Modeling and Simulation of Hydrocracking of Fischer–Tropsch

Article pubs.acs.org/IECR

Modeling and Simulation of Hydrocracking of Fischer−Tropsch Hydrocarbons in a Catalytic Microchannel Reactor M. Irfan Hosukoglu, Mustafa Karakaya, and Ahmet K. Avci* Department of Chemical Engineering, Bogazici University, Bebek 34342, Istanbul, Turkey ABSTRACT: Product quality of Fischer−Tropsch synthesis is improved by catalytic hydrocracking which converts heavy hydrocarbon fractions (wax) to commercially valuable fuels. The process is highly exothermic and requires strict temperature control; high temperatures cause overcracking to lower, commercially undesired hydrocarbons, whereas low temperatures reduce the conversions. Running hydrocracking in microchannel reactors is promising, since submillimeter dimensions lead to significant compaction that favors robust temperature control. This work investigates modeling and simulation of hydrocracking in a heatexchange-integrated microchannel reactor involving parallel groups of square-shaped cooling and catalyst-coated reaction channels. Effects of material type and thickness of the wall separating the channels, and operating parameters (reactant and coolant feed temperatures and space velocity of the reactant stream) on reaction temperature and product distribution are investigated. Mole fractions of the products in the diesel cut (C19−C22) and jet cut (C11−C18) ranges are highly sensitive to operating parameters due to fast heat transport. The process suffers from overcooling and reduced conversions in reactors characterized by thicker walls with high thermal conductivities, whereas hot spots may exist in reactors characterized by thinner walls with low thermal conductivities. Temperature and product distributions in hydrocracking can be optimized within the pertinent operating window by careful configuration of the reactor. improving cold flow properties of the middle distillates by their isomerization.5,6 Hydrocracking of a linear paraffin occurs after a variety of successive reactions. The catalyst is bifunctional and characterized by the presence of acidic sites, which are active in cracking function, and of a metallic site for the hydrogenation/ dehydrogenation steps. The first step is the dehydrogenation of alkanes to produce olefin intermediates, which then forms carbenium ions that are adsorbed on the acid sites of the catalyst. The adsorbed carbenium ions are catalytically converted by isomerization and later by cracking with respect to the positively charged carbon atom. After the conversion, the products (i.e., isomerized and cracked species) desorb from the acid site and undergo hydrogenation on the metal site to form the corresponding paraffins. As a result of this mechanism, the number of carbon atoms in paraffins is reduced and degree of branching is increased.7,8 Various metals such as Pt, Ni, and Mo, and different acidic supports such as zeolites, sulfated zirconia, silicoaluminaphosphates, and amorphous silica− alumina are reported to be involved in the structures of hydrocracking catalysts. The degree of isomerization and cracking can be controlled by changing the acidic strength or by the shape selectivity of some zeolite-structured materials.9 Apart from the proper choice of the catalyst, temperature control is essential for the success of FT wax hydrocracking. The network of reactions releases high amounts of exothermic heat which needs to be removed effectively from the catalyst

1. INTRODUCTION Petroleum crude oil is the primary raw material of the petrochemical industry and of transportation fuels. However, depletion of the quality and level of its reserves is making crude oil refining more difficult and expensive for use as a primary fuel in the future. Although alternatives such as solar and wind energies offer sustainable solutions to energy needs, they cannot replace fossil fuel use completely mainly due to the compatibility of liquid products with the existing infrastructure. Production of liquid fuels via Fischer−Tropsch (FT) process, a synthetic route involving conversion of syngas (CO + H2) to a wide range of hydrocarbons, seems to be a promising option since it does not depend on crude oil and utilizes alternative resources such as natural gas, coal, and biomass.1−3 FT synthesis is generally operated at the low temperature (LT) mode for decreasing the selectivity toward light hydrocarbons (C1−C4). In LTFT synthesis, long chain (C22+) paraffins with boiling points in excess of 370 °C are mainly produced and the yield of middle distillate range (predominantly C9−C22 hydrocarbons in the boiling point interval of 150−370 °C) is limited. Although middle distillates are characterized by cetane numbers (above 75) much higher than the required limit of 51, their cold flow properties are poor which cause problems both in their direct use in internal combustion engines and in their indirect use as a blending component in diesel pool. Improving the catalysts and process conditions in FT synthesis allows adjustment of selectivity toward desired hydrocarbon fractions, but only at limited levels.4 Hydrocracking of FT product stream (also called wax mixture, mainly characterized by C5−C70 n-paraffins) is the method for increasing the yield of middle distillates by cracking of high boiling point (>370 °C) C22+ components and for © 2012 American Chemical Society

Special Issue: CAMURE 8 and ISMR 7 Received: Revised: Accepted: Published: 8913

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equations, all of partial differential type, using the finite element method (FEM) under COMSOL Multiphysics software.

bed. Insufficient cooling followed by high temperatures leads to products lighter than the desired specifications, and, in extreme cases, thermal deactivation of the catalyst bed.10 In the cases of overcooling, however, the process can suffer from reduced reaction rates due to the low temperatures. A novel way of meeting precise heat transfer management and temperature control can be through the use of microchannel reactors, the units involving flow paths with characteristic dimensions in the 10−4 to 10−3 m range, which can lead to very high surface area to volume ratios up to ca. 50 000 m2/m3.11,12 Microchannel units are usually made of metallic substrates and catalysts are generally in the form of thin layers coated on the inner walls of the channels. Combination of these properties ends up with very high rates of heat transfer allowing robust temperature control. The resulting uniform temperature distribution throughout the reactor helps in efficient use of the catalyst. Moreover, small channel volumes ensure inherent safety, while well-defined microflow paths lead to narrow and low residence time distribution.12,13 In contrast to the investigations evaluating microchannel reactor performance for FT synthesis,14−18 the number of studies addressing FT wax hydrocracking in microchannel reactors in the open literature is very limited. One recent study on this topic by Branco et al.19 reported the feasibility of running hydrocracking in the context of an offshore microchannel GTL plant producing synthetic crude oil. A big portion of the research related to microchannel hydrocrackers is being conducted by Velocys Inc. (www. velocys.com) and is in the form of classified information. This work involves parametric analysis of and aims to provide insights for FT wax hydrocracking in a heat exchange integrated microchannel reactor through a series of computerbased simulations. FT wax is assumed to be composed of normal paraffins up to 70 carbon atoms. The reactor system, shown in Figure 1 and explained in the next section, is

2. MODELING OF THE MICROCHANNEL REACTOR SYSTEM The microchannel reactor heat-exchanger system is illustrated in Figure 1. The system consists of rectangular microflow paths either used as reaction (hydrocracking) channels, whose inner walls are considered to be coated with a porous Pt on silica− alumina catalyst layer, or as cooling channels for absorbing the exothermic heat released from the reactions. Parallel groups of reaction and cooling channels are placed consecutively and separated by metallic walls. Since properties of the fluids flowing within the channels of the same (horizontal) group do not differ, it is assumed that heat transfer between the channels of the same group (i.e., in z-direction) is negligible compared to the heat flow between the channels of different groups (i.e., in the y-direction). Therefore, gradients in the z-direction are eliminated. This simplification allows the calculations to be based on a two-dimensional unit cell, which is made up of the domain between the centerlines of two channels through the channel length, as shown in Figure 1. The repeating pattern of the channels in the y-direction and the resulting symmetry allow the consideration of the half-channel heights in the unit cell as the characteristic section of the multichannel unit. It is reported that the use of a three-dimensional model and a simplified two-dimensional model in the simulation of a heat exchange integrated microchannel reactor system similar to the one shown in Figure 1 gave nearly identical results.20 In this study, hydrocracking reactions are considered to run over a silica−alumina supported Pt catalyst. The feed stream is composed of excess hydrogen and a mixture of C5−C70 linear paraffins, whose compositions are adapted from the study of Pellegrini et al.21 The reaction network involves the conversion of linear paraffins to iso-paraffins, described by 67 isomerization reactions, and hydrocracking of the resulting iso-paraffins to smaller hydrocarbons, described by 66 cracking reactions. Mole fraction of hydrogen in the feed mixture is known to be around 95% and does not decrease below 93% within the investigated conversion range. Therefore, the reactive stream is assumed to be in vapor phase only since the vapor fraction of the reactive volume does not fall below 96% in typical hydrocracking conditions.21 Similar to the reaction channel, no condensation occurs in the cooling channel as the coolant (steam) is fed above its boiling point and is heated by the exothermic heat of cracking reactions along the channel. On the basis of these facts, construction of the mathematical model is based on a gas−solid type of operation with fluid (gas) and washcoat (catalyst) phases. Modeling and simulation of hydrocracking reaction network and heat transfer between the microchannels require simultaneous solution of momentum, mass, and energy conservation equations in four domains; the hydrocracking channel, catalytic washcoat layer, solid wall, and the cooling channel. It is assumed that the reactions take place only within the washcoat domain and homogeneous gas-phase reactions are neglected. The fluids, reactive mixture and steam, which are arranged to flow countercurrently, are assumed to be of incompressible Newtonian type and are treated as ideal gases. The entire operation is considered to be at steady state. Momentum, mass, and energy conservation equations used to model the reaction system are given in Table 1. Conservation of momentum for the gas phase is described by Navier−Stokes

Figure 1. Heat-exchange integrated microchannel reactor and the unit cell

considered to be a microchannel network of square-shaped adjacent cooling and wall-coated catalytic reaction channels in the form of parallel arrays, which are arranged such that reactant and coolant streams flow in successive channels in countercurrent mode. Parametric analysis is conducted by investigating the effects of material type and thickness of the wall separating the channels, and operating parameters (reactant and coolant feed temperatures, space velocity of the reactant stream) on conversion and product distribution (both defined in the next section) and on reaction temperature. The results, presented in section 3, are obtained by the simultaneous solution of the momentum, mass, and energy conservation 8914

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Table 1. Model Equations Used to Describe Transport and Reactions in Microchannels Fluid Phase equation of continuity

∂vx , i ∂xi

+

∂vy , i ∂yi

=0

⎛ ∂ 2v ⎛ ∂v ∂p ∂vx , i ⎞ ∂ 2vx , i ⎞ x ,i ⎟ ⎟⎟ = − i + μ ⎜⎜ x , i + ρf, i⎜⎜vx , i + vy , i 2 i ∂yi ⎠ ∂xi ∂yi 2 ⎟⎠ ⎝ ∂xi ⎝ ∂xi

equations of motion

⎛ ∂ 2v ⎛ ∂vy , i ∂vy , i ⎞ ∂ 2vy , i ⎞ ∂p ⎟ ⎟⎟ = − i + μ ⎜⎜ y , i + ρf, i⎜⎜vx , i + vy , i 2 i ∂yi ⎠ ∂yi ∂yi 2 ⎟⎠ ⎝ ∂xi ⎝ ∂xi equation of species continuity

vx ,1

∂cj ,1 ∂x1

+ vy ,1

⎛ ∂ 2c ∂ 2cj ,1 ⎞ j ,1 ⎟ = DAB,1⎜⎜ + 2 ∂y1 ∂y12 ⎟⎠ ⎝ ∂x1

∂cj ,1

⎛ ∂ 2T ⎛ ∂T ∂T ⎞ ∂ 2Ti ⎞ ⎟ ρf, iC Pf, i⎜⎜vx , i i + vy , i i ⎟⎟ = k f , i⎜⎜ 2i + ∂ x ∂ y ∂ x ∂yi 2 ⎟⎠ ⎝ i ⎝ i i ⎠

equation of energy

catalytic washcoat phase equation of continuity

∂vx ,1 ∂x1

equations of motion

+

∂vy ,1 ∂y1

=0

⎛ μ ⎞⎛ ∂ 2v ∂p ∂ 2vx ,1 ⎞ ⎛ μ1 ⎞ 1,eff ⎟ ⎟⎟⎜⎜ x ,1 + ⎜ ⎟v = − 1 + ⎜⎜ 2 ⎝ κ ⎠ x ,1 ∂x1 ⎝ εp ⎠⎝ ∂x1 ∂y12 ⎟⎠ ⎛ μ ⎞⎛ ∂ 2vy ,1 ∂ 2vy ,1 ⎞ ∂p ⎛ μ1 ⎞ 1,eff ⎟ ⎟⎟⎜⎜ ⎜ ⎟v = − 1 + ⎜⎜ + 2 ⎝ κ ⎠ y ,1 ∂y1 ∂y12 ⎟⎠ ⎝ εp ⎠⎝ ∂x1

equation of species continuity

vx ,1 equation of energy

∂cj ,1 ∂x1

+ vy ,1

⎛ ∂ 2c ∂ 2cj ,1 ⎞ j ,1 ⎟ + ρ R j ,1 = DAB,eff ⎜⎜ + 2 s ∂y1 ∂y12 ⎟⎠ ⎝ ∂x1

∂cj ,1

N ⎛ ∂ 2T ⎛ ∂T ∂T ⎞ ∂ 2T1 ⎞ ⎟ + ρ ∑ (−ΔHm)(rm) ρsC Ps⎜⎜vx ,1 1 + vy ,1 1 ⎟⎟ = keff ⎜⎜ 21 + 2⎟ s ∂y1 ⎠ ∂y1 ⎠ ⎝ ∂x1 ⎝ ∂x1 m=1

Solid Wall Phase

⎛ ∂ 2T ∂ 2Tw ⎞ ⎟=0 k w ⎜⎜ w2 + ∂yw 2 ⎟⎠ ⎝ ∂x w

conservation of energy

equations from the remaining equations. Binary diffusivity data in the form of diffusion of species of interest in air are adapted from the literature, which are then corrected for the pertinent temperature, pressure, and diffusion in hydrogen.28 Effective diffusivity appearing in the mass conservation equation for the porous medium is calculated from binary diffusivity data by taking tortuosity, taken as 2, and porosity of the catalyst into account.29,30 Effective thermal conductivity of the fluid in the washcoat is calculated by the following equation:26

equations including contributions of diffusive and convective mechanisms. The Darcy number for the porous washcoat (Da = κ/δs2) is equal to 4, which is large enough to make the flow non-Darcian. Therefore, flow in the washcoat is described using the Brinkman-extended Darcy model, which accounts for viscous and pressure effects and makes use of an effective viscosity. Studies in the literature report varying and sometimes inconsistent results as to the magnitude of the effective viscosity (μ1,eff < 1 and 1 < μ1,eff < 10).22,23 Moreover, its value strongly depends on the geometry of the medium.24 Due to the present work’s parametric nature, it is taken to be equal to the fluid viscosity in the bulk. Inertial losses during flow in the porous medium can be accounted for by the Brinkman−Forchheimer extension of the Darcy equation through the (dimensionless) term CF/Da1/2|v1|v1.25 However, due to the high value of the Darcy number, inertial effects, thus the Forchheimer term, can be neglected. Permeability and porosity of the washcoat domain are taken as 1 × 10−8 m2 and 0.5, respectively.26,27 All the fluid physical properties (e.g., density, viscosity, thermal conductivity) are assumed to be the same as those of hydrogen since mole fraction of hydrogen is reported to be ca. 94% in the stream.21 Fluid density and viscosity are assumed to be independent of temperature and composition and are calculated at the reactor inlet. This allows for decoupling of the momentum transport

⎛ k ⎞0.280 − 0.757logε− 0.057log(k w / k f,1) keff = ⎜⎜ w ⎟⎟ k f,1 ⎝ k f,1 ⎠

(1)

Density and thermal conductivity of the solids are assumed to be independent of temperature. Coolant is assumed to be pure steam, whose properties are taken directly from the material library of COMSOL Multiphysics. Heat capacity of the reactive stream is calculated according to the feed composition and is assumed to remain constant along the reaction channel. The kinetic model used to describe hydrocracking reactions over Pt supported on amorphous silica−alumina catalyst is adapted from the study of Pellegrini et al.31 The model, presented in Table 2, considers all of the linear hydrocarbons and lumps the isomers according to their carbon number. Activation energies and frequency factors for isomerization and cracking reactions 8915

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In this study the effects of operating conditions (feed temperatures of the reactant and coolant streams, weight hourly space velocity (WHSV) defined as the ratio of mass flow rate of wax to the mass of the catalytic washcoat layer (kgwax/kgcat·h)) and of microchannel reactor configuration (type of reactor material, thickness of the solid wall separating the channels) on the conversion of n-hydrocarbons with more than 22 carbon atoms (eq 1), mole fraction of diesel fuel cut (normal and isohydrocarbons in the C19−C22 range) (eq 2), and mole fraction of jet fuel cut (normal and iso-hydrocarbons in the C11−C18 range) (eq 3) are investigated:

Table 2. Kinetic Model Used in Simulations (Adapted from the work of Pellegrini et al.31) risomer‑C,i = (kisomer‑C,i(pn‑C,i − piso‑C,i/Keq,i))/(ADSORB) rcracking‑C,i = (kcracking‑C,ipiso‑C,i)/(ADSORB) 70 ADSORB = pH2[1 + ∑70 i=1 KLn‑C,ipn‑C,i + ∑i=1 KLiso‑C,ipiso‑C,i] kisomer‑C,i = Aiso,ie−Eiso,i/(RT); Aiso,i = 1.39 × 1016i7.9241; Eiso,i = 22253 ln(i) + 118350 kcracking‑C,i = Acr,ie−Ecr,i/(RT); Acr,i = 5.56 × 1017i5.7677; Ecr,i = 29806 ln(i) + 86006 KLn‑C,i = 0.1e0.39i; KLiso‑C,i = 0.1e0.38i; Keq,i = 1.1268i2 − 3.1914i + 3589.1

are defined separately. Equilibrium constant for each reversible isomerization reaction is also specified. Conservation equations given in Table 1 are solved according to the boundary conditions outlined in Table 3. No-slip

22+n‐hydrocarbon (n‐C22 +) conversion =

(Fn ‐ C22 +)in − (Fn ‐ C22+)out (Fn ‐ C22 +)in

Table 3. Boundary Conditions Related to the Equations in Table 1

(2)

mole fraction of diesel fuel cut =

Channel Inlets u1 = uin1 ; cj,1 = uinj,1; T1 = Tin1 u2 = uin2 ; T2 = Tin2

reaction channel, x = L cooling channel, x = 0 Channel Outlets

× 100 Fn ‐ C19 − C22 + Fiso ‐ C19 − C22 Ftotal (3)

mole fraction of jet fuel cut =

n·(−DAB∇cj,1) = 0; n·(−kf,1∇T1) = 0 reaction channel, x = 0 p1 = cooling channel, x = L p2 = n·(−kf,2∇T2) = 0 Line of Symmetry (Center Line of the Channels, Figure 1) pout 1 ; pout 2 ;

Fn ‐ C11− C18 + Fiso ‐ C11− C18 Ftotal

(4)

Default dimensions and properties of the unit cell (Figure 1) as well as inlet properties of the reaction mixture and cooling fluid are given in Table 4. The operating conditions and

reaction n·v1 = 0; n·(−DAB∇cj,1 + v1.cj,1) = 0; n·(−kf,1∇T1 + v1ρf,1CPf,1T1) = 0 channel cooling n·v2 = 0; n·(−kf,2∇T2 + v2ρf,2CPf,2T2) = 0 channel Fluid−Solid Wall Interface

Table 4. Default Values of Microchannel Reactor Configuration and the Feed Conditions

n·v1=0; n·(−DAB∇cj,1 + v1.cj,1) = 0 n·(−kf,1∇T1 + v1ρf,1CPf,1T1) = n·(−kw∇Tw) cooling channel n·v2 = 0; n·(−kf,2∇T2 + v2ρf,2CPf,2T2) = n·(−kw∇Tw) Solid Wall Boundaries at x = 0 and x = l reaction channel

microreactor configuration

n·(−kw∇Tw) = 0

condition is used on the channel walls. Concentrations of the species are specified at the inlet boundaries, and convective fluxes defined at the exit are taken as zero. Similarly, for the energy conservation equations, boundary conditions at the exit are assumed to be of convective type and set as zero. Heat transfer between the solid wall and the fluid flowing in porous washcoat is handled by means of heat flux continuity at the interface. At the free fluid−washcoat interface, continuity of velocity32 and heat and mass fluxes are imposed. The set of partial differential equations that represents simultaneous transport and reaction in the heat exchange integrated microchannel reactor is solved using the finite element method (FEM) under the COMSOL Multiphysics environment. The solutions are obtained by using an unstructured meshing with ca. 8300 triangular elements, which provide results similar to the ones involving higher number of grid points. A HP xw8600 workstation with 8 × 2.6 GHz processors and 16 GB of memory is used for the execution of the computer codes. It should be noted that the momentum transport equations in the porous washcoat lack inertial (convective) terms while the heat and mass transport equations do. Solution to this system of equations that contain pseudoconvective heat and mass transport terms in the washcoat is obtained by first solving for the momentum equations that are already decoupled from the rest, and then by substituting the velocity components into the heat and mass transport equations.

−4

side length of reaction channel thickness of the catalyst layer thickness of the separating wall side length of cooling channel channel length

5 × 10

wall material

AISI steel

cooling fluid

steam

feed conditions m

5 × 10−5 m 3 × 10−4 m 6 × 10−4 m 1 × 10−1 m

temperature/reaction channel pressure/reaction channel WHSV/reaction channel temperature/cooling channel pressure/cooling channel superficial velocity/ cooling channel H2/wax ratio (kgH2/ kgwax)

460 °C 53.75 bar 160 kgwax/ kgcat·h 390 °C 30 bar 0.05 m/s 0.1275

microchannel reactor configuration are changed on a systematic basis. The parameter under investigation is changed according to the values presented in Table 5, whereas remaining parameters are kept at their default values given in Table 4. Table 5. Values Investigated in the Parametric Analysis reactant feed temperature (°C) coolant feed temperature (°C) WHSV (kgwax/kgcat·h) type of wall materiala wall thickness (× 106 m)

440, 450, 460 300, 330, 360, 390 100, 120, 140, 160, 180, 200, 220 alumina (27), AISI steel (44.5), silicon carbide (87) 200, 300, 400, 500, 600

a

Thermal conductivity of each material in units of watts per meter kelvin is given in parenthesis.

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Figure 2. Effects of WHSV (kgwax/kgcat·h) and coolant feed temperatures (C) on 22+ n-hydrocarbon conversion at reactant feed temperatures (R) of (a) 440, (b) 450, and (c) 460 °C. Refer to Table 4 for the values of other geometric and operational parameters.

Figure 3. Effects of WHSV (kgwax/kgcat·h) and coolant feed temperatures (C) on mole fraction of diesel cut at reactant feed temperatures (R) of (a) 440, (b) 450, and (c) 460 °C. Refer to Table 4 for the values of other geometric and operational parameters.

Figure 4. Effects of WHSV (kgwax/kgcat·h) and coolant feed temperatures (C) on mole fraction of jet cut at reactant feed temperatures (R) of (a) 440, (b) 450, and (c) 460 °C. Refer to Table 4 for the values of other geometric and operational parameters.

3. RESULTS AND DISCUSSION Effects of reactant and coolant feed temperatures and space velocity of the reactant stream on the conversion of 22+ nhydrocarbons are given in Figure 2. It can be observed that reactant and coolant feed temperatures positively affect conversion, which is found to vary in the ranges of 64−83%, 77−93%, and 90−99% for the reactant feed temperatures of 440 °C (Figure 2a), 450 °C (Figure 2b), and 460 °C (Figure 2c), respectively, due to positive dependence of reaction kinetics on temperature. The impact of space velocity on conversion, however, is found to be slightly different. For all combinations of feed temperatures, conversion showed a decreasing trend with WHSV values in the 140−220 kgwax/ kgcat·h range, in which more hydrocarbons remained unconverted as a result of insufficient contact with the coated catalyst layer. However, this phenomenon is not clear in the 100−140 kgwax/kgcat·h WHSV range, in which, as the space velocity is lowered, conversion values either increase at coolant feed temperatures of 360 and 390 °C, or decrease at coolant feed temperatures of 300 and 330 °C. This behavior can be

explained by the degree of cooling of the reaction channel. At lower WHSV values, the amount of reactant entering to and the exothermal energy released within the reaction channel decrease, and the impact of the cooling channel becomes more significant through effective removal of the generated heat, ending up with temperature drops as observed for coolant feed temperatures of 300 and 330 °C (Figures 2a−c). For coolant feed temperatures of 360 and 390 °C, however, temperature in the reaction channel stays high enough to sustain the reactions and, as a result, conversions keep on increasing at lower WHSV values. Nevertheless slight decreases in the slopes of the conversion profiles are observed (e.g., at 360 °C), indicating that effective cooling has still impacts on conversion (Figures 2a−c). Responses of mole fractions of diesel fuel cut (C19−C22) at the microchannel reactor outlet against reactant and coolant feed temperatures, and space velocity of the reactant stream are shown in Figure 3. The results show that mole fractions of diesel cut are sensitive to and decrease with reactant feed temperature; at 440 °C, the values are clustered around 0.138, 8917

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regardless of the WHSV and coolant feed temperatures, whereas they become less than 0.135, inlet mole fraction of diesel cut, when the reactant feed temperature is increased to 460 °C. At 450 °C, the net increase in diesel cut fraction is limited to certain combinations of coolant feed temperatures and WHSVs (all WHSV, between 300−360 °C; WHSV > 140 kgwax/kgcat·h, 390 °C) (Figure 3b). These trends can be explained by the changes in conversion of 22+ n-hydrocarbons explained above. It can be observed that the highest diesel cut mole fraction (∼0.1385) is obtained at 440 °C reactant and 390 °C coolant feed temperatures and at WHSV of 180 kgwax/ kgcat·h that give ca. 73% conversion (Figures 2a and 3a). Since reactant temperatures of 450 and 460 °C favor cracking to lower hydrocarbons and lead to conversions higher than 73% and (Figure 2b and c), the respective diesel cut mole fractions become lower than the maximum value and further decrease at higher coolant feed temperatures (Figure 3b and c). The results also show that 460 °C is not suitable for enhancing diesel yield since the mole fraction of the diesel cut is lower than the inlet value of 0.135. Changes in the mole fraction of jet fuel cut (C11−C18) at the microchannel reactor outlet against reactant and coolant feed temperatures and space velocity of the reactant stream are shown in Figure 4. Similar to diesel cut, the jet cut is also an intermediate product, and its responses against process parameters can be explained by the changes in conversion. For all coolant feed temperatures, the mole fraction of jet fuel cut is found to increase with reactant feed temperatures of 440 and 450 °C (Figure 4a and b), at which conversions are lower than 97%, the value corresponding to maximum mole fraction of the jet fuel cut, 0.485 (Figures 2c and 4c). In other words, the conditions yielding 22+ n-hydrocarbon conversions up to 97% also favor the production of C11−C18 hydrocarbons. When reactant stream is fed at 460 °C, a maximum value of jet fuel cut is observed for every coolant feed temperature, whose locus depends on the WHSV employed (Figure 4c). At coolant feed temperatures of 300, 330, and 360 °C, left-hand sides of the maximum points show decrease in mole fraction of jet cut with WHSV. This can be explained by the cooling of the reaction channel, which becomes easier due to the low quantities of reactive flow, followed by subsequent reduction in reaction rates, conversions, and jet cut mole fraction (Figures 2c and 4c). It is worth noting that this effect becomes less pronounced when a warmer coolant is used in the process. At 390 °C, however, the sharp drop in jet cut fraction at WHSVs less than ∼170 kgwax/kgcat·h is due to the insufficient heat removal from the reaction channel involving temperatures high enough to get near complete hydrocracking to