Chemical Engineering and Processing 37 (1998) 117–124

Mass transfer influences on the design of selective catalyticreduction (SCR) monolithic reactors

A. Santos a, A. Bahamonde b, M. Schmid c, P. Avila b, F. Garcıa-Ochoa a,*a Departamento de Ingenierıa Quımica, Facultad de CC. Quımicas, Uni6ersidad Complutense, 28040 Madrid, Spain

b Instituto de Catalisis y Petroleoquımica, CSIC, Campus Uni6ersidad Autonoma, Cantoblanco, 28049 Madrid, Spainc ABB, Max-Hogger Strasse 2, CH 8048 Zurich, Switzerland

Received 2 January 1997; received in revised form 21 November 1997

Abstract

A selective catalytic reduction (SCR) monolithic reactor based on Ti–V–W–Sepiolite catalyst for removal of nitrogen oxidesfrom Power Plant was simulated by using an approaching model which took into account interfacial and intraphase gradients.Different correlations for mass transfer between the solid and the fluid phase were compared, the best obtained results were thecorresponding results for the experimentally determined correlations of Ullah (U. Ullah, P. Waldram, C.J. Bennet, T. Truex,Monolithic reactors: mass transfer measurements under reacting conditions, Chem. Eng. Sci. 47 (1992) 2413–2418) and Votruba(J. Votruba, J. Sinkule. V. Hlavacek, J. Skrivanek, Heat and mass transfer in honeycomb catalysts I, Chem. Eng. Sci. 30 (1975)117–123; J. Votruba, O. Mikus, K. Nguen, V. Hlavacek, J. Skrivanek, Heat and mass trasfer in honeycomb catalysts II, Chem.Eng. Sci. 30 (1975) 201–206). For the considered reaction, internal behavior seems to be essential. A refined study of theintraparticle behavior, calculating De from an experimental tortuosity factor previously obtained, was accomplished. It was shownthat the amount of catalyst could be drastically reduced without changing the performance of the monolithic reactor. © 1998Elsevier Science S.A. All rights reserved.

Keywords: Monolith; Modeling; Ti–V–W–Sepiolite; SCR; Monolithic reactor; External mass transfer coefficients; Effectiveness factor

1. Introduction

Selective catalytic reduction (SCR) is the best-devel-oped and most widely used catalytic NOx removaltechnology for cleaning up power plant stack gas [4,5].

The first wave of SCR application to Power Plantemissions occurred in Japan [6], which in the 1970saddressed its concern with serious and growing airpollution problems in populous metropolitan districts[7]. SCR was identified as a suitable approach forcontrolling NOx and applied the technology on all threetypes of fossil-fueled Power Plant: gas, oil, and coal.Several types of catalyst forms were tried: beads, tubes,wash coated ceramic monoliths or metal plates, andextruded catalytic monoliths. Extruded catalytic mono-liths and coated metal plates became the dominanttechnologies [8], because of the very low pressure drop

and high geometric area per unit of volume of thesecatalytic structures.

A second wave of SCR application is now takingplace in Central Europe. Scientists, environmentalists,and the Green movement focused on a phenomenoncalled ‘dying forests’ in the mid-1980s caused by acidrain. Japanese technology of both the plate and mono-lith types was imported into Germany.

The emergence of the Central European market inthe 1980s was a strong driving force for the develop-ment of catalysts that were both lower in cost andlonger lived than those originally developed in Japan.To determine the potential for major improvements,Grace undertook fundamental studies of the limits ofSCR catalyst performance (NOx conversion, ammoniaconversion, SO2 oxidation, back pressure, sensitivity topoisoning) as a function of the process operating condi-tions (temperature, flow rate, feed conditions) and theproperties of the catalyst [9]. Reaction engineering forcatalyst design is rapidly gaining application [10–12].

* Corresponding author. Tel.: +34 1 3944176; fax: +34 13944176/14; e-mail: CHAPU@EUCMAX.SIM.UCM.ES

0255-2701/98/$19.00 © 1998 Elsevier Science S.A. All rights reserved.

PII S 0255 -2701 (97 )00057 -3

A. Santos et al. / Chemical Engineering and Processing 37 (1998) 117–124118

The results of this study indicated that a substantialreconfiguration of the pore structure of the catalystcould increase NOx conversion activity by about 50%,while simultaneously increasing resistance to poisoningand thereby extending catalyst life. The model showedthat the optimum balance between surface catalyticactivity and diffusivity is provided by a bimodal porestructure with a substantial percentage of macropores[13]. The developed monolithic catalyst for this work isa type of this second generation catalyst and it is alow-dust SCR monolith.

To maintain the selectivity, the catalyst has a properoperating temperature, which must be high enough toprovide useful activity but low enough to avoid theparasitic oxidation of ammonia. Nowadays, commer-cial employed catalysts are fundamentally based onTi–V–W oxides [4] and the technology as typicallypracticed removes 80–90% of the NOx, a higher per-centage than that available by any other commerciallyproven technology [7]. Typical operating conditions forlow dust SCR monoliths are temperatures in the rangeof 300–380°C and gas hourly space velocities (GHSV)are in the 2000–7000 h−1 (STP) range, being areavelocities of 6–10 N m h−1 [9].

Monolithic structures and reactor are relatively newin comparison with other traditional heterogeneous re-actors, as fixed or fluidized beds. In the last twodecades a considerable effort has been accomplished inthe analysis of monolithic catalysts and reactors, focusmainly in the automotive emission control [14,15] andcatalytic combustors [16–18] and less attention havebeen paid for SCR process.

In the monolithic reactor the flow is usually laminarand the modeling of heat and mass transfer with het-erogeneous reaction has been realized employing dis-tributed or lumped parameter models, considering ornot, respectively, the profile of velocity, temperatureand concentration in the radial coordinate of the chan-nel. It seems that only little differences are found whenthe lumped model is used [19]. Furthermore, for SCRreactors other simplifying assumptions that can bejustified under power plant conditions are: isothermally,because of the reactants are diluted in the power plantcombustion gases (NO=200–1200 ppm); negligible ax-ial dispersion (Peclet numbers\50), [15] and fully de-veloped laminar flow in the monolithic channels.

Some questions are opened yet in the SCR reactormodeling as the description of interphase mass transferand the accounting of internal diffusion inside themonolith wall:

(i) External mass transfer of reactants from the bulkphase to the channel surface seems to be an importantresistance to the global reaction rate. Different correla-tions have been proposed [1–3,20] and employed arbi-trarily in simulations but it has been shown [1,14], thatdiscrepant results are achieved when those correlationsabove are employed.

(ii) Because of the internal part of the pellet containsso much more area than the external part most of thereaction takes place within the catalytic solid itself andtherefore effectiveness factor should be employed. Fre-quently some simplifications are assumed in literaturesuch as to ignore the transport inside the monolith walland consider that the reaction occurs only on theexternal catalyst surface [21]. First order kinetic expres-sions are also frequently assumed, then a fixed effective-ness factor along the reactor, and entry an arbitrarytortuosity factor value to calculate the effective diffusiv-ity [22]. Nevertheless, it seems to be proved that tortu-osity factors can vary in a wide range, usually from 2 to13. Taking into account the significant control of theinternal diffusion to the reaction rate, an accuratedetermination of effective diffusivity (or tortuosity) isnecessary; which probably should be done experimen-tally. Thus, in this way a correct evaluation of theeffectiveness factor for the model would be reported.

In this paper results obtained by Bahamonde [23] ina SCR monolithic reactor are compared with thosesimulated employing difference correlations from litera-ture to evaluate the external mass transfer coefficientsand considering internal diffusion inside the monolithwall. The monolithic catalyst (Ti–V–W-oxides on sepi-olite) has been manufactured and previously describedelsewhere [24]. Kinetic equation and experimental tor-tuosity factor for this catalyst have been obtained previ-ously and respectively presented in early papers [25,26].

1.1. Formulation of the model

Neglecting the axial dispersion and entrance effectsphenomena in the channel, and considering isothermalconditions, the mass balance in the fluid bulk can bewritten as:

−ndCb

NO

dx=kgas(Cb

NO−C sNO)=hrbr s

NO

= −asDe

dC sNO

dz(1)

and the mass balance in solid results:

De

d2CNO

dz2 =rNOrp (2)

where the local reaction rate [25], in the solid is givenby:

rNO=ksCNOCNH3

CO2KNH3

KO2

(1+CNH3KNH3

)(1+CO2KO2

)(3)

with the boundary conditions:

x=0: CbNO=Co

NO (4)

z=0: kgas(CbNO−C s

NO)= −asDe

dC sNO

dz(5)

A. Santos et al. / Chemical Engineering and Processing 37 (1998) 117–124 119

Fig. 1. Cross-sectional area of the manufactured monolith.

2. Experimental

The monolithic catalyst used in this study was manu-factured from vanadium-tungsten/titanium dioxidesand sepiolite, following a direct preparation methodreported elsewhere [24], to obtain an incorporated-typemonolithic catalyst, where all the components of thecatalytic system were added to the ingredient mixture toget the precursor paste of the monolith before gettingthe green-body. Subsequently, after drying and a calci-nation temperature program, it was obtained monolithswith a square cell shape at a cell density of 8 cellscm−2.

Fig. 1 shows the cross-sectional area of the monolithwhere it is indicated wall thickness and cell size. Cata-lytic properties, kinetic parameters and tortuosity factorfor this catalyst have been reported in previous papers,[25,26], and are summarized in Table 1.

Effective diffusivity value at the monolithic reactorconditions has been calculated from bundle pore modeland cross-linked pore introducing the experimental tor-tuosity factor and they are also shown in Table 1.

The catalytic activity test were carried out in a mono-lithic reactor described previously [13], giving the oper-ating conditions in Table 2. The gas-feed compositionwas established to reproduce the conditions of the gasesfrom Power Plant, operating in an integral regime andisothermal conditions.

The pore size distribution was made by mercuryintrusion in a Micromeritics 9300 porosimeter and BETsurface area was determined by nitrogen adsorption-desorption in a Micromeritics 2100 D instrument.

The general structural aspect of the catalyst surfaceof the monolith was carried out by scanning electronmicroscopy (SEM), and in this way by wavelengthdispersive spectroscopy (WDS) a mapping of the ele-ments inside the catalysts have been obtained.

z=Lp :dC s

NO

dz=0 (6)

Three correlations have been employed to estimatethe interphase mass transfer coefficient, kg:(i) Hawthorn [20],

Sh=kg ·DDNO

=Sh��

1+0.095ReScDL�0.45

(7)

Sh� being the corresponding asymptotic Sherwoodnumber, calculated by Shah and London [27], andYoung and Finlayson [28],(ii) Votruba et al. [2,3], fitted their results to the equa-tions:

Sh=0.705�

ReDL�0.43

Sc0.56 (8)

(iii) Ullah et al. [1], proposed the following correlation:

Sh=0.766�D

LReSc

�0.483

(9)

Table 1Morphological properties, chemical composition, transport properties and kinetic parameters for the SCR catalyst

Chemical compositionMorphological characteristics

Sg (BET), (m2 g−1) 93 43.21TiO2 (wt.%)4.22V2O5 (wt.%)125rm (A)

260rM (A) WO3 (wt.%) 0.16Vp (m3 kg−1) 0.56×10−3 SiO2 (wt.%) 27.2

MgO (wt.%) 9.30Vpm (m3 kg−1) 0.2856×10−3

Al2O3 (wt.%) 2.60VpM (m3 kg−1) 0.2744×10−3

1.90SO3 (wt.%)rp (kg m−3) 1100op Rest. (wt.%)0.616 21.41

Transport properties Kinetic parameters

ks (m3 kg−1 s−1) 1.59×106 exp(−8786/T)t :3.0 Bundle pore model:5.6 Cross-linked pore model KNH3

(m3 kmol−1) 7.26 exp(6311/T)56.09 exp(3363/T)KO2

(m3 kmol−1)DeNO · 107 (m2 s−1) 9.77 (B.P.M)

9.51 (C-L.P.M.)

A. Santos et al. / Chemical Engineering and Processing 37 (1998) 117–124120

Table 2Operating parameters used for the simulations

Catalyst type Ti–V–W on sepiolite

21Number of channelsCell side (mm) 2.7Wall thickness (mm) 0.85

0.06Channel length (m)857Geometrical surface (m2 m−3)640Monolith density (kg m−3)

Temperature (K) 623Pressure (Pa) 120 000

1NH3/NOx inlet ratio1000NOx inlet concentration (ppm)1000NH3 inlet concentration (ppm)

O2 inlet concentration (%) 3Linear gas velocity (m s−1) 0.12

Table 3Atomic relations for titanium, vanadium and silicium

Global compositionAtomic relation Values of XPS

0.138 0.391V/Ti0.808Ti/Si 2.973

V/(Ti+Si) 0.2930.062(V+Ti)/Si 0.902 4.135

impregnated on titania particles but not on sepiolitefibres, it is said vanadium is always associated withtitania particles.

Taking into account that information can be ob-tained from the proportion of the different elements onthe catalyst surface by XPS, the comparison of theseobtained values with the global composition enablesone to analyze the distribution of the active compo-nents on the catalyst surface [23].

Thus, given these values corresponding to the atomicrelations Ti:V:Si in Table 3, some differences are foundwith respect to the global composition of the catalyst.Thus, it is appreciated that the XPS values are alwayshigher than those corresponding to their global compo-sition. This is indicated when the vanadium is impreg-nated on the titania particles which are principallyfound on the external surface.

Finally, these results imply that the titania particles,impregnated with the vanadium and wolframium salts,are dispersed among the sepiolite fibres, in the way thatthe active phases of the catalytic system were moreaccessible to the reactant gases.

3.2. Simulation studies

Simulation of the monolithic reactor has been per-formed and results compared with the experimentaltests. Orthogonal collocation [29] was applied to calcu-late the intraparticle gradients and in axial direction aRunge–Kutta algorithm of fourth order was employed.A study of the number of collocation points pointedout that 9 collocation points is sufficient to obtainreliable results.

The three cited correlations above—Eqs. (7)–(9)—for estimation of interphase transport coefficient havebeen employed. Because the channel has a square ge-ometry and constant wall temperature is assumed, theasymptotic Sherwood number in Hawthorn correlationgives a value of 2.98. As is shown in Fig. 3, thecorrelations of Ullah [1] and Votruba [2,3] slightlyunderestimate conversion while Hawthorn’s correlation[20] overestimates it. Votruba’s [2,3] and Ullah’s [1]correlations allow the calculation mean transfer num-bers along the reactor and do not take into account thelocal variation of Nu and Sh due to developing concen-tration and velocity profiles. Therefore, underestima-

For analyzing the distribution of the active phase onthe catalyst surface and deeplying into the knowledgeof the nature of the components X-ray photoelectronspectroscopy (XPS) studies were achieved in a Leybold-Heraus LHS-10 spectrometer.

3. Results and discussion

3.1. Characterization studies

From the morphological and textural properties thisstudied catalyst presented a bimodal pore size distribu-tion in the range of micro and mesopores as is given inTable 1.

The use of a natural magnesium silicate, sepiolite, asan inorganic binder in the manufacture of this type ofmonolithic catalyst not only gave a good mechanicalproperties, but also generated a bimodal size poredistribution which became the diffusion of the reactantgases inside the monolithic wall.

Fig. 2 shows the results from WDS. It can be ob-served that Ti and V concentrations are parallel, mean-while Si is abundant where they are scarcing. Thismatter establishes that the vanadium salt has been

Fig. 2. Mapping of elements of the monolithic catalysts by wave-length dispersive spectroscopy.

A. Santos et al. / Chemical Engineering and Processing 37 (1998) 117–124 121

Fig. 3. Comparison of the results with different mass transfer correla-tions. Rigorous intraparticle profile calculations.

and the entry of h as a constant value to use in alumped parameter model can be questioned.

The changes of the profiles of bulk, surface andintraparticle concentrations with the axial position inthe reactor are shown in Fig. 4 for a monolithic lengthof 0.06 m. It can be observed that these profiles aresharper at the entrance of the reactor, thereby with alower effectiveness factor.

From results in Table 4 and Fig. 4, it can be noticedthat a strong internal mass transfer resistance is takingplace. This limitation could be reduced by decreasingthe wall thickness of the monolith. In a series ofsimulations the influence of the wall thickness wasanalyzed. Therefore, the wall thickness was changedkeeping constant the duct shape and the interstitialvelocity. Fig. 5 shows that for a wall thickness greaterthan 0.2 mm no differences in conversion were ob-served. A continuous decrease of conversion occurredonly for wall thickness lower than 0.2 mm. Hence theincrease of the monolith wall from 0.2 to 0.85 mm isnot useful for enhancing the reaction extension and thiszone behaves as an inert material. This is due to astrong internal mass transfer resistance taking placeand therefore the concentration of the reactants isalmost negligible at this region. Table 5 shows theresults for the effectiveness factor at the inlet and outletof the reactor and the nitrogen oxides conversionchanging the wall thickness for a reactor length of 0.06m and an interstitial velocity of 0.24 m s−1.

However, mechanical requirements and manufactura-tion can limit the decreasing of the wall thickness. Inthis way, the manufacture of monolithic supports withsmaller wall thickness could lend to weak structureswhich will not have enough mechanical properties inorder to be useful at industrial conditions where isnecessary to treat large gas flows with solid particlesand ashes in suspension.

The relevance of considering a constant value foreffectiveness factor in the reactor model has also beenanalyzed. It has been studied the relevance of entry aconstant value for effectiveness factor in the reactormodel. This was calculated by Bahamonde [23] in asimple and non strict way from experiments carried outat 350°C, using particles of 1.2 mm in diameter fromthe same monolithic catalyst prepared in this work. Theapparent kinetic constant for a pseudo first order ki-netic equation was obtained and compared with thatone determined under kinetic regime control obtainingthe effectiveness factor as the ratio. Then, the corre-sponding Thiele modulus was obtained from the plot ofthe h values versus Thiele modules for a first orderreaction [30] and thus the value of (ka/De)0.5 was calcu-lated. At the same temperature, 350°C, the value ofThiele module for the flat plate of the same thickness ofthe monolith wall was calculated and then, the corre-sponding effectiveness factor, yielding a value of h=0.1[11]. When this estimated effectiveness factor is used to

tion was to be expected. However, the shape of thecurve approximates quite satisfactorily with experimen-tal data, while the curve calculated by Hawthorn’s [20]equation shows a different behavior.

To investigate the significance of external mass trans-fer resistance it has been calculated the curve XNO

versus 1/GHSV considering plug flow and negligibleexternal gradient of concentration. This curve has beendrawn in Fig. 3 with PFR legend and it can be observedthat conversion values are lower than predicted withPFR.

The choice of the correlations mentioned above maybe justified since working at low Reynolds numbers(B100) where the theoretical models reach an asymp-totic value which could not be supported by the exper-imental works of Ullah [1] and Votruba [2,3].

By now a theoretical calculation of mass transfercoefficients is not yet available and values obtainedexperimentally are within a factor of two from those ofthe theoretical calculations presented by Young andFinlayson [28].

For the wall thickness of the monolithic catalystemployed (0.85 mm) the effectiveness factors have beencalculated, changing the reactor length, at the inlet, atthe exit and a mean value from the h profile in thereactor and results are shown in Table 4. It can be seenunder different conditions h can appreciably changealong the reactor, due to the complex kinetic model,

Table 4Mean effectiveness factor for different reactor lengths, wall thick-ness=0.85 mm

hinletReactor length (cm) houtlet h XNOx

0.0501 0.400.0520.0530.830.0630.0756 0.052

12 0.054 0.105 0.075 0.930.05630 0.21 0.116 0.98

A. Santos et al. / Chemical Engineering and Processing 37 (1998) 117–124122

Fig. 4. Interphase and intraphase gradients (CNO/CNOo ) along the reactor (L=0.06 m).

account for intraparticle diffusion resistance, the math-ematical model of the monolithic reactor can be nota-bly simplified because of Eq. (2) can be omitted.

Results obtained from simulation are compared withthose experimentally found in Fig. 6. Again the experi-mentally obtained correlations from Ullah [1] andVotruba [2,3], and correspond better to the experimen-tal data than correlation from Hawthorn [20]. Experi-mental results are reasonably fitted with this approach.Nevertheless, relatively significant differences in thepredicted NO conversion curves can be found fromcomparison of Figs. 3 and 6, due to the non rigorous h

calculation.

4. Conclusions

SCR reactor behavior can be described accurately byusing a one-dimensional model, employing experimen-tal transfer correlations for the interphase mass trans-fer, being the correlations of Ullah [1], Votruba [2,3],and those which best fit the experimental results.

When the effectiveness factor is calculated solving thedifferential equation describing the concentration gradi-ents inside the wall, refined results were obtained. Thus,effective diffusivity value was calculated from a tortuos-ity factor found experimentally for this catalyst underinert conditions [26]. Introducing this value into theIntraphase Gradient Model, calculated conversion val-ues were close to those experimentally found for thiswork [23], employing a monolithic reactor.

Estimating an effectiveness factor, which is one of thegoverning parameters for the calculations in the one-di-mensional model, could not be done in a reliable wayfor the complex kinetic model of this reaction. It wasfound that the effectiveness factor raises with axialcoordinate and reaches an asymptotic value when thewall thickness increases. For the wall thickness of thismonolith and because of the low effectiveness factorvalue, most of the catalyst weight is not reached by thereactants.

Then, a good balance between the catalyst morphol-ogy and mechanical properties of this monolithic cata-lyst ought to be carried out to reduce the stronginternal mass transfer limitations in order to get in-creasing the reaction extension of NOx removal.

On the other hand, when the description of thetransport inside the wall is simplified by means of aconstant effectiveness factor along the channel, an ac-

Fig. 5. Influence of wall thickness at a constant interstitial velocity(6i=0.24 ms−1).

Table 5Effectiveness factor for different wall thickness at the reactor inletand outlet and its corresponding conversions

hinlet houtletWall thickness (mm) XNOx

0.660.05 0.70 0.7680.1 0.8160.550.46

0.250.2 0.34 0.8330.85 0.05 0.075 0.836

Reactor length: 0.06 m and a interstitial velocity of 0.24 ms−1.

A. Santos et al. / Chemical Engineering and Processing 37 (1998) 117–124 123

Fig. 6. Comparison of results employing a constant effectivenessfactor (h=0.1).

axial coordinate (m)xXO2

conversion of NOx

z radial coordinate (m)

Greek symbolsrb bed density (kg m−3)rg gas density (kg m−3)rp catalyst density (kg m−3)h effectiveness factoro porosity of bedop porosity of catalystoM porosity of macroporesom porosity of micropores

Superscriptso reactor inlets solid surfaceb fluid bulk

SubscriptsM referred to macroporesm referred to micropores

Acknowledgements

This work was supported by CAM-6704 contract.

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Appendix A. Nomenclature

geometrical surface (m−1)as

CNO concentration of NOx (kmol m−3)CNH3

concentration of NH3 (kmol m−3)CO2

concentration of O2 (kmol m−3)D channel diameter=4�hydr.radius (m)DNO molecular diffusion coefficient for NO (m2

s−1)effective diffusion coefficient (m2 s−1)De

activation energy (J mol−1)Ea

DHNH3adsorption enthalpy (J mol−1)adsorption enthalpy (J mol−1)DHO2

ka pseudo first order kinetic constant (s−1)kg mass transfer coefficient (m s−1)KNH3

adsorption constant (m3 kmol−1)KO2

adsorption constant (m3 kmol−1)ks reaction rate constant (m3 kg−1 s−1)L channel length (m)rm mean pore radius macropores (A)rNO reaction rate (kmol kg−1 s−1)rp mean pore radius (A)Re Reynolds number (6i.D.rg/m)Sc Schmidt numberSg pore area (m2 g−1)

Sherwood numberShT temperature (K)6 linear velocity (m s−1)

interstitial velocity at the channel (m s−1)6Imacro pore volume (cm3 g−1)Vm

pore volume (cm3 g−1)Vp

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