Ind. Eng. Chem. Res. 2004, 43, 8207-8216
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Catalytic Hot Gas Cleaning with Monoliths in Biomass Gasification in Fluidized Beds. 2. Modeling of the Monolithic Reactor Jose´ Corella,* Jose´ M. Toledo, and Rita Padilla Chemical Engineering Department (Faculty of Quı´micas), University Complutense of Madrid (UCM), 28040 Madrid, Spain
Nickel-containing monoliths can be used to eliminate tar and ammonia in a real biomass gasification gas. They can work with a fuel gas containing important amounts of particulates, as in the case of the fuel gas produced in fluidized-bed gasifiers. The use of monoliths is a very recent and promising technology that has not yet reached a commercial level and requires experimental studies at pilot scale. Those studies indicate that tar and ammonia conversions (eliminations) with these monoliths depend on so many experimental variables that a model is needed to understand and correlate the results obtained with these monolithic reactors. A model is developed in this paper for the monolithic reactor that has two very different zones: the gas reheating zone and the monolith itself. The model is developed according to the basic rules of chemical reaction engineering, and it includes two microkinetic models for the tar and NH3 elimination reactions, mass balances for tar and NH3, and a heat balance in the monolith. Several important facts appear and are confirmed in this study, such as the control of the external mass transfer in the channels of the monolith, which generates a dependence of the overall kinetic constants on the temperature of potential, not Arrhenius, type. The ∆T across the monolith is also analyzed in detail. The equations developed for the model are easy to handle, allow a correct analysis of experimental data, and may be used to design new monolithic reactors for this application. Introduction As is well-known, biomass gasification in a fluidized bed generates a useful fuel gas (a mixture of H2, CO, CO2, CH4, light hydrocarbons, N2, H2O, etc.), which contains some impurities such as tar, NH3, and particulates. These impurities have to be eliminated for the most promising and advanced applications of the fuel gas. Among the possible gas cleaning methods, the hot cleaning is usually preferred because it really destroys the tar and transfers its energetic content to the fuel gas mainly as H2 and CO. Commercial steam-reforming, nickel-based, catalysts have proven to be very active for tar elimination in biomass gasification with air,1-5 with pure steam,6 and with steam-oxygen mixtures.7 Nevertheless, they need a gas without particulates; they do not accept particulates in the fuel gas. The fuel then has to be filtered before the catalytic reactor. That filtration can be avoided by using nickel-based steam-reforming catalysts but in the form of monoliths, with honeycomb structures. They are not commercialized yet, but some companies, such as BASF AG, can manufacture them. There are abundant studies about different applications of the monoliths that can be found, for instance, in the good review by Cybulski and Moulijn.8 Nevertheless, nickel-containing monoliths for fuel gas cleaning in biomass gasification in a fluidized bed is a quite new technology. To our knowledge, apart from the authors at the University Complutense of Madrid (UCM), only one Finnish9 and one German institution10-12 are active in this field. No more significant data can be found in the public domain literature about this process. * To whom correspondence should be addressed. Fax: +3491-394 41 64. E-mail:
[email protected].
Tar and NH3 elimination with monoliths, in a biomass gasification gas, depends on so many experimental variables that a model was needed to understand and correlate the results obtained with monoliths. The presentation of this model is the main objective of this paper. With the model developed here, the results on tar and NH3 conversions with monoliths will be analyzed and explained in detail. The feasibility, performance, and usefulness of nickelbased monoliths for catalytic tar and NH3 elimination, parts 1 and 3 of this work, are being published in separate papers (refs 13 and 14). In those studies, the gasification or fuel gas to be cleaned was obtained in a atmospheric bubbling fluidized-bed biomass gasifier operating under conditions very similar to existing ones at demo and commercial scales. Something else has to be pointed out: the model presented here was not only generated on a table, using mathematical and physical chemical principles, but also, and mainly, based on results simultaneously obtained in a small pilot plant. These results served as a guide for the developments presented in this paper. Description of the Monolithic Reactor To Be Modeled. (a) Location of the Monolithic Reactor in the Plant. The facility used was a small pilot plant that was fully described previously.13 The gasifier used was a bubbling fluidized bed of 15 cm diameter, at the bottom zone, and 5.2 m height. It was continuously fed with biomass directly into the bed, near the gas distributor. Besides the primary air, there is a second air flow in the upper part of the gasifier. After two in-series cyclones, the main flow is sent, without filtration, to the catalytic (monolithic) reactor. There is also a slip flow, and with valves and rotameters at the end of these two exit lines, the gas flow in each line can be varied from
10.1021/ie040148h CCC: $27.50 © 2004 American Chemical Society Published on Web 11/19/2004
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Figure 1. Effect of the third or reheating air flow on the profile of temperature in the monolithic reactor (notice how the ash softening temperature may be overpassed before of the monolith).
Figure 2. Longitudinal profiles of temperature in the monolith: (i) calculated for adiabatic operation and measured; (ii) the two-layer proposed solution for a further second-generation reactor.
test to test. Hence, the gas flow to the monolith is made independent of the total gasification gas flow produced. The gas hourly space velocity (GHSV) in the monolithic reactor becomes, in this way, independent of the upstream gasifier operating conditions. The main experimental conditions in the gasifier were total air flow, expressed as the total equivalence ratio (ER)Total, partitioning of the air between the primary and secondary flows, temperatures in the bed and in the upper part (after the addition of the second air flow), superficial gas (air) velocity at the inlet of the gasifier (uo), and type and flow rate of biomass, given as its weight hourly space velocity (WHSV) [)kg of biomass/ h‚kg of in-bed material (S + D or S + O)]. These
experimental conditions were given in part 1 of this work.13 (b) Monolithic Reactor. The monolithic reactor to be modeled was shown in detail in Figures 1 and 2 appearing in part 1 of this work.13 This reactor has two very different zones: (1) the preheating and/or gas reheating zone and (2) the monolith itself. Preheating and/or reheating of the fuel gas is made in two ways: with a big (10 kW) external oven and by using a third (or reheating) air flow that burns a part of the fuel gas while heating it up to 1050 °C if needed. This heated gas then enters directly into the monolith (Figure 1). The reheating or third air flow is introduced by a nozzle, specially designed to get a good mix between the two
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flows (the fuel gas and the third air), thus avoiding zones without or with too much oxygen. The reactor has five thermocouples. Two are on the face of the monolith, at the axis and at the periphery of the monolith. Two are just at the exit of the monolith, at the axis and near the wall. The fifth thermocouple enters from the top of the reactor and is movable to measure the axial profile of the temperature from the third air feeding point to the face or front of the monolith. A rough profile of the temperature of the gas in the monolithic reactor is shown in Figure 1. This profile can be varied by modifying the partitioning of the air between the primary (at the bottom of the gasifier), secondary (at the upper zone of the gasifier), and tertiary (at the inlet of the monolithic reactor) flows. The vessel for the monolith was designed to obtain an adiabatic operation in it. There were two different insulating zones: one inside the reactor (see Figure 1) and another one, 15 cm wide, outside the reactor. Despite this, subsequent measurements of the temperature and heat balances indicated that the reactor was not really adiabatic and that some heat was being lost by the big (70 × 70 × 8 cm) flange located (Figure 2 in ref 13) at the bottom of the reactor, next to the exit of the monolith. For this reason, the temperature at the monolith exit was always lower than that corresponding to an adiabatic operation. This fact generated some problems, which have already been solved in a “secondgeneration” monolithic reactor very recently built and set up at UCM. (c) Monoliths Used. Monoliths were obtained initially from different manufacturers, but this study is based only on monoliths from BASF AG. The monoliths were always located in the catalytic reactor with their full size (15 × 15 × 30 cm). Except in two tests, each monolith was new (one different monolith) for each test run. Usually all channels of the monolith were open and used in each test. In some tests, and to get relatively high gas velocities in the monolith, some (2, 3, and 4) peripheral rows of channels were blocked with a mask and filled with small particles of SiC. In this way, the face velocity of the gas (uf) and the Reynolds number in the remaining open channels were increased. (d) Sampling. Sampling was made before and after the monolithic reactor. Notice that, because of the very high temperature (around 920 °C) at the front of the monolith, sampling could not be made there. That sampling had to be made just before the monolithic reactor, not at the monolith itself. Sampling was made at different times on stream. That is to say, several samples were taken before and after the monolith in each test. (e) Analysis, Characterization, and Lumping of Tar. The oldest, easiest, and most used approach is to consider tar as only one single lump and a first-order (for tar) reaction for its kinetics of elimination.2,15 That approach was enlarged to a more correct one based on two lumps for tar and four kinetic constants16 and, further, to two very advanced models based on consideration of tar as a continuous mixture and as a reacting network composed of six lumps and described by 11 kinetic constants.1 In the present model for the monolithic reactor, the simplest microkinetic model, with only one lump for tar, is used. There are two main reasons for this choice: (1) there are not many experimental data available with
monoliths, and (2) some experimental errors with monoliths are still important to use microkinetic models that need accurate data. In the future, with more abundant and accurate data, the more advanced microkinetic models mentioned above could be introduced in the overall model developed here for the monolith. These would replace the one-lump first-order reaction for tar used in this paper. Modeling of the Monolithic Reactor. (a) Level of Confidence or Description. There are several mathematical models for monoliths in the literature. Forzatti et al.17,18 Gupta and Balakotaiah,19 Cybulski and Moulijn,8 Votruba et al.,20 and Uberoi and Pereira21 are only a few examples of good work on monolithic modeling. One-, two-, and three-dimensional (1D, 2D, and 3D) models already exist in the literature on monoliths. There is no doubt by now concerning the level of confidence for the model to be developed for the process considered in this paper. Going from 1D to 3D models, their degree of mathematical sophistication increases and they seem more “beautiful”, rigorous, and scientific, as occurred with the above-mentioned case about the microkinetic model for the catalytic tar abatement. Nevertheless, there are some conclusive facts that have also to be considered: (1) The 2D and 3D models developed by Forzatti et al.18 for combustion in a monolith of a biomass gasification gas contain a strongly nonlinear system consisting of about 1000 equations. Even with 1000 equations, they “neglected (did not consider) the presence of small amounts of tars and of N- and S-containing compounds”. In the present work, where tar (and NH3) cannot be “neglected”, a 2D or 3D model would then need more than 1000 equations. Such a set of equations was of no interest for this work, and the idea of a 2D or 3D model was abandoned. A 1D model was therefore preferred instead of 2D and 3D models. (2) From the first experiments with these monoliths, it was seen how the main problem and bottleneck of the monolith for its use on a commercial scale was its deactivation. Deactivation of these monoliths may occur by several simultaneous mechanisms and causes and may be relatively fast. The problem of the monolith for further commercial use is not going to be its activity, but its life. It therefore made no sense for the authors to develop a sophisticated set of mathematical equations for the activity of the monolith. (3) “The monolithic reactor” and “the monolith” are, in fact, two different reactors. The monolithic reactor, as shown in detail in Figure 1, has two very different and in-series reaction zones: “zone 1”, in which the gas is reheated by means of a third air flow, which burns a part of the tar, NH3, H2, CO, CH4, etc., present in the fuel gas, and “zone 2”, the monolith itself. The temperature of the fuel gas along the axis of the monolithic reactor increases by the third air flow, as Figure 1 shows. This profile of the temperature has been measured by the mobile thermocouple located at the top of the reactor. The temperature at the monolith front/face is also measured with two more thermocouples. Temperatures between 900 and 1050 °C were measured in the space at the front of the monolith. Gas sampling cannot be made in a zone with these high temperatures. Samplings for gas, tar, and NH3 therefore have to be made just before the full monolithic reactor and not before the monolith itself. The gas composition, includ-
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Table 1. Relative Importance at 900 °C of the Chemical Step (1/k′tar), as a Percentage of the Total Resistance (Chemical Resistance/Total Resistance) × 100, for Different Gas Velocities at the Front of the Monolitha k′eff,tar,900°C (s-1, reactor cond.)
total resistance 1/k′eff,tar,900°C (s, reactor cond.)
(chemical resistance/ total resistance) × 100tar elimination
uf(inlet monolith) (cm/s)
Re
Votruba et al.20
this work13
Votruba et al.20
this work13
with ref 20
with ref 13
30 80 150 200 400
8.8 24 44 59 118
6.5 10 13 15 20
2.6 4.4 5.9 6.8 9.6
0.15 0.10 0.077 0.068 0.050
0.38 0.23 0.17 0.15 0.10
9.3 14 18 20 28
3.7 6.1 8.2 9.5 13
a
[k′tar,900°C (from ref 4): 72 (s-1, reactor cond.). Chemical resistance, 1/k′tar,900°C: 0.014 (s, reactor cond.)].
ing the tar and NH3 contents, just in the front of the monolith, cannot be known with accuracy. To develop a complex model for the monolith itself makes no sense when there is this uncertainty on the gas composition at its entrance. For the above reasons, it was decided to develop and use a 1D model for the monolithic reactor with two reaction zones in it. (b) Reaction Network. The network considered for the simultaneous tar and ammonia elimination reactions, in the reactor shown in Figure 1, is
8 H2 + CO + ... tar (+ ?) 9 cat. kcat.,NH
3
2NH3 98 N2 + 3H2 3
(1)
This network considers both thermal and catalytic reactions for the tar and NH3 elimination. (c) Microkinetic Model for Tar Elimination. In the monolithic reactor, if first-order (for tar) elimination reactions are considered in each zone, the overall rate of tar elimination is
(-rtar) ) (-rtar)thermal,zone 1 + (-rtar)cat., zone 2 ) kth,tar,zone 1Ctar + kcat.,tar,zone 2Ctar ) (kth,zone 1 + kcat.,zone 2)tarCtar ) keff,tarCtar (2) where
keff,tar ) (kth,zone 1 + kcat.,zone 2)tar
∫XX
dXtar
tar,exit
tar,inlet
kcat.,tar,zone 2Ctaro(1 - Xtar)
(5)
or
τmon )
∫XX
dXtar
tar,exit
tar,inlet
τreactor )
kth,tar
kth,NH
W ) QoCtaro
kcat.,tar,zone 2(1 - Xtar)
(6)
A mass balance for tar in the whole monolithic reactor would lead to
kcat.,tar
tar + H2O 9 8 H2 + CO + ... cat. CO2
4NH3 + O2 + ... 98 N2 + 2NOx + 6H2O
a mass balance for tar in it leads to
(3)
In the monolith itself, there may be different solids that may have some but different activity on tar elimination. So, to understand some results, kcat.,tar,zone 2 includes different contributions:
kcat.,tar,zone 2 ) kNi cat. + kth,monolith + kunimpregnated wall of the monolith + kcoke on the monolith (4) For the overall NH3 elimination, a kinetic equation similar to eq 2, with another affective kinetic constant, keff,NH3, is postulated and used in the next calculations. Mass Balances for Tar and for NH3. Considering piston flow in the monolith and supposing that there is not an appreciable change in the total volume of the gas,
∫0X
tar,exit
dXtar
keff,tar(1 - Xtar)
(7)
Analogous equations are obtained and used for NH3. If the values of keff and space time (τ) are known, the tar and NH3 conversions at the monolithic exit can be calculated with these equations. If, on the contrary, XNH3,exit and Xtar,exit are fixed, the required space time can be calculated. When using eqs 6 and/or 7, and the similar corresponding ones for NH3, it has to be taken into account that the monolith is not isothermal at all; it works under a near-adiabatic situation. To solve eqs 6 and 7, the variations of kcat.,tar and/or keff,tar with temperature and of the temperature with Xtar are needed then. Dependence of the Kinetic Constants in the Monolith with Temperature. (a) Importance of the Mass Transfer in the Channels of the Monolith. The Reynolds number in the channels of the monolith calculated for the experimental conditions of this work ranged from 11 to 60. In all experiments, the monolith clearly operates under laminar flow. For this reason, mass transfer from the bulk of the gas flow to the wall of the channels plays an important role.21 To determine how important this mass-transfer limitation is in this process, both the total resistance and the resistance of the chemical step, for tar decomposition, were calculated. They are described below. The total resistance, as is well-known in chemical reaction engineering, is the inverse of the overall kinetic constant (keff,tar). For the catalytic tar decomposition in the monolith, the overall or effective kinetic constant (keff,tar) was calculated for all test runs. keff,tar values obtained (ref 13) at 900 °C for different uf values are shown in the fourth column of Table 1. The inverse or reciprocal of these values is the total resistance for the tar conversion. This is shown in the sixth column of Table 1. The keff values calculated with Votruba et al.’s20 equation, used as a reference, were also simultaneously calculated. They are also shown in Table 1.
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The intrinsic resistance of the chemical step for the catalytic tar elimination can be calculated given the abundant and previously available data for this purpose. Corella et al.4 found and published the intrinsic kinetic constants for tar elimination with these Ni/R-Al2O3 catalysts at 900 °C and without diffusional control. They were obtained with crushed catalyst particles, of 1.01.6 mm (η ) 1), and with superficial gas velocities higher than 50 cm/s, without external mass-transfer limitations. They were also obtained with (H2O/C*) > 2.0. These kinetic constants can now be used and compared with the ones obtained with monoliths. From the data of Corella et al.,4 the intrinsic (η ) 1 and no external diffusion control) kinetic constant at 900 °C (with a similar Ni/R-Al2O3 catalyst) for tar elimination is
ktar(900 °C) ) 395 m3/kg‚h
(8)
k′tar(900°C) ) 72 s-1
(9)
equivalent to
The inverse of this value, shown at the bottom of Table 1, is the resistance of the chemical step in the tar elimination reaction. From the sixth column in Table 1 and the value shown at its bottom, the relative importance of the chemical step, with respect to the whole or total resistance, has been calculated for several gas face velocities in the monolith. This value, given as a percentage, is shown in the last column of Table 1. One very important conclusion is deduced from these results: with a gas face velocity (uf) of 80 cm/s, with the BASF monolith, the resistance of the chemical step is only about 6-14% of the total resistance. At 80 cm/s, the external mass transfer controls the process to an extent of 86-94%. In other words, and as was already foreseen from previous studies with monoliths (i.e., ref 21), the controlling step is the external mass transfer in the channels of the monolith and not the chemical step itself. In Table 1, results using Votruba et al.’s20 equation are also presented. With Votruba et al.’s equation, not so drastic results appear, but they are due to the fact that they used monoliths with much higher cell density than the one of the monolith used in this work. Their dH value ranged from 0.1 to 1 mm only. The external mass transfer was not as important as in this work, as the last column in Table 1 shows. From the previous paragraph, and if the pitch (dH) of the channels are not varied, it may be deduced that the tar and NH3 conversions can be increased somewhat by increasing the gas face velocity. Nevertheless, notice that even at 2.0 m/s the chemical step means only about 20% of the total resistance. The overall process is and will be (for values of uf below 10 m/s approximately) limited by the external mass transfer. This is the main reason the overall or effective kinetic constants for tar and NH3 removal obtained with monoliths13,14 are quite lower than the corresponding values for the nickel catalysts in granular/particulate form. Correlating the keff,tar and keff,NH3 Values. With the correlations for monoliths existing in the bibliography8,19-21 and after several initial attempts, the authors decided to use the keff,tar and keff,NH3 parameters introduced in Sherwood (Sh) numbers and to correlate them with the Reynolds and Schmidt numbers. Because only one k value has to be introduced in the Sherwood
Table 2. Gas Composition (Dry and Wet Basis, for ER ) 0.40) at the Inlet of the Monolith Reactor (Gasifier Exit) Used as a Reference in the Model vol %
vol %
species
dry basis
wet basis
species
dry basis
wet basis
H2 CO CO2 CH4
11.5 16.5 15.5 4
10.4 14.8 14.0 3.6
C2H4 N2 H2O total
1.2 51.3
1.1 46.1 10 100
100
Table 3. Molecular Weight of the Main Species in the Fuel Gas and Values of Ei, EAi, kT/EAi (for T ) 1093 K), ΩD, σi, σAi, and DAi (for T ) 1093 K) species NH3
H2
Mi (g/mol) 17 i × 1021 (J) 7.70 Ai × 1021 (J) kT/Ai ΩD σi (Å) 2.9 σAi (Å) DAi (cm2/s)
CO
CO2
CH4
C2H4
N2
H2O
2 28 44 16 28 28 18 0.82 1.26 2.69 2.05 3.10 0.98 11.2 2.52 3.12 4.55 3.97 4.89 2.75 9.29 6.0 4.8 3.3 3.8 3.1 5.5 1.6 0.81 0.85 0.93 0.89 0.94 0.83 1.17 2.8 3.7 3.9 3.8 4.2 3.8 2.6 2.9 3.3 3.4 3.3 3.5 3.4 2.8 7.5 2.2 1.8 2.4 1.8 2.2 2.5
number, it was decided to use the k value for a given temperature of reference. Because 900 °C (1173 K) is the most recommended temperature for the front of the nickel-based monolith,13 the targeted k values were the kinetic constants calculated at that temperature (900 °C). The values of the kinetic constants at this temperature will be referred to from now on as keff,tar,900°C and keff,NH3,900°C. This work was done to analyze and further predict tar and NH3 conversions. For this reason, two different Sherwood numbers have been used:
Sheff,tar ) βdH/Dnaph-m ) k′eff,tar,900°CdH/AvDnaph-m ) k′eff,tar,900°CdH2/4Dnaph-m (10) Sheff,NH3 ) βdH/DNH3-m ) k′eff,NH3,900°CdH/AvDNH3-m ) k′eff,NH3,900°CdH2/4DNH3-m (11) To calculate the Sherwood values for tar, naphthalene was the targeted species because it is one of the most abundant species present in tar and because its elimination or destruction is the key factor in the overall elimination of tar.1 Nevertheless, other species, if preferred, can be used as representatives for the tar; it would not change the next development and conclusions. According to eqs 10 and 11, the diffusion coefficients of naphthalene and NH3 in the complex reacting mixture (m) have to be calculated. To carry out these calculations, the gas composition (dry and wet basis) used as the reference is shown in Table 2. It corresponds to a fuel gas very often obtained at the exit of a fluidizedbed biomass gasifier. The diffusion coefficient of A (naphthalene or NH3) in i (H2, CO, CO2, N2, H2O, etc.) (DAi) was calculated with the well-known equation22
DAi )
0.001858T3/2[1/MA + 1/Mi]1/2 Pσ1i2ΩD
(12)
Table 3 presents the molecular weights and calculated values for A, Ai, (kBT/Ai), ΩD, σi, σAi, and DAi (at 1093 K) for the main species present in the fuel gas. The diffusion coefficient of NH3 in the gas mixture (DNH3-m)
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Table 4. Values of DNH3-m and of the Density and Viscosity of the Gas at the Inlet of the Monolith (Tinlet) test run Tinlet (°C) DNH3-m (cm2/s) Fg (kg/m3) µg × 105 (kg/m‚s) N-78 N-79 N-80 N-81 N-82 N-83 N-84 N-85 N-86 N-87 N-88 N-89 N-90 N-91
820 865 890 835 855 956 912 937 961 855 910 832 875 880
2.7 2.9 3.1 2.8 2.9 3.4 3.2 3.3 3.4 2.9 3.2 2.8 3.0 3.0
0.285 0.267 0.258 0.279 0.271 0.232 0.249 0.239 0.230 0.271 0.250 0.280 0.264 0.262
3.90 4.04 4.11 3.94 4.00 4.31 4.18 4.25 4.32 4.00 4.17 3.93 4.06 4.08
CT,NH3 ) CT,NH3,reference(1173)
∑DAixi
(13)
The calculated value for DNH3-m, at 820 °C and for the realistic gas mixture shown in Table 1, is
DNH3-m(820 °C) ) 2.7 cm2/s
T (1173 )
2(1
(19)
CT,NH3 and CT,tar do not follow the Arrhenius dependence on T. The reason is that the overall phenomenom is mass-transfer-limited, as was demonstrated above. Using eq 18, for instance, at two different temperatures, T and 900 °C (1173 K) as the temperature of reference, and using eq 19 for CT,NH3, it can be deduced that
is further calculated by
DNH3-m )
experimental part of this work.14 From this calculation, it was deduced that m ) m′ ) 0.5 and n ) n′ ) 0.5, although some error has to be admitted for these exponents because of the small amount of experiments and the nonisothermicity of the monolith, which introduced some uncertainty in the analysis of the data. For the said exponents, it was also deduced14 that
( )( ) ( )(
(keff,NH3)T
T ) keff,NH3,900°C 1173
Dnaph-m(900 °C) ) 0.87 cm2/s
(15)
Scnaph(900 °C) ) 1.87
(16)
To calculate the Reynolds number at the monolith inlet, the density and viscosity of the gas were carefully calculated for the experimental conditions in each test. These values are given in Table 4. With these values, the Reynolds number was calculated for different gas face velocities (uf) and for two temperatures (820 and 920 °C). Once the values for keff, NH3, DNH3-m, Tinlet, uf, and Re were measured or calculated for each experiment, the dimensionless numbers Sheff,NH3,900°C, ScNH3, and (dH/L)RemScm′NH3, and the equivalent dimensionless numbers for the simultaneous tar elimination, were easily calculated (in the case of tar, the Schmidt number was calculated with the diffusion coefficient of naphthalene). The next step was to correlate the experimental Sheff,NH3,900°C and Sheff,tar,900°C numbers against (dH/L)Re(ScNH3) and (dH/L)Re(Scnaph), respectively. Taking into account the work of Uberoi and Pereira,21 the following types of correlations were used:
Sheff,tar ) CT,tar(dH/L)RenScn′naph
(17)
Sheff,NH3 ) CT,NH3(dH/L)RemScm′NH3
(18)
and
where CT,tar and CT,NH3, defined by these equations, are parameters that vary with temperature. They can be calculated in all tests. This calculation is shown in the
m′
ScNH3,1173
T 1173
(14)
DNH3-m has also been calculated for the temperatures at the front of the monolith in the different tests. These DNH3-m values are given in Table 4. The calculated values for the diffusion coefficient at 900 °C of naphthalene in the same mixture (Dnaph-m) and for the Schmidt number are
ScNH3,T
2(1
DNH3-m,T
DNH3-m,1173
DNH3-m,T
2(1
)
≈
1-m′
DNH3-m,1173
(20)
Taking into account that, according to eq 12, DNH3-m varies with the temperature as T3/2, for m′ ) 0.5 eq 20 becomes
(keff,NH3)T keff,NH3,900°C
)
T T (1173 ) (1173 ) 2(1
0.75
)
T (1173 )
2.75(1
(21)
from which
keff,NH3,T ) keff,NH3,900°C
T (1173 )
2.75(1
(22)
The same dependence on T was found for keff,tar. Equation 22 and the similar one for keff,tar are the basic keff-T relationships used in the overall model for the monolithic reactor. Another approach might also be used: an averaged (with respect to temperature) value of the kinetic constant (k h ) for the whole monolith can also be considered. In this way, eq 6 becomes
τmon )
∫X k h cat.,tar,zone 2 X 1
tar,exit
tar,inlet
dXtar
(1 - Xtar)
(23)
or in integrated form
h cat.,tar,zone 2τmon (24) ln[(1 - Xtar,inlet)/(1 - Xtar,exit)] ) k Equation 24 may be used to calculate k h cat.,tar,zone 2. This k h cat.,tar,zone 2 can be further correlated with the whole interval of temperature (from Tinlet to Texit) existing in the monolith in each test. To use eq 24, Xtar,inlet has to be known, and this can be made with tests without monolith in the reactor. In this case, Xtar at the exit of the reactor nearly corresponds to the Xtar existing at the inlet of the monolith (Xtar,inlet) when the tests are made under the same experimental conditions. When k h cat.,tar,zone 2 is correlated with the interval of temperature in the monolith, a dependence similar to eq 22 was found.
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With eq 22, and the other similar one for tar, the mass balance equations for the tar (eq 6) and NH3 elimination in the monolith become
τmon )
dXtar tar,inlet T 2.75(1 keff,tar,900°C (1 - Xtar) 1173
∫XX
tar,exit
(
)
(25)
∆TCH4 ≈ 100XCH4
and
τmon )
∫XX
dXNH3
NH3,exit
NH3,inlet
keff,NH3,900°C
T 2.75(1 (1 - XNH3) 1173
(
)
is not possible to know the fraction or extent of CH4 reacted by each jth reaction, y1,j. So, the effective or overall ∆T1 value (∆TCH4) because of the CH4 conversion cannot be exactly known from eq 30. As a rough estimate, it might be proposed that
(26)
To solve these equations, the variations of T with Xtar and XNH3 are now needed. These relationships will appear from the following heat balance in the monolith. Heat Balance in the Monolith. Call j the jth reaction in the network (1, reaction with H2O; 2, reaction with CO2; 3, thermal degradation; etc.) and i the ith gas-phase compound or reactant in the reacting fuel gas (1, CH4; 2, naphthalene; etc.). The heat balance on an element dV of the monolith, for adiabatic operation, is
∑j [∑i (-ri,j) dV (-∆Hr )] ) (∑i m˘ icpi) dT ) m˘ Tcjp dT
(32)
The same applies to all other reactants (tar, C2Hn, H2, etc.). For instance, considering i ) 2 ) C10H8 (naphthalene, key species in tar) and j ) 1 ) steam reforming:
C10H8 + 10H2O f 10CO + 14H2 ∆Hr2,1 ) 9.1 MJ/kg ) 1164.8 kJ/molnaphthalene (33) In test run N-84 (ref 13), as just one example, with C2,0 ) 5230 mg/N‚m3 and F2,0 ) 0.39 molnaphthalene/h, the calculated ∆T2,1 was 14.6 °C. Several other exothermic (as the CO shift) and endothermic reactions have been also considered to occur in the monolith. Overall, from eq 31
Texit ) Tinlet -
∑i ∆TiXi,exit
(34)
i,j
(27)
or
Texit ) Tinlet - [∆TCH4XCH4,exit + ∆TnaphXnaph,exit +
or
∑j ∑i (Fi)0 dXi,j (-∆Hr )yi,j ) m˘ Tcjp dT i,j
(28)
where yi,j is the fraction of i that reacts in the jth reaction. Integrating eq 28 from the monolith inlet to a point or length in which there is a given conversion and temperature
∑j ∑i Fi,o(-∆Hr )Xiyi,j i,j
T ) Tinlet -
m ˘ Tcjp
(29)
or at the exit of the monolith
∑j ∑i Fi,o(-∆Hr )(Xiyi,j)exit i,j
Texit ) Tinlet -
m ˘ Tcjp
(30)
when only one reactant and one reaction, the jth one, is considered (yi,j ) 1), eq 30 becomes
Texit ) Tinlet -
Fi,o(-∆Hri,j) Xi,exit ) Tinlet - ∆Ti,jXi,exit m ˘ Tcjp (31)
CH4 is the species that mainly determines the decrease of T in the monolith. So, if i ) 1 ) CH4, then ∆Ti,j ) ∆T1,j. The calculated ∆T1,j values, under real experimental conditions, for the most relevant reactions in which CH4 disappears are (i) for the reaction of CH4 with CO2, 135 °C, (ii) for the reaction of CH4 with H2O, 112 °C, and (iii) for the thermal decomposition of CH4, 41 °C. The total CH4 conversion (disappearance) in the monolithic reactor is experimentally measured, but it
∆TC2HnXC2Hn,exit + ∆TCOXCO,exit + ...] (35) The following facts now have to be pointed out: (1) The activity of the monolith is not the same for all reactions in the network (see, for example, ref 14). So, Xnaph,exit * XC2Hn,exit * XCH4,exit * XNH3,exit .... (2) Conversions of some reactants may be difficult to estimate. (3) The temperature measured at the monolith exit was lower than that if the monolith had been really adiabatic. The UCM monolith did not have an adiabatic operation because some heat was lost through (i) the big flange located at the bottom of the reactor and (ii) the mask used in some tests to block/close some channels of the monolith. For this reason, Texit - Tinlet in the reactor at UCM has always been higher than the adiabatic ∆T (∆Tadiab) calculated by eq 34 as ∑i∆Ti). For these reasons, eq 34 or eq 35 was replaced by
Texit ) Tinlet - ∆Texp,tarXtar,exit
(36)
or for a point or length in the monolith
T ) Tinlet - ∆Texp,tarXtar
(37)
∆Texp,tar appearing in these equations is a parameter that has to be calculated in each experiment. It is made in the following way. In each test run, the temperature is measured in two points: at the inlet/front/face of the monolith (Tinlet) and at its exit (Texit). Texit - Tinlet and Xtar,exit are known for each test, so ∆Texp,tar can be easily calculated by
∆Texp,tar )
Texit - Tinlet Xtar,exit
(38)
The variation of T with Xtar in the monolith will then
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Ind. Eng. Chem. Res., Vol. 43, No. 26, 2004
be calculated by eq 37, with the auxiliary eq 38 to calculate the parameter appearing in it. The same can be applied to the simultaneous ammonia elimination reaction. For this reaction:
keff,NH3,900°C) are not difficult to deduce and may be used by all people or institutions working in this field.
T ) Tinlet - ∆Texp,NH3XNH3
A 1D model has been presented for a one-layer monolithic reactor for the simultaneous tar and ammonia elimination in a real biomass gasification gas. The macrokinetic model includes (1) two first-order microkinetic models for the two tar and ammonia removal reactions, (2) mass balances for tar (eqs 6 and 7) and for NH3, (3) heat balances in the monolith that lead to T-Xtar and T-XNH3 relationships given by eqs 37 and 39, respectively, and (4) variation of the effective kinetic constants with temperature of the potential type, as indicated by eq 22, because the overall process is controlled by the external mass transfer in the channels of the monolith, as the data in Table 1 confirms. When all of the above-mentioned equations are combined, the overall or macrokinetic model given by eqs 41 and 42 is deduced. These equations may be used in two, at least, different ways. In parts 1 and 3 of this work,13,14 these equations are used to calculate the keff,tar,900°C and keff,NH3,900°C parameters, which are a good indexes of the activity of the monolith for tar and NH3 elimination. These parameters have been used13,14 to compare the activity of the monoliths with those of competitive catalysts for hot gas cleaning such as commercial (rings) nickel-based steam-reforming catalysts and calcined dolomites or related materials. A different or second use of eqs 41 and 42 is when keff,tar,900°C and keff,NH3,900°C are known. In this case, eqs 41 and 42 may be used to calculate (a) the space time required in the monolithic reactor to reach fixed tar and NH3 conversions or (b) the tar and NH3 conversions at the exit of the monolithic reactor if its space time is fixed.
(39)
and
∆Texp,NH3 )
Texit - Tinlet XNH3,exit
(40)
Because of the relatively very low temperature at the exit of the monolith (because of the high ∆Texp), some filamentous whisker-type coke may be formed there, as shown in part 1 of this work.13 This type of coking is very well documented in the literature on steam reforming of natural gas and naphthas, and it must be avoided in this gas cleaning process. To avoid this coking, a temperature higher than 800 °C has to be obtained and maintained at the exit of the monolith. A reasonable technical solution for this problem is to locate the monolith(s) in two layers, one after the other. Remember that DeNOx units in waste incineration plants usually have four layers. So, to have several layers of monoliths is not a problem today. In the first layer of monoliths, XCH4 should be kept to no more than 40%; approximately ∆Tadiab in each layer would then only be 50-60 °C, as Figure 2 shows, and the whisker-type coke would therefore not be formed. Some reheating between the two layers would be needed with a relatively small fourth air flow, as Figure 2 shows. It is the technical solution proposed for further application of this technology on a commercial scale. Overall or Macrokinetic Model (Mass Balance + Heat Balance + Kinetic Equations). When eqs 25 and 37 are combined, the resulting equation for tar becomes
τreactor )
∫0X
tar,exit
keff,tar,900°C
dXtar Tinlet - ∆Texp,tarXtar 1173
(
)
2.75(1
(41) (1 - Xtar)
and, from eqs 26 and 39, for the ammonia decomposition reaction
τreactor )
∫0X
NH3,exit
keff,NH3,900°C
(
dXNH3
)
Tinlet - ∆Texp,NH3XNH3 1173
Conclusions
Acknowledgment The UCM team expresses their gratitude to the monolith manufacturer BASF AG by providing samples of their monoliths. Scientific discussions with Dr. Markus Ising from Fraunhofer UMSICHT in Oberhausen, Germany, and with Dr. Pekka Simell from VTT in Espoo, Finland, helped to develop the model work. Alicia Laguna, Eduardo Torcal, Laura Guijarro, and Jorge Ruiz-Peinado, students of chemical reactors (Chemical Engineering Department) at UCM, helped a lot in several mathematical calculations. This work was carried out under the Ammonia Removal Project No. ERK5-CT99-00020 of the EC, DG Research, Directorate J. The authors thank the European Commission for its financial support.
2.75(1
(1 - XNH3) (42)
These equations are the core of the so-called overall or macrokinetic model. They are rigorous equations that are used in refs 13 and 14 to calculate the keff,tar,900°C and keff,NH3,900°C parameters. Once Xtar,exit and XNH3,exit are measured and known in each test, and once ∆Texp,tar and ∆Texp,NH3 are calculated by eqs 38 and 40, respectively, keff,tar,900°C and keff,NH3,900°C can be calculated. They can be further correlated with the experimental conditions in each test. These two parameters (keff,tar,900°C and
Nomenclature Av ) area to volume of the channels of the monolith (cm2/ cm3) BFB ) bubbling fluidized bed Ctar ) concentration of tar in the fuel gas (mg of tar/mn3) Ctar,0 ) concentration of tar at the inlet of the monolithic reactor (mg of tar/mn3) Cp ) (kcal/kg‚K) D ) calcined dolomite DAi ) diffusion coefficient of A in i (cm2/s) DNH3-m ) diffusion coefficient of ammonia in the reacting mixture (cm2/s)
Ind. Eng. Chem. Res., Vol. 43, No. 26, 2004 8215 Dnaph-m ) diffusion coefficient of naphthalene in the reacting mixture (cm2/s) daf ) dry, ash free dH ) hydraulic diameter of the channels of the monolith (mm) ERTotal ) total equivalence ratio, defined as the total airto-fuel ratio used in the gasifier divided by the air-tofuel ratio for the stoichiometric combustion Fi,o ) molar flow rate of the ith reactant at the inlet of the monolithic reactor (mol of i/h) GHSV ) gas hourly space velocity [m3 nc/h‚m3 cat, h-1] ∆Hi,j ) heat of reaction, referred to as the ith reactant in the jth reaction, kJ/mol keff,tar, keff,NH3 ) effective or apparent kinetic constant for the overall tar, or ammonia, respectively, elimination, [m3(Tcat, wet)/kgcat‚h] k′eff,tar, k′eff,NH3 ) effective or apparent kinetic constant for the overall tar, or ammonia, respectively, elimination (s-1) kth, kcat ) kinetic constants for the tar elimination by thermal (zone 1) and catalytic (zone 2) reactions, respectively [(m3(Tcat, wet)/kgcat‚h] k h cat ) averaged value of the kinetic constant for the whole monolith, defined by eq 24 [(m3(Tcat, wet)/kgcat‚h] ktar ) intrinsic (η ) 1) kinetic constant for the overall tar elimination [m3(Tcat, wet)/kgcat‚h] MA, Mi ) molecular weights of A and i m ˘ T ) total mass flow rate of the fuel gas (kg/h) n, n′ ) exponents in eqs 17 and 18 O ) olivine P ) pressure (atm) PAH ) polyaromatic hydrocarbons Q0 ) gas flow rate at the inlet of the monolithic reactor [m3(Tcat,wet)/h] Q′0 ) gas flow rate at the inlet of the monolithic reactor (m3 nc/h) ri,j ) reaction rate of the ith reactant by the jth reaction rtar ) reaction rate of tar elimination {(mg of tar/kgcat‚h)[m3(Tcat, wet)/mn3]} S ) silica sand T ) temperature (K) Tinlet, Texit ) temperature at the inlet and exit of the monolith (K) T′, T′′,T′′′, Tiv ) temperatures at different points in the gasification plant, as indicated in Figure 1 (K) ∆Texp,tar, ∆Texp,NH3 ) parameters defined by eqs 38 and 40, respectively (K) Vreactor ) internal volume of the monolithic reactor (m3) u0 ) superficial gas (air) velocity at the inlet (bottom) of the gasifier (cm/s) uf ) gas velocity at the face or front of the monolith (cm/s) W ) weight of the monolith (kg) WHSV ) weight hourly space velocity for the biomass in the gasifier [kg of biomass/h‚(kg of gasifier bed material)] xi ) mole fraction of the ith compound in the reacting mixture Xtar ) conversion of the tar Xtar,inlet, Xtar,exit ) conversion of the tar at the inlet and exit of the monolith, respectively yi,j ) fraction of the ith reactant that reacts by the jth reaction Greek Symbols σAi ) Lennard-Jones function force constant, calculated as 1/ (σ + σ ) (Å) 2 A i ΩD ) collision integral, a function of kBT/Ai, with kB as Boltzmann’s constant and Ai as another force constant in the Lennard-Jones potential function (Ai is calculated by Ai ) xAi) τmon ) space time of the gas in the monolith, defined as W/Q0 [kg‚h/m3(Tcat,wet)]
τreactor ) space time of the gas in the monolithic reactor, defined as Vreactor/Q0 [mreactor3‚h/m3(Tcat,wet)] η ) effectiveness factor
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Received for review May 18, 2004 Revised manuscript received October 4, 2004 Accepted October 4, 2004 IE040148H
