Successful Scale-up of an Industrial Trickle Bed Hydrogenation Using

Publication Date (Web): March 27, 2013 ... This work validates the appropriate application of chemical reaction engineering ... The scale factor from ...
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Successful Scale-up of an Industrial Trickle Bed Hydrogenation Using Laboratory Reactor Data Daniel 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 design of a full-scale industrial trickle bed reactor for a proprietary hydrogenation reaction over a palladium catalyst. After identifying an effective catalyst formulation in a continuous laboratory scale trickle bed reactor, the project team used the same small-scale reactor to generate kinetic data for scale-up. The scale factor from the laboratory to the final design was about 3 × 106. The development 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) catalyst deactivation.

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INTRODUCTION

Reactor System. The integral laboratory reactor operated in cocurrent downflow (trickle flow) or cocurrent upflow depending on the configuration of multiple feed delivery and product 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 the catalyst (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 empty space in the reactor above the catalyst bed. The reactor consisted of a 1/2-in. or 1/4-in. nominal outer diameter metal tube inside an oil jacket constructed of 1-in. tubing. An oil bath with circulation pump circulated thermostatically controlled oil through the jacket to ensure a uniform reactor wall temperature. A positive displacement pump delivered the liquid feed mixture to the reactor from a feed reservoir on a balance. A mass flow controller continually delivered hydrogen. A cooled vapor−liquid separator provided the means to separate the vapor and liquid effluent from the reactor. A control valve maintained the liquid level in the phase separator, and a second control valve on the vapor effluent controlled the system pressure. Online gas chromatography provided separate analyses of the vapor and liquid products, enabling closure of the system mass balance for each atomic species. For both reactors, the feed mixture was representative of the composition expected in the commercial scale reactor. We varied the reactor pressure and temperature systematically to cover the entire range expected in the full-scale reactor, and we varied the liquid and hydrogen feed rates from 0.35 to 7.0 mL/ min and 100 to 400 sccm, respectively.

An important objective for the industrial reaction engineer is to design a commercial scale reactor that achieves the target performance parameters, including production rate and product yield, while minimizing the investment of resources and the time elapsed. The scale-up risks encountered by the engineer vary depending on the nature of the chemistry and the reactor system. In practice, no single work process can be universally applied to all reactor scale-up projects. However, certain classes of problems provide opportunities to apply reaction engineering fundamentals to enable an efficient scale-up program while sufficiently mitigating risk factors associated with the scale-up process. In this paper, we describe a specific program in which a new trickle bed hydrogenation process was scaled directly from the laboratory to the commercial scale reactor without building and operating intermediate scale reactors. This program succeeded by properly accounting for the relevant interactions between transport phenomena and kinetics1 while scaling by a factor of about 3 × 106. Many previous authors have highlighted the importance of properly designing a continuous laboratory scale fixed bed reactor to avoid axial dispersion,2−7 wall effects,7,8 incomplete catalyst wetting,8−10 and nonisothermal bed temperatures.11 In this work, we applied those principles to enable generation of apparent kinetic data, or reaction kinetics that lump the effects of pore diffusion and the intrinsic rates, in an integral reactor with fines.12 During this scale-up program we encountered and resolved three particular problems, resolutions of which are summarized in this paper. While we do not discuss the proprietary chemistry, we offer this case study to illustrate the application of basic reaction engineering principles to the scaleup process and to provide insights into the nature of some of the typical problems encountered in the scale-up of an industrial trickle bed reactor. We also use this case study to defend our assertion that an intermediate-scale pilot plant is not always necessary to achieve successful scale-up and commercialization of new trickle bed reactor technology. © XXXX American Chemical Society

EXPERIMENTAL METHODS

Special Issue: NASCRE 3 Received: February 18, 2013 Revised: March 22, 2013 Accepted: March 27, 2013

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Catalyst Deactivation Experiments. Following an extended (3600 h) lab reactor run, the catalyst was unloaded and separated into two samples based on color: a tan sample from the beginning (top) of the catalyst bed and a black sample from the end (bottom) of the bed. We also collected samples of fresh catalyst and catalyst used in a different experiment for 600 h for analysis by various analytical techniques. The goal of the analyses was to understand the deactivation mechanisms and the extent that each played in the observed catalyst deactivation process. We label the samples as follows: (1) fresh, (2) 600 h, (3) 3600 h bottom, and (4) 3600 h top. Portions of both 3600 h samples were subsequently separately loaded and rerun in the lab reactor to determine the activity of each type of catalyst and compared with fresh catalyst. The fresh sample, the bottom sample, and the top sample gave relative rates of 1.3, 1.0, and 0.5, respectively. X-ray Photoelectron Spectroscopy. Catalyst extrudates were affixed onto a metal plate using carbon tape for each sample and spectra were recorded from three different areas for statistical analysis. Samples were examined in the as-received state, and it was necessary to solvent extract the residual reaction liquids prior to the analysis. Samples were initially examined by low-resolution survey scans followed by high-resolution spectra of specific elements in order to determine the binding energy (chemical state) and concentration of the elements detected in the survey scans. The quantification of the elements was accomplished by using the atomic sensitivity factors for a Kratos model HSi XPS spectrometer, using monochromatic Al−Kα as the X-ray source. 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-resolution spectra. Intensity due to aromatic carbon−carbon bonding was shifted to 284.8 eV. The resulting offset was measured and applied to the other high-resolution spectra for the same point of analysis. Each point of analysis was energy corrected independent of the others. Microprobe Sample Preparation. Cross-section samples of the fresh and used catalysts were prepared by embedding the catalysts in acrylic resin followed by microtoming the samples. These samples were used for both electron microprobe and TEM (transmission electron microscopy) experiments. The procedure for embedding the catalyst samples was as follows: First, several of the fresh extrudates were placed “edgeon” into gelatin capsules, filled with LR White acrylic resin, and placed under house vacuum (∼1 Torr) for 20 min to remove trapped air from the catalyst pores. The capsules were then dried at room temperature for 7 days. The samples were cured at room temperature to eliminate any artificial sintering by heating 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) of the used catalysts were soaked overnight in acetone to extract the excess oligomers that were trapped in the porous catalyst that would inhibit the acrylic resin from curing. The acetonesoaked catalysts were then cured at room temperature and embedded in the LR White acrylic resin. The samples were microtomed (dry) using a diamond knife to prepare crosssection blocks. Electron Microprobe Experiments. The microtomed blocks were carbon coated and examined in a Cameca SX50 electron 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 measured using wavelength dispersive spectrometers (WDS) using the Fe−Kα and Pd−Lα peaks. Element maps were collected at 15 keV, 50 nA, and 50 μs/pixel using WDS. The standard map size was 800 μm × 200 μm, although higher resolution maps of 100 μm × 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 the pellet exterior. This was done by collecting WDS peaks for each element (and background) at discrete positions along a line (∼200 μm long) starting at the rim of the pellet and traversing into the pellet interior. The distance between data points was 10 μm in the bulk and reduced to 2 μm closer to the acrylic resin/catalyst surface. Aberration-Corrected High-Resolution TEM. Aberration corrected (AC) TEM experiments were conducted at Oak Ridge National Laboratory (ORNL) using the JEOL AC2200FS. The JEOL 2200FS-AC was equipped with a CEOS GmbH aberration corrector. The AC-TEM formed extremely small probe sizes of less than 1.2 Å diameter and was optimized for 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 was 2.8 Å/pixel. The probe size for the ORNL AC-TEM instrument was 1.2 Å diameter using a 35-μm aperture (25 pA probe current). Before the scanning-TEM analysis, the area of interest was exposed to an electron dose (“beam shower”) by removing the condenser aperture and defocusing the probe over a ∼100μm area for several minutes. The catalysts for the TEM analysis were prepared using the same procedure as for the microprobe experiments. Thin sections ∼70 nm thick were cut using a Riechart Ultracut microtome at room temperature using a diamond knife, floated onto DI water, and collected onto Cu grids with a lacey carbon support. Note that the samples that were prepared for electron microprobe studies were cut dry with a diamond knife as only a polished block face was needed. Before the TEM analysis, the area of interest was exposed to an electron dose (“beam shower”) by removing the condenser aperture and defocusing the probe over a ∼100-μm area for several minutes. This step minimized carbon contamination from the acrylic resin and did not introduce artifacts into the TEM analysis.

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RESULTS AND DISCUSSION Catalyst Wetting. The earliest experiments in our research used the 1/2-in. reactor tube. In preliminary experiments to generate kinetic data, we found that the reaction order was first order with respect to the organic reactant. We determined this by varying the reactant concentration in the feed by a factor of 2 at a fixed space time (liquid volumetric feed rate), temperature, and pressure; the fractional conversion was the same in both cases. However, when we varied the space time at several different temperatures with a fixed feed composition, the apparent reaction order with respect to the reactant was near 1.5. These contradictory results (reaction order of 1.55 instead of first order) led to the hypothesis that the experiments gave fractional wetting efficiencies that increased with increasing flow rates. In support of this hypothesis, a report in the open literature shows that complete wetting is not guaranteed when the catalyst is diluted with small, inert particles.13 We tested this B

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increase of about 140% (a factor of 2.4), although some of that difference could be attributed to catalyst deactivation since the data point from the 1/2-tube was for a catalyst with significantly more time on stream. We also performed experiments comparing the effect of the flow direction with the 1/4-in. reactor. According to the literature, a lab reactor in which the catalyst bed is diluted with fines should give identical results regardless of the flow direction (cocurrent upflow or cocurrent downflow) if the fines have effectively decoupled the hydrodynamic effects from the reaction kinetics and intraparticle effects.14,15 In other words, since complete wetting of the catalyst surface is ensured during upflow because the continuous fluid phase is the liquid phase, then achieving identical results in both upflow and downflow implies that essentially complete wetting of the catalyst 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. These experiments were conducted after the catalyst had been in operation for about 4000 h. The conversion was 53.6% with downflow and 52.6% with upflow under otherwise identical conditions, supporting our assumption that the catalyst was fully wetted during downflow experiments. With confidence that these data were not compromised by incomplete catalyst wetting, we used this same catalyst load to generate data for development of models of reaction kinetics and deactivation kinetics. Catalyst Productivity. Early catalyst development work used catalyst extrudates with a diameter of 1/8-in. Calculations predicted that pore diffusion significantly limited the rate of reaction, even in smaller catalyst particles. For example, for extrudates 1/16-in. in diameter and 1/8-in. long, based on the observed reaction rate and assuming a first-order reaction, calculations gave an estimated Thiele modulus of 2.5. For a Thiele modulus above 2.0, the effectiveness factor is inversely proportional to the Thiele modulus and is therefore inversely proportional to the characteristic pore length. Consequently, for particles approximately 1/16 in. and larger, we predicted and observed experimentally (Table 1) that the activity per

hypothesis by loading a similar quantity of catalyst and silicon carbide diluent into the 1/4-in. tube and repeating the variable space time experiments using the same range of volumetric flow rates and, thus, higher superficial velocities. In this case, the results fit very nicely to a first order model over the entire range of flow rates (Figure 2). This result supported the hypothesis that the higher apparent reaction order in the experiments in the 1/2-in. tube was the consequence of incomplete catalyst wetting, with the fractional wetting efficiency increasing with increasing flow rate. All subsequent experiments for catalyst testing and kinetic model development used the 1/4-in. tube reactor, where the superficial velocities were high enough to give complete catalyst wetting for range of flow rates employed in 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

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

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

mass of catalyst (g) time on stream (h) liquid feed rate (mL/min) hydrogen feed rate (sccm) first-order rate constant (mL/g-min)

1/8-in. extrudates

1/16-in. extrudates

2.05 6.2 3.04 400 2.46

2.00 8.2 3.27 400 4.85

mass of catalyst was inversely related to the particle diameter. This motivated the use of catalyst particles as small as possible within the limits of reactor pressure drop constraints and the limitations of the catalyst production process. By switching from 1/8-in. to 1/16-in. extrudates, we effectively doubled the reactor productivity, providing a positive boost to the economic attractiveness of the process. Catalyst Deactivation. Our conceptual model of the deactivation of the supported Pd catalyst in this reaction system involved four distinct mechanisms: (1) sintering of the initially highly dispersed Pd particles, (2) iron poisoning of the catalyst, which results in the formation of Fe−Pd alloy particles with

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 different temperatures (T1 and T2) and various total volumetric flow rates (F). Each temperature curve was fit separately assuming a reaction order of 1.0 with respect to the reactant.

(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, and 763 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 C

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lower or perhaps no activity for hydrogenation, (3) Pd loss near the surface, and (4) fouling. We observed these modes of deactivation 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) of the reactor bed were discolored on the outer surfaces (with a distinct “tan” color) compared to the fresh, 600-h, and the 3600-h bottom samples, which were all black. The interior of the discolored “tan” extrudates was still black. XPS revealed a substantial increase in the concentration of surface carbon with time on stream, strong evidence for significant catalyst fouling (Figure 3). A second notable difference identified by the XPS

Figure 4. Line scan of fresh catalyst.

Figure 3. XPS surface concentration (atom %) of C and Fe in fresh and used catalysts. Figure 5. Line scan of 3600-h bottom sample.

analyses was a sharp increase in the concentration of iron in the 3600-h top sample. Iron was not expected or detected in the fresh catalyst material. Among the most informative electron microprobe experiments in this study were quantitative line scans of the fresh and used catalysts in cross-section. The line scans provided a more accurate measurement of the Pd and Fe concentration compared to the element maps since a longer dwell time/ pixel could be used. Line scans were generated by collecting a spectrum at discrete points (for example, every 2 μm) on the cross-sectional sample from the edge of the pellet to ∼200 μm into 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 Fe concentration was below the detection limit (0.05 wt %) in the fresh catalyst. Similar line scans were done for two different 3600-h bottom pellets, and a representative profile is shown in Figure 5. The Pd concentration (∼2.0 wt %) was uniform across most of the cross-section, although a small decrease in the Pd concentration was observed near the pellet edge. The Fe concentration was near 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 concentration of the first pellet was ∼1.7 wt % Pd (as shown in the figure) 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 decreased toward the rim of the pellet but with spikes to ∼1−2 wt % at

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

the areas with high Fe contamination. The line profiles clearly show Pd depletion in the outer ∼30 μm of the pellet to as low as ∼0.3 wt % Pd (data not shown). In the pellet interior, the Fe concentration was low, ∼0.1 wt %. The line profiles clearly showed Pd depletion in the outer 30 μm of the pellet, with Fe preferentially depositing onto Pd rich areas near the surface, giving local Fe concentrations approaching 20 wt %. Fuentes and Figueras previously reported similar iron poisoning of supported Pd catalysts.16 Based on further analyses of the fresh catalyst samples, the Pd-rich areas appear to be the D

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due to sintering, leaching, and fouling in the absence of iron poisoning was determined to be sufficiently slow to provide an economically viable catalyst lifetime.

consequence of the catalyst preparation process rather than a result 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 and distribution. Surprisingly, after 600 h, the average Pd size distribution had decreased (near the pellet surface), and many small Pd particles were observed. The median particle size for the 600-h sample was ∼17 Å, compared to ∼34 Å in the fresh sample (Table 2). The downward trend in Pd particle size

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CONCLUSIONS By increasing the aspect ratio of the catalyst bed in order to increase the liquid superficial velocity, the development team overcame the catalyst wetting problem in the laboratory reactor. Furthermore, decreasing the characteristic diffusion length by choosing smaller catalyst particles increased the productivity to an economically attractive conversion rate. By using raw materials representative of the expected commercial plant feed stream rather than synthetic feeds from the beginning of the experimental program, the team identified catalyst deactivation as an important problem. The team then managed this problem using state-of-the-art analytical techniques to identify several parallel modes of catalyst deactivation, followed by the identification of appropriate steps to minimize the rate of catalyst deactivation to an economically viable rate. Then, using the appropriately sized and loaded lab reactor, the team generated data for development of models of reaction kinetics and deactivation kinetics. The team used these models to design the commercial-scale reactor, successfully scaling the reactor by a factor of about 3 × 106 from these laboratory scale experiments to the full-scale reactor design. Finally, effective technology transfer from the development team to the process design and construction team, coupled with implementation of a strict operating discipline in the plant, resulted in the successful startup and operation of the hydrogenation plant.

Table 2. Palladium Particle Size near the Pellet Surface fresh 600-h 3600-h bottom 3600-h top

median

average (±1 SD)

34 Å 17 Å 17 Å little Pd detected

45 ± 32 Å 23 ± 16 Å 18 ± 9 Å little Pd detected

continued with the 3600-h bottom catalyst, and few Pd particles were observed near the surface of the 3600-h top sample. From these observations, the leaching of Pd into the liquid reaction mixture 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 particle size and distribution as a function of time on stream. The median 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.

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Table 3. Palladium Particle Size in the Pellet Interior (∼100 μm from the Pellet Surface) median fresh 600-h 3600-h bottom 3600-h top

14 21 62 21

Å Å Å Å

average (±1 SD) 18 26 58 30

± ± ± ±

12 18 40 26

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]; phone: 989-636-2165.

Å Å Å Å

Notes

The authors declare no competing financial interest. § Retired.

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ACKNOWLEDGMENTS We thank Timm Richardson and Cliff Todd of Analytical Sciences, Dow Chemical, for their expertise in conducting the electron microprobe and XPS experiments. We also thank Doug Blom and Larry Allard at ORNL for assistance with the AC-TEM experiments (High Temperature Materials Laboratory, Microscopy, Microanalysis, Microstructures Group, Oak Ridge National Laboratory, PO Box 2008, 1 Bethel Valley Road, Oak Ridge, TN 37831-6064). This portion of the research was sponsored by the Asst. Sec. for Energy Efficiency and Renewable Energy, Office of FreedomCAR and Vehicle Technologies, as part of the High Temperature Materials Laboratory User Program, ORNL, managed by UT-Battelle LLC for the U.S. DOE.

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

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These efforts to understand the likely mechanisms for catalyst deactivation led the scale-up team to focus on controlling the one mechanism that could be controlled: iron poisoning. To prevent or minimize this poisoning, we designed the process and operating discipline to prevent introduction of an iron-contaminated feed to the catalyst bed. The deactivation E

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