Ind. Eng. Chem. Res. 2001, 40, 2801-2809
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Optimization of Geometric Properties of a Monolithic Catalyst for the Selective Hydrogenation of Phenylacetylene Theo Vergunst,† Freek Kapteijn,* and Jacob A. Moulijn Delft University of Technology, Faculty of Applied Sciences, Section on Industrial Catalysis, Julianalaan 136, 2628 BL Delft, The Netherlands
Two major characteristics of monolithic catalysts, which are often ignored in performance considerations, are the cell density and the thickness of the catalytic layer. The impact of these two parameters becomes apparent from a monolithic reactor simulation applied to the selective hydrogenation of phenylacetylene to styrene. The reactor model takes kinetics, diffusion, and mass transfer into account. As functions of the cell density and coating thickness, the required reactor volume for a phenylacetylene conversion level of at least 92% and the yield of undesired ethylbenzene were calculated.With increasing coating thickness, the conversion of phenylacetylene becomes limited no longer by the amount of catalyst present in the reactor but by diffusion. External mass transfer was never found to be rate-limiting. The required reactor volume shows a minimum at the transition from kinetically limited control to internal diffusion-limited control. A higher cell density (smaller channels) favors a smaller reactor. To minimize the conversion of styrene, a very thin coating layer suffices, but this maximizes the required reactor volume. Introduction Monolithic catalysts are examples of structured catalysts and are characterized by long parallel channels separated by thin walls. A monolithic catalyst consists of either a ceramic structure that is coated with a thin oxide1,2 or carbon3,4 layer that acts as a catalyst support material (coated type) or a structure made completely of the support material (integral type).4,5 The highly open structure of monoliths causes low catalyst content per unit reactor volume as compared with other fixedbed reactors. As with other reactor types, the reactor and catalyst dimensions of a monolithic catalyst reactor can be adjusted to the kinetics of the reaction, the desired production rate, and the desired selectivity of the reaction. By varying the coating thickness of a monolithic catalyst, the amount of catalyst per unit reactor volume can be varied, independently of the reactor volume. However, a change in the coating thickness not only changes the amount of catalyst in the reactor, but also affects the linear velocity in the channels at constant feed flow rates and reactor dimensions, the degree of conversion, the diffusion distance, and thereby, the product distribution in the case of a complex reaction network. Moreover, the channel size of the monolithic supports can be changed as well, so that the amount of catalyst per unit volume reactor can be changed without changing the coating thickness. This paper describes the optimization of a monolithic catalyst using a selective hydrogenation reaction in a gas-liquid solid-catalyzed system, the selective hydrogenation of phenylacetylene to styrene (Figure 1), as a model reaction. The selective hydrogenation of phenylacetylene was chosen because it is a fast reaction operated in practice, the intermediate product is desired (selectivity problem), and kinetic data for this reaction * Corresponding author:
[email protected]. † Present address: Avantium Technologies BV, P.O. Box 2915, 1000 CX Amsterdam, The Netherlands.
are available.6-9 For a specified feed flow rate and desired level of conversion, the reactor dimensions can be optimized to minimize the consecutive hydrogenation of styrene into ethylbenzene, taking into account kinetics, diffusion, and mass transfer. Model Description Process. In the production of styrene through the dehydrogenation of ethylbenzene, phenylacetylene is formed as a side product. For several applications, e.g., polymerization reactions, this impurity has to be removed with minimal hydrogenation of styrene.6 For the selective hydrogenation of phenylacetylene, an eggshell Pd/Al2O3 catalyst is usually used.7,8 Because phenylacetylene reacts rapidly, a temperature of only 298-353 K and a pressure of 0.1-0.4 MPa are required. Table 1 shows the typical feed composition and operating conditions used in the selective hydrogenation of phenylacetylene.7,9 A typical styrene production plant produces 500 kt of styrene per year6 (∼ 62.5 mL3 h-1). Reaction Kinetics. Mochizuki and Matsui8 developed kinetic expressions for the selective hydrogenation of phenylacetylene (eq 1). The first reaction represents the hydrogenation of phenylacetylene, and the second reaction represents the hydrogenation of styrene (Figure 1). Both reactions are based on surface reactions between adsorbed hydrocarbon and two adsorbed hydrogen atoms competing for the same sites. Ethylbenzene is assumed not to adsorb on the surface.
KPACPAKH2CH2 rVC,1 ) k1 (1) (1 + KPACPA + KSTCST + xKH2CH2)3 rVC,2 ) k2
KSTCSTKH2CH2 (1 + KPACPA + KSTCST + xKH2CH2)3
The kinetic parameters used are listed in Table 2. The model values8 were confirmed by independently per-
10.1021/ie000712y CCC: $20.00 © 2001 American Chemical Society Published on Web 05/23/2001
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Figure 1. Reaction scheme of the consecutive hydrogenation of phenylacetylene to styrene and ethylbenzene. Table 1. Feed Compostion9 and Operating Conditions Used in Selective Hydrogenation of Phenylacetylene T P φv CPA,bulk CST,bulk CEB,bulk ξPA
323 0.4 1.74 × 10-2 80 3440 1475 0.92
(K) (MPa) (m3 s-1) (mol m-3) (mol m-3) (mol m-3)
Table 2. Kinetic Parameters for the Selective Hydrogenation of Phenylacetylene8,a at 323K k1 k2 KPA KST KH2
1522 311.1 7.3 × 10-3 2.9 × 10-4 3.6 × 10-3
(mol mc-3 s-1) (mol mc-3 s-1) (mL3 mol-1) (mL3 mol-1) (mL3 mol-1)
a
Parameter values8 were converted to meet dimensions used in this model using the following data: R ) 0.4, dp ) 3.5 × 10-3 m, δPd ) 50 × 10-6 m.
formed experiments and optimization procedures.9 The values are only available at 323 K. Reactor Model. For the monolithic reactor, isothermal plug-flow conditions are assumed. The distribution of gas and liquid across the diameter of the reactor is uniform. Gas and liquid flow concurrently downward in so-called Taylor or bubble-train flow1,10 through the monolithic channels, that is, gas bubbles are separated from each other by small liquid plugs. The pressure difference across the monolithic catalyst is neglected. All physical properties are assumed to be constant throughout the reactor. The catalytic activity is uniformly distributed through the coating layer and is constant. The monolithic reactor is modeled using a onedimensional catalyst model, which means that the monolithic catalyst is modeled using slab geometry, with mass transfer to and from the catalyst layer and reaction and diffusion occurring in the catalyst layer. Because of the presence of gas bubbles, a continuous local supply of hydrogen is secured, so hydrogen is not a limiting component along the reactor length. Otherwise, multiple feed points along the length of the reactor are assumed to maintain a constant gas hold-up and concentration. In this model, monolithic supports with squareshaped channels are considered. The structure of a monolithic catalyst is depicted in Figure 2 (following ref 11). The geometric parameters4,11 of monoliths used in the simulations, such as cell density, geometric surface area, void fraction, hydraulic diameter, volume fraction catalyst, and average coating thickness, can be calculated with the equations and data presented in Table 3. The data used for the optimization are included in Table 3. The mathematical description of the monolithic reactor is depicted in Figure 3 and Table 4 and explained here. The change in bulk concentration of the reactant and products is calculated as the mass transfer to the coating (eq 8). The surface concentration Ci|y)1 is
Figure 2. Definition of geometric properties for a monolithic support, following ref 11. dch is the channel diameter, δw is the wall thickness, δc is the coating thickness, and R is the fill radius (the radius of a circle that fits in the corner of the coating).
required for this calculation but is unknown. It can be calculated from a mass balance over a slice of catalyst at axial position z and a mass balance across the catalyst surface. The mass balance across a slice of coating at axial position z results in a second-order differential equation in the y direction (eq 10). The mass balance across the coating surface states that the transport to the surface by mass transfer equals the transport into the coating (eq 12). The mass-transfer coefficients are calculated using Sherwood correlations12-14 from the literature. For hydrogen, the expressions used are adapted to include mass transfer involving the gas phase as well. The catalyst effectiveness is defined according to eqs 2123. Optimization Procedure The monolithic reactor considered in this modeling is a single-pass reactor with no recirculation of reactants. The reactor is built from pieces that are 1 m in length. The preparation of longer catalysts is impractical. Both the gaseous and liquid reactants are fed to the top of the reactor with constant compositions and flow rates. At the entrance of the monolith, the liquid is completely saturated with gas. The volumetric ratio gas/ liquid is 1:1 (at 0.4 MPa), with hydrogen as a pure gasphase reactant. The physical properties of the liquid are those of pure styrene and are listed in Table 5. In the model, the cell density, n, and the coating thickness, δc, were varied. The coating thickness was varied between 20 and 140 µm, with the fill radius, R, set equal to the coating thickness. A fractional conversion of phenylacetylene of at least 0.92 was required. Using this requirement, the reactor dimensions (dm and Lm) and the amount undesired ethylbenzene produced could be calculated as a function of cell density and
Ind. Eng. Chem. Res., Vol. 40, No. 13, 2001 2803 Table 3. Geometric Parameters for Monolithic Structures with Square-Shaped Channels4,11 cell density
(m-2)
geometric surface area
(m-1)
n)
(2)
R 2 m ) n[(dch - δw - 2δc)2 - (4 - π)R2] m dh ) 4 am
[
]
am ) 4n (dch - δw - 2δc) - (4 - π)
void fraction hydraulic diameter
1 dch2
(3) (4)
( )
(m)
volume fraction of catalyst
(5)
2
average coating thickness
xc ) n(dch - δw) - m xc Lc ) am
(m)
(6) (7)
cell density n (cpsi/cm-2)
channel diameter dch (mm)
wall thickness δw (µm)
200/31 400/62 600/93 1100/170
1.796 1.270 1.037 0.766
270 165 112 64
Table 4. Model Equations for the Monolithic Reactor Mass Balances
∂Ci,bulk kLSam(Ci|y)1 - Ci,bulk) Lm ) ∂z mL v ∂2Ci ∂y
)-
Lc2 De,i
2
∑
(8)
νi,jrvc,j
z ) 0 w Ci,bulk ) Ci,bulk,0
(9)
∂Ci )0 ∂y
(11)
y)0w
(10)
j
kLC(Ci|y)1 - Ci,bulk) ) -
De,i ∂Ci | Lc ∂y y)1
(12)
Mass Transfer
( ) ( )
kLSdh Re Sc Sh ) ) 3.51 D Ls/dh
( |
0.44
Lplug dh
-0.09
(13)
Gas Phase
)
kov CH2
De,H2 ∂CH2 P )| y)1 HeH2 Lc ∂y y)1
(14)
kovam )
kGLaGLkLSam + kGSaGS kGLaGL + kLSam
(15)
δf ) dh0.18[1 - exp(-3.1Ca0.54)] kGLdh Sh ) ) 0.69(1 + 0.724ReL0.48Sc1/3) DH2
kGS )
DH2
(16)
δf aGS ) (1 - L)am
(17)
(18) 2
(19)
aGL )
π(dh - 2δf) Lm Lplugdh2
(20)
Catalyst Effectiveness
ηi,PA ) ηi,H2 ) ηe,PA )
kLSam(CPA,bulk - CPA|y)1) xcrvc,1(Ci|y)1) kovam(CH2,bulk - CH2|y)1) xc[rvc,1(Ci|y)1) + rvc,2(Ci|y)1)] rvc,1(Ci|y)1) rvc,1(Ci,bulk)
(21) (22) (23)
coating thickness, setting the linear superficial velocity to v ) 0.2 m s-1. The calculations were performed using Matlab. Simulation Results Figure 4 shows the changes in geometric properties of the monolithic supports as a function of coating thickness and cell density. An increase in the coating thickness results in a decrease in the geometric surface area, as well as a decrease in the void fraction. On the
Figure 3. Coordinate system for modeling of a monolithic reactor and depiction of the mass-transfer pathways for gaseous compounds.12 z ) 0 is the reactor inlet.
other hand, the volume fraction of catalyst increases with increasing coating thickness.
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Table 5. Properties6,15,16 of Styrene at 323 K µST γST FST HeH2,ST DSTa DH2a a
0.452 × 10-3 29.5 × 10-3 879 23.9 × 103 2.89 × 10-9 7.26 × 10-9
(Pa s) (N m-1) (kg m-3) (Pa mL3 mol-1) (m2 s-1) (m2 s-1)
De ) 0.1D.
Figure 5. (a) Required reactor length and (b) resulting reactor volume for a fractional conversion of g0.92 of phenylacetylene at constant feed flow rate as a function of coating thickness for monolithic supports with cell densities of (() 200 cpsi, (2) 400 cpsi, (b) 600 cpsi, and ()) 1100 cpsi.
Figure 6. Ratio of ethylbenzene produced to phenylacetylene converted at a fractional conversion of 0.92 as a function of coating thickness for monolithic supports with cell densities of (() 200 cpsi, (2) 400 cpsi, (b) 600 cpsi, and ()) 1100 cpsi. Figure 4. (a) Geometric surface area, (b) void fraction, and (c) volume fraction of catalyst as functions of the coating thickness for monolithic supports with cell densities of (() 200 cpsi, (2) 400 cpsi, (b) 600 cpsi, and ()) 1100 cpsi.
The required diameter of the monolithic reactor increases with coating thickness. With increasing coating thickness and decreasing cell density, the void fraction of the monolithic catalyst decreases at a constant volumetric feed flow rate for every geometry. This results in an increase in the reactor diameter to maintain a constant linear velocity. Figure 5 shows the reactor length and reactor volume required to obtain a fractional conversion of phenylacetylene of at least 0.92. With increasing cell density and increasing coating thickness, the required reactor length decreases, and the reactor volume shows a shallow minimum at a coating thickness of 80 µm. An important parameter in this process is the ratio between ethylbenzene produced and phenylacetylene
converted. This ratio is plotted in Figure 6 for the various geometries. The ratio increases with the coating thickness and with the cell density as well. Two parameters that indicate diffusion and masstransfer limitations are the internal and external effectiveness factors (eqs 21-23). These are shown in Figure 7. The internal effectiveness factor decreases with increasing coating thickness and increasing cell density. The external effectiveness factor for reaction 1 shows a minimum that shifts with the cell density. As an example, the concentration profiles along the length of the reactor for a 600 cpsi monolithic catalyst with a coating thickness of 80 µm are presented in Figure 8. Phenylacetylene is initially converted rapidly, but the conversion is slower at the end of the reactor. There is a net production of styrene in the first part of the reactor, but the styrene is converted further down the length of the reactor. Ethylbenzene is produced over the whole reactor length.
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Figure 7. (a) Internal catalyst effectiveness factor and (b) external effectiveness factor for the selective hydrogenation of phenylacetylene as a function of coating thickness for monolithic supports with cell densities of (() 200 cpsi, (2) 400 cpsi, (b) 600 cpsi, and ()) 1100 cpsi.
Figure 8. Concentration profiles relative to inlet concentrations in the axial direction for a monolithic catalyst of 600 cpsi and a coating of 80 µm: ∆, phenylacetylene (C0,PA ) 80 mol m-3); 0, styrene (C0,ST ) 3440 mol m-3); and O, ethylbenzene (C0,EB ) 1475 mol m-3).
The concentration profiles of the reactants and products in the catalyst layer are shown in Figure 9 for a 600 cpsi monolithic catalyst with a coating thickness of 80 µm at different axial positions in the reactor. The concentrations of phenylacetylene and hydrogen decrease, whereas the concentration ethylbenzene increases toward the inner part of the coating layer (y ) 0) for all axial positions. The concentration profile of styrene shows that styrene is produced in the initial reactor section. Further downstream, the concentration profiles of styrene show decreasing trends similar to the profiles of phenylacetylene. Discussion Assumptions. At v ) 0.2 m s-1, the assumption of plug flow is valid. It was shown17 that, under the conditions modeled, Pem > 20, indicating plug-flow behavior. In hydrodynamic experiments, it was shown18
Figure 9. Concentration profiles across the catalytic layer at different positions along the length of the reactor for a 600 cpsi monolithic catalyst with a coating thickness of 80 µm: 4, phenylacetylene; 0, styrene; O, ethylbenzene; and (, hydrogen.
Figure 10. Pressure difference as a function of coating thickness for down-flow Taylor flow for a monolithic support with cell densities of (() 200 cpsi, (2) 400 cpsi, (b) 600 cpsi, and ()) 1100 cpsi, following Heiszwolf.19 The pressure difference is defined in the direction of flow.
that the operation of monolithic supports with negligible pressure differences is possible. The total pressure difference consists of two parts: friction with the wall and gravity forces. In every case, he contribution of gravity will be the same because of the constant liquid hold-up; the friction with the wall becomes more important for higher cell densities because of the higher geometric surface area. For down-flow-operated monolithic reactors, these factors counteract, resulting in a small pressure difference (Figure 10, according to Heiszwolf19). For low cell densities, the pressure increases in the axial direction (∆P/L > 0), whereas for high cell densities, the pressure decreases in the axial direction (∆P/L < 0). In the former situation, gravity dominates the pressure difference, whereas in the latter friction, with the wall dominates the pressure difference. The operating window of monolithic reactors with respect to superficial gas and liquid velocity is discussed in more detail elsewhere.18,19 Rather high superficial velocities
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are required for a hydrodynamically stable operating regime. Using literature data,20 an adiabatic temperature rise between 14.6 and 17.4 K is calculated, which is too high for isothermicity to be assumed. A multiple-bed reactor with interstage cooling could be necessary to ensure this assumption. However, for practical applications, this temperature rise could be acceptable. Because the reactant and product molecules are very much alike, it is valid to assume similar and constant physical properties for the liquid phase in the reactor. A one-dimensional model was used for the catalyst. This model uses an average coating thickness, defined as the volume of coating divided by the external surface area of the coating. A more complex two-dimensional catalyst model could be applied12,21 to account for the changing layer thickness in the corners of the channel. However, as an approximation, the one-dimensional model performed well enough12,21 to show the effect of the coating thickness on the product distribution and the reactor dimensions. The conversion of hydrogen is assumed to be negligible. The amount of hydrogen fed to the reactor is 2.6 mol s-1. In the case of the lowest selectivity (n ) 1100 cpsi, δc ) 140 µm) the amount of hydrogen required is 3.2 mol s-1. This amount is not negligible, and hydrogen must therefore be fed to the reactor at multiple feed points along the reactor length, for instance, between two stacks of monoliths. This ensures a constant hydrogen concentration in the liquid phase over the length of the reactor. Kinetics. The selective hydrogenation of phenylacetylene has not often been studied with respect to the kinetic rate expressions. The model developed and the parameters used8 were confirmed with independently performed experiments and optimization procedures.9 Chaudhari et al.22 studied the selective hydrogenation of phenylacetylene and found a similar ratio for the reaction rate constants k1/k2. Although all of their experiments were either performed at lower reactant concentrations8 than used in this modeling or without addition of styrene to the feed mixture,22 the agreement between the data was good enough to show the reliability of the data. It is stated that the reaction only occurred at the external surface of the catalyst pellets8,9 (eggshell). For calculation reasons, a Pd-containing layer of 50 µm has been assumed, which is a thin eggshell layer. Assuming a different layer thickness would change the activity per unit volume of catalyst and, as a consequence, the optimal thickness of the catalyst layer resulting in the smallest reactor volume. Simulation Results. The simulation results show a decrease in reactor length required to obtain a phenylacetylene fractional conversion of 0.92 with increasing coating thickness (Figure 5). A sharp decrease in required length for thin coating layers is followed by a less steep decrease for thicker coating layers. The required reactor volume, therefore, shows a minimum for all cell densities. This minimum is caused by the decrease in reactor length and the increase in reactor diameter with increased coating thickness. The dependency of the reactor length on the coating thickness is caused by the increased influence of internal diffusion limitation. The increase in coating thickness is therefore not completely utilized for the conversion of phenylacetylene, so that the required reactor length is not
inversely proportional to the coating thickness, but rather levels off for thicker coating layers. The exact position of the minimum (∼80 µm) is a function of the reaction rate constants. The selectivity of the reaction drops with increasing coating thickness and increasing cell density (Figure 6). With increasing coating thickness, the styrene present can more easily be converted to ethylbenzene, which results in a lower selectivity to styrene. As shown in Figure 9, the concentration of phenylacetylene decreases toward the catalyst center (y ) 0), so that reaction 1 runs more slowly and reaction 2 is less suppressed by the presence of phenylacetylene. A thicker coating layer results in a larger decrease in phenylacetylene concentration than a thin layer, so that more styrene can be converted. For higher cell densities and equal coating thicknesses, the average coating thickness, Lc, is higher, because of the fill radius, resulting in an even lower selectivity. With increasing coating thickness, the amount of ethylbenzene produced increases. This occurs only up to a certain level, as shown in Figure 6 by the flattening of the curves, which shows the amount of ethylbenzene produced per amount of phenylacetylene converted. For thick coating layers, hydrogen is exhausted within the catalyst layer, so that no conversion can occur in the center of the catalytic layer. This can also be concluded from the fact that the product ηi,H2Lc, which can be called an effective coating layer, approaches a constant value for thick coating layers. The latter is evident as the internal catalyst effectiveness factor is inversely proportional to the Thiele modulus when the Thiele modulus becomes large.23 The Thiele modulus is proportional to the characteristic length of a catalyst particle, so that ηi ∼ 1/φ|φf∞ ∼ 1/Lc and ηiLc becomes constant. Increasing the coating layer does not lead to further reaction because of hydrogen depletion, so that the selectivity becomes constant and the required reactor length does not decrease further. Phenylacetylene is converted throughout the catalytic layer, although the concentration profiles become flat for y e 0.2. Both the hydrogen and phenylacetylene concentrations become sufficiently low that the rate of reaction 1 drops to almost zero. Ethylbenzene is the final reaction product and is produced throughout the catalytic layer. Even at y ) 0, the concentration styrene is high enough to produce some ethylbenzene. The intermediate product styrene shows an increasing concentration at high phenylacetylene concentration. At lower phenylacetylene concentration, styrene is (only) converted. The internal effectiveness factor as a function of coating thickness shown in Figure 7 displays a trend similar to the results of Sie.24 With increasing coating thickness, the internal effectiveness factor decreases. The transport of reactants to the interior of the coating cannot cope with the reaction rate. The catalyst effectiveness generally drops faster with decreasing diffusion rate and increasing reaction rate. At a coating thickness of ∼80 µm when ηi,PA < 0.6 (Figure 7), internal mass-transfer limitations become pronounced.24 Before that, the amount of catalyst limits the conversion rate. The effect of cell density on the internal catalyst effectiveness, as shown in Figure 7, is caused by the definition of the characteristic length, Lc. At a given coating thickness δc, the characteristic length, Lc, is larger for a higher cell density because of the larger
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Figure 11. Styrene formation as a function of phenylacetylene conversion for a 600 cpsi monolith with a coating of 80 µm. (9) No diffusion limitation, (2) diffusion limitation of both phenylacetylene and hydrogen, and (b) diffusion limitation of phenylacetylene.
contribution of the filled corners. At a given Lc, the internal catalyst effectiveness is the same for different cell densities. Although monolithic catalysts are often idealized to have small diffusion distances,1 diffusion limitation can still occur. This diffusion control is also observed in gas-phase applications.21,25 The external effectiveness factor is always larger than 0.90, indicating no or little external mass-transfer limitation. This is due to the excellent mass transfer in monolith channels in Taylor-flow mode.2,13,14 The external effectiveness factor as a function of coating thickness (Figure 7) shows a minimum. This minimum is caused by two opposing factors. The first factor is that, with decreasing channel diameter (and so, with increasing coating thickness), the mass-transfer coefficient increases. On the other hand, the internal effectiveness is high foe thin coating layers and becomes smaller at higher coating thicknesses, so that the total local activity does not increase linearly with the amount of catalyst present. The external effectiveness factor, therefore, first becomes smaller (increased total activity in the layer) and then becomes larger (constant total activity and only enhanced mass transfer). Often, diffusion limitation is said to have a negative effect on the selectivity of a reaction.24 In the previous section, however, it was concluded that, because of hydrogen exhaustion, the ethylbenzene production was limited. The effect of the diffusion limitation was therefore studied by calculating the amount of styrene produced as a function of the conversion of phenylacetylene for a 600 cpsi monolith with a coating of 80 µm for three different cases: (i) diffusion limitations of both hydrogen and phenylacetylene, (ii) no diffusion limitation of hydrogen, and (iii) no diffusion limitation at all (Figure 11). Without diffusion limitations, the least styrene is converted to ethylbenzene; without diffusion limitation for hydrogen only, the most styrene is converted to ethylbenzene. This demonstrates the usefulness of a thin coating layer (no diffusion limitations, so highest selectivity) and the use of an optimized low hydrogen pressure, which causes diffusion limitation of hydrogen and therefore lower conversions of styrene. The Taylor-flow mode thereby ensures a constant hydrogen concentration along the length of the reactor. In this respect, complete hydrogen depletion in the catalyst layer might not be desirable, as it might lead to undesired oligomerization of styrene and so deactivation. Here, another advantage of the monolithic reactor becomes apparent, as thin coatings allow this regime to be avoided, increasing the catalyst’s time-on-stream.
Figure 12. Layout of a monolithic reactor for the selective hydrogenation of phenylacetylene, consisting of four segments. Between the segments, hydrogen can be fed, and heat can be exchanged.
The Monolithic Reactor. In this part of the discussion, the construction of a monolithic reactor and some general considerations for using a monolithic reactor are addressed. The layout of a monolithic reactor used for the selective hydrogenation of phenylacetylene could be envisaged as shown in Figure 12. To obtain a reasonable length-to-diameter ratio, the reactor is constructed of four parts (other numbers are also possible). The complete feed is passed through segment 1, followed by passage through segments 2, 3, and 4. Between the segments, fresh hydrogen is supplied at the top of the reactor, and interstage cooling between the segments is possible as well. Using multiple segments, the catalyst in each segment can be different too. A thinner coating layer in the latter part(s) of the reactor could diminish the conversion of styrene into ethylbenzene and avoid deactivation. The simulation presented in this paper displays two optima. The first is related to the reactor volume. The minimum reactor volume is obtained with a monolithic catalyst of 1100 cpsi and a coating thickness of 80 µm. At this point, diffusion limitation becomes important. This optimum represents the point with the lowest investment costs for a reactor (assuming cost ∝ volume). The second optimum corresponds to a minimum production of ethylbenzene. From Figures 5 and 6, it can be deduced that this optimum is obtained for a monolithic reactor with the thinnest possible coating layer. Obviously, the reactor will then be unfeasibly large. Extrapolation of the obtained simulation results can provide more general information about the application of gas-liquid solid-catalyzed reactions in monolithic reactors. The simulation results showed two conversionlimiting factors: the amount of catalyst and diffusional transport. For slow reactions, diffusion is relatively fast, and the amount of catalyst will be conversion-limiting. For moderately fast catalytic reactions, such as that considered in this paper, diffusion can become conversion-limiting when the catalytic layer is too thick. For fast reactions, diffusion limitation and external masstransfer limitations can occur. Especially in the case of consecutive reactions, diffusion limitations should be minimized if an intermediate product is desired. In the current case, one could adjust the operating conditions such that, in the catalytic layer, the hydrogen becomes depleted whenever the phenylacetylene is converted, avoiding the consumption of styrene by residual hydrogen. In the absence of hydrogen, however, undesired
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reactions might occur, such as oligomerization and “coking”, which should be avoided in fixed-bed reactors. A slight excess of hydrogen during operation in the transition between kinetic and diffusion control in a high-cell-density monolith is therefore the preferred choice for this reaction system. Conclusions The simulated monolithic reactor for the selective hydrogenation of phenylacetylene yields results that can be well explained from chemical engineering principles. The two variables, cell density and coating thickness, that were changed during the optimization can have a pronounced effect on the reactor size and product distribution. A high cell density is more favorable, because both the geometric surface area and the volume fraction of the catalyst are largest for the highest cell density when a constant coating thickness is considered. Because of diffusion and reaction inside the catalytic layer, the amount of ethylbenzene produced undesirably increases with the coating thickness. Depending on the optimization criterion considered, two optima can be found. The first minimizes the required reactor volume; the second minimizes the amount of side product. The smallest reactor volume is found for a catalyst with a coating thickness of 80 µm, corresponding to operation in the transition between kinetic and diffusion control. This thickness is a function of the reaction rate constants, as well as other parameters. External mass transfer was not found to be ratelimiting for the current system. A thin coating layer (
