Ind. Eng. Chem. Res. 1994,33,3025-3030
3025
Is the Catalytic Monolith Reactor Well Suited to Environmentally Benign Processing? Jacques Villermaux' Laboratoire des Sciences du GLnie Chimique-CNRS, Ecole Nationale Supdrieure des Industries Chimiques-ZNPL, B.P. 451, 54001 Nancy Cedex, France
Daniel Schweich Laboratoire de Gdnie des ProcddLs Catalytiques-CNRS, CPE-Lyon, B.P. 2077, 69616 Villeurbanne Cedex, France
The potential use of monolithic reactors for carrying out gas-phase catalytic reactions is discussed.
A comparison is made between the monolith and the classical fixed bed reactor, from which it is concluded that the monolithic structure offers a convenient way to put a given amount of catalytic material into contact with a flowing reactant avoiding pressure drop and internal masstransfer limitations. Heat- and mass-transfer effects are discussed, and it is shown that unusually low Sherwood numbers for external mass transfer reported in the literature might result from a residual contribution of heterogeneous reactions, even under well established masstransfer control.
General Background Intrinsically clean processes require very high atomic selectivity [251,i.e., optimal use of reactants t o make the desired products and almost no byproducts. This can be achieved by organometallic catalysis in solution with the disadvantage that noble compounds and solvent have to be carefully recycled to avoid losses. Vaporphase heterogeneous catalysis offers an alternative possibility as a solvent-freeoperation in which the active material is immobilized. On the other hand, achieving high selectivity requires reduced mass-transfer limitations and very precise process technology [301. The requirements for an efficient reactor may be listed as follows: (1)plug flow; (2) high active catalyst loading; (3) possibility of using a large variety of catalyst materials; (4) low manufacturing cost; ( 5 ) low pressure drop; (6) easy stability and temperature control; (7) no external heat- and mass-transfer resistance from the bulk fluid to the catalyst surface; (8) no internal heatand mass-transfer limitations within the catalyst; (9) efficient radial heat and mass transfer; (10) possibility of efficient heat exchange with the environment; (11) mechanical resistance and durability; (12) easy catalyst regeneration. Obviously, no existing reactor exhibits all these advantages simultaneously. The classical fixed bed, with an appropriate design, can meet most of these requirements except points 5 and 8 which are contradictory when very fine catalyst particles have to be used. In order to overcome this difficulty several alternate designs have been proposed, namely the honeycomb reactor, the parallel passage reactor, the lateral flow reactor, and the bed string reactor, which are discussed in ref 3. The so-called honeycomb reactor or monolithic reactor consists of a honeycomb structure whose parallel channels (typically 30-60 per square centimeter) are coated with a thin layer of porous catalyst carrier impregnated with the active catalytic material. The main advantages of this kind of reactor are points 5 and 8 above-low pressure drop and absence of internal mass-transfer limitation. Its main disadvantages at least for ceramic monoliths are related to points 9 and 10, and perhaps 12. Other points of the list above can
be met with monolithic reactors, in spite of what has been sometimes stated [31. In particular, sophisticated methods have been developed t o achieve controlled coating of the channels with various materials [181. Monolithic reactors are very popular for pollution abatment requiring dust-proof, low pressure drop selective heterogeneous catalysts [6,21,22,261.They have been extensively studied and modeled for automobile emission control [ l o ,19,321. Irandoust and co-workers have reviewed the potential interest of monolithic catalysts for nonautomobile applications [131 and have published a series of papers showing that monolithic reactors could also be used for heterogeneously catalyzed gas-liquid reactions [7, 1 1 , 12, 14-161. These reactions may be carried out in segmented slug flow produced by alternatively feeding the channels with gas and liquid from an appropriate distributor. Mass transfer from gas to liquid and from liquid to solid is enhanced across the thin liquid films which are forming between the slugs and within these slugs as they move along the channels [151. Scale-up therefore seems easy by simple increase of the honeycomb diameter, keeping the channel size constant. However, the liquid distributor is a key issue of the design. This renders the monolithic reactor attractive for wastewater treatment, and several applications have been described in this area and for bioprocessing [4,17,241. Recently, Cybulski et al. have proposed t o use a monolithic reactor for liquid-phase methanol synthesis at commercial scale with distinct advantages over more conventional slurry columns, autoclaves and trickle-bed reactors El. The aim of the present paper is to discuss two features of the monolithic reactor for vapor-phase catalytic reactions. In the first section, we shall compare the behavior of a monolithic and of a fixed bed reactor for a "low demanding" quasi-isothermal reaction. In the second section, we shall discuss a paradox which arose in the study of more demanding reactions controlled by mass transfer from the gas phase: the apparently very low values found for the Shenvood numbers characterizing this transfer. In this study, we shall limit ourselves to the case of ceramic monoliths, commercially available as cordierite (2Mg0*2Al203*5Si02) or y-alumi-
0888-5885/94/2633-3025$04.50/00 1994 American Chemical Society
3026 Ind. Eng. Chem. Res., Vol. 33, No. 12, 1994 Table 1 Monolithic Reactor (See Figure 1) 1.12 x 1.12 mm square channels d = 0.92 mm free passage (e = 0.1 mm) €1 = 0.525 global porosity (free volume) wash-coat (catalyst) volume fraction ccatl = 0.253 a1 = 2282 m2 m-3 specific external surface area L1=lm length ecat= 1500 kg m-3 catalyst (wash coat) Fixed Bed Reactor d, = 1.8 mm spherical pellets packing porosity €2 = 0.4 catalyst volume fraction ceat2 = 0.6 a2 = 6d,-l(l - €2) = specific external surface area 2000 m2 m-3 length Lz=lm ecat= 1500 kg m-3 catalyst (pellets) internal macroporosity cp = 0.5 tortuosity tp= 2 Gas-PhaseProperties (Operating Conditions,p x 1 bar) e = 1.20 kg m-3 density ,u = 1.8 x Pa-s dynamic viscosity Y = 1.5 x m2 kinematic viscosity GJ = 1.5 x m2 molecular diffisivity (Sc = 1)
time defined from the interstitial velocity u at the reactor inlet. The equality of the WSV imposes that
be the same for both reactors, from which one deduces that t1 = 1.245 s and t 2 = 0.4 s. The corresponding velocities at reactor inlet are, for the monolith u 1 = 0.803 m s-l (interstitial), u01 = 0.42 m s-l (superficial) and for the fixed bed u2 = 2.5 m s-l (interstitial), u02 = 1 m s-l (superficial). Let us first calculate the pressure drop per unit length a t reactor inlet. For the monolith
-(dz), - fZi e 4d dp
-
The Reynolds number is Re1 = uldlv = 49 indicating that the flow is laminar. Under these conditions, Leclerc [201 has shown that Redl = 57. This leads to
-E),
= 487 Pa m-l
which is very small and illustrates the fact that the monolith is a very low pressure drop reactor. In the case of the fxed bed, we may use the KozenyCarman equation
f2
=
1T 1.75 ( + 150E2
(4)
€2
The Reynolds number based on the superficial velocity is Re2 = uz,&fv= 120 from which it is deduced that c
1.27 mm
-,
Figure 1. Sketch of a channel showing the wash-coat deposit.
na. Metallic monoliths-should we say “monosiderits”?-have also been manufactured by rolling up a corrugated sheet of metal [ 2 3 , 2 n . This provides axial heat conductivity which may be exploited to control thermal stability or heat storage in transient regime such as the periodic reversed flow operations [81. This constitutes a distinct advantage over conventional fixed beds. In order to discuss the respective merits of both types of reactors on a concrete basis, let us compare their performances in the case of a first-order isothermal reaction exhibiting no thermal effects. Table 1 summarizes the features of both reactors, designed to operate with the same weight space velocity WSV = 900-’ m3(NTP)kgCat-ls-l and to achieve 99% conversion; these values were arbitrarily chosen for the sake of illustration. The size of the channels is typical of a monolith with 62 channels/cm2, the walls are 0.15 mm thick, and the wash coat is 0.10 mm thick (Figure 1). The fixed bed is packed with spherical porous pellets of the same material as the wash coat, designed to exhibit nearly the same external surface area per unit reactor volume as the monolith, i.e., 2000 m2 m-3, which corresponds to d, = 1.8 mm. Let t = W u be the space
-(g)2
= 97656 Pa rn-l
based on NTP conditions
(5) which is much higher than for the monolith. Accounting for the compressibility of the gas, one finds an inlet pressure of 1.7 x 105Pafor an outlet pressure of 105Pa. This is a first disadvantage of the fixed bed. Under ideal conditions, 99% conversion for an isothermal first-order reaction without volume change requires that the rate constant in chemical regime be equal to
k, = WSV In 1 = 5.12 x
l-x
m3 kgcat-1 s-1 (6)
The rate constant referred to unit catalyst material volume is thus
k, = k d C a t= 7.675 s-l
(7)
Is there the possibility of external mass-transfer limitation? The value of the Shenvood number in square channels in laminar regime has been the subject of extensive work [IO,191. Abnormally low values are often reported in the literature [27,29]. This controversial point will be further considered in the next section. For usual
Ind. Eng. Chem. Res., Vol. 33, No. 12, 1994 3027 applications, it is generally recommended t o adopt Shl = kDld/@ % 3. From this, the mass-transfer coefficient can be estim s-l and the fraction of mated to kD1 = 4.9 x external resistance calculated as fel
= k$/kDl = 1.57 X
(8)
which indicates that external mass transfer is not limiting. In the fixed bed, the standard Ranz-Levenspiel correlation may be used:
Sh, = 2
+ 1.8Re21’2Sc1‘3= 21.7
(9)
This leads to k D 2 = 0.181 and 0.105 m s-l a t p = lo5 and 1.7 x lo5 Pa, respectively. The fraction of external resistance is then fez
= k.$,,f6kD2 = 1.27 x IOV2 and 2.18 x
at respectivelyp = lo5 a n d p = 1.7 x
lo-’ lo5 P a
(10)
which indicates that external mass transfer is not limiting either in the fixed bed. Estimating the effective diffusivity as De = ~,G7/i/t~= 3.75 x 10-6 m2 s-l, Thiele moduli can be calculated. The characteristic dimension for the monolith is the washcoat thickness e. Is the concept of effective diffusivity still meaningful and is the standard theory for diffusion and reaction in porous media still applicable in the case of thin layers of irregular texture whose thickness may be only a few times larger than the size of pores or micrograins? This is certainly a question. However, let us assume that the usual criteria apply and let us calculate
qSl2= kg21De= 2 x
lo-’
= 0.184
effective values for the monolithic matrix considered globally, close to ambient temperature Ar = 0.2W K-’ A, = 0.4W K-I eapp= 410 kg m-3 cp,app= 840 J kg-l K-l average thermal diffusivity F;: 7 x 10-7 m2 s-1 a=
bution-problems with the flow distributor-increased external mass-transfer resistance and deviation to plug flow. Alternatively, one might choose a larger particle size. With d, = 5 mm, one finds (dp/dz)z = 4550 Pa and 4),z2 = 1.42 (qsz = 0.6). The m-l, fez = 6.1 x reduction of the pressure drop (although still 10 times higher than in the monolith) is counterbalanced by the intrusion of internal mass-transfer resistance. Consequently, it seems that the main advantage of the monolith over the packed bed is reduced pressure drop and more favorable mass-transfer conditions.
Thermal Effects The example above is concerned with an isothermal reaction. In many applications however exo- or endothermic reactions have to be carried out. Are heattransfer processes efficient in a monolith? In order to discuss this point, let us introduce characteristic times for radial heat transfer: t,, =
axial heat transfer: t,, =
(ecp)appR2
(12)
which corresponds to an effectiveness factor of qSz % 0.90. This indicates that the fixed bed operates close to the chemical regime. We may draw from this very simple illustration a conclusion which could be expected: the monolithic structure offers a very convenient way to put a given amount of catalytic material into contact with a flowing reactant, avoiding pressure drop and mass-transfer limitations. Under equivalent conditions, a packed bed of small particles is less performing when both criteria are simultaneously considered. Up to this point, the advantage of the monolith seems to be the low pressure drop only. However, the ratio ~ ) ~ 2 ~ is / ~close ) ~ t1o~9 and this indicates that a faster reaction would be responsible for an effectiveness factor in the monolith that is 3 times larger than in the packed bed when the internal masstransfer regime prevails. Of course, it might be argued that instead of setting L1 = L2,one might have designed a shorter but wider fixed bed (L2 < LI), keeping r2 = 0.4 s the same. This would result in smaller interstitial velocity and consequently lower pressure drop. However, packed beds in the shape of “flat boxes” are not recommended as the decrease in axial velocity may lead to flow maldistri-
(13)
4
(ec,)a,pL2
(11)
This is an indication that the wash coat operates in the chemical regime. Conversely, in the case of the spherical pellets
q,: = k.$:/36De
Table 2. Thermal Parameters for Cordierite (62 Channeldcm2)
l a
overall (radial) heat transfer:
rh =
(ecp)appR
2u
(15)
In these expressions, &, ,la and, U are the effective radial and axial thermal conductivities and the overall heat transfer with the environment, respectively. The monolith is assumed of cylindrical shape, of length L and radius R. For gas-phase reactions, thermal conductance is mainly due to the solid matrix. Table 2 gives typical values of thermal parameters from commercially available data for cordierite monoliths. Assuming U = 20 W m-2 K-l if the monolith is not insulated from the outside (the case of automobile converters for instance), the heat-transfer time constants are
This shows that axial conduction is usually negligible and radial conduction across the channels in the radial direction is very ineffective for reactions taking place on time scales of a few seconds or less. Each channel will therefore operate for its own and there is no radial equalization of temperatures. This may be a disadvantage when hot spots are developing in the monolith, especially in the case of flow maldistribution at the inlet. Wendland [311 and Leclerc E201 have studied this problem in catalytic monoliths for automotive pollution control. Leclerc measured velocity distributions among the channels and showed that the
3028 Ind. Eng. Chem. Res., Vol. 33,No. 12, 1994 design of the inlet zone was of crucial importance to ensure radial flow uniformity. In designs with a feed pipe of small diameter, the flow is concentrated in the central channels. Starting from a cold monolith, this results in light-off times which are different from one channel to the next. Therefore, the reactor may be partially ignited. Paradoxically, this may be a favorable feature for provoking early light-off of a few channels followed by the extension of the ignition zone. Considering for instance a reaction time t~ = 0.5 s and the value of the thermal diffusivity in Table 2, we see that the thermal diffusion time Z2/ais equal to t~ on a length scale Z x 0.5 mm comparable to the wall thickness between adjacent channels. This proves that ignition can propagate from one channel to the next. However, as i!hr >> Thl, the monolith reactor at steady state should be usually considered as globally adiabatic.
example. We consider the oxidation of carbon monoxide (molar fraction XAO = 5 x lov3)by oxygen (XBO = 2.5 x in an inert gas flowing in a monolith with square channels of size d = 1.1 mm coated with a standard catalyst (wash-coat thickness e = 0.01 mm). After ref 10, the rate of reaction referred to unit wash-coat volume is assumed to be (16) with
(- 1 2 ~o ) m ~ (17) ~ s
k, = 4.14 x 10l6exp and
External Mass-TransferResistance In the case of highly intensive reaction with large heat evolution, light-off takes place close to the channel inlet and the temperature rises abruptly. Steep temperature gradients build up between the gas and the catalyst, the difference being on the order of the adiabatic temperature rise multipled with the remaining fraction of the key reactant in the gas phase. The reaction generally becomes controlled by mass transfer from the bulk to the wash coat. Under fully developed masstransfer control, the reactant concentration at the catalyst surface is nil and the solid temperature remains constant equal to the inlet temperature plus the adiabiatic temperature rise up to the end of the channel. This is why it is important to understand the mechanism of this transfer and to be able to predict relevant values of Sherwood (Nusselt) numbers in such a regime. This problem has been the subject of many investigations since the first papers of Hawthorn [91,Heck et al. [lo], and Vortruba et al. [291. The analysis of external heat and mass transfer to the wall channel in laminar flow is a classical Graetz problem [281 whose results may be expressed in the form of a global Sherwood (or Nusselt) number. This number is higher in the entrance region and goes down to an asymptotic value downstream in the fully developed regime. Evidence for this effect was provided by Leclerc [201, who determined light-off times with a monolith partially coated with active material from one extremity. Light-off was found t o occur earlier when the active extremity was set upstream than when it was set downstream, and this could be ascribed to enhanced mass transfer in the establishing zone. To take advantage of this phenomenon, it was proposed to cut the monolith into segments placed in series with intermediate remixing of the fluid in order to increase the number of establishing zones [311. In standard monoliths, contradictory results can be found in the literature. Some authors found values close to the theoretical asymptotic values [lo]; others propose correlations showing a dependence of Sherwood number on flow rate (on Reynolds number) and yielding abnormally small values, sometimes lower by an order of magnitude than theoretical predictions [29]. Recent data and discussions are reported for instance in refs 1,2,and 27. When such data are obtained under reacting conditions, even a small residual contribution of chemical reaction to the overall rate of the process, essentially controlled by mass-transfer resistance, may also be sufficient to explain low values of the apparent Sherwood number. Let us illustrate this point by an
K = 65.5 exp (961/T)
(18)
The concentration in the channel section is assumed homogeneous (XA,XB). The rate of mass transfer to the wash coat is expressed as kDCT(xA - x h ) and kDCdxB - X B ~ )with kD&l@= Sh = 3 = constant (assumption of long channels) and G = 1.9 x m2 s-l (STP). The WH based on the monolith volume is WH = 75000 h-l NTP and the inlet temperature is TO= 673 K. The mass- and heat-balance equations for this very simple model are
-u
dx 4kD d z d
- = +x
- x,) = --r5 4e 5eT,
d
P
(19)
(20)
euc
d T 4h 4e - = -Vs - r ) = -r,(-AH)
P d z
d
d
(21)
where z = 0, x = XAO = 5 x and T = TO= 673 K. Owing to the stoichiometric composition of the mixture, X A = 2Xg = n. Subscript s refers to the wash coat. It may be generally assumed that Sh = Nu Le, which implies that h = @CpkD. When mass transfer is fully controlling, xs
