Successful Scale-up of an Industrial Trickle Bed Hydrogenation UsingLaboratory Reactor DataDaniel A. Hickman,*,† Michael T. Holbrook,‡,§ Samuel Mistretta,‡ and Steven J. Rozeveld†

†The Dow Chemical Company, Midland, Michigan 48674, United States‡The Dow Chemical Company, Plaquemine, Louisiana 70765, United States

ABSTRACT: This work validates the appropriate application of chemical reaction engineering principles in the successful designof a full-scale industrial trickle bed reactor for a proprietary hydrogenation reaction over a palladium catalyst. After identifying aneffective catalyst formulation in a continuous laboratory scale trickle bed reactor, the project team used the same small-scalereactor to generate kinetic data for scale-up. The scale factor from the laboratory to the final design was about 3 × 106. Thedevelopment effort identified and resolved three important problems: (1) incomplete catalyst wetting of the small catalyst bed,even though the catalyst was diluted with inert fines; (2) lower than economically attractive catalyst productivity; and (3) catalystdeactivation.

■ INTRODUCTION

An important objective for the industrial reaction engineer is todesign a commercial scale reactor that achieves the targetperformance parameters, including production rate and productyield, while minimizing the investment of resources and thetime elapsed. The scale-up risks encountered by the engineervary depending on the nature of the chemistry and the reactorsystem. In practice, no single work process can be universallyapplied to all reactor scale-up projects. However, certain classesof problems provide opportunities to apply reaction engineer-ing fundamentals to enable an efficient scale-up program whilesufficiently mitigating risk factors associated with the scale-upprocess. In this paper, we describe a specific program in which anew trickle bed hydrogenation process was scaled directly fromthe laboratory to the commercial scale reactor without buildingand operating intermediate scale reactors. This programsucceeded by properly accounting for the relevant interactionsbetween transport phenomena and kinetics1 while scaling by afactor of about 3 × 106.Many previous authors have highlighted the importance of

properly designing a continuous laboratory scale fixed bedreactor to avoid axial dispersion,2−7 wall effects,7,8 incompletecatalyst wetting,8−10 and nonisothermal bed temperatures.11 Inthis work, we applied those principles to enable generation ofapparent kinetic data, or reaction kinetics that lump the effectsof pore diffusion and the intrinsic rates, in an integral reactorwith fines.12 During this scale-up program we encountered andresolved three particular problems, resolutions of which aresummarized in this paper. While we do not discuss theproprietary chemistry, we offer this case study to illustrate theapplication of basic reaction engineering principles to the scale-up process and to provide insights into the nature of some ofthe typical problems encountered in the scale-up of anindustrial trickle bed reactor. We also use this case study todefend our assertion that an intermediate-scale pilot plant is notalways necessary to achieve successful scale-up and commerci-alization of new trickle bed reactor technology.

■ EXPERIMENTAL METHODS

Reactor System. The integral laboratory reactor operatedin cocurrent downflow (trickle flow) or cocurrent upflowdepending on the configuration of multiple feed delivery andproduct collection valves. An excess volume of silicon carbide(100−140 mesh, or about 0.2 mm diameter) or glass beads(60−80 mesh, or about 0.4 mm in diameter) diluted thecatalyst (0.4−4.1 g of extrudates with a nominal diameter of 1/16 in. or 1/8 in., or about 1.6 or 3.2 mm) and filled the emptyspace in the reactor above the catalyst bed. The reactorconsisted of a 1/2-in. or 1/4-in. nominal outer diameter metaltube inside an oil jacket constructed of 1-in. tubing. An oil bathwith circulation pump circulated thermostatically controlled oilthrough the jacket to ensure a uniform reactor wall temper-ature. A positive displacement pump delivered the liquid feedmixture to the reactor from a feed reservoir on a balance. Amass flow controller continually delivered hydrogen. A cooledvapor−liquid separator provided the means to separate thevapor and liquid effluent from the reactor. A control valvemaintained the liquid level in the phase separator, and a secondcontrol valve on the vapor effluent controlled the systempressure. Online gas chromatography provided separateanalyses of the vapor and liquid products, enabling closure ofthe system mass balance for each atomic species.For both reactors, the feed mixture was representative of the

composition expected in the commercial scale reactor. Wevaried the reactor pressure and temperature systematically tocover the entire range expected in the full-scale reactor, and wevaried the liquid and hydrogen feed rates from 0.35 to 7.0 mL/min and 100 to 400 sccm, respectively.

Special Issue: NASCRE 3

Received: February 18, 2013Revised: March 22, 2013Accepted: March 27, 2013

Article

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Catalyst Deactivation Experiments. Following anextended (3600 h) lab reactor run, the catalyst was unloadedand separated into two samples based on color: a tan samplefrom the beginning (top) of the catalyst bed and a black samplefrom the end (bottom) of the bed. We also collected samples offresh catalyst and catalyst used in a different experiment for 600h for analysis by various analytical techniques. The goal of theanalyses was to understand the deactivation mechanisms andthe extent that each played in the observed catalyst deactivationprocess. We label the samples as follows: (1) fresh, (2) 600 h,(3) 3600 h bottom, and (4) 3600 h top. Portions of both 3600h samples were subsequently separately loaded and rerun in thelab reactor to determine the activity of each type of catalyst andcompared with fresh catalyst. The fresh sample, the bottomsample, and the top sample gave relative rates of 1.3, 1.0, and0.5, respectively.X-ray Photoelectron Spectroscopy. Catalyst extrudates

were affixed onto a metal plate using carbon tape for eachsample and spectra were recorded from three different areas forstatistical analysis. Samples were examined in the as-receivedstate, and it was necessary to solvent extract the residualreaction liquids prior to the analysis.Samples were initially examined by low-resolution survey

scans followed by high-resolution spectra of specific elements inorder to determine the binding energy (chemical state) andconcentration of the elements detected in the survey scans. Thequantification of the elements was accomplished by using theatomic sensitivity factors for a Kratos model HSi XPSspectrometer, using monochromatic Al−Kα as the X-raysource. Charge compensation was used for all spectra.The carbon (1s) photoline was used as the calibration

reference for the binding energy axis of all high-resolutionspectra. Intensity due to aromatic carbon−carbon bonding wasshifted to 284.8 eV. The resulting offset was measured andapplied to the other high-resolution spectra for the same pointof analysis. Each point of analysis was energy correctedindependent of the others.Microprobe Sample Preparation. Cross-section samples

of the fresh and used catalysts were prepared by embedding thecatalysts in acrylic resin followed by microtoming the samples.These samples were used for both electron microprobe andTEM (transmission electron microscopy) experiments.The procedure for embedding the catalyst samples was as

follows: First, several of the fresh extrudates were placed “edge-on” into gelatin capsules, filled with LR White acrylic resin, andplaced under house vacuum (∼1 Torr) for 20 min to removetrapped air from the catalyst pores. The capsules were thendried at room temperature for 7 days. The samples were curedat room temperature to eliminate any artificial sintering byheating the samples during the curing process.The used catalysts were prepared using a slightly different

method from the fresh catalyst. Several pieces (0.5 in. long) ofthe used catalysts were soaked overnight in acetone to extractthe excess oligomers that were trapped in the porous catalystthat would inhibit the acrylic resin from curing. The acetone-soaked catalysts were then cured at room temperature andembedded in the LR White acrylic resin. The samples weremicrotomed (dry) using a diamond knife to prepare cross-section blocks.Electron Microprobe Experiments. The microtomed

blocks were carbon coated and examined in a Cameca SX50electron microprobe (serial no. SX401) run by SAMx software.Quantitative microanalysis was done at 15 keV and 50 nA. The

concentrations of iron (Fe) and palladium (Pd) were measuredusing wavelength dispersive spectrometers (WDS) using theFe−Kα and Pd−Lα peaks. Element maps were collected at 15keV, 50 nA, and 50 μs/pixel using WDS. The standard map sizewas 800 μm × 200 μm, although higher resolution maps of 100μm × 100 μm were also recorded at the pellet rim.Quantitative line scans were used to determine the weight

percent of iron and palladium as a function of distance from thepellet exterior. This was done by collecting WDS peaks for eachelement (and background) at discrete positions along a line(∼200 μm long) starting at the rim of the pellet and traversinginto the pellet interior. The distance between data points was10 μm in the bulk and reduced to 2 μm closer to the acrylicresin/catalyst surface.

Aberration-Corrected High-Resolution TEM. Aberrationcorrected (AC) TEM experiments were conducted at OakRidge National Laboratory (ORNL) using the JEOL AC-2200FS. The JEOL 2200FS-AC was equipped with a CEOSGmbH aberration corrector. The AC-TEM formed extremelysmall probe sizes of less than 1.2 Å diameter and was optimizedfor scanning-TEM experiments on catalyst materials.We captured images at a resolution of 1024 × 1024 pixels

using a 100 μs/pixel dwell time. At a magnification of 200kx,the resolution was 6.9 Å/pixel and at 500kx, the resolution was2.8 Å/pixel. The probe size for the ORNL AC-TEM instrumentwas 1.2 Å diameter using a 35-μm aperture (25 pA probecurrent). Before the scanning-TEM analysis, the area of interestwas exposed to an electron dose (“beam shower”) by removingthe condenser aperture and defocusing the probe over a ∼100-μm area for several minutes.The catalysts for the TEM analysis were prepared using the

same procedure as for the microprobe experiments. Thinsections ∼70 nm thick were cut using a Riechart Ultracutmicrotome at room temperature using a diamond knife, floatedonto DI water, and collected onto Cu grids with a lacey carbonsupport. Note that the samples that were prepared for electronmicroprobe studies were cut dry with a diamond knife as only apolished block face was needed.Before the TEM analysis, the area of interest was exposed to

an electron dose (“beam shower”) by removing the condenseraperture and defocusing the probe over a ∼100-μm area forseveral minutes. This step minimized carbon contaminationfrom the acrylic resin and did not introduce artifacts into theTEM analysis.

■ RESULTS AND DISCUSSIONCatalyst Wetting. The earliest experiments in our research

used the 1/2-in. reactor tube. In preliminary experiments togenerate kinetic data, we found that the reaction order was firstorder with respect to the organic reactant. We determined thisby varying the reactant concentration in the feed by a factor of2 at a fixed space time (liquid volumetric feed rate),temperature, and pressure; the fractional conversion was thesame in both cases. However, when we varied the space time atseveral different temperatures with a fixed feed composition,the apparent reaction order with respect to the reactant wasnear 1.5.These contradictory results (reaction order of 1.55 instead of

first order) led to the hypothesis that the experiments gavefractional wetting efficiencies that increased with increasingflow rates. In support of this hypothesis, a report in the openliterature shows that complete wetting is not guaranteed whenthe catalyst is diluted with small, inert particles.13 We tested this

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hypothesis by loading a similar quantity of catalyst and siliconcarbide diluent into the 1/4-in. tube and repeating the variablespace time experiments using the same range of volumetric flowrates and, thus, higher superficial velocities. In this case, theresults fit very nicely to a first order model over the entire rangeof flow rates (Figure 2). This result supported the hypothesisthat the higher apparent reaction order in the experiments inthe 1/2-in. tube was the consequence of incomplete catalystwetting, with the fractional wetting efficiency increasing withincreasing flow rate. All subsequent experiments for catalysttesting and kinetic model development used the 1/4-in. tubereactor, where the superficial velocities were high enough togive complete catalyst wetting for range of flow rates employedin the experimental program.For the experiments in Figures 1 and 2, the total mass of

catalyst loaded for each of these experiments was the same

(2.05 g). In the 1/2-in. tube experiments, we observed 60%conversion at temperature T2, a flow rate of 1.24 mL/min, and763 h on stream. In the 1/4-in. tube, we obtained 62%conversion at the same temperature, a flow rate of 2.95 mL/min, and 145 h on stream. This corresponds to a productivity

increase of about 140% (a factor of 2.4), although some of thatdifference could be attributed to catalyst deactivation since thedata point from the 1/2-tube was for a catalyst with significantlymore time on stream.We also performed experiments comparing the effect of the

flow direction with the 1/4-in. reactor. According to theliterature, a lab reactor in which the catalyst bed is diluted withfines should give identical results regardless of the flowdirection (cocurrent upflow or cocurrent downflow) if thefines have effectively decoupled the hydrodynamic effects fromthe reaction kinetics and intraparticle effects.14,15 In otherwords, since complete wetting of the catalyst surface is ensuredduring upflow because the continuous fluid phase is the liquidphase, then achieving identical results in both upflow anddownflow implies that essentially complete wetting of thecatalyst surface is obtained during downflow.These experiments used 0.82 g of supported Pd, 1/16-in.

extrudates, diluted with 60/80 mesh glass balls. Theseexperiments were conducted after the catalyst had been inoperation for about 4000 h. The conversion was 53.6% withdownflow and 52.6% with upflow under otherwise identicalconditions, supporting our assumption that the catalyst wasfully wetted during downflow experiments. With confidencethat these data were not compromised by incomplete catalystwetting, we used this same catalyst load to generate data fordevelopment of models of reaction kinetics and deactivationkinetics.

Catalyst Productivity. Early catalyst development workused catalyst extrudates with a diameter of 1/8-in. Calculationspredicted that pore diffusion significantly limited the rate ofreaction, even in smaller catalyst particles. For example, forextrudates 1/16-in. in diameter and 1/8-in. long, based on theobserved reaction rate and assuming a first-order reaction,calculations gave an estimated Thiele modulus of 2.5. For aThiele modulus above 2.0, the effectiveness factor is inverselyproportional to the Thiele modulus and is therefore inverselyproportional to the characteristic pore length. Consequently,for particles approximately 1/16 in. and larger, we predictedand observed experimentally (Table 1) that the activity per

mass of catalyst was inversely related to the particle diameter.This motivated the use of catalyst particles as small as possiblewithin the limits of reactor pressure drop constraints and thelimitations of the catalyst production process. By switchingfrom 1/8-in. to 1/16-in. extrudates, we effectively doubled thereactor productivity, providing a positive boost to the economicattractiveness of the process.

Catalyst Deactivation. Our conceptual model of thedeactivation of the supported Pd catalyst in this reaction systeminvolved four distinct mechanisms: (1) sintering of the initiallyhighly dispersed Pd particles, (2) iron poisoning of the catalyst,which results in the formation of Fe−Pd alloy particles with

Figure 1. Fit of 1.55 order model (curves) to fractional conversiondata (symbols) for 1/8-in. extrudates in 1/2-in. tube at five differenttemperatures (T1−T5) and various total volumetric flow rates (F).Each temperature curve was fit separately assuming a reaction order of1.55 with respect to the reactant.

Figure 2. Fit of 1.0 order model (curves) to fractional conversion data(symbols) for 1/8-in. extrudates in 1/4-in. tube at two differenttemperatures (T1 and T2) and various total volumetric flow rates (F).Each temperature curve was fit separately assuming a reaction order of1.0 with respect to the reactant.

Table 1. Comparison of Apparent Activity of 1/16-in. and 1/8-in. Extrudates

1/8-in.extrudates

1/16-in.extrudates

mass of catalyst (g) 2.05 2.00time on stream (h) 6.2 8.2liquid feed rate (mL/min) 3.04 3.27hydrogen feed rate (sccm) 400 400first-order rate constant (mL/g-min) 2.46 4.85

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lower or perhaps no activity for hydrogenation, (3) Pd loss nearthe surface, and (4) fouling. We observed these modes ofdeactivation by comparing fresh and spent catalyst using XPS,electron microprobe, and TEM.Following extended (3600 h) lab reactor runs, we discovered

that a portion of the catalyst extrudates from the inlet (top) ofthe reactor bed were discolored on the outer surfaces (with adistinct “tan” color) compared to the fresh, 600-h, and the3600-h bottom samples, which were all black. The interior ofthe discolored “tan” extrudates was still black. XPS revealed asubstantial increase in the concentration of surface carbon withtime on stream, strong evidence for significant catalyst fouling(Figure 3). A second notable difference identified by the XPS

analyses was a sharp increase in the concentration of iron in the3600-h top sample. Iron was not expected or detected in thefresh catalyst material.Among the most informative electron microprobe experi-

ments in this study were quantitative line scans of the fresh andused catalysts in cross-section. The line scans provided a moreaccurate measurement of the Pd and Fe concentrationcompared to the element maps since a longer dwell time/pixel could be used. Line scans were generated by collecting aspectrum at discrete points (for example, every 2 μm) on thecross-sectional sample from the edge of the pellet to ∼200 μminto the pellet interior.Line scans from the fresh sample, shown in Figure 4,

indicated that the Pd loading was uniform across the pellet(∼2.2 wt % Pd) with no depletion near the rim. The Feconcentration was below the detection limit (0.05 wt %) in thefresh catalyst.Similar line scans were done for two different 3600-h bottom

pellets, and a representative profile is shown in Figure 5. ThePd concentration (∼2.0 wt %) was uniform across most of thecross-section, although a small decrease in the Pd concentrationwas observed near the pellet edge. The Fe concentration wasnear or below the detection limit (0.05 wt %).Line scans from the 3600-h top sample were recorded from

two different pellets (Figure 6). The maximum Pd concen-tration of the first pellet was ∼1.7 wt % Pd (as shown in thefigure) but only ∼1 wt % Pd for the other pellet. In both cases,the Pd concentration was highest away from the pellet surface.The Pd concentration in the 3600-h top sample decreasedtoward the rim of the pellet but with spikes to ∼1−2 wt % at

the areas with high Fe contamination. The line profiles clearlyshow Pd depletion in the outer ∼30 μm of the pellet to as lowas ∼0.3 wt % Pd (data not shown). In the pellet interior, the Feconcentration was low, ∼0.1 wt %. The line profiles clearlyshowed Pd depletion in the outer 30 μm of the pellet, with Fepreferentially depositing onto Pd rich areas near the surface,giving local Fe concentrations approaching 20 wt %. Fuentesand Figueras previously reported similar iron poisoning ofsupported Pd catalysts.16 Based on further analyses of the freshcatalyst samples, the Pd-rich areas appear to be the

Figure 3. XPS surface concentration (atom %) of C and Fe in freshand used catalysts.

Figure 4. Line scan of fresh catalyst.

Figure 5. Line scan of 3600-h bottom sample.

Figure 6. Line scan of 3600-h top sample.

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consequence of the catalyst preparation process rather than aresult of Pd migration with time on stream.Scanning TEM (dark field) experiments of the used catalysts

were conducted at ORNL to determine the Pd particle size anddistribution. Surprisingly, after 600 h, the average Pd sizedistribution had decreased (near the pellet surface), and manysmall Pd particles were observed. The median particle size forthe 600-h sample was ∼17 Å, compared to ∼34 Å in the freshsample (Table 2). The downward trend in Pd particle size

continued with the 3600-h bottom catalyst, and few Pd particleswere observed near the surface of the 3600-h top sample. Fromthese observations, the leaching of Pd into the liquid reactionmixture likely drives the loss of Pd near the surface.Scanning TEM (dark field) images were also recorded ∼100

μm away from the pellet surface to determine the Pd particlesize and distribution as a function of time on stream. Themedian Pd particle size was 14 Å for the fresh catalyst (Table 3)and increased linearly over time on stream (Figure 7),providing clear evidence of Pd sintering.

These efforts to understand the likely mechanisms forcatalyst deactivation led the scale-up team to focus oncontrolling the one mechanism that could be controlled: ironpoisoning. To prevent or minimize this poisoning, we designedthe process and operating discipline to prevent introduction ofan iron-contaminated feed to the catalyst bed. The deactivation

due to sintering, leaching, and fouling in the absence of ironpoisoning was determined to be sufficiently slow to provide aneconomically viable catalyst lifetime.

■ CONCLUSIONSBy increasing the aspect ratio of the catalyst bed in order toincrease the liquid superficial velocity, the development teamovercame the catalyst wetting problem in the laboratoryreactor. Furthermore, decreasing the characteristic diffusionlength by choosing smaller catalyst particles increased theproductivity to an economically attractive conversion rate. Byusing raw materials representative of the expected commercialplant feed stream rather than synthetic feeds from thebeginning of the experimental program, the team identifiedcatalyst deactivation as an important problem. The team thenmanaged this problem using state-of-the-art analytical techni-ques to identify several parallel modes of catalyst deactivation,followed by the identification of appropriate steps to minimizethe rate of catalyst deactivation to an economically viable rate.Then, using the appropriately sized and loaded lab reactor, theteam generated data for development of models of reactionkinetics and deactivation kinetics. The team used these modelsto design the commercial-scale reactor, successfully scaling thereactor by a factor of about 3 × 106 from these laboratory scaleexperiments to the full-scale reactor design. Finally, effectivetechnology transfer from the development team to the processdesign and construction team, coupled with implementation ofa strict operating discipline in the plant, resulted in thesuccessful startup and operation of the hydrogenation plant.

■ AUTHOR INFORMATIONCorresponding Author*E-mail: dahickman@dow.com; phone: 989-636-2165.

NotesThe authors declare no competing financial interest.§Retired.

■ ACKNOWLEDGMENTSWe thank Timm Richardson and Cliff Todd of AnalyticalSciences, Dow Chemical, for their expertise in conducting theelectron microprobe and XPS experiments. We also thankDoug Blom and Larry Allard at ORNL for assistance with theAC-TEM experiments (High Temperature Materials Labo-ratory, Microscopy, Microanalysis, Microstructures Group, OakRidge National Laboratory, PO Box 2008, 1 Bethel ValleyRoad, Oak Ridge, TN 37831-6064). This portion of theresearch was sponsored by the Asst. Sec. for Energy Efficiencyand Renewable Energy, Office of FreedomCAR and VehicleTechnologies, as part of the High Temperature MaterialsLaboratory User Program, ORNL, managed by UT-BattelleLLC for the U.S. DOE.

■ REFERENCES(1) Dudukovic, M. P. Relevance of multiphase reaction engineeringto modern technological challenges. Ind. Eng. Chem. Res. 2007, 46,8674−8686.(2) Saroha, A. K.; Khera, R. Hydrodynamic study of fixed beds withcocurrent upflow and downflow. Chem. Eng. Process. 2006, 45, 455−460.(3) Sie, S. T.; Krishna, R. Process development and scale up: III.Scale-up and scale-down of trickle bed processes. Rev. Chem. Eng.1998, 14, 203−252.

Table 2. Palladium Particle Size near the Pellet Surface

median average (±1 SD)

fresh 34 Å 45 ± 32 Å600-h 17 Å 23 ± 16 Å3600-h bottom 17 Å 18 ± 9 Å3600-h top little Pd detected little Pd detected

Table 3. Palladium Particle Size in the Pellet Interior (∼100μm from the Pellet Surface)

median average (±1 SD)

fresh 14 Å 18 ± 12 Å600-h 21 Å 26 ± 18 Å3600-h bottom 62 Å 58 ± 40 Å3600-h top 21 Å 30 ± 26 Å

Figure 7. Pd median and average particle size versus time on stream inthe pellet interior (100 μm from catalyst surface).

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(4) van Herk, D.; Kreutzer, M. T.; Makkee, M.; Moulijn, J. A. Scalingdown trickle bed reactors. Catal. Today 2005, 106, 227−232.(5) Wanchoo, R. K.; Kaur, N.; Bansal, A.; Thakur, A. RTD in tricklebed reactors: Experimental study. Chem. Eng. Commun. 2007, 194,1503−1515.(6) Mears, D. E. The role of axial dispersion in trickle-flow laboratoryreactors. Chem. Eng. Sci. 1971, 26, 1361−1366.(7) Mederos, F. S.; Ancheyta, J.; Chen, J. Review on criteria to ensureideal behaviors in trickle-bed reactors. Appl. Catal., A 2009, 355, 1−19.(8) Al-Dahhan, M. H.; Dudukovic, M. P. Catalyst bed dilution forimproving catalyst wetting in laboratory trickle-bed reactors. AIChE J.1996, 42, 2594−2606.(9) Gierman, H. Design of laboratory hydrotreating reactors: ScalingDown of Trickle-flow Reactors. Appl. Catal. 1988, 43, 277−286.(10) Ruecker, C. M.; Hess, R. K.; Akgerman, A. Effect of PartialWetting on Scale-up of Laboratory Trickle-Bed Reactors. Chem. Eng.Commun. 1987, 49, 301−315.(11) Mears, D. E. Diagnostic criteria for heat transport limitations infixed bed reactors. J. Catal. 1971, 20, 127−131.(12) Hickman, D. A.; Weidenbach, M.; Friedhoff, D. P. A comparisonof a batch recycle reactor and an integral reactor with fines for scale-upof an industrial trickle bed reactor from laboratory data. Chem. Eng. Sci.2004, 59, 5425−5430.(13) Tsamatsoulis, D.; Al-Dahhan, M.; Larachi, F.; Papayannakos, N.The effect of particle dilution on wetting efficiency and liquid filmthickness in small trickle beds. Chem. Eng. Commun. 2001, 185, 67−77.(14) Wu, Y.; Khadilkar, M. R.; Al-Dahhan, M. H.; Dudukovic, M. P.Comparison of Upflow and Downflow Two-Phase Flow Packed-BedReactors with and without Fines: Experimental Observations. Ind. Eng.Chem. Res. 1996, 35, 397−405.(15) Chander, A.; Kundu, A.; Bej, S. K.; Dalai, A. K.; Vohra, D. K.Hydrodynamic characteristics of cocurrent upflow and downflow ofgas and liquid in a fixed bed reactor. Fuel 2001, 80, 1043−1053.(16) Fuentes, S.; Figueras, F. Hydrogenolysis of cyclopentane andhydrogenation of benzene on palladium catalysts of widely varyingdispersion. J. Chem. Soc., Faraday Trans. 1978, 1, 174−181.

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