Chemid Engineering Science. Vol. 47, No. 9-11, pp. 2313-2318, 1992. Printed in Great Britain.

0009-2509/92 $%oO+O.OO 0 1992 Pergamon Press Ltd

FLUIDIZED BED REACTORS FOR PARAFFINS DEHYDROCENATION

Domenico Sanfilippo, Franc0 guonomo, Giorgio Fusco, Maria Lupieri and Ivano Niracca

Snamprogetti S.p.A. Via Maritano 26, 20097 S.Donato Milsnase (Ml), Italy

Abstract

The development of processes to dehydrogenate =3-=4 pareff ins to the corresponding mono-olefins has been pushed by the market demand. A technology based on a bubbling fluidized bed reactor-regenerator system is proposed to solve the problems involved in the industrial application of such processes.

1. fntroduction

In the Last decade various situations in the market of chemicals have focused interest in the catalytic dehydrogenation of light paraffins to the corresponding mono-olefins. The increasing demand for methyt tert-butyl ether (MTBE) as gasoline pool component has induced such a heavy use of the traditional sources of isobutene (steam cracking and FCC), that its availability has become a bottle-neck in the expansion of MTBE market. The demand for polypropylene is running faster than the demand for polyethylene: the tno starting monomers are usual Ly coproduced in steam cracking furnaces and, even feeding pure propane, a substantial conversion to ethylene is achieved. Such problems have led to the development and commercialization of various processes able to dehydrogenate with high yields C3-C5 paraffins. The nature of these react ions and of the avai table catalysts set the problems to be solved in the development of dehydrogenation technologies.

2. Thermodynamic and kinetic constraints.

The reactions of dehydrogenation are strongly endothermic and the conversion per pass is Limited by equilibrium. For C2-Cl5 paraffins, temperatures in the range from 800 to 1000 “K are

needed to reach equilibrium conversions around 50%: lighter paraffins need higher temperatures. Being the energy required by the splitting of an hydrogen molecule almost independent from the molecular weight of the paraffin (113-134 kJ/mol), the heat per unitary weight of product to be supplied increases significantly decreasing the molecular ueight of the feedstock. The high temperatures required to reach acceptable levels of conversion per pass f evour side therma 1 cracking react ions to tighter paraffins and coke, decreas i ng the selectivity to the desired product: the catalyst needs to be periodicaLLy regenerated burning the coke deposited on its surface. Finally it is clear from stoichiometry that high pressure adversely affects the reaction.

2313

2314 DoMENTC~ SANFILIP~O et al. Al9

3. Technological implementation

The main problems to be solved in the development of the reaction section of an industrial dehydrogenation process are: heat supply to the reacting system at temperatures uell above 800 K, careful control of temperature to minimize by-products and meximi re yields, necessity of catalyst regeneration. The constraint is to approach es close as possible a continuous process. TUO kinds of catalyst have found industrial application: chromia- or platinum-based ones. Apart from the tendency of platinum to some skeletal isomerization, the two catalysts show comparable activities in the experimental field of interest. Different solutions have been proposed, depending mainly on the know-how of the various companies in related fields. Typical examples are reported: - Adiabatic fixed bed reactors in cyclic operation, quickly alternating

reaction and regeneration periods (Penny et al., 1984). Reaction heat is supplied to the catalyst in the regeneration step conveying hot air in excess. Temperature in the reaction zone iS never stationary but decreases during the reaction period and its profile along the bed is continuously changing. Being each reactor batchwise operated from the view-point of the solid phase, to achieve cant i nuous operation, severa 1 reactors in parallel are used.

- Adiabatic mov i ng bed reactors in series with catalyst downflow (Pujado and Vora.1990). Heat of reaction is supplied to the gas in furnaces by intermediate heating between reactors: very high Wall temperatures are required to SUPPLY the necessary amount of heat. The catalyst is pneumatically conveyed to the regenerator and then back to the first reactor.

- Tubular fixed beds placed in a furnace for heat supply in a way very similar to steam-reforming processes <Rock and Dunn,1988). Severa 1 reactors in pare1 Let are used, one of these being alternatively in regeneration.

- The fluidized bed snamprogetti-Yarsintez process below described.

4. The Snamprogetti/Yarsintez process.

A technology based on the principle of bubbling f luidired bed has been originally developed in Soviet Union and, recent Ly, jointly improved by Snamprogetti and the Russian company Yarsinter. The catalyst used is promoted chromi a alumina < Buonomo et ISOBUTANE DEHYDROGENATION PROCESS al., 1985) having a microspheroidal shape with average diameter Less than 2 lo-4 m and with density less than 2000 kg/m3, classifiable in the tA’ group according to s Geldart <Geldart,1973); IA’ or aeratable powders show the best f luidizetion features: FCC catalyst is also classifiable in this group. A typical process scheme iS

reported in fig. 1 for the case of isobutane feedstock. _ Frg. 1

Fresh feed is vaporized, mixed with a recycle from an <e.g. MTBE), preheated in a gas-gas feed effluent heat exchanger and fed to the reactor on the bottom of the catalyst bed through a distributor giving

A19 Fluidized bed reactors for paraffins dehydrogenation 2315

a pressure drop able to ensure a perfect distribution of the gas in the whole cross section. Reaction products are separated from the entrained catalyst powder by high efficiency cyclones and after a complete dust elimination in Fl, go to compression and purification section to separate Cq stream from hydrogen and byproducts. The heart of the process is the reactor/regenerator system uhich is described in fig. 2. The reaction is carried out in

ISOBUTANE DEHYDROGENATION CYCLE a catalytic bubbling fluidized 4 7

bed operating at 800-900 *K atnllu t l-t- fd and 120-150 kPa. 2 ~~riCpi~ ni trcQm 3 lift- Reaction heat is supplied by 4 r_tOr Dutullt 5 raca-mm~ml the heat capacity of the solid e. -cat 7 catalyst, continuously

flla gas 8 r-rcltiol air D ,._mI

circulating from the bottom of IO ‘Irlpplng “‘trqp the reactor to the top of the I L-p* regenerator and vi ceversa R r-l!=

q dDoorp,lDl through an U-bend piping c trum,*r llrm 0 roQDmrPllCn system, originating a E rac%cllon F -rptsCIl countercurrent movement of gas t *Ia-m,o~ llre and solid either in the

H cyla-eD#c%1a.9 Fig. 2

reactor or in the regenerator. In the regenerator, catalyst restores its initiat pet-f ormances by burning the small quantity of coke formed on it during the reaction: additionally some fuel is burned directly on the catalyst to satisfy the overall thermal balance, as uill be detailed Later. Before being conveyed to the regenerator the catalyst is stripped with nitrogen to avoid loss of adsorbed products. The same operation is performed on the bottom of the regenerator to avoid transport of oxygen to the reactor with loss of selectivity.

5. The fluidized bed choice.

Conceptually, a bubb 1 i ng fLuidired bed system has particularly desirable features for dahydrogenation processes. The chance to organize reaction and regeneration in two separate vessels uith pneumatic conveying of solids betueen them, allows a truly continuous process. In this way there is no need for a Purge phase be tneen regeneration and reaction and contact of the feedstock with oxygen is avoided. Heat can be supplied to the reaction directly by the hot regenerated catalyst in a very efficient way. No heat exchange surf aces are needed: the extremely high area of sol ids exposed to the fluidizing gas makes almost instantaneous the heat transfer from the hot particles to the surrounding gas and its dispersion through the whole bed. A constant temperature is achieved throughout the bed, aasi ly controllable varying the hot catalyst flourate. Being the dehydrogenation reaction limited by equilibrium, the achievement of optimal yields and reactor volumes would be favoured by a plug-f Lou movement of the gas phase. The solid recirculation induced by the rising dilute phase bring about a certain degree of gas backmixing: it is mainly function of the degree of adsorption of the gas on the particle (gohle and van Suaaj,1978) and of the ratio betueen the superficial gas velocity and the minimum fluidization velocity <Nguyen et a1.,1978>. Uith Galdart’s aeratabla powders, minimum fluidization velocity is very Lou <usually less than 1 cm/s) and gas backmixing not negligible: a CSTR-like behavi our is approached also from the view-point of the gas phase.

2316 DOMENICO SANFILIPH) er al. Al9

The problem has been solved by the insertion of internals in the bed. limiting the fully free movement of solids. In this way the bed is no longer isotherm and an axial temperature prof i le is established ui th the highest temperature on top where hot catatyst arrives from the regenerator and the Lowest on the bottom where fresh gas is fed: this profile, along with countercurrent overal 1 flow of 98s and sol id, allows optimal yields. In fact, a temperature difference of 25-50 degrees from top to bottom in the reactor is established: so, fresh gas meets catalyst at relatively lou temperatures reacting with maximum selectivity to isobutene: in the final part of the bed, at higher temperature, a big push to conversion is given, uith thermal conversion of some byproducts. Fig. 3 shows a qualitative ISOBUTANE DEHYDROGENATION comparison of the different BUAtLTflTIVE TEI-IPERATURE PROFILES temperature profiles in the various kinds of reactors ‘a previously described: it is on 0 clear that the bubbling 5 .?’ __.’ f luidired bed with internals ~ : o: equilibrium ourva 8110~s an easier achievement ,_:. b: fluidized bed

of high levels of conversion, c: fixed bed approaching the equilibrium yp. d: ,ul tibed

line at higher temperature. The temperature prof i le also f0- allous the sol ids circulation 650 ax, 750 to be considerably reduced u3nperoture c Fog. 3 uith respect to the isothermal bed: the mean residence time of the solid particles in the reactor is increased and more heat per unitary weight of sol id is transferred to the reacting gas: consequently, the mechanical life of catalyst is increased. Let us formulate a thermal balance around the reactor uith the assumptions summarized below: -Feed: 100% isobutane -Temperature variations between reactor inlet and outlet: -Gas: 50 K -Catalyst: 100 K -Catalyst specific heat: 0.9 kJ/ K/kg The overal 1 reaction heat (dehydrogenation plus undesirable reactions) to be supplied is 62.9 kJ/mol of isobutane; neglecting thermal losses, the necessary cataiyst circulation rate is:

C=62.9/0.9/100=0.7 kg/mot isobutane=lZ kg/kg isobutane. Conveying of aerated catalyst allows such huge circulation, o imposing lower density in the $ upward part of the transfer g tines. It means that heat i requirements (not catalyst deactivation) dictate catalyst P

recirculation rate (fig. 4); ! owing to the different 2 feedstock, coke formation is 8 much slower then in FCCU: it .? also means that heat produced ‘,

; burning superficial coke is by 3 far louer than the heat requirements of the overal 1 reaction:it is therefore

,D CATALYST TO FEED RATIO

0

necessary to burn some fuel in the regenerator to transfer further heat to the catalyst _ Once again this is achieved in the efficient uay alloued by

A19 Fluidized bed reactors~ for paraffins dehydrogenation 2317

f luidired bed technology. These concepts can be quantified through a thermal balance around the regenerator: the burning of coke and fuel must SUPPlY, always neglecting thermal losses, the heat required by reactor plus the cnthalpy increment of gases through the regenerator itself: the tote1 amount is 69.5 kJ/mol of isobutane. Dehydrogenating isobutane, the selectivity to coke is around 2X by weight: burning this coke, 18.5 kJ/mo,l of isobutane are supplied, corresponding to about 25% of the therma 1 requirements. The remaining 75% of thermal contribution must be supp1ie.d burning a fuel gas:. using pure methane, due to its high combustion enthalpy on(y 0.06 mq$& _ of iethane/nole of isobutane are needed. I kg=+&;. + Another much favourable .fea ture of fluidized. l&as “it& the chance to substitute part of the catalyst while the plant i rr’unfli ng ,’ In this way the problem of satalyst agei ng with var i at;i on of yields is eliminated: periodicai substitutions are made, Of!30 altouing for the fines lost through the cyclones, so -that the average_ age &--he catalyst remain unchanged: other solut i or proposed

‘. every suffers &f catalyst decay. Reactor sizing requi res that the aspect Ila+-)O (fluidited bed height/internal diameter) be high enough to ensure a good gas distribution over the entire section (Geldart and Baeyens, 1985). But higher aspect ratio means higher gas velocity with the risk of excessive entrainment and overloading of the cyclones. T h e’ compromise must allou for the gas-sol id contact time required to reach the desired yield. The reactor is usually designed for the maximum predictable gas f lourate, allowing a remarkable flexibility toward turndoun rates. Very high capacity can be reached on a single train.

6. Process demonstration unit runs.

The improved catalyst has been thoroughly tested in Yaroslavl (USSR) in a process demonstration unit hydro-dynamically homologous to an industrial plant. The inf Luence of the main operating parameters on conversion and selectivity is briefly reported in fig. 5 for isobutene production. The figure shous how conversion and selectivity change by varying one by one some reactor operating parameters from a central reference point. Reaction temperature has the heaviest.influence, as stressed before; space

ISOBUTANE DEHYDROGENATION INFLUENCE OF OPERPTING PFlRAMETERS

ON PERFORMRNCES

a: temper-o ture C K 1 b: space velocity c : pressure

velocity has the same kind of V.W 0.80 0.65 0.90 0.85 1 .m* 1.05 I .I0 1.15 1.20

influence than temperature CONVERSION (orb, trory -its)

even though damped, calling for a compromise between conversion and selectivity. Pressure must be kept as Lou as possible to enhance selectivity.

7. Conclusions.

Utilizing the features of bubbling fluidized beds, it has been possible to achieve a very efficient and economical solution to the problems concerned with the development of paraffins dehydrogenetion processes.

23i8 DonseNlco SANFILI~~~ et al. Al9

Summarizing. the Snamprogetti/Yarsintez process allows: - Stssdv tC?S”D‘kratUre profile

with the highest values at DEHYDR6GENATlON PLANT EMISSIONS the reactor outlet uhere the highest conversion is required. ContTnuos production <no cyclic operations). Emissions environmentally compatible with practical absence of nitrogen gxides, CO <no fired heaters used), and perticulate <fig. 6). H i gh on-stream factor and steady operating condtitions by periodical substitution of catalyst.

Fig. 6

- Intrinsic safety, bci ng the oxidation anY reaction zones physically separated and the operation at positive prerrure:

ast - .045 m!yNm-3 m - <lOwm Mx - Tppn He - c loo ppn M OTI-ER Cl-EtdI~s

References

Bohle, U. and van Suaaj, W.P.U., 1978, Fluidiration CEds. Oevidson,J.F. and Keairns.0.L.). Cambridge Univ. Press, 967. 0uonomo.F. et al., USP 4,746,643, 1985, to Snamprogetti. Geldart. D _, 1973, Powder Technol.. 1, 285. Geldart, D. and Baeyens, J., 1985, The design of distributors for gas-fluidized beds, Powder Technol. 42, 67-78. Nguyen, H.W., Whitehead, A.B. and Potter, O.E.,1978, Fluidization <Eds. Davidson,J.F. and Keeirns,D.L.), Cambridge University Press, 140. Penny, S-J., Clark, R.G. and Spence, 0-C.. LPGS - Putting a Canadian resource to a nen use - Part 2 - LPG dehydrogenat ion - optimizing LPG, 1984, Can. Chem. News 34 C3), 19-24. Pujado, P.R. and Vora, B-V., 1990, Hake C3-C4 olefins selectively, Rydroc. Proc. 69 (3), 65-70. Rock, K.L. and Dunn, R.G.. 1988, Economics of producing butene intermediate feedstocks from butane, Ener. Prog. 8 (4). 191-195.