Na-X aud CS-X. T = 308 K; p = IIJ~mbar. the molecules axis [11]. The higher 281 wuvenumber of the symmchic stretching vibration (2856 cm-’), comp~red to free (gas phase) mcthwrol (2841 cm” [12]), is attributed to a slight strengthening of the C-H bonds in response to the elongation of the C-O bond. Furthermore, two rotational b~nds can bc observed at 3012 (physical adsorption of methanol on silanol groups) and 2993 cm-’.These results indicate that for H-ZSM5 the intemctions between the lattice oxygen and the methyl group of methanol were minimal. With increasing temperature, elimination of water from methanol occurred (not shown here), resulting in a surface methoxy species as is concluded from the significant increase of the asymmetric and symmetric stretching vibration of the methyl group (2980 and 2868 cm-i, respectively [13]), leading [o the formation of dimethylether [ 13]. Table 2: W~venumbers of the Vo.lland VC.ll(stretching vibration) and the 6..,1 and 6C.1[ (deforrmuion) bands of sorbed MeOH (i.r. bands of gas phase methanol, see ref. [12]), Zcolitc Si/Al Vo.,, (cm”’) bo.,l(cm”’) Vc.li (cm-’) 6C.,,(cm-’) -60 [5] H-ZSM5 35 3545 16~0 2958, 2856 not resolved 0.154 Na-X 1.~8 3485,3345 1411 2957,2841 1475, 1453 0.317 CsNa-X 1.25 3244 1428 2941, ~8~0 1481, 1451 0.384 Methanol(g) - 3681 1340 2956,2844 1477, 1465, 1454 - After adsorption of methanol on NJ-X, an asymmetric broad band at 3354 cm-’ of the hydroxy group and a broadened CH asymmetric and symmetric stretching vibration band at 2957 and 2841 cm-’, respccti vely, is observed (see Fig. 1). Upon methanol adsorption on CS-X, a further lowering of tbe stretching vibration band of the hydroxy and methyl group is observed (3244, 2941 imd 2820 cm-’, respectively).The broudcning of the methyl and the hydroxy group of methanol, when adsorbed on Na+and Cs+exchanged X, indicates significant hydrogen bonding of the methyl and the hydroxy group with the b~sic framework oxygens as a result of the zeolitc composition and clectronegativity of the cation (see also Table 2). Note in this respect also the absence of the rotational bands of the methyl group at 2993 cm-’. Gravimetric studies have shown that broadening of the hydroxy bands is not caused by steric hindrance or methanol clustering [14,15]. Also the O-H and C-H deformation bands (see Table 2) indicate a large influence of the polarity of the zeolite on the wavenumbcr of the hydroxy and methyl group. Additionally, lateral hydrogen bonding is believed to exist between methanol hydroxygroups adsorbed on neighboring cations in the supcrcagc of the metal exchanged X zeolites, because of the high concentration of cations in these samples [14]. This strong hydrogen bonding, as observed with CS-X, is believed to bc responsible for the dchydrogenation of methanol at elevated temperatures. It should bc ;. emphasized that under the sorption conditions used, methanol decomposition into surface #.. . mcthoxy groups or formaldehyde w~s not observed on the used zeolites [14, 16]. Toluene adsorbed via its rc-ring on the proton and the alkali cations, as indicated by the Iargc rcd shift of the acidic OH vibration of H-ZSM5 and the blue shift of the momatic C-H stretching vibration upon adsorption (see Table 3), in agreement with Jentys and Lercher [17]and Mielczwski and Davis [18], respectively. The asymmetric shape of the perturbed OH stretching band of H-ZSM5 upon tohrcne adsorption suggests the existence of two different perturbations of the SiOHAl groups. A Iw-gcfraction of the acidic hydroxyl groups of H-ZSM5 is expcclcd to 282 , ,.’:,L.‘,‘;- !.,, ..-, ,,, ,.,’ .,-. ““,..;, .,-.. ,, :,, .. . . . be located at the channel intersections [19] and it is expectedthat the large void allows toluene to adapt an energetically more favorable structure than sites in the zig-zag channels. Adsorption of toluene on H-ZSM5 and Na-X induces a significant upward shift of the aromatic C-H stretching band compared to gas phase toluene [20] (see Table 3). In sharp contrast, upon adsorption on CS-X a significant smaller increase of the aromatic C-H stretching vibration is found resulting from the fact that(i) the large Cs+ cations overlap better with the x-system of tohtene than the small Na+and H+cations (which should lead to even higher wavenumbers than observed with H-ZSM5 and Na-X) and (ii) the methyl and aromatic hydrogens of toluenestrongly interact with the zeolite lattice oxygens (which Ieads to a decrease of the C-H stretching wavenumbers). This leads to a high heat of adsorption on CS-X [21,22]. Therefore, we conclude that strong interactions between the methyl group of tohtene and the basic Iattice oxygens of the zeolite and a large polarization of the toluene ring occum on CS-X, resulting in facilitated hydrogen abstraction from the methyl group of tohrene at elevated temperature. Table 3: Wavenumbers of the VW,,and v~ll (stretching vibration) bands of adsorbed toluene (i.r. bands of gas phase toluene, see ref. [20]). ,, Zeolite SilAl vO}{ (cm”’) VCII (cm-l) VC}I (cm”l) -60 [5] zeolite aromatic a]iDhatic H-ZSM5 35 3220 3086,3060,3030 2925,2877 0.154 .,. Na-X 1.28 - 3055,3024 2921,2859 0.317 ,“.,> CsNa-X 1.25 - 3048,3021 2916,2857 0.384 ., toluene (g) - 3080,3040,3010 2960,2930,2880 - ,,, Methanol Toluene ‘0i@eeYe7 ‘if-?--w -’. Q 0 002 004 0C6 0c2 01 0 02 a4 06 Q8 1 pressure (mbar) pressure (InL@ Fig. 2: Adsorption isotherms for the aa?ro~tion of methanol (a) and toluene (b) on the used zeolites. T = 323 K; p = lU~-1 mbar. (H-Z$M5: as deducedfrom i.r. data) [../.. t., ,,, The adsorption isotherms for methanol and tohrene are shown in Fig. 2. From these figures it can be seen that the sorption per site on H-ZSM5 for the adsorption of toluene and methanol is significantly larger at 1 and 101 mbar, respectively, compared to Na-X and CS-X, as also found for methanol on Na-ZSM5 [14,15]. And, while the sorption capacity for toluene on Na+and Cs+exchanged X is the same (at 1 mbar), differences are observed for the adsorption of ,. methanol on these two samples. 283 —- .’, ,, . . . ,: ‘. ..-, Toluene adsorbs up to 0.3 molecules per site (at 13mbar (not shown here)) in CS-X and Na-X, suggesting that primarily the cations located in site III’ are responsible for the adsorption (=40% of the Cs+and Na+cations per unit cell, respectively [23,24]), in agreement with the heat of adsorption (ie., see ref [21,22]). FIJI1loading of the cations in site IIf’ cm not be obtained at 13mbar due to steric hindrance. Interaction of the cations located in site II with adsorbed toluenc can not be excluded. Whercm methanol adsorption on CS-X reaches equilibrium quickly (at 0.03 mbar), the amount adsorbed on Na-X at low methanol partial pressures is much smaller and reaches equi Iibnum only at higher partial pressures than observed for CS-X. This is related to the fact that the smaller Na+ctttions are better protected by the zeolite framework. At 0.03 mbar only 337oof the Cs+cations in CS-X intemct with methanol. These results also indicate that methanol does not cluster inside the zcolite pores under the conditions used in the i.r. adsorption experiments. Thus, clustering of methanol can definitely be excluded as the reason for the broadening of the methanol hydroxy and methyl bands with CS-X and Na-X. Co-adsorption of tolucne and methanol show that toluene is mainly Absorbmcc (IU ) adsorbed on CS-X, while on H-ZSM5 and Na-X the main adsorbed species is methanol (see Fig. 3). Note also that upon co-adsorption of methanol and toluene on H-ZSM5, a large red shift (Av = 85 cm-’) of the methanol OH band WJS observed suggesting a bimolecular sorption complex between the reactants. The asymmetric . shape of the McOH band indicates that not all molecules interact with tohsene. This can be explained by steric hindrance, induced by adsorbed methanol ~~pertureof I I H-ZSM5: 5.6x5.3~; kinetic diameter 3s00 3000 2500 2000 1s00 toluene: 6.1A). As the positions of the \Vavcnumber (cm ‘) absorption bands of the perturbed zeolite hydroxy b~nds (2900, 2450 and 1690cm”’) Fig. 3: Co-mLsorptiou of to[mme und me fhanol are not altered upon adsorption of tolucne, 011H-ZSM5, Nu-X and CS-X. T = 308 K; wc conclude that this intemction is hardly p =5x 10”~mbar. influenced by the wcttk co-adsorption. On Na+and Cs+exchanged X zeolites a mutual influence of adsorb~d methanol and tohsene was not observed (see FIR.3). The wavenumbers of the single adsorbed components and the co-adsorbed molecules wer~identical. Note, however, that the stoichiomctry depended on the zeolite composition. In a detailed previous study we reported that the sorbed phase on Na-X contained 80% methanol and 20% toluenc, CS-X 3370 methanol and 677o tohsene [22]. Fig. 4 displays the difference i.r. spcctm during the reaction as a function of the temperature on (a) H-Z.SM5 and (b) CS-X. Kinetic studies in combination with i.r. spectroscopy have shown that the reacting species on CS-X is formaldehyde/formate [22] formed by dehydrogcnation of methanol at elevated temperatures (as observed by the appearance of a band at 1610 cm-’). On H-ZSM5 reaction seems to occur from the bimolecular methoxonium-tol uene adsorption complex being stable up to 473 K and disappearing with the simultmreous appemwrce of xylene in the gas phase. Alkylation might occur in a concerted step or be preceded by decomposition of adsorbed mcthwrol, resulting in asuti~ce mcthoxy species, which subsequently 284 alkylates the aromatic ring [25]. During the temperature programmed reaction of methanol and tohtene on CS-X (see Fig. 4b) and Na-X (not shown here, see ref. [4]), a band at 1661 and 1687 cm-[ respectively indicates the formation of water, which results from methanol decomposition, in agreement with Bertsch and Habgood [26]. The disappearance of this band at 523 and 513 K for CS-X and Na-X, respectively, is correlated to the onset of the formation of dimethyl ether (surface methoxy species react with gas phase methanol) and formaldehyde (dehydrogenation of physisorbed methanol by the framework oxygens). As can be deduced from these results, formaldehyde is also formed on Na-X (appearance of a band at 1616 cm-l), in good agreement with Philippou and Anderson [27]. However, the sorption stoichiometry (hardIy any toluene is adsorbed on the surface) prevents the side chain:alkylation (selectivity ring aIkylation/side chain alkylation on Na-X = 95/5). For CS-X, the ratio ring alkylatiorrkide chain alkylation was found to be 10/82. At higher temperatures further decomposition of formaldehyde to H. and CO is observed under re~ction conditions [4,22]. - 1651 15S9-1585 1610 G-x J,.;{:, ., :’.,, ., ., , r. .: ,, ,, -., ,, .,.- ,,; 1 .’, .’. ‘, . .,..,,,:<.’ . :,, ..,., ,. ), .-,-. .,,’ . . . . ,:, : ,,.., . ,, . ,,. :, ,: 3s00 3600 34W 3203 3tS13 28(N181XI 1700 1602 1503 1400 ,, .,,., ,. ,, ... ,,” Wawmders [cm-l] Wavcnumtas [cm-l] .-. . .,, ..,, ,.. Fig. 4: Methylation of toluene caialysed over H-ZSM5 and CS-X as a fltnction of the .-.“. ,;:: ,.: temperature, shown as dl~erence i.r. spectra. T= 373-773 K. ,,’< ,...,. :> .~., ,.. , ,“.’’.,, ., .,-. . “,, ,’,.” ~ ,,. ,.’ .. 4. Conclusion -, From these results, we conclude that for a successful side chain alkylation catalyst a substantial interaction between toluene and methanol with the cation and the basic lattice oxygens (resulting in methyl proton abstraction at elevated temperatums) and a balanced sorption stoichiometry has to exist. However, as a result of the large basicity of the framework oxygens of CS-X, further decomposition of formaldehyde to Hz and CO is observed under reaction conditions. This decomposition of methanol/formaldehyde currently limits the selective use of methanol as methylating agent for side chain alkylation. In sharp contrast, H-ZSM5 and Na-X catalyze ring alkylation, which results from a high preference of the catalyst for methanol t ,. adsorption, i.e. the smaller the cation, the higher the preference for methanol adsorption. ,,, ;!,, ,,:j ; . . ;,,.”,;- ‘.: .,.~:’,:::l Acknowledgement ,.. ,’!’/,, ,.:,. $; Financial support of SONiNWO for the benefit of the workgroup 336-017 is gratefully ,, :;.:.,,, ,::,: / ,!‘,,+:};:,;,, .1 acknowledged. We thank Drs J.A.Z. Pieteme and Dr G. Nivarthy for helpful discussions ,, ?):,.:’., “-l ,,, ::,. .,,;, ,,: r,,..,. -i .::i;,”;s,’.;’” ,I,:,,if,.,.,,, , t,,-,’,>,,,;, r,.:I 285 ,,.,L:,~,.,~ ,. ,> ,,,: ‘ ,< concerning the gmvirnetric measurements and for mmsuring the adsorption of tolucne. This work was performed under the auspices of NIOK, The Netherlands Institute for Catalysis and PIT, Process Technology Institute Twente, The Netherlands. References 1.P. Ciambelli, R.P. Dasio imdM. Turco,Proc. 9’hInt. ZeoL Conf., Montreal, Eds., Van Ballmoos et al., 535 (1992). 2. P.G. Smimiotis and E. Ruckenstein, Ind. Eng. Chem. Res., 341517 (1995). 3. K. T~nabe and W.F. Holderich, Appt. C~tal. A: General, 181399 (1999). 4. A.E. Palomares, G. Eder and J.A.Lercher, J. Catal., 168442 (1997). 5.(@ W.J. Mortier, J. CJtaL, 55 (1978) 138. (b) R.T. SanderSon, J. Am. Chem. Sot., 1052259 (1983). 6. G. Mirth and J.A. Lercher, J. CataL, 132244 (1991). 7. R.G. Argauer and R.G. Landolt, U.S. Patent 3702886 (1972). 8.. M. Derewinski, J. Haber, J. Ptaszynski, J.A. Lcrcher and G. Rumplmayer, Stud. Surf. Sc. Card., 28957 (1986). 9. S.R. Blaszkowski and R.A. van Santen, J. Phys. Chem., 9911728 (1995). .’. ” 10. F. Hxme and J. .%uer, J. Am. Chem. Sot., 1173780 (1995). 6,,. 11. A.E. Pelmenschikov, G. Morosi and A. Gamba, J. Phys. Chem., 962241 (1992). 12. A. Serrallach, R. Meyer and Hs.H. Gunthard, J. Molec. Spectrosc., 5294 (1974). 13.T.R. Forester and R.F. Howe, J. Am. Chem. Sot., 1095076 (1987). 14. M. Rep, A.E. Palomares, G. Eder-Mirth, J.G. van Ommen and J.A. Lercher, to be published. 15. M. Rep, A.E. Palomares, J.G. van Ommen and J.A. Lercher, Proc. 12* ICC, submitted for publication. 16. G. Mirth , J.A. Lercher, M.W. Anderson, and J. Klinowski, J. Chem. Sot., Faraday Tmns., 863039 (1990). 17. A. Jentys and J.A. Lercher, Stud. Surf. Sc. Catd., 46585 (1989). 18. E. Mielczarski and M.E. Davis, Ind. Eng. Chem. Res., 291579 (1990). 19. J.-C. Lin, K.-J. Chao and Y. Wang, Zeolites, 11376 (1991). ~().J.K. Wijmshumt and H.J. Bernstein, Can. J. Chem., 35911 (1957). 21. J. Xie, M. Bousmina, G. Xu and S. Kaliaguine, J. Molec. Cattd. A: Chemical, 135187 (1998). ~~. A.E. pajomurfi, G. ~er-Mirth, M. Rep~nrJJ.A. Urcher, J. Catal., 18056 (1998). ~3. Yu,F. Shepelev,I.K.Butikova,andYu.L Srnolin, Zcolites, 11287 (199 1). 24. D. Olson, Zeolites, 15439 (1995). 25.1,1. IvanOva and A. Corms, J. Phys. Chem. B, 101547 (1997). 26. L. Bertsch and H.W. Habgood, J. Phys. Chem., 671621 (1963). 27. A. Philippou and M,W. Anderson, J. Am. Chem. Sot., 1165774 (1994). 286