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Cover
Title Page
List of contributors
Preface
PART 1: Principles of catalytic reaction engineering
1. CHAPTER 1: Catalytic reactor types and their industrial significance
  1. 1.1 Introduction
  2. 1.2 Reactors with fixed bed of catalysts
  3. 1.3 Reactors with moving bed of catalysts
  4. 1.4 Reactors without a catalyst bed
  5. 1.5 Summary
  6. References
2. CHAPTER 2: Microkinetic analysis of heterogeneous catalytic systems
  1. 2.1 Heterogeneous catalytic systems
  2. 2.2 Intrinsic kinetics of heterogeneous reactions
  3. 2.3 External (interphase) transport processes
  4. 2.4 Internal (intraparticle) transport processes
  5. 2.5 Combination of external and internal transport effects
  6. 2.6 Summary
  7. Nomenclature
  8. Greek letters
  9. References
PART 2: Two‐phase catalytic reactors
1. CHAPTER 3: Fixed‐bed gas–solid catalytic reactors
  1. 3.1 Introduction and outline
  2. 3.2 Modeling of fixed‐bed reactors
  3. 3.3 Averaging over the catalyst particle
  4. 3.4 Dominant fluid–solid mass transfer
  5. 3.5 Dominant fluid–solid mass and heat transfer
  6. 3.6 Negligible mass and thermal dispersion
  7. 3.7 Conclusions
  8. Nomenclature
  9. Greek letters
  10. References
2. CHAPTER 4: Fluidized‐bed catalytic reactors
  1. 4.1 Introduction
  2. 4.2 Key hydrodynamic features of gas‐fluidized beds
  3. 4.3 Key properties affecting reactor performance
  4. 4.4 Reactor modeling
  5. 4.5 Scale‐up, pilot testing, and practical issues
  6. 4.6 Concluding remarks
  7. Nomenclature
  8. Greek letters
  9. References
PART 3: Three‐phase catalytic reactors
1. CHAPTER 5: Three‐phase fixed‐bed reactors
  1. 5.1 Introduction
  2. 5.2 Hydrodynamic aspects of three‐phase fixed‐bed reactors
  3. 5.3 Mass and heat transfer in three‐phase fixed‐bed reactors
  4. 5.4 Scale‐up and scale‐down of trickle‐bed reactors
  5. 5.5 Trickle‐bed reactor/bioreactor modeling
  6. Nomenclature
  7. Greek letters
  8. Subscripts
  9. Superscripts
  10. Abbreviations
  11. References
2. CHAPTER 6: Three‐phase slurry reactors
  1. 6.1 Introduction
  2. 6.2 Reactor design, scale‐up methodology, and reactor selection
  3. 6.3 Reactor models for design and scale‐up
  4. 6.4 Estimation of transport and hydrodynamic parameters
  5. 6.5 Advanced computational fluid dynamics (CFD)‐based models
  6. 6.6 Summary and closing remarks
  7. Acknowledgments
  8. Nomenclature
  9. Greek letters
  10. Subscripts
  11. References
3. CHAPTER 7: Bioreactors
  1. 7.1 Introduction
  2. 7.2 Basic concepts, configurations, and modes of operation
  3. 7.3 Mass balances and reactor equations
  4. 7.4 Immobilized enzymes and cells
  5. 7.5 Aeration
  6. 7.6 Mixing
  7. 7.7 Heat transfer
  8. 7.8 Scale‐up
  9. 7.9 Bioreactors for animal cell cultures
  10. 7.10 Monitoring and control of bioreactors
  11. Nomenclature
  12. Greek letters
  13. Subscripts
  14. References
PART 4: Structured reactors
1. CHAPTER 8: Monolith reactors
  1. 8.1 Introduction
  2. 8.2 Design of wall‐coated monolith channels
  3. 8.3 Mapping and evaluation of operating regimes
  4. 8.4 Three‐phase processes
  5. 8.5 Conclusions
  6. Nomenclature
  7. Greek letters
  8. Superscripts
  9. Subscripts
  10. References
2. CHAPTER 9: Microreactors for catalytic reactions
  1. 9.1 Introduction
  2. 9.2 Single‐phase catalytic microreactors
  3. 9.3 Multiphase microreactors
  4. 9.4 Conclusions and outlook
  5. Nomenclature
  6. Greek letters
  7. Subscripts
  8. References
PART 5: Essential tools of reactor modeling and design
1. CHAPTER 10: Experimental methods for the determination of parameters
  1. 10.1 Introduction
  2. 10.2 Consideration of kinetic objectives
  3. 10.3 Criteria for collecting kinetic data
  4. 10.4 Experimental methods
  5. 10.5 Microkinetic approach to kinetic analysis
  6. 10.6 TAP approach to kinetic analysis
  7. 10.7 Conclusions
  8. References
2. CHAPTER 11: Numerical solution techniques
  1. 11.1 Techniques for the numerical solution of ordinary differential equations
  2. 11.2 Techniques for the numerical solution of partial differential equations
  3. 11.3 Computational fluid dynamics techniques
  4. 11.4 Case studies
  5. 11.5 Summary
  6. Nomenclature
  7. Greek letters
  8. Subscripts/superscripts
  9. References
PART 6: Industrial applications of multiphase reactors
1. CHAPTER 12: Reactor approaches for Fischer–Tropsch synthesis
  1. 12.1 Introduction
  2. 12.2 Reactors to 1950
  3. 12.3 1950–1985 period
  4. 12.4 1985 to present
  5. 12.5 The future?
  6. References
2. CHAPTER 13: Hydrotreating of oil fractions
  1. 13.1 Introduction
  2. 13.2 The HDT process
  3. 13.3 Fundamentals of HDT
  4. 13.4 Process aspects of HDT
  5. 13.5 Reactor modeling and simulation
  6. Nomenclature
  7. Greek letters
  8. Subscripts
  9. Non‐SI units
  10. References
3. CHAPTER 14: Catalytic reactors for fuel processing
  1. 14.1 Introduction—The basic reactions of fuel processing
  2. 14.2 Theoretical aspects, advantages, and drawbacks of fixed beds versus monoliths, microreactors, and membrane reactors
  3. 14.3 Reactor design and fabrication
  4. 14.4 Reformers
  5. 14.5 Water-gas shift reactors
  6. 14.6 Carbon monoxide fine cleanup: Preferential oxidation and selective methanation
  7. 14.7 Examples of complete fuel processors
  8. Nomenclature
  9. References
4. CHAPTER 15: Modeling of the catalytic deoxygenation of fatty acids in a packed bed reactor
  1. 15.1 Introduction
  2. 15.2 Experimental data for stearic acid deoxygenation
  3. 15.3 Assumptions
  4. 15.4 Model equations
  5. 15.5 Evaluation of the adsorption parameters
  6. 15.6 Particle diffusion study
  7. 15.7 Parameter sensitivity studies
  8. 15.8 Parameter identification studies
  9. 15.9 Studies concerning the deviation from ideal plug flow conditions
  10. 15.10 Parameter estimation results
  11. 15.11 Scale‐up considerations
  12. 15.12 Conclusions
  13. Acknowledgments
  14. Nomenclature
  15. Greek letters
  16. References
Index
End User License Agreement

List of Tables

Chapter 02
1. Table 2.1 Individual terms of LHHW rate equations for the surface reaction‐controlling cases of various catalytic reactions.
Chapter 03
1. Table 3.1 Fixed‐bed reactor mathematical model (definition of timescales τ in Table 3.2).
2. Table 3.2 Timescales of the different processes in the model for a fixed bed.
Chapter 04
1. Table 4.1 Some typical key characteristics of catalytic fluidized‐bed reactors compared with those of alternative types of reactor.
2. Table 4.2 Solid catalyzed gas‐phase reactions which have been carried out in commercial fluidized‐bed reactors.
3. Table 4.3 Typical operating ranges and features of catalytic fluidized‐bed reactors.
4. Table 4.4 Comparison of typical properties of catalytic and gas–solid fluidized‐bed reactors.
5. Table 4.5 Common instrumentation required in fluidized‐bed reactors.
6. Table 4.6 Complementary components of fluidized‐bed reactor system.
Chapter 05
1. Table 5.1 Model parameters.
2. Table 5.2 Values of kinetic parameters.
3. Table 5.3 Parameters used in simulations.
4. Table 5.4 Two‐bed reactor operating conditions.
Chapter 06
1. Table 6.1 Three‐phase catalytic reactors.
2. Table 6.2 Comparison of multiphase reactors (qualitative rating: more stars mean superior performance on the pertinent metric).
3. Table 6.3 Some illustrative applications of three‐phase slurry reactors.
4. Table 6.4 Overview of vessel designs and performance attributes of three‐phase slurry reactors.
5. Table 6.5 Idealized flow and axial dispersion models (steady state).
6. Table 6.6 Mixing cell model.
7. Table 6.7 Correlations for estimation of k_la in three‐phase slurry and fluidized beds.
8. Table 6.8 Liquid holdup, mass transfer coefficients, and effective interfacial area in gas–liquid reactors.
9. Table 6.9 Correlations for liquid dispersion coefficients in three‐phase slurry and three‐phase fluidized beds.
10. Table 6.10 Governing equations for Euler–Euler formulation.
11. Table 6.11 Turbulence closures.
12. Table 6.12 Closures for interphase momentum exchange.
13. Table 6.13 Closures for solid phase.
Chapter 08
1. Table 8.1 Monolith reactors classified according to flow, materials, and operation features.
2. Table 8.2 Typical dimensions and pressure drop in some monolith reactor processes [46].
3. Table 8.3 Friction factors and transfer coefficients for some common cross‐sectional shapes in monolith channels [39].
4. Table 8.4 Solution strategies for the Graetz–Lévêque problem with wall reaction.
5. Table 8.5 Bridging the gap between convection and diffusion regimes.
6. Table 8.6 Calculation of the effectiveness factor in uniform and nonuniform washcoats.
7. Table 8.7 Accuracy of effectiveness factor calculation methods for nonuniform geometry (circle‐in‐square shape with a first‐order reaction).
8. Table 8.8 Experimental ranges used in the development of empirical mass transfer correlations.
9. Table 8.9 Comparison between models for catalytic combustion in a monolith.
10. Table 8.10 Vertices in the Damköhler–Graetz plot with reaction–transport and profile development regimes.
Chapter 09
1. Table 9.1 Asymptotic Sherwood number for constant reactant concentration.
2. Table 9.2 Two‐phase laminar–laminar frictional pressure drop correlations.
3. Table 9.3 Comparison of mass transfer parameters in different gas–liquid contactors [81].
Chapter 11
1. Table 11.1 Linear system solvers available in COMSOL Multiphysics.
2. Table 11.2 Model equations used to simulate heat‐exchange integrated microchannel reactor operation shown in Figure 11.8.
3. Table 11.3 Boundary conditions related to the mathematical model given in Table 11.2.
Chapter 12
1. Table 12.1 Comparison of reaction engineering models for the Fischer–Tropsch synthesis in slurry bubble column reactor.
2. Table 12.2 Kinetic models for Table 12.1.
3. Table 12.3 Existing operating Fischer–Tropsch plants.
Chapter 13
1. Table 13.1 Properties of various Mexican crude oils (obtained at IMP analytical lab).
2. Table 13.2 Standard enthalpies and equilibrium constants of representative HDT reactions.
3. Table 13.3 Typical process conditions of various hydrotreating and hydrocracking processes.
4. Table 13.4 Properties of the feedstocks.
Chapter 15
1. Table 15.1 The values of input variables in laboratory‐scale unit [15, 20].
2. Table 15.2 Model equations.
3. Table 15.3 Physical properties and variables of the model system.
4. Table 15.4 Parameter values of the sensitivity study.

List of Illustrations

Chapter 01
1. Figure 1.1 Outline of a chemical process.
2. Figure 1.2 Schematic presentation of a packed‐bed reactor.
3. Figure 1.3 Furnace configurations for multitubular packed‐bed reformers.
4. Figure 1.4 Side‐fired tubular reformer design by Haldor‐Topsøe.
5. Figure 1.5 Heat transfer strategies in multitubular packed‐bed reactors. (a) Cross‐flow, (b) parallel flow, and (c) boiling‐water cooling.
6. Figure 1.6 Various configurations of vessel‐type packed‐bed reactors. (a) Single‐bed adiabatic packed‐bed reactor, (b) adiabatic reactor with interstage gas injection, and (c) multiple adiabatic beds with interstage heat exchange.
7. Figure 1.7 Packed‐bed reactor with multiple adiabatic beds for ammonia synthesis.
8. Figure 1.8 Packed‐bed reactor configuration for autothermal reforming of methane to synthesis gas.
9. Figure 1.9 Schematic presentation of a monolith reactor.
10. Figure 1.10 Comparison of pressure drop in various configurations of monoliths and packing structures.
11. Figure 1.11 Radial flow reactor concept.
12. Figure 1.12 Radial flow ammonia synthesis converter by Haldor‐Topsøe.
13. Figure 1.13 Disk reactor concept.
14. Figure 1.14 Schematic presentation of a microchannel reactor. (a) Machined plates with microchannels, (b) microchannel reactor block obtained after bonding the plates, and (c) characteristic section of the multichannel reactor.
15. Figure 1.15 Riser cracking process by UOP. (a) Reactor, (b) stripper, (c) riser, (d) slide valve, (e) air grid, and (f) regenerator.
16. Figure 1.16 High‐temperature Fischer–Tropsch synthesis reactors. (a) Sasol Synthol circulating fluidized‐bed reactor.(b) Sasol Advanced Synthol turbulent fluidized‐bed reactor.
17. Figure 1.17 Slurry bubble column reactor for low‐temperature Fischer–Tropsch synthesis.
18. Figure 1.18 Stirred‐tank reactor with typical dimensions.
19. Figure 1.19 Heat transfer strategies in stirred‐tank reactors. (a) Jacket, (b) internal coils, (c) internal tubes, (d) external heat exchanger, (e) external reflux condenser, and (f) fired heater.
Chapter 02
1. Figure 2.1 Potential energy curves representing the action of a solid catalyst.
2. Figure 2.2 Effect of temperature on amount of gas adsorbed for simultaneous physical adsorption and activated chemisorption.
3. Figure 2.3 Reactant concentration profiles in different global rate regimes: I, external mass transfer limitation; II, pore diffusion limitation; III, both external and internal mass transfer limitations; IV, no mass transfer limitations on the intrinsic rate.
4. Figure 2.4 Total pressure dependence of initial rates for the reaction A → B + C.
5. Figure 2.5 Types of laboratory catalytic reactors mostly used: (a) plug‐flow packed‐bed reactor and (b) spinning basket reactor mimicking CSTR performance.
6. Figure 2.6 The effect of external mass transport limitations on conversion rates.
7. Figure 2.7 Correlation of experimental data on external mass transport between bulk fluid and particle surface in packed beds by various researchers.
8. Figure 2.8 Steady‐state multiplicity caused by external temperature gradients in simple nonporous surface‐catalyzed exothermic reactions when external mass transfer resistance is either considerable or negligible.
9. Figure 2.9 Representation of a catalyst particle having micro‐ and macropore structure.
10. Figure 2.10 Internal effectiveness factor as a function of Thiele modulus for porous particles of various shapes.
11. Figure 2.11 Reactant concentration and temperature profiles inside a porous spherical catalyst particle with exothermic reaction occurring at steady state.
12. Figure 2.12 Nonisothermal internal effectiveness factors for first‐order reactions in porous spherical particles.
Chapter 03
1. Figure 3.1 Fixed‐bed reactors. Schematic representation of several operating configurations and arrangements: (a) Cylindrical with axial flow (a.1‐single, a.2‐multistage, a.3‐multitubular); (b) cylindrical with radial flow; (c) spherical reactor (c.1‐axial and c.2‐radial flow).
2. Figure 3.2 Transport–reaction mechanisms in a fixed‐bed reactor at different scales.
3. Figure 3.3 Effectiveness factor curve (η vs. φ) for a spherical catalyst particle with Dirichlet boundary condition and first‐order exothermic reaction. Numerical results from Ref. [99]. A curve for is depicted (with ), while the others were calculated for and several values of β.
4. Figure 3.4 Asymptotic forms of the effectiveness factor () as a function of the Thiele modulus (φ) for a slab catalyst with a first‐order exothermic reaction ( and ) for . Numerical values in the intermediate region are given in Ref. [80].
5. Figure 3.5 Relative error (δ %) associated with the calculation of conversion from a pseudo‐homogeneous model for a first‐order isothermal reaction as a function of , with or 0.5.
6. Figure 3.6 Relative error (δ %) associated with the calculation of conversion from a pseudo‐homogeneous model for an mth‐order isothermal reaction as a function of , with or 0.5.
7. Figure 3.7 regime diagram for a second‐order reaction occurring in an isothermal fixed bed. Iso‐δ lines given by Equation 3.85.
8. Figure 3.8 Cross‐sectional average concentration (a) and temperature (b) profiles according to a pseudohomogeneous model (full lines: two‐dimensional; dashed lines: one‐dimensional). After Ref. [133] with °C and °C.
9. Figure 3.9 Deviation in the reaction rate between pseudo‐homogeneous and heterogeneous models in the partial oxidation of methanol to formaldehyde, as calculated in Ref. [135].
10. Figure 3.10 Comparison of the axial temperature (a) and methanol concentration (b) according to several models (cross‐sectional averages for 2D models) for a fixed‐bed reactor to produce formaldehyde from methanol [135].
Chapter 04
1. Figure 4.1 Schematic of principal features of a modern fluid catalytic cracking (FCC) reactor system.
2. Figure 4.2 Schematic of main features of fluidized‐bed reactor for production of acrylonitrile showing immersed serpentine heat transfer tubing, separate distributors for air and ammonia + propane or propene, and one representative cluster of internal cyclones in series.
3. Figure 4.3 Diagrammatic representation of flow regimes of gas fluidization with increasing superficial gas velocity, with transition velocities indicated on the arrows between adjacent flow regimes. Dashed rectangular box encloses the five flow regimes of primary interest with respect to catalytic fluidized‐bed reactors.
4. Figure 4.4 Pictorial representation of particles and void regions in flow regimes of principal interest for catalytic fluidized‐bed reactors.
5. Figure 4.5 Simple single‐phase model predictions for first‐order irreversible catalytic ozone decomposition reaction in comparison with experimental fluidized‐bed reactor data of Sun and Grace [44].
6. Figure 4.6 Schematic of two‐phase model representation of bubbling or slugging fluidized‐bed reactor.
Chapter 05
1. Figure 5.1 Fixed‐bed reactors for gas–liquid–solid catalyzed systems. (a) Two‐phase downflow fixed‐bed reactor (trickle‐bed reactor). (b) Fixed‐bed reactor with countercurrent flow. (c) Two‐phase upflow fixed‐bed reactor (packed‐bed bubble flow reactor).
2. Figure 5.2 Hydrodynamic regimes for two‐phase downflow in fixed beds—nonfoaming (a) and foaming (b) liquids.)
3. Figure 5.3 Hydrodynamic regimes for two‐phase upflow in fixed beds: transition zones (/////) from Ref. [30] and transition lines from Ref. [28].)
4. Figure 5.4 Representation of connectivity between subregions in the representative elementary volume: V_g, V_ℓ, and V_s for the bulk phases and A_gℓ and A_ℓs for the interfaces (a), interaction of interface A_gℓ with V_g and V_ℓ, bulk phases in creating and destroying gas–liquid interfacial area (b).)
5. Figure 5.5 Mass transfer resistances in three‐phase fixed‐bed reactors.
6. Figure 5.6 Liquid and gas loads in typical trickle‐bed reactors for gas oil hydroprocessing.)
7. Figure 5.7 Sketch of trickle‐bed reactor (a) and local‐scale representation of the two‐phase flow (b). Sketch showing the resistances involved in the G/L reaction on the catalyst particle (c, d).)
8. Figure 5.8 Two‐phase pressure drop ratio (a) and bed volume average specific deposit (b) versus time at different values of fine particle concentration (d_f = 5 µm, Tⁱⁿ = 598 K, Pⁱⁿ = 10 MPa).
9. Figure 5.9 Effect of fine particle deposition process on dibenzothiophene conversion at t = 485 min (Tⁱⁿ = 598 K, Pⁱⁿ = 10 MPa): (a) d_f = 5 µm, (b) d_f = 1 µm.
10. Figure 5.10 Biofilm thickness versus time at different values of inlet cell concentration (only cell detachment was considered): (a) ; (b) (, , , , α = 0.1, D_f = 2.0).
11. Figure 5.11 Number of aggregates versus axial distance ( mg/l, mg/l, mg/l, g/l, α = 0.1, D_f = 2.0, t = 2000 min, surface 1—, surface 2—).)
12. Figure 5.12 Two‐phase pressure drop ratio. (a) Cell + aggregate volume average specific deposit and biofilm thickness at z/H = 0.11 (b) versus time in the presence or absence of biomass detachment (, , , , , α = 0.5, D_f = 1.5).
13. Figure 5.13 Schematic diagram of two‐bed reactor system for DME production from glycerol via an integrated process involving aqueous‐phase glycerol reforming coupled with DME synthesis process.
14. Figure 5.14 Steady‐state axial profiles of the volumetric gas flow rate (a) and temperature (b) for aqueous‐phase glycerol reforming and DME synthesis without in situ H₂O removal.
15. Figure 5.15 Steady‐state axial profiles of the volumetric gas flow rate for aqueous‐phase glycerol reforming and DME synthesis with in situ H₂O removal. (a) . (b) .
16. Figure 5.16 Steady‐state axial profiles of the temperature for aqueous‐phase glycerol reforming and DME synthesis with/without in situ H₂O removal. (a) . (b) .
Chapter 06
1. Figure 6.1 Schematic diagrams of industrial three‐phase slurry reactors. (a) Slurry bubble column. (b) Three‐phase fluidized bed. (c) Three‐phase agitated vessel.
2. Figure 6.2 Schematic of the overall methodology for relating performance of the reactor with the hardware and the operating protocol.
3. Figure 6.3 Possible ideal contacting patterns in three‐phase slurry reactors. (a) Countercurrent (gas and liquid in plug flow). (b) Co‐current (gas and liquid in plug flow). (c) Mixed (gas and liquid in mixed flow).
4. Figure 6.4 Reactant profiles in neighborhood of gas (bubble) and catalyst particle.
5. Figure 6.5 Hydrodynamic simulation of methanol synthesis in gas–liquid slurry bubble column reactors. (a) Reactor operating conditions and simulation grid. (b) Simulated instantaneous distribution of gas and solid holdup. (c) Simulated instantaneous distribution of temperature.
6. Figure 6.6 Predicted instantaneous distribution of various species in (a) gas and (b) slurry phases at catalyst concentration of 20%.
Chapter 07
1. Figure 7.1 An estimate on the current use in industry of bioreactors with different basic characteristics regarding aeration and stirring.
2. Figure 7.2 Typical configuration of commonly used bioreactors.
3. Figure 7.3 Modes of bioreactor operation: (a) batch, (b) continuous, and (c) fed batch.
4. Figure 7.4 Flow through a tubular reactor of total length L, where a differential element of volume, dV = Adz, is highlighted.
5. Figure 7.5 Different stages of cell growth in batch processes: (1) lag phase, (2) exponential growth phase, (3) stationary phase, and (4) death phase.
6. Figure 7.6 CSTR with recycle.
7. Figure 7.7 Variation of substrate concentration, C_A, from bulk liquid medium to center of support particle.
Chapter 08
1. Figure 8.1 Representation of the several scales in a catalytic monolithic reactor. (a) Monolith honeycomb [1] . (b) Single washcoated monolith channel. (c) Catalytic washcoat.
2. Figure 8.2 Conversion calculated according to Graetz series with different number of terms. Equation 8.26 is plotted with 1, 5, and 12 terms as X_Graetz for intermediate mass transfer control (). The solution with only one term corresponds to the fully developed concentration profile. A uniformly valid solution for conversion X (plotted as a solid line) is given by Equation 8.37 and is in good agreement with the numerical solution.
3. Figure 8.3 Normalized dependence of the first eigenvalue on the Damköhler number: versus curve. The analytical result given by Equation 8.34b was derived in Lopes et al. [40]. Numerical results for laminar flows are also shown (dashed lines).
4. Figure 8.4 Fully developed concentration profile given by Equation 8.34. Numerical (full lines) and analytical (dashed lines) refer to laminar flow in a circular channel. The approximate solution uses eigenvalues and coefficients calculated from Equations 8.34b and 8.34c for several values of Damköhler number.
5. Figure 8.5 Reactant concentration (mixing cup) as a function of the rescaled Damköhler number (Da^*) given in Equation 8.36b for laminar flow between parallel plates. The dashed lines are plotted according to Equation 8.35b for several values of the Graetz parameter (αPe_m/z) in the ranges of Da^* leading to less than 5% relative error.
6. Figure 8.6 Conversion as a function of the dimensionless axial distance for laminar flow in a circular channel with instantaneous wall reaction. Three regions can be identified in different ranges of the Graetz parameter: Graetz and Lévêque descriptions, separated by a transition regime.
7. Figure 8.7 Convection–diffusion limiting cases for channel mass transfer with wall reaction in a Damköhler–Graetz parametric map (). Four regions can be depicted: fully developed profile (for αPe_m/z below the prediction shown by the full lines for given values of ε_fd), developing profile (for Graetz numbers above the ones estimated by dashed lines for given e_dev), transition region (separating the two previous limits at high Da), and overlapping region (where similar predictions from both models are obtained at low Da). All boundaries were obtained analytically and the respective expressions can be found in Ref. [43].
8. Figure 8.8 Calculation approaches for the effectiveness factor in thin catalytic coatings. The distinction between the conventional method and the one proposed in Lopes et al. [94] is illustrated. The Thiele modulus for a first‐order reaction is given by . The ratio between the catalyst and channel transverse length scales is given by and σ is the catalyst shape factor.
9. Figure 8.9 Damköhler–Graetz plot for linear kinetics with mapping of reaction–transport and convective–diffusive regimes in a washcoated monolith. Two values of the diffusion ratio are examined, and laminar flow inside a circular channel with is considered. Regime boundaries are plotted for given values of the criteria (internal control, ; external control, ; no internal limitation, ; no external limitation, ; developed profile, ; developing profile, ). Vertices (V) delimiting the intermediate regime are also shown.
10. Figure 8.10 operating map for a circular channel with laminar flow coated with an annular catalyst layer () and a first‐order reaction. External and internal mass transfer control observed for and , respectively. Negligible resistances for and . Full lines associated with conditions in channel are plotted for values of the Graetz parameter. The area for overall kinetic control (Ov. KC) is depicted as well as the ones for mass transfer control: overall (Ov. MTC) and internal (Intraph. MTC).
11. Figure 8.11 Damköhler–Graetz diagram for nonlinear kinetics. A second‐order reaction occurs in the washcoat, but the remaining conditions are identical to previous representations.
12. Figure 8.12 Damköhler–diffusion ratio diagram for nonlinear kinetics (second‐order reaction). Same conditions from previous representations apply.
13. Figure 8.13 Design in the presence of a constraint on conversion. Two distinct problems are depicted: (i) increment in pressure drop due to the existence of a maximum reaction rate above which operation is not feasible and (ii) reduction in energetic requirements by increasing the allowed pressure drop.
14. Figure 8.14 Taylor flow in washcoated monoliths. Gas–liquid–solid mass transfer routes are represented schematically. The relevant geometrical dimensions can be also depicted.
Chapter 09
1. Figure 9.1 (a) Influence of channel diameter variations on the reactor conversion; ideal reactor (thick line), reactor with = 0.1 (thin line). (b) Influence of the parameter γ on the conversion in microreactor with = 0.01; γ = 15, 30, and 60.
2. Figure 9.2 Typical ranges of force ratios in multiphase flow for packed beds, micropacked beds, and fine porous media. is the gravity/surface tension ratio and . The term u is the superficial velocity, ρ is the density, l is the characteristic hydrodynamic length, that is, wetted area over wetted perimeter, g is gravity, μ is the viscosity, and σ is the surface tension. The subscripts L and G refer to liquid and gas, respectively.
3. Figure 9.3 Flow transition for cyclohexene/H₂ cocurrent flow in a micro trickle bed reactor (■) versus superficial mass velocity [23]. Charpentier flow map (♦) is presented for comparison [24]. G is the superficial mass velocity.
4. ²²²
5. Figure 9.8 Mechanisms of the formation of gas bubbles in the microchannels: (a, b) squeezing and (c) dripping.
6. Figure 9.9 Schematic of Taylor flow showing the definitions of the unit cell, gas bubble length L_b and the liquid slug length L_s. The lengths of the nose L_nose and tail L_tail sections of the gas bubble are also indicated. Mass transfer contributions in the Taylor flow are indicated with the numbers: (1) bubble to wall through film, (2) bubble to slug, and (3) slug to wall.
Chapter 10
1. Figure 10.1 Conceptual diagram of a molecular beam scattering (MBS) experiment comprised of pulsed beam source, sample target, mass spectrometer detector, and a differentially pumped ultrahigh‐vacuum system.
2. Figure 10.2 Plot showing the scattered CO intensity as a function of time from an initially clean Pd(111) surface. The initial sticking coefficient can be determined from the mass spectrometer signal I(t) using the relationship S(t) = 1 − I(t)/I(∞).
3. Figure 10.3 S₀ for O₂ on Pd (111) as a function of substrate temperature.
4. Figure 10.4 Adaptation of the “pressure/materials” correlation proposed by Bonzel [37] indicating pressure regimes and types of materials found in various types of catalyst characterization experiments.
5. Figure 10.5 Key components of a TAP experiment.
6. Figure 10.6 (a) Schematic of TAP single particle microreactor configuration. The 400 µm diameter Pt particle is packed within a bed of inert quartz particles with diameters between 210 and 250 µm. (b) Image comparing a 400 µm Pt particle to a pencil point. (c) SEM image showing the complex surface structure of a polycrystalline Pt particle. (d) Higher magnification (15 000×) of the particle shown in (c), which shows the surface is nonporous [86].
7. Figure 10.7 Illustration of a TAP pump‐probe experiment in which O₂/Ar and CO₂/Ar are pulsed in an alternating sequence and the CO₂ transient response is measured during each pulse.
8. Figure 10.8 Comparison of CO₂ produced during TAP vacuum pump‐probe experiments and atmospheric flow experiments for CO oxidation over single Pt particle with the same composition of reactants. (a) A typical set of pump‐probe CO₂ responses (m/e = 44) for reaction at 140, 170, and 350°C. There is a shift in the amount of CO₂ produced during both CO and oxygen pulses as temperature increases. (b) CO₂ production observed from atmospheric flow experiment. The CO₂ produced while increasing reactor temperature is less than the CO₂ produced during reactor temperature decrease as shown by the counterclockwise hysteresis loop. (c) CO₂ production observed from vacuum pump‐probe experiment. The black line represents the total CO₂ yield. The circle and diamond points represent the CO₂ yield on the oxygen pulse and CO pulse, respectively.
9. Figure 10.9 (a) Oxygen uptake over reactor‐equilibrated VPO at 800 torr O₂. (b) O₂ desorption spectrum under vacuum from ¹⁸O₂‐treated (VO)₂P₂O₇.
10. Figure 10.10 (a) n‐Butane pulse response curves over oxygen‐treated VPO at various temperatures. Each curve is obtained for the same initial VPO oxidation state. (b) Arrhenius plot obtained from the temperature dependence of the n‐butane conversion giving an activation energy of 12 kcal/mol.
11. Figure 10.11 (a) Arrhenius plots for n‐butane oxidation over a single VPO sample reduced by a long series of n‐butane pulses. As the surface is reduced, the activation energy increases. (b) Apparent equilibrium constant for various products as a function of VPO oxidation state.)
12. Figure 10.12 Pump‐probe data showing CO₂ production as a function of temperature and pump‐probe interval. At 150°C the CO₂ production is essentially independent of the pump‐probe interval out 9 s separation. At 350°C the CO₂ production drops as the pump‐probe interval increases, indicating a drop in the active oxygen concentration with time.
13. Figure 10.13 Normalized CO₂ production on the CO pulse calculated from the zeroth moment of the pulse response curve. CO₂ production is constant at 150°C and drops to ca. 0.5 times its value in 9 s at 350°C.
Chapter 11
1. Figure 11.1 Structure of a finite element grid.
2. Figure 11.2 One‐dimensional finite elements. (a) A linear element, (b) a quadratic element, and (c) linear and (d) quadratic variation of field variable ϕ over an element. Black circles represent the nodes, while the white, numbered circles represent the particular element.
3. Figure 11.3 Schematic presentation of a monolith reactor with porous catalyst‐coated channels
4. Figure 11.4 Steady‐state reactor temperature profiles obtained at different values of molar CH₄/O₂ and H₂O/CH₄ ratios at the inlet.
5. Figure 11.5 Transient reactor temperature profile obtained at CH₄/O₂ = 2.24 and H₂O/CH₄ = 1.17.
6. Figure 11.6 Transient reactor temperature profile obtained upon changing CH₄/O₂ from 2.24 to 1.89 (H₂O/CH₄ = 1.17).
7. Figure 11.7 Transient reactor temperature profile obtained upon changing H₂O/CH₄ from 1.17 to 1.56 (CH₄/O₂ = 1.89).
8. Figure 11.8 Frontal view of the heat‐exchange integrated microchannel reactor concept and the two‐dimensional unit cell (shown at the lower inset).
9. Figure 11.9 The first one‐tenth of the computational grid used in the 2D simulation of the microchannel reactor for coupling catalytic methane combustion and i‐octane steam reforming.
10. Figure 11.10 Temperature profiles obtained from the solution of 3D model (side length of the microchannel = 5.6 × 10⁻⁴ m, thickness of the wall separating the channels = 2 × 10⁻⁴ m, reactor material: AISI‐steel).
11. Figure 11.11 Profiles of average bulk temperatures along the steam reforming channel obtained at different wall thickness values by the solution of the 2D model.
12. Figure 11.12 Hydrogen yields and i‐octane conversions (in parentheses) obtained for different (a) wall thickness values, (b) wall materials, and (c) channel side lengths by the solution of the 2D model.
13. Figure 11.13 Two‐dimensional unit cell with (a) straight‐through (ST) and (b) microbaffled (BF) channel wall configurations.
14. Figure 11.14 The first one‐third of the computational grid used in the 2D simulation of the microchannel reactors with ST and BF wall configurations for coupling catalytic methane combustion and i‐octane steam reforming.
Chapter 12
1. Figure 12.1 Scale of catalytic materials and reactors [2].
2. Figure 12.2 Schematic of some possible forms of catalyst pellets, the smallest chemical reactor [9].
3. Figure 12.3 Schematic of German reactors used in the pre‐1955 period according to Tramm [10].
4. Figure 12.4 Schematic of the slurry reactor with multilayer beds [11].
5. Figure 12.5 Possible reactors for Fischer–Tropsch synthesis. (a) Slurry bubble column reactor, (b) multitubular trickle bed reactor, (c) circulating fluidized‐bed reactor, and (d) fixed fluidized‐bed reactor [15].
6. Figure 12.6 Relative activity profile of an precipitated iron catalyst after different periods of time on stream. Times are in relative units (: 1 tu; , 50 tu; , 270 tu; x, 1000 tu (tu = time units)).
7. Figure 12.7 Flow regime maps for trickle‐bed reactors [17].
8. Figure 12.8 Examples of monolith reactors illustrating cell size and shape possibilities [41].
9. Figure 12.9 The Velocys commercial‐scale Fischer–Tropsch reactor.
10. Figure 12.10 Comparison of catalyst productivity for Shell, Sasol, and Velocys Oxford Catalysts Group (OCG) reactors (courtesy of Velocys).
11. Figure 12.11 Productivity of commercial Velocys microchannel, fixed‐bed, and slurry reactors compared to that of the reactor scale (courtesy of Velocys).
12. Figure 12.12 Profile of reactor temperature at the top, middle, and bottom of a small channel reactor during Fischer‐Tropsch synthesis [48].
13. Figure 12.13 CO conversion, hydrocarbon productivity, and methane selectivity for Fischer–Tropsch synthesis at 214°C with a Pt–Co–alumina catalyst [48].
14. Figure 12.14 Model for structured bed of catalyst in a straight flow channel reactor [50].
15. Figure 12.15 Steady‐state performance of a structured catalyst bed (25 bar) [50].
16. Figure 12.16 Illustration of possible types of slurry bubble column reactors. (a) Simple bubble column, (b) cascade bubble column with sieve trays, (c) packed bubble column, (d) multishaft bubble column, and (e) bubble column with static mixers [61].
17. Figure 12.17 Schematic of an Exxon slurry reactor (left) [68]. Schematic of Sasol slurry bubble column reactor (right) [69].
18. Figure 12.18 Illustration of some concepts for the sparger for the bubble column reactor. Left, from [61]: (a) dip tube, (b) perforated plate, (c) perforated ring, and (d) porous plate. Right: closeup view of a sparger from an Exxon patent [68].
19. Figure 12.19 Typical particle trajectories within three different flow regions around a rising bubble in a slurry bubble column reactor [70].
20. Figure 12.20 Model of the flow scheme in the slurry bubble column reactor [73].
21. Figure 12.21 Hydrodynamic model of a slurry bubble column reactor in the heterogeneous flow regime [82].
22. Figure 12.22 Experimental data for gas holdup in a 0.1 m diameter bubble column operating with air–water system spanning both the homogeneous and heterogeneous flow regimes [113].
23. Figure 12.23 Examples of structured packing: (a) open cross flow structure (OCFS), (b) closed cross flow structure (CCFS), (c) knitted wire, and (d) aluminum foam [117].
24. Figure 12.24 Schematic representation of structured catalytic packing made by BUCT: (a) cross section, (b) side view, and (c) actual photography [118].
25. Figure 12.25 Membrane reactors for FT synthesis from the literature: (a) distributed feeding of reactants A and B, (b) in situ water removal by selective membrane (F, feed; S, sweep), (c1) plug‐through contactor membrane (PCM) with wide transport pores, (c2) forced‐through flow membrane contactor, product and heat removal by circulated liquid product, (d) zeolite encapsulated FT catalyst, P′, modified product [123].
26. Figure 12.26 Density versus pressure of the mixture of 55% hexane and 45% pentane at 220°C [131].
27. Figure 12.27 Schematic of a laboratory‐scale supercritical fixed‐bed reactor [131]. T/C: thermocouple.
28. Figure 12.28 CO conversion versus time on stream on a 25% Co/γ‐Al₂O₃ slurry‐phase impregnation catalyst in a fixed‐bed reactor with varying partial pressure of supercritical fluid (SCF).
29. Figure 12.29 Olefin selectivity as an function of carbon number for a variety of reactors.
30. Figure 12.30 Cost of plant versus the plant capacity [135].
31. Figure 12.31 Schematic of a FT reactor suitable for a mobile biorefinery [136].
Chapter 13
1. Figure 13.1 Flow diagram of a typical hydrotreating unit.
2. Figure 13.2 Hydrotreating reactor technologies.
3. Figure 13.3 Process scheme of a typical petroleum refinery.
4. Figure 13.4 World consumption of oil refined products from 1990 to 2010.
5. Figure 13.5 Crude oil production in Mexico.
6. Figure 13.6 Reaction pathways of dibenzothiophene HDS.
7. Figure 13.7 Reaction scheme of quinoline HDN.
8. Figure 13.8 Simplified representation of the HDM of metalloporphyrins.
9. Figure 13.9 Examples of saturation reactions.
10. Figure 13.10 Reaction mechanism of paraffin hydrocracking.
11. Figure 13.11 First‐order rate coefficient values for the HDS of various sulfur compounds.
12. Figure 13.12 Multibed hydrotreating reactor with quenching.
13. Figure 13.13 Hydrogen circuit in a hydroprocessing unit.
14. Figure 13.14 Representation of the trickle‐flow regime in a hydroprocessing reactor.
15. Figure 13.15 Concentration profiles in a TBR reactor.
16. Figure 13.16 General representation of liquid quenching‐based processes. (a) Multiple feeding. (b) Product recycling.
17. Figure 13.17 Hydroprocessing reactor internals. (—) Axial delta‐T. (‐‐‐) Radial delta‐T.
18. Figure 13.18 Distributor tray design parameters. (a) Effect of tray spacing on liquid distribution, (b) comparison of tray spacing, and (c) discharge pattern of several tray designs.
19. Figure 13.19 Onstream catalyst replacement reactor.
20. Figure 13.20 H‐Oil ebullated‐bed reactor.
21. Figure 13.21 Slurry‐phase hydroprocessing reactor.
22. Figure 13.22 Sulfiding conditions of an HDS catalyst.
23. Figure 13.23 IMP heavy oil upgrading process.
24. Figure 13.24 Details of the experimental program.
25. Figure 13.25 Sulfur, nitrogen, and metals molar concentration profiles. Simulated: (—) liquid phase, (‐‐‐) solid phase; (■) experimental.
26. Figure 13.26 Hydrogen and hydrogen sulfide molar concentration profiles. Simulated: (—) liquid phase, (‐‐‐) solid phase.
27. Figure 13.27 Evolution of the gas‐phase composition. (—) Simulated. (■) Experimental.
28. Figure 13.28 H₂/oil ratio and superficial gas velocity profiles. Simulated: (—) 380°C, (‐‐‐) 400°C; (■) experimental.
29. Figure 13.29 Process performance during time‐on‐stream: (symbols) experimental, (—) simulated.
30. Figure 13.30 HDM performance and MOC: (symbols) experimental, (—) simulated. The experimental values (■) of MOC at the end of the run were estimated based on the metals balance between feed and products.
31. Figure 13.31 Simulated axial MOC profiles. The MOC units are referred to the total amount of fresh catalyst in the reactors.
32. Figure 13.32 Catalyst wetting efficiency as function of superficial liquid mass velocity: (■) bench scale, () semi‐industrial scale, (—) simulated.
33. Figure 13.33 Comparison of semi‐industrial and bench‐scale performances.
34. Figure 13.34 Simulation of the proposed industrial reactor configuration. Evolution of reactor temperature, H₂/oil ratio, and concentration of the chemical lumps. (‐‐‐) Average temperature and H₂/oil ratio.
35. Figure 13.35 Average H₂/oil ratio as function of the number of hydrogen quenches.
36. Figure 13.36 Evolution of the axial temperature profiles.
Chapter 14
1. Figure 14.1 Ethylene formation versus time on stream; (a) autothermal reforming of synthetic diesel (O/C = 1.12, S/C = 1.25 [36]).(b) autothermal reforming of commercial diesel (S/C=3, O/C=1.6 [35]).
2. Figure 14.2 Different feed injection systems; (a) ATR‐5, (b) ATR‐8 [37].
3. Figure 14.3 Different feed injection systems; (a) first generation [42].(b) improved injection system without recirculation zone [43].
4. Figure 14.4 Methane conversion and reactor exit temperature versus GHSV for a fixed bed compared to a metallic monolith (S/C = 3) [63].
5. Figure 14.5 (a) Jellyroll monolith prepared from platinum/ceria catalyst, (b) comparison of propane conversion, hydrogen yield and H₂/CO ratio for fixed bed, jelly roll, and monolithic reactors as determined for different GHSV [64].
6. Figure 14.6 Hot spot formation and methane and oxygen concentration profiles in two ceramic monoliths of different cell density for partial oxidation of methane at O/C = 0.54 and at a feed inlet temperature of 362°C; (a) temperature profile for 400 cpsi, (b) temperature profile for 115 cpsi, (c) concentration profiles for 400 cpsi, (d) concentration profiles for 115 cpsi [65].
7. Figure 14.7 Ethylene formation versus temperature at different S/C ratio for autothermal reforming of low sulfur diesel (O/C = 0.45) [34].
8. Figure 14.8 Effect of feed inlet temperature and O/C ratio (here shown as O₂/C ratio) on reforming efficiency (O/C = 2.5, GHSV = 9 900–10 800 h⁻¹) [44].
9. Figure 14.9 (a) System start‐up concept, (b) reformer start‐up time [43].
10. Figure 14.10 Species concentration versus O/C ratio (here expressed as O₂/C) for an autothermal diesel reformer at S/C = 2.3. Symbols represent experimental data, the lines represent simulation results [76].
11. Figure 14.11 Simulation of a cocurrently operated plate heat exchanger for diesel steam reforming coupled to a catalytic diesel burner. Left: flow arrangement, second left: fuel mass fraction in the reformer, third left: fuel mass fraction in the burner, right: temperature in the plate.
12. Figure 14.12 Scheme of the membrane reactor for methanol reforming [109].
13. Figure 14.13 Conversion in a membrane reactor applying different membrane types. For comparison, thermodynamic equilibrium conversion and conversion as achieved in a fixed catalyst bed are added [113].
14. Figure 14.14 Scheme of the membrane reactor for methanol oxidative steam reforming [115].
15. Figure 14.15 Scheme of the fluidized bed membrane reactor [117].
16. Figure 14.16 Scheme of the electrochemical membrane reactor for partial oxidation of methane [119].
17. Figure 14.17 Glycerol conversion at different WHSV in a membrane reactor (MR) and in a traditional reactor (TR) [125].
18. Figure 14.18 Schematic of the staged‐separation membrane reactor [126].
19. Figure 14.19 (a) Integrated water–gas shift reactor/heat exchanger designed for 5 kW fuel cell system, (b) internal temperature profiles determined during operation of this reactor; the different profiles correspond to different load (100% load is the 5 kW equivalent). The values shown for 100% load were determined at the fresh reactor (0 h) after 24 h under operation (24 h) and after an accidental exposure to reformate of too low temperature for the water–gas shift reaction (after LT) [141].
20. Figure 14.20 Thermal conductive catalyst plates integrated into a water–gas shift reactor. The dotted line indicates the flow paths of reformate and air [143].
21. Figure 14.21 (a) Axial, (b) centrifugal, and (c) centripetal flow arrangements of the water–gas shift reactors [144].
22. Figure 14.22 CO conversion as a function of reaction temperature for a membrane reactor (MR) and the traditional process at different GHSV for 350°C feed temperature and a reaction pressure of 15 bar [145].
23. Figure 14.23 Volume reduction by application of membrane technology for water–gas shift for different feed pressure and reaction temperature [145].
24. Figure 14.24 Cocurrently liquid‐cooled folded‐plate reactor for preferential oxidation [166].
25. Figure 14.25 (a) Cocurrently operated microreactor for preferential oxidation cooled by water evaporation, (b) CO concentration in the off‐gas of the reactor at different total reformate flow rates, CO inlet concentrations and O/CO values [167].
26. Figure 14.26 Temperature profile of a cocurrently cooled preferential oxidation reactor with 5 kW_el,net power equivalent.
27. Figure 14.27 Flow scheme of a natural gas fuel processor/fuel cell system [168].
28. Figure 14.28 Gas temperature, carbon monoxide, and hydrogen content in the gas phase at different components of the natural gas fuel processor/fuel cell system [168].
29. Figure 14.29 Start‐up energy demand of the different components of the natural gas fuel processor/fuel cell system [168].
30. Figure 14.30 Sankey diagram of the 2 kW_el CHP PEM fuel cell/methane fuel processor system [171].
31. Figure 14.31 Integrated methanol fuel processor with 100 W power equivalent [16, 86].
32. Figure 14.32 (a) Fuel processor of the VeGA system [16].(b) 250 W_el fuel cell/fuel processor system VeGA developed by a cooperation of TRUMA and IMM [106].
Chapter 15
1. Figure 15.1 Simplified reaction scheme of stearic acid deoxygenation.
2. Figure 15.2 The experimental, the first‐order power law, and the Langmuir batch reaction curves for stearic acid [20]. Parameters: p₂ = 0.011 l/mol, p₃ = 25 l/mol.
3. Figure 15.3 Dimensionless initial concentration of stearic acid and heptadecane inside spherical catalyst particles located near the reactor inlet.
4. Figure 15.4 Effect of the rate constant of the main reaction on stearic acid concentration.
5. Figure 15.5 Effect of the rate constant of coking on stearic acid concentration.
6. Figure 15.6 Effect of the final activity on stearic acid concentration.
7. Figure 15.7 Stearic acid profiles inside the reactor.
8. Figure 15.8 Correlation between the main reaction rate and the final activity (Pe = 3).
9. Figure 15.9 Correlation between the deactivation rate and the final activity (Pe = 3).
10. Figure 15.10 Correlation between the main and coking reaction rate parameters (Pe = 3).
11. Figure 15.11 Identification between the rate parameter of the main reaction and the final activity.
12. Figure 15.12 Identification between reaction rate parameters of the main reaction and coking.
13. Figure 15.13 The effect of Peclet number on stearic acid consumption (catalyst: (Pd/C (Aldrich))).
14. Figure 15.14 The dependence of parameter estimation results on the Pe number.
15. Figure 15.15 Estimated and experimentally observed concentrations of stearic acid (catalyst: (Pd/C (Aldrich))).
16. Figure 15.16 Estimated and experimentally observed concentrations of stearic acid (catalyst: (Pd/C (Sibunit))).
17. Figure 15.17 Simulated stearic acid concentration change with time in the pilot reactor (T = 633°C, p = 20 bar, residence time, and active metal concentration same as in the laboratory scale).
18. Figure 15.18 Simulated stearic acid concentration change with time in the pilot reactor (T = 633°C, p = 10 bar, active metal loading same as in the laboratory‐scale unit).
19. Figure 15.19 Cumulative production of the pilot reactor.
20. Figure 15.20 Stearic acid conversion (%) as a function of time on stream.