Hydrodynamics and mass transfer in three-phase composite minichannel fixed-bed reactors

Chemical Engineering Science 94 (2013) 224–236

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Hydrodynamics and mass transfer in three-phase composite minichannel fixed-bed reactors$ S. Haase, M. Weiss, R. Langsch, T. Bauer, R. Lange n Technische Universitaet Dresden, Chair of Chemical Reaction Engineering and Process Plant, 01062 Dresden, Germany

H I G H L I G H T S c c c c

We examine a reactor consisting of minichannels with dumped catalyst particles. We analyse hydrodynamics and mass transfer at a simultaneously occurring reaction. We propose correlations to predict hydrodynamic parameters and mass transfer rates. The reactor exploits dissipated energy for mass transfer processes superiorly.

a r t i c l e i n f o

a b s t r a c t

Article history: Received 29 October 2012 Received in revised form 16 January 2013 Accepted 27 January 2013 Available online 14 February 2013

The improved mass transfer and reduced macroscopic backmixing of segmented flow regimes in mini and microchannels favour its application in three-phase reaction processes. Therefore, industrial available and standardised catalyst particles or pellets may benefit from these microfluidic phenomena if they are packed into inert minichannels. Such packings form the key components of composite minichannel reactors. In order to evaluate this reactor concept, hydrodynamic phenomena, mass transfer, and pressure drop will be examined for a reactor consisting of a ceramic minichannel packing with a hydraulic diameter of 1.0 mm and dumped spherical catalyst particles of 0.8 mm in diameter. The experimental data, achieved in a setup combining hydrodynamic observation and chemical reaction, were used to derive universal applicable correlations to predict mass transfer coefficients and friction factors from Reynolds, Schmidt, and Sherwood numbers. The work concludes with an extensive comparison of composite minichannel reactors with conventional multiphase reactors and developing packed-bed reactors in terms of mass transfer capability, power consumption, and contacting efficiency. At identical power consumption, the investigated composite minichannel reactor offered a remarkably higher overall mass transfer rate for the gaseous compound than conventional trickle-bed, slurry bubble column, or slurry stirred tank reactors. Similar rates or even higher rates were achieved in miniaturised packed-bed reactors with particles less than 1.0 mm in diameter. Consequently, it is expected that structured and miniaturised packed-bed reactors are a promising concept to intensify multiphase reaction processes, e.g. by switching from batch to continuous processing. & 2013 Elsevier Ltd. All rights reserved.

Keywords: Hydrodynamics Mass transfer Minichannel Multiphase reactors Packed bed Pellet-string

1. Introduction A promising concept in catalyst screening and testing as well as in chemical production is the composite minichannel fixed-bed reactor in continuous mode of operation. In this concept, the inert flow channels of a minichannel array are packed with catalytic particles. Such types of integrated packings have been reported by Kapteijn

$ ¨ This paper is dedicated to Professor Werner Weisweiler (former Universitat Karlsruhe, Germany) on the occasion of his 75th birthday on 18th of February 2013. n Corresponding author. Tel.: þ49 351 46335181; fax: þ 49 351 46337757. E-mail address: [email protected] (R. Lange).

0009-2509/$ - see front matter & 2013 Elsevier Ltd. All rights reserved. http://dx.doi.org/10.1016/j.ces.2013.01.050

et al. (2001) as ‘‘structured trickle-bed reactor’’ as well as by Dautzenberg and Mukherjee (2001) as ‘‘composite structured packing’’. If a single flow channel is employed, which was pioneered by Satterfield et al. (1969), the term spiral reactor (Kallinikos and Papayannakos, 2007a) or pellet string reactor (Hipolito et al., 2010) often encounters. Within this work, the term ‘‘composite minichannel reactor’’ will be used to emphasise that the dumped reactor packing contains several straight flow channels with characteristic dimensions in the millimetre scale. The term ‘‘structured trickle-bed’’ is avoided because the flow inside such structures may significantly differs from that in trickle-flow. Composite minichannel reactors have been conventionally used to investigate the performance of commercial catalyst particles.

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The aim of applying such apparatuses is the testing of catalyst particles at industrially relevant liquid hourly space velocities (LHSV) and superficial liquid velocities, which is usually not possible in standard fixed-bed reactors of lab or pilot scale units. For example, Kallinikos and Papayannakos (2007b) hydrodesulphurated heavy gas oil fractions. They compared the sulphur conversion in a diluted bed reactor, in a pilot scale reactor, and in three different composite minichannel reactors at identical LHSVs (0.5–2 1/h) and at identical superficial liquid velocities (o0.001 m/s). The examination demonstrated that the composite reactors of their study are able to reproduce the behaviour of the pilot plant reactor. The same authors hydrogenated benzene (2010) at liquid superficial velocities below 0.0014 m/s to analyse mass transfer limitations. Overall hydrogen mass transfer coefficients were below 0.15 1/s and correlated reasonably with the gas-to-liquid ratio. Hipolito et al. (2010) studied the fast hydrogenation of allene for LHSVs between 67 1/h and 200 1/h. At these conditions, the liquid phase flowed preferentially at the bottom of the horizontally orientated flow channel. Nevertheless, they found that mass transfer was not limiting and the composite minichannel reactor represents an attractive pilot plant reactor. Contrary to the previous researchers, Bauer and Haase (2011) considered the composite minichannel reactor as an alternative apparatus for the production of chemicals which enables industrial available and standardised catalyst particles or pellets to benefit from microfluidic phenomena. In detail, Taylor-flow-similar regimes with a segmented flow characteristic may provoke high interfacial areas, intensified mass transfer rates and may provide a nearly uniform residence time for the feed molecules, which enhances process selectivity. This research concentrated on maximally achievable reaction rates and the comparison with that of monolithic reactors (wall-coated minichannel reactors) and trickle-bed reactors. Two different configurations were tested: (i) inert minichannel substrates packed with catalytic particles and (ii) wall-coated minichannel catalysts packed with catalytic particles. The hydrogenation of alpha-methylstyrene was performed in a multichannel system. The highest space-time yields were observed for liquid superficial velocities of 0.09 m/s and 0.13 m/s, respectively, depending on the configuration. Lower or higher liquid velocities effectuated a decline in the space-time-yield. The superficial gas velocity should be higher than 0.10 m/s. In the experiments of Bauer and Haase (2011), maldistribution of gas and liquid is highly probable because the liquid was distributed with one conventional spray nozzle across the multichannel system and such nozzles will not operate perfectly for a broad range of liquid throughputs. Nevertheless, the study clearly illustrates that composite minichannel reactors outperform conventional trickle-bed reactors with respect to spacetime yield. It can be concluded that the full potential of composite minichannel reactors is still not well defined. Reasons include studies without optimal hydrodynamic regimes or with maldistribution of the liquid phase – both could have lowered measured reaction rates. In order to analyse the maximum capability, detailed investigations on the complex hydrodynamics and on mass transfer under reacting conditions will be performed. In detail, flow regimes, specific interfacial areas, mass transfer rates and pressure drop will be examined. The generated data will be generalised in form of empirical correlations to estimate mass transfer rates and power consumption in any gas–liquid–solid reaction process. Additionally, the obtained data will be employed to compare the composite minichannel reactor with other multiphase reactor concepts in order to identify beneficial areas of application.

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2. Experimental 2.1. Reactor concept The composite minichannel reactor was created by inserting spherical catalyst particles into a single representative flow channel of a empty cordierite monolith segment with a cell density of 400 cpsi. The minichannel was provided by Corning Incorporated and had a hydraulic channel diameter of 1.0 mm. The spherical and catalytically active particles with a mean diameter of 0.8 mm were allocated by Degussa AG (now Evonic Industries AG) and contained palladium on an alumina support. The channel-to-particle diameter ratio was 1.25. Depending on the filling method, different packing configurations are possible. The two extreme cases are presented in Fig. 1: a single string packing or the densest packing. In vertical arrangement, a single string packing will not occur in the entire minichannel. Packing experiments in a glass capillary reveal that the spheres are also not arranged in the densest form. In some packings for instance, two adjoining spheres are arranged in a line with contact to two identical channel walls whereas the next sphere has contact to the opposing two channel walls. As a result, the porosity of the achieved packing will establish between both maxima, i.e. between the densest and the single string packing. However, the difference between both is very small. The detailed analysis of the catalytic spheres shows that they have a mesoporous texture with a sharp pore size distribution and a median pore diameter of 8.4 nm. Palladium is incorporated in egg-shell configuration and the shell has a thickness of at least 120 mm.

2.2. Experimental setup The main component of the experimental setup was the modular flow reactor consisting of: (i) the gas-liquid contacting module, (ii) the inlet observation module, (iii) the reaction module,

Fig. 1. Configurations and dimensions of the composite minichannel reactor packing (a) densest packing (eBed ¼0.500; without consideration of the channel wall), (b) single string packing (eBed ¼0.504; without consideration of the channel wall).

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and (iv) the outlet observation module. Inside the reaction module, the composite minichannel reactor packing was fixated. The gas was supplied from a compressed gas cylinder (Air Liquide, purity499.999%), and the flow rate was adjusted by three different mass flow controllers. The reaction mixture was supplied from a reservoir, pumped by a pulsation-free HPLC pump, and preheated to the reaction temperature by an electrically heated micro heat exchanger. Temperature loss, especially at low liquid flow rates, was minimised by electrically heated connection pipes. Inside the observation module, the gas was fed into the co-flowing liquid by using capillary injectors. The occurring two-phase flow was recorded by a high-speed video camera system before entering and after passing the reaction module. The temperature along the reaction section was adjusted by electrical heating. Two separation vessels were installed to avoid any liquid droplets in the gas outlet. The liquid samples were collected directly at the outlet of the reactor and analysed by a refractometer or by a GC/MS system. The pressure drop across the bed was detected by using two pressure indicators. A more detailed description of the applied setup can be found elsewhere (Haase and Bauer, 2011). 2.3. Experimental procedure In the experiments, the palladium-catalysed hydrogenation of alpha-methylstyrene to cumene was investigated because the intrinsic reaction rates are relatively high enabling to study mass transfer of hydrogen inside the reactor. Before the start of an experimental run, the minichannel reactor, micro heat exchanger, and pipes of the liquid supply were preheated to the desired reaction temperature; in all experiments 343 K. During this process, pure hydrogen was flowing into the system and a system pressure of 1.0 MPa was set. When the desired temperature was reached, the reaction mixture was pumped into the reactor. The reaction mixture consisting of alpha-methylstyrene in cumene (20% w/w; both Sigma Aldrich Co. LLC., purity498%) was treated with activated alumina and molecular sieves to remove traces of water and 4-tert-butylcatechol, which can cause catalyst deactivation (Kreutzer et al., 2001; Meille and de Bellefon, 2004). The desired flow rates for the gas and liquid phase were adjusted, and the system was allowed to reach a steady state, controlled by analysing liquid samples of the reaction mixture at the reactor outlet every 10 min. The experimental plan covered four different two-phase velocities in the range from 0.01 m/s to 0.20 m/s with gas-holdups at the inlet of 0.25, 0.50, and 0.75. The adjusted superficial two-phase velocities correspond to interstitial velocities, i.e. only the void space inside the composite packing is considered, between 0.015 m/s and 0.30 m/s, respectively. For most conditions, the experiments were repeated several times with an experimental error of less than 5%. In the following, only the mean values for each condition will be presented. After a slight reduction in the catalyst activity within the first few hours on stream, no further catalyst deactivation was detected in the experiments.

3. Results and discussion 3.1. Reaction kinetics To examine the reaction kinetics without external mass transfer limitations, concentration vs. reaction time plots were investigated experimentally in a stirred tank reactor system of Parr Instrument Company. The autoclave was a 300 ml doublejacket steel vessel. To provide maximum turbulence and gas entrainment, a stirrer combining a pitched-blade turbine and a

gas entrainment impeller was used and effectively prevented an accumulation of the uncrushed catalyst particles (mean diameter of 0.8 mm) at the bottom of the vessel, as separate experiments in a glass vessel validated. The reaction temperature was varied between 333 K and 353 K at a fixed system pressure of 1.0 MPa. In pre-experiments, the palladium mass was varied between 2 mg and 14 mg. The linear dependency between the mass of active component and the observed reaction rate excluded external mass transfer limitations inside the vessel. From the data, an activation energy of 27.7 kJ/mol was computed that is remarkably smaller than the activation energy reported for the kinetically controlled region which is about 38 kJ/mol as reported by Turek and Lange (1981) or by Meille et al. (2002). Such a characteristic indicates the presence of internal mass transfer limitations that result from the very fast reaction and the comparably large thickness of the active layer (at least 120 mm). For a hydrogen overpressure of 1.0 MPa and a reaction temperature of 343 K, a reaction rate of 2.2 mol/(kgPd s) was detected. Analysing the change in alpha-methylstyrene concentration with time, the experiments also validate that the hydrogenation is zeroth order with respect to alpha-methylstyrene if its concentration is above 10% w/w. This behaviour corresponds with the literature (Kreutzer et al., 2001; White et al., 1974) and may be explained by a strong adsorption of alpha-methylstyrene on the active sites. 3.2. Hydrodynamic phenomena A fundamental understanding of the appearing flow phenomena inside the sphere packing is required to characterise mass transfer and to develop an adequate mass transfer model. It is of general importance how the flow is segmented, how the bubble will curl around the spheres, as well as if the flow characteristic will change with reactor length. As discussed in detail in another paper (Bauer and Haase, 2011) there are flow conditions wherein the spherical particles are enclosed either completely or partially by the gas bubble. The first phenomenon occurred at low superficial two-phase velocities, as illustrated in the left two pictures of Fig. 2. At higher velocities, the gas bubble shows the tendency to curl only around a part of the particles. However, the flow around the spheres also depends on the local packing, i.e. the alignment between the neighbouring particles. For the highest packing density and displaced particles, the gas bubble penetrated into the entire open frontal area (see right picture for uTP,s ¼0.13 m/s in Fig. 2). At identical process conditions, the bubble flowed only around a part of the spheres if the particles were arranged in line with each other (see middle picture for uTP,s ¼0.13 m/s in Fig. 2). A transformation of the flow inside the composite reactor packing was detected near the entrance as well as inside. Characteristic photographs of the observed phenomena are depicted in Fig. 3. Gas bubble segregation inside the packing was obtained at superficial two-phase velocities above 0.1 m/s and gas holdups of 0.25. At these conditions, the generated gas bubble in the empty part of the channel is marginally longer than the channel in diameter. When these small bubbles enter the packing, they form a continuous segment in the first part of the reactor, but break up rapidly due to high shear stresses affected by the acceleration and deceleration of the two-phase flow inside the packing. Gas bubble coalescence was detected above the sphere packing as well as inside the packing. The first phenomenon occurred only when the superficial two-phase velocity was between 0.05 m/s and 0.13 m/s at a gas holdup of 0.75. The merging of two bubbles was promoted by short liquid slugs which are smaller than 0.5 mm for the above mentioned conditions. However, a break-up of the gas bubbles inside the sphere packing

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Fig. 2. Photographs of the gas–liquid flow around individual spheres of the composite minichannel reactor for different sphere configurations and two-phase velocities.

Fig. 3. Photographs of the observed flow phenomena above the packing or inside the packing of the composite minichannel reactor for different operating conditions.

occurred very regularly within the first particles. This leads to the macroscopic effect that the gas bubble length in the packing was not greatly influenced by gas bubble coalescence at the inlet. The coalescence inside the packing was only observed for a twophase velocity of 0.01 m/s and a gas holdup of 0.50 and 0.75, i.e. the combination of low inertial forces and short liquid slug lengths will contribute to the generation of this phenomenon. On the whole, a strong modification of the gas bubble and liquid slugs due to the sphere packing was sporadically observed. 3.3. Hydrodynamic parameters In composite minichannel reactors, the two-phase flow is generally generated in an empty channel before entering the composite packing. Therefore, the transferability of hydrodynamics in empty and composite minichannels is examined. Within the evaluation, the open frontal area in the flow channel has to be regarded. In the

empty minichannel, a constant clearance for the flow is provided. This is completely different for the composite minichannel. In this case, the conduit is determined by the difference in the area of the channel and of the sphere, which is a function of the axial position. By passing one sphere of the ideal pellet-string, the open frontal area first reduces from virtually 1.0 mm2 to around 0.5 mm2 before rising to the initial value. By using the different void volumes of the empty minichannel and the composite minichannel, an enlargement factor of 1.50 is computed. This factor will be used in the following. It is important to keep in mind that the local flow velocity inside the packing will constantly change according to the continuity equation for plug flow with a varying open frontal area. To calculate the mean bubble length and the mean bubble velocity, both values have been determined for the flow in the empty channel and been multiplied by a factor of 1.50 which represents the ratio between the mean void volumes. In Figs. 4 and 5, these values are opposed to those which have been experimentally measured.

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Fig. 4. Parity plots for the experimentally measured gas bubble lengths in the composite minichannel and the prediction by using the gas bubble lengths in the empty channel and the void fraction factor.

Fig. 6. Parity plots for the experimentally measured liquid slug lengths in the composite minichannel and the prediction by using the gas bubble lengths in the composite channel and the gas holdup.

Fig. 5. Parity plots for the experimentally measured gas bubble velocities in the composite minichannel and the prediction by using the gas bubble velocity in the empty channel and the void fraction factor.

Fig. 4 clarifies that the bubble lengths in the experiments were slightly higher than the predicted ones. This trend can be explained by the fact that the gas bubble does not always fill the entire cross section and, consequently, enlarges in length as shown in Fig. 2. Fig. 5 illustrates that the bubble moves remarkably faster than predicted if the bubble velocity inside the packing is higher than 0.15 m/s. An increased deviation between bubble and two-phase velocity for rising Capillary numbers has been previously reported by Thulasidas et al. (1995) and Liu et al. (2005) who analysed the flow in empty channels. In agreement with their expectations, the determined bubble velocities in the work at hand are almost independent of the gas holdup. Additionally, the length of the liquid slug was experimentally and theoretically examined. The liquid slug length was calculated using the previously determined gas bubble length and the dynamic gas holdup. These theoretical values are depicted against the experimentally measured values in Fig. 6. Only minor deviations have been detected within the studied conditions. Summarising, knowledge about the hydrodynamic parameters of gas bubble and liquid slug lengths in the empty flow channel can be used to predict both parameters in the investigated composite minichannel. In mass transfer modelling, macroscopic parameters such as gas holdup should be preferred to precise values such as gas bubble or liquid slug lengths because of the observed flow modifications inside the channel. From the above mentioned parameters the interfacial areas can be computed as will be presented in Section 3.5.

Fig. 7. Alpha-methylstyrene conversion in the flow reactor packed with the composite minichannel catalyst for different two-phase velocities and gas holdups (p ¼ 1 MPa, TR ¼ 343 K).

3.4. Alpha-methylstyrene conversion and reaction rates The observed alpha-methylstyrene conversion for different flow conditions inside the composite minichannel reactor is depicted in Fig. 7. It is obvious that the conversion at the reactor outlet increases by decreasing the superficial velocity or by elevating the gas holdup in the channel. A smaller superficial velocity will cause a higher residence time in the reaction channel or contact time with the catalyst and will provoke a higher conversion per reactor pass. A larger gas holdup at one specific two-phase velocity is generated by feeding more gas and less liquid into the channel. A lower liquid throughput will generally affect a higher alpha-methylstyrene conversion if the reaction rate is not a strong function of the gas holdup. The unexpected decrease in conversion at a two-phase velocity of 0.01 m/s and a gas holdup of 0.75 may be affected by observed flow pulsations inside the system. At such liquid flow rates, it is highly expected that droplets will be formed inside the liquid feeding which provoke the mentioned flow pulsations. They resulted in inhomogeneous residence times and flow conditions, i.e. bubble coalescence inside the packing. Both effects may reduce the overall mass transfer rate and, as a result, educt conversion. The determined reaction rates per mass of palladium are opposed to the flow conditions in Fig. 8. Analysing the observed reaction rate, it can be concluded that large reaction rates are

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in which csat H2 is the saturation concentration of hydrogen in the liquid phase. The overall mass transfer resistance consists of diffusion of the species inside the porous catalyst specified by the catalyst effectiveness factor ZCat and an external mass transfer resistance between gas and solid phase. In this approach, the gas adsorption in the liquid and the diffusion of dissolved gas to the catalyst surface are lumped into one overall mass transfer resistance (kGLSaGLS) and mass transfer resistances within the gas phase are neglected because pure hydrogen was employed. A correction factor was introduced to transform the coefficients related to the catalyst volume or to the reactor volume. It has to be mentioned that the reactor volume corresponds to the void volume of the channel, i.e. the channel wall was not respected. The intrinsic reaction rate constant kintr,v per catalyst volume was calculated from the experimental data on reaction rates in the stirred tank respecting the Thiele modulus. It is important to note that a temperature correction was introduced because calculations show that adiabatic temperature

enlarges between 5 K and 20 K depending on the liquid flowrate. Unfortunately, such a temperature rise could not be validated experimentally because commercially available temperature sensors had to be assembled with metallic parts of the setup to ensure pressure resistance, which caused a high heat loss to the surrounding because total enthalpy flow is very small due to the low flow rates. As a consequence, the system could not be adequately characterised by a heat balance based on experimentally measured temperatures. To include temperature effects, an reduced reaction rate was computed which represents the expected rate under isothermal conditions. In the procedure, the adiabatic temperature rise was calculated by the observed reaction rate in the flow reactor and the reaction enthalpy (  109 kJ/mol). The kinetic rate law assuming isothermal conditions was related to the kinetic rate law assuming adiabatic conditions which resulted in a factor used to calculate the reduced reaction rate. In the rate law, the arithmetic average of the activation energy for the intrinsic kinetics and that for the diffusion process was employed because internal transport limitations dominated for the studied catalysts. By using the reduced reaction rates, the overall mass transfer coefficients reduced by 5%, 10% and 25% for gas holdups of 0.25, 0.50, and 0.75, respectively. The authors are aware that the system may not behave strictly adiabatically due to the high ratio of heat exchange surface to reactor volume and the resulting large heat losses. However, the smaller mass transfer coefficients (computed from the reduced reaction rates assuming isothermal conditions) will be used in the following in order to avoid the presentation of too optimistic values for the novel reactor concept. The computed overall mass transfer coefficients for hydrogen are depicted in Fig. 9. In general, the overall mass transfer coefficient is elevated if the two-phase velocity or the gas holdup is enlarged. The shift with two-phase velocity agrees with the general behaviour of packed-bed reactors in which mass transfer intensifies with the volumetric throughput due to enhanced vortices or turbulences inside the system at higher flow velocities. The absolute values are about one order of magnitude higher than those reported by Kallinikos and Papayannakos (2010) which is addressed to the different fluid dynamic conditions and flow regimes in their spiral reactor. For interstitial two-phase velocities above 0.08 m/s in the packing, the strong impact of the gas holdup on the overall transfer coefficients may result from the specific interfacial areas for gas–solid and liquid–solid mass transfer (see Fig. 10). Details on the method to compute the interfacial areas are given in the

Fig. 8. Determined reaction rate per palladium mass in the flow reactor packed with the composite minichannel catalyst for different two-phase velocities and gas holdups (p¼ 1 MPa, TR ¼ 343 K).

Fig. 9. Computed overall mass transfer coefficient for hydrogen in the flow reactor packed with the composite minichannel catalyst for different two-phase velocities and gas holdups.

achieved if the two-phase velocity and the gas holdup are high. Fig. 8 also validates that there is no linear dependency between liquid throughput and reaction rate. To illustrate, let us consider a two-phase velocity of 0.19 m/s inside the packing. By elevating the gas holdup from 0.25 to 0.50 or 0.75, the liquid throughput is reduced to 67% or 33%, respectively. However, the corresponding reaction rates decline only to 73% or 60%, respectively. This validates the above postulated assumption that the rate is not a strong function of the gas holdup. The impact of the gas holdup even reduces for smaller two-phase velocities. For example, only an insignificant deviation between the reaction rates for different gas holdups was found if the two-phase velocity inside the packing was below 0.08 m/s. The maximum detected reaction rate was about 1.5 mol/(kgPd s), which is about 30% smaller than the reaction rate without external mass transfer limitations identified in the stirred tank reactor. 3.5. Mass transfer rates As discussed before, the chemical reaction in the flow reactor with the composite minichannel packing is limited by external and internal mass transfer. According to Losey et al. (2001), the apparent reaction rate for the first order reaction rapp,v with respect to the catalyst volume corresponds with: r app,v ¼

csat H2 1=kGLS aGLS þ 1=ZCat kint,v

ð1Þ

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For simplification, the GS and LS specific interfacial areas aGS and aLS are related to the mean gas holdup bG,m which represents an arithmetic average between the holdup at the inlet and at the outlet of the reactor: aGS ¼ f a,CMC bG,m

pdS

aLS ¼ f a,CMC ð1bG,m Þ

Fig. 10. Measured mean interfacial areas between liquid–solid and gas–solid phases in the flow reactor packed with the composite minichannel catalyst for different two-phase velocities and gas holdups.

next section. For a gas holdup of 0.50, both interfacial areas are almost equal, whereas for gas holdups of 0.25 and 0.75, the gas– solid and liquid–solid specific areas differ by a factor of 3 to 4. The superior overall mass transfer rates for hydrogen at a gas holdup of 0.75, which exhibits high gas-solid interfacial areas, reveal that the transfer of hydrogen molecules is more intense from the gas bubble to the solid catalyst than the transfer of dissolved hydrogen molecules from the liquid slug to the catalyst surface. It has to be noted that the hydrogen transport into the liquid slug, i.e. the gas-liquid mass transfer, may also limit the liquid-solid transfer because the equivalent transfer resistance for both consecutive steps is represented by the resistance in series model. In other words, a small gas-liquid mass transfer coefficient will cause a reduced hydrogen concentration inside the liquid slug which decreases the total amount of transported hydrogen molecules at the liquid–solid interphase although the liquid–solid mass transfer resistant remains high.

3.6. Modelling mass transfer In an effort to obtain equations to predict the mass transfer rates in a composite minichannel reactor, the occurring mass transfer steps have to be respected. According to the mass transfer in unpacked minichannel reactors (Kreutzer et al., 2001), it is assumed that gas molecules are transported through the gas–solid (GS), the gas-liquid (GL) and the liquid–solid (LS) interphases. The latter two can be described by the model of resistances in series in parallel to the gas–solid transfer resistance. As a result, the overall volumetric mass transfer coefficient of the gaseous compound kGLSaGLS was expressed to:  1 1 1 kGLS aGLS ¼ kGS aGS þ þ ð2Þ kGL aGL kLS aLS in which kGS, kLS, and aGS, aLS are the mass transfer coefficients and the interfacial areas per empty channel volume for GS and LS transfer, and kGLaGL characterises the volumetric transfer coefficient between the gas and liquid phase. In this approach, the overall volumetric GL transfer coefficient was not separated into the mass transfer coefficient and the volume specific interfacial area because the latter could not be detected precisely in the performed experiments. Additionally, these areas will vary to a great extent with the axial position because less interfacial area is available in the middle of a sphere than at the top or bottom of a sphere.

ð3Þ

2

dh

pdS 2

dh

ð4Þ

in which dS and dh indicate the diameter of the catalyst sphere and the channel. The type of packing configuration is characterised by the form factor fa,CMC. This factor specifies how much solid surface is provided in the desired packing in comparison with the single pellet string packing. In consequence, the form factor is equal to 1.00 for the single pellet string, i.e. the lowest packing density, whereas it will be 1.02 for the densest packing. The packing density of the investigated bed will be between both values but due to the small difference, the form factor was set to 1.00. The mean gas holdup inside the reactor bG,m represents the mean gas holdups of all unit cells inside the composite flow channel. The gas holdup of one unit cell bG,UC can be computed from the gas bubble and liquid slug lengths in the empty minichannel:   LB,EMC þ p6 1 dh bG,UC ¼ ð5Þ LB,EMC þLS,EMC According to the generalised description of the transfer of matter in packed-beds, the GS and LS mass transfer coefficients were correlated to the corresponding Sherwood, Schmidt, and Reynolds numbers using the particle diameter as the characteristic length. Respecting only the GS and LS transfer of hydrogen, the approach included the following equations: kGS,H2 ¼

ShG,H2 UDH2,G G3 with ShG,H2 ¼ cG1 þ cG2 RecS,G USc0:33 G,H2 ds

ð6Þ

kLS,H2 ¼

ShL,H2 UDH2,L L3 with ShL,H2 ¼ cL1 þ cL2 RecS,L USc0:33 L,H2 ds

ð7Þ

in which ShG,H2, ShL,H2, ReS,G, ReS,L, ScG,H2, and ScL,H2 represent the dimensionless Sherwood, Reynolds, and Schmidt numbers of hydrogen in the gas and liquid phase. The empiric coefficients cG1, cG2, cG3, cL1, cL2 , and cL3 were estimated by a Matlabs script using the method of least squares. In a first approach, the minimal Sherwood numbers cG1 and cL1 were set to zero and both exponents of the Reynolds numbers cG3 and cL3 were set at identical values. The parameter estimation gave an exponent of 0.90. This is slightly higher than values reported in the literature for packed-bed reactors, which are usually between 0.5 and 0.8 (Gnielinski, 1978; Hipolito et al., 2010; Romkes et al., 2003; Wakao and Funazkri, 1978). The specific characteristics of the Taylor-flow-similar regimes with internal recirculation patterns may cause the observed phenomenon. All experimental data for mass transfer of hydrogen can be described adequately (R2 ¼0.97) by the following correlations: ShG,H2 ¼ 3:3U103 ShL,H2 ¼ 0:37

0:33 Re0:90 S,G USc G,H2

0:33 Re0:90 S,L UScL,H2

kGL,H2 aGL ¼ 1:0 1=s

ð8Þ ð9Þ ð10Þ

More details on the computation of the coefficient of determination are given in the nomenclature. The relatively large difference in the pre-exponential factors between the GS and LS transport, indicated in ShG,H2 and ShL,H2, may arise from the fact that even if the catalyst particle is surrounded

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by the gas phase, a thin liquid film will wet the porous surface, which hinders the exchange of matter. In the packed channel, a volumetric GL mass transfer coefficient of 1 1/s was estimated. This value agrees with coefficients observed in empty channels which are in the range of 0.5 1/s to 2 1/s within the studied conditions (Heiszwolf et al., 2001; van Baten and Krishna, 2004). From the physics, it is expected that the gas–liquid coefficient enhances if the flow velocity enlarges, which is not part of the developed model. To analyse this transfer route in more detail, separate experiments, e.g. the absorption of oxygen in water, should be performed to correlate this coefficient on the relevant parameters. Nevertheless, the derived mass transfer model reasonably described all experimental data. In extended parameter estimations, more variables were added to the mass transfer model, i.e. the minimal Sherwood numbers (cG1 and cL2 Þ were implemented or the exponents of the Reynolds numbers for the gas and liquid phase (cG3 and cL3) were not set identically. However, these model variations did not improve its prediction accuracy. The comparison between the experimental data and the predicted values using Eqs. (8)–(10) is shown in Fig. 11. Larger discrepancies emerge for a two-phase velocity of about 0.001 m/s. At this low flow velocity, the alpha-methylstyrene conversion is high (see Fig. 7) and could cause a divergence in the mean reaction temperature which is not precisely described by the applied temperature correction. A detailed inspection of the mass transfer shows that the ratio between the volumetric mass transfer coefficient from the gas bubble to the solid catalyst kGSaGS and the volumetric mass transfer coefficient from the gas bubble via the liquid slug to the solid catalyst kGL þ LS aGL þ LS is about 1, 2.4 and 4.5 for gas holdups of 0.25, 0.50, and 0.75, respectively. The obtained volumetric mass transfer coefficients at the gas–solid interphase are depicted in Fig. 12. Ergo, composite minichannel reactors should be operated at high gas holdups and two-phase velocities to achieve intense mass transfer for the gaseous component. However, the corresponding immense throughputs will limit the conversion per pass due to the reduced residence time and cause a high pressure drop. 3.7. Pressure drop The experimentally measured pressure loss across the complete reactor was corrected by the pressure loss of the gas–liquid feeding and the reactor outlet to give the total pressure loss across the composite minichannel packing ðDp=LÞtot . The frictional pressure loss was calculated by subtraction of the static pressure head (1 bG,m)rLLRg from the total pressure loss. The contribution

231

Fig. 12. Computed gas–solid mass transfer coefficient for hydrogen in the flow reactor packed with the composite minichannel catalyst for different two-phase velocities and gas holdups.

of the gas phase was assumed to be negligible because of its small density. The friction factor fF,CMC was calculated as following (Kreutzer, 2003):   Dp   LR tot ð1bG,m ÞrL LR g 4 f F,CMC ¼ Þ ð11Þ dh,CMC ð1bG,m Þð12 rL u2TP,is in which uTP,is is the mean interstitial two-phase velocity inside the packing and dh,CMC represents the mean hydraulic diameter of the composite minichannel which is about 0.47 for the studied configuration. The hydraulic diameter of the composite flow channel dh,CMC was computed with the generic equation: dh,CMC ¼

4 ACMC Pm CMC

ð12Þ

with ACMC as mean open frontal area of the channel and Pm CMC as the mean wetted perimeter of the channel. The friction factor shows inversely proportional dependence on the Reynolds number in the packing ReTP,is , which is based on the interstitial two-phase velocity: f F,CMC ¼

133 ReTP,is

ð13Þ

The coefficient in Eq. (13) is about two times higher than the Darcy friction factor for laminar flow, which is 64 (see Fig. 13). Such behaviour may be explained by the periodic flow acceleration and deceleration during passing the composite packing. For a detailed discussion of energy dissipation due to pressure drop, the reader is referred to the next subsection. Two-phase pressure drop in a composite minichannel reactor was previously reported by Hipolito et al. (2010). In their investigations, spherical particles with diameters of 2.0 mm, 3.0 mm, and 4.0 mm have been applied. The particle-to-channel-diameter ratio was varied between 0.5 and 1.0, respectively. A modified Ergun equation was proposed. If using their correlation, the predicted pressure drop differs significantly from the experimental data of this work. The deviation may result from a different hydrodynamic regime in the reactor of Hipolito et al. (2010). 3.8. Comparison of composite minichannel reactors with conventional multiphase reactors

Fig. 11. Parity plot for the overall volumetric mass transfer coefficients of hydrogen in the composite minichannel reactor.

In a real production process, the decision which system performs best is complex because process selectivity, productivity,

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The specific energy dissipation (P/V)dyn which is required to force gas and liquid through the packing can be written as:   Dpdyn,R _ P ¼ ðV L þ V_ G Þ V dyn VR

Fig. 13. Friction factor as a function of the Reynolds number in the composite minichannel reactor.

and process safety have to be related to investment and operational costs as well as to production downtimes and technical risks of implementation. Within this work, the comparison is made quantitatively for the power requirement and the mass transfer capability. Other important aspects will be discussed qualitatively. With respect to the chemistry, two general cases of multiphase reaction processes can be distinguished. Firstly, the reaction is slow and the overall rate is limited by the activity and the mass of the catalyst. Therefore, the reactor has to provide enough catalyst mass and large residence times. In the composite minichannel reactor configurations, the mass of catalyst can be fitted to the requirements of the chemistry by modifying the thickness of the active layer in the particle as well as by choosing another mean channel or particle diameter. However, the overall mass of catalyst will be lower than in a conventionally packed-bed reactor because packing density is smaller in this case. Due to the expected almost ideal plug flow behaviour, the mean residence time in the composite minichannel reactor can be adjusted by varying the flow rates and the reactor length. Secondly, the reaction is fast and the maximum productivity of the apparatus is related to its mass transfer capability, whereas the performance is usually imposed by the mass transfer of the gaseous component due to its low solubility in the liquid phase. Power is a useful tool in intensifying the transfer of matter. To keep operation costs low, the reaction system, however, should exploit the spent energy as efficient as possible. Therefore, the mass transfer capability of the reactor is related to the energy dissipation inside the system. Within this evaluation, the experimentally obtained mass transfer data which have been given before will be used. The specific energy dissipation is computed from the power input per unit reactor volume, which depends on static as well as dynamic pressure losses across the column. The static head (P/V)stat is dominated by the liquid feed, which gives the following expression (Kreutzer, 2003):   DpL,feed V_ L ðRL g LR ÞðAq uL,s Þ P ¼ ¼ ¼ RL uL,s g Aq LR V stat VR

ð14Þ

with rL as liquid density, g as gravitational constant, and uL,s as superficial liquid velocity. It has to be mentioned that the feeding tube is expected to be large enough to neglect frictional pressure losses in the piping.

ð15Þ

in which V_ G and V_ L represent the volumetric gas and liquid flow rates, VR is the reactor volume, and Dpdyn,R is the dynamic pressure loss in the column. It is noteworthy to mention that the reactor volume represents the volume of the flow channels without considering the volume of the channel walls. This approach was chosen because the wall thickness can vary for different minichannel structures. If the wall thickness amounts to one-tenth of the diameter, the volumetric power consumption of the complete packing will be about 18% less. Of course, the volumetric mass transfer coefficient will also be reduced by the same factor. A conventional trickle-bed reactor was integrated in the comparison as a typical industrial fixed-bed reactor. Losey et al. (2001) reported data for such a reactor which contains catalyst pellets with diameters between 3 mm and 5 mm. For a fair comparison, the listed mass transfer coefficients in cyclohexene were corrected by a factor employing the unequal diffusion coefficients of hydrogen in alpha-methylstyrene and in cyclohexene. The computed values are higher than the data of Reynders and Nicol (2011) who hydrogenated octene in a conventional trickle-bed although octene has similar properties as alphamethylstyrene. As a result, the mass transfer coefficients for a conventional trickle-bed, which are presented in the following, may be too optimistic. With respect to slurry reactors, a stirred tank and a bubble column reactor will be included. To ensure a comparison on the production scale, experimental results were selected which are obtained in systems with liquid volumes in the range of 1 m3. ¨ Schluter and Deckwer (1992) investigated gas–liquid mass transfer as well as energy dissipation in stirred vessels with volumes of up to 3 m3. Linek et al. (1993) correlated oxygen–water mass transfer with energy dissipation for different superficial gas velocities in a bubble column with volumes of up to 1 m3. The volumetric mass transfer coefficients of hydrogen in alphamethylstyrene have been calculated from the reported volumetric mass transfer coefficients of oxygen in water and the different diffusivities. To make a fair comparison between all reactor types, the LS mass transfer in the slurry reactors was included by a resistance-in-series model. Tschentscher et al. (2010) suggested volumetric LS mass transfer coefficients for hydrogen in the organic phase of a stirred tank between 0.3 1/s and 1.5 1/s. The maximum value was used for the estimation of the overall transfer coefficient. In a bubble column reactor, the LS mass transfer coefficient was computed according to the correlation of Dutta and Pangarkar (1996). To estimate the interfacial area, Baerns et al. (2006): p. 192 specified a typical catalyst volume fraction of 0.05 and a particle size of 200 mm. As the maximal value, a volumetric liquid–solid transfer coefficient of 0.56 1/s was obtained and used in the further calculations. In Fig. 14, the overall kGLS aGLS values for hydrogen transfer are plotted versus the volume specific power consumption. In all systems, the mass transfer rates increase if more energy is dissipated inside the apparatus. As a matter of fact, the additionally spent energy will improve internal circulation and mixing or create larger interfacial areas, e.g. due to the generation of smaller gas bubbles. In trickle-bed reactors the overall mass transfer coefficient correlates with the specific energy consumption to the power of about 0.70, which was estimated from the slope of the corresponding lines in Fig. 14, whereas it is smaller for the stirred tank reactor (about 0.40) and the bubble column reactor

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233

minichannel reactors is contra-productive because the maximum mass transfer capability of these packings is not used efficiently, but they will consume more energy per unit reactor volume. Nevertheless, the segmented flow characteristic in the composite minichannel reactor will provoke a sharper residence time distribution in comparison with a conventional trickle-bed. This may lead to improvements of process selectivity if composite minichannel reactors instead of trickle-bed reactors are employed although its mass transfer capability is not fully utilised.

3.9. Comparison of composite minichannel reactors with miniaturised random packed-bed reactors

Fig. 14. Volumetric mass transfer coefficients for hydrogen versus power input per unit reactor volume for the composite minichannel reactor of this work, for a trickle-bed reactor (Losey et al., 2001) and for different slurry reactors [(Linek ¨ et al., 1993; Schluter and Deckwer, 1992), corrected with LS transfer by (Baerns et al., 2006: p.192; Dutta and Pangarkar, 1996; Tschentscher et al., 2010)].

(about 0.50). In the investigated composite minichannel reactor, a remarkably higher amount of gas molecules can be transported to the solid catalyst per time in a specific reactor volume than in the mentioned multiphase reactors at similar power consumption levels. Overall mass transfer coefficients above 3 1/s were generated. The exponent to correlate power consumption and mass transfer rate is up to 0.90 for large gas holdups. The efficiency of power consumption depends on the superficial two-phase velocity and the gas holdup. By increasing the flow velocity, the enlarged energy dissipation inside the system will cause a rise in the volumetric mass transfer coefficient. Considering a constant two-phase velocity, the gas contacting is more efficient for a high gas holdup because the pressure drop is drastically reduced whereas the mass transfer rate does not change to the same extent. In consequence, the exponents of energy dissipation drops to 0.66 at low gas holdups or even below in the studied composite minichannel reactor. The high mass transfer capability of the composite minichannel reactor packing favours the application in processes which have fast reaction kinetics or which suffer from a low concentration of the gaseous component at the surface of the catalyst, e.g. by elevated catalyst deactivation or by reduced product selectivity. Ergo, the application of such structured catalyst packings may intensify multiphase reaction processes, i.e. lead to a drastic reduction of the reactor size or an increased productivity at identical reactor size. On the other hand, the improved mass transfer capability is paid by large energy dissipation inside the apparatus. Fig. 14 also illustrates that there is no need to use composite minichannel reactors if the chemical reaction is slow. For example, if the catalyst is able to transform only such an amount of gaseous molecules per time, which are transported by an overall volumetric mass transfer coefficient for the gas of 0.05 1/s, a conventional trickle-bed will perform just as well as the minichannel reactors. In this case, the application of composite

In order to investigate the performance of miniaturised random packed-bed reactors, two different concepts were analysed with respect to achievable mass transfer rates and required power consumption. Concept A was a mini packed-bed consisting of catalyst particles with a diameter of 0.8 mm filled in a steel pipe with an inner diameter of 6 mm. The performance of the catalyst bed was experimentally studied in the setup which has been previously used for analysis of the composite minichannel reactor packing. The flow conditions covered trickle as well as pulse flow according to the Excel Worksheet Simulator for Packed-Bed Reactors of Larachi and Grandjean (2010). In the evaluation, axial dispersion inside the packing was neglected according to the findings of Turek and Lange (1981) for Reynolds numbers above 1. Table 1 summarises the experimental conditions and results. Concept B was a micro packed-bed, i.e. the particles had a mean diameter of about 50 mm and were inserted into a flow channel with nearly rectangular cross section of 300 mm  650 mm (Losey et al., 2001). In the reactor, the authors hydrogenated cyclohexene and reported average reaction rates, overall volumetric mass transfer rates, as well as the corresponding power input per reactor volume. The detailed mass transfer rates were computed by using linear interpolation between the given reaction rates and the minimum and maximum of achieved mass transfer rates. Such a simple treatment does not respect internal diffusion but the results will not change drastically. The calculated mass transfer coefficients were corrected by the different diffusivities. The detailed values are listed in Table 2. The studied composite minichannel reactor as well as the mini packed-bed reactor provide similar maximum mass transfer rates. The latter outperforms the composite minichannel reactor especially if the gas holdup is high and the energy dissipation is low. This means that the mass transfer rates of a composite minichannel reactor may also be generated in a mini packed-bed. In other words, the periodic variation of the external wetting due to the periodic flow regime, as discussed before, does not seem to be as effective as expected. Overall mass transfer rates which are an order of magnitude higher are generated in a micro packed-bed at similar energy Table 1 Experimental conditions and results observed in the mini packed-bed reactor (dh ¼9.0 mm, dS ¼ 0.8 mm). uL,s [m/s]

uG,s [m/s]

Flow regime

P/V [W/m3]

kGLS aGLS (for hydrogen) [1/s]

0.002 0.006 0.012 0.020 0.002 0.006 0.012 0.020

0.020 0.020 0.020 0.020 0.040 0.040 0.040 0.040

Trickle flow Trickle flow Trickle flow Pulse flow Trickle flow Trickle flow Trickle flow Pulse flow

92 211 726 1887 446 1480 2214 4433

0.31 0.80 1.36 1.62 0.33 1.00 1.66 2.63

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Table 2 Experimental data of Losey et al. (2001) for a micro packed-bed and computed volume specific power input and mass transfer rates. _L m [mg/min]a

V_ G [std cm3/min]a

Average reaction rate [10  4 mol/ (min  gCat)]a

P/V [W/m3]b

kGLS aGLS (for hydrogen) [1/s]b

15 75 153 75 75 75 55

5.0 5.0 5.0 3.2 5.0 6.6 10.0

8.6 10.0 14.0 90. 10.0 13.0 14.0

2000 3304 5000 3304 3304 3304 2870

5.0 7.6 15.0 5.7 7.6 13.1 15.0

a b

Original data of Losey et al. (2001). Computed values within this work.

dissipation as in the composite minichannel reactor. As expected, a reduction of the characteristic dimensions enhances mass transfer capability, e.g. by higher interstitial velocities, increased interfacial areas, shorter diffusion distances, or intensified mixing inside the phases. 3.10. Evaluation of mass transfer and power dissipation in all reactors Adopting a more general approach, the necessary power input per reactor volume in the several reactors is related to the corresponding overall mass transfer coefficients. By such a treatment, the size of the apparatus becomes irrelevant. A large device with low mass transfer rates and low power consumption can have a similar contacting efficiency as a small device with high mass transfer rates which are affected by high energy dissipation. The resulting minima and maxima of the ratio between volumetric power input and mass transfer coefficient are depicted in the upper diagram of Fig. 15. The achievable overall volumetric mass transfer coefficients will dictate the investment costs for the reactor installation as well as the costs for the catalyst. A fixed productivity is achieved in a smaller vessel if the absolute volumetric mass transfer coefficient is high. The reduced volume will cause fewer expenses for the catalyst as well as for the reactor itself, especially if high operation pressures are desired. Furthermore, the amount of reactive substances is reduced in small apparatuses which decreases expenditures for safety equipment. The minima and maxima of the mass transfer coefficient are shown in the lower diagram of Fig. 15. Based on both mentioned parameters, the process of a proper reactor selection for an industrial hydrogenation process will be discussed briefly. In the evaluation, a fast chemical reaction is assumed which is limited by the external transport of the gaseous molecules and not by reaction kinetics. Inspecting the ratio between power consumption and the generated volumetric mass transfer coefficient, the minimal values in the upper diagram of Fig. 15 verify that micro and mini packed-bed reactors can exploit the energy most efficiently for the transfer of matter. About four times less power efficiency is provided in the studied composite minichannel reactor and in a conventional trickle-bed reactor whereas it is more than one order of magnitude smaller in slurry stirred tank and bubble column reactors. In consequence, the lowest operational costs in terms of pumping power will be needed by micro and mini packed-bed reactors whereas they are slightly higher in a composite minichannel reactor. The evaluation clarifies that a conventional trickle-bed reactor exploits the dissipated energy less effectively and suffers from low overall mass transfer rates. As a result, the apparatus will be relatively large and contain a high mass of catalyst. For processes

Fig. 15. Contacting efficiencies and overall volumetric mass transfer coefficients for hydrogen versus power input per unit reactor volume for the composite minichannel reactor of this work, for several packed-bed reactors [(Losey et al., ¨ 2001), this work] and for different slurry reactors [(Linek et al., 1993; Schluter and Deckwer, 1992), corrected with LS transfer by (Baerns et al., 2006: p.192; Dutta and Pangarkar, 1996; Tschentscher et al., 2010)].

with very selective catalysts and without any volume restrictions, this established fixed-bed technology may be a reasonable option, but scale-up from the lab to production scale is still a challenging issue because of phase maldistribution, flow channelling, and hot spot formation due to incomplete catalyst wetting. To improve catalyst wetting as well as to reduce axial dispersion, small-sized structures seem to be an interesting approach. In particular, the composite minichannel reactors with Taylor-flow-similar regimes ensure complete wetting and nearly plug-flow behaviour. The latter is especially important if a non-uniform residence time reduces the yield or the selectivity of the process. If the mini and micro packedbed reactors show a similar characteristic as the composite minichannel reactor should be analysed in further investigations. From theoretical expectations, the residence time distribution for fluid molecules is narrowed if the particle diameter of a random packedbed is reduced, i.e. by employing micro or mini packed-beds instead of the packed-beds of a conventional trickle-bed reactor. However, it has to be mentioned that small sized flow paths promote fouling or mechanical blocking which may limit the industrial application in particular if non-purified feedstocks are used. Comparing the performance of mini-structured reactors with that of slurry reactors and especially stirred tank reactors, which are extensively used to produce fine chemicals and pharmaceutics, it becomes clear that mini fixed-bed reactors provide a similar or even higher mass transfer capacity than the slurry reactors and consume less energy. Furthermore, they contain a fixed catalyst and there is no need for catalyst filtration. In addition, the vertical downflow inside the fixed-bed will generate significantly smaller backmixing of the liquid phase which may enhance conversion and selectivity. Finally, the time of the batch cycle needed for the charging and discharging of the stirred tank can be used for production if continuous flow reactors such as the presented micro and mini-structured devices are used. For processes without rapid catalyst deactivation, these reactor types are an interesting alternative to intensify multiphase reaction processes. These potential benefits explain the efforts taken in industry to switch from a batchwise mode of operation to a continuous mode of operation, known as batch-to-conti transfer. It is important to mention that in technical applications the maximal heat transfer capacity may also limit the reactor productivity. To express it differently, the system has to be operated at process conditions in which the cooling system can remove the reaction heat and ensure a safe operation. The potentially high

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mass transfer rates in the composite minichannels or the miniaturised packed-beds in particular will cause an enormous production of thermal energy if the reaction is highly exothermic. Conventional fixed-bed reactors suffer from low axial heat transfer and operation is nearly adiabatical. This will also be the case for the applied ceramic minichannel structures. To enhance the heat removal, metallic substrates which have a better thermal conductivity in combination with multitubular reactors can be applied. Cross flow structures and reactors operated at high liquid throughputs or with external heat exchangers form further supplemental features in the design of a reliable reactor concept.

4. Conclusions This work presents a systematic analysis of a composite minichannel reactor consisting of square minichannels packed with spherical catalyst particles. The hydrodynamic investigations revealed that important hydrodynamic parameters inside the packing such as gas bubble and liquid slug lengths and, consequently the gas holdup and the interfacial areas for the gas–solid and liquid–solid interface can be computed from generated data in an empty channel. From experimental results for the hydrogenation of alpha-methylstyrene, overall volumetric mass transfer coefficients for hydrogen have been computed and an adequate mass transfer model for the composite minichannel reactor is proposed. Additionally, friction factors in such reactors were correlated with the liquid interstitial Reynolds numbers. By using the developed set of empirical correlations, the performance of a composite minichannel reactor can be roughly estimated for any reaction system a priori. An extensive comparison between the composite minichannel reactor with conventional and developing reactor technologies was performed. The results demonstrate that composite minichannel reactors as well as mini and micro packed-bed reactors offer drastically larger overall mass transfer rates than a conventional trickle-bed reactor. In consequence, significantly smaller fixed-bed reactors may be created by adopting this concept, which contain less amount of catalyst and may reduce the costs of the catalyst packing and the safety equipment. In comparison to slurry reactors, composite minichannel reactors as well as mini and micro packed-bed reactors provide a similar or even higher mass transfer capacity, consume less energy, redundantise catalyst filtration, and provoke lower backmixing. In particular, composite minichannel reactors with their segmented flow structure significantly reduces the intermingling of the liquid phase along reactor length. Therefore, the application of miniaturised and structured packed-bed reactors may be a virtual concept in switching from batch to continuous processing. The mass transfer capacity of miniaturised reactors increases by decreasing the characteristic particle diameter while the energy consumption per transported molecule does not change drastically. In consequence, the dimension should be chosen as small as tolerable by other process restrictions created e.g. by fouling, blocking, or catalyst handling.

Nomenclature aGS [m2/m3] Gas–solid interfacial area per unit flow reactor volume aGL [m2/m3] Gas–liquid interfacial area per unit flow reactor volume aLS [m2/m3] Liquid–solid interfacial area per unit flow reactor volume Aq [m2] Cross sectional area of the reactor ACMC [m2] Mean open frontal area of the channel 3 out cin AMS ,cAMS [mol/m ] Alpha-methylstyrene concentration at the reactor inlet and outlet

235

cG1 . . .cG3 Empirical constants to describe gas–solid mass transfer cL1 . . .cL3 Empirical constants to describe liquid-solid mass transfer 3 csat H2 [mol/m ] Saturation concentration of hydrogen in the liquid phase dh [m] Hydraulic channel diameter dh,CMC [m] Hydraulic diameter of the composite minichannel dS [m] Diameter of the catalyst spheres D [m2/s] Diffusion coefficient fa,CMC Form factor for packing configuration fF,CMC Friction factor in the composite minichannel g [m/s2] Gravitational constant ( ¼9.81) I Index kGS [m/s] Gas–solid mass transfer coefficient kGL [m/s] Gas–liquid mass transfer coefficient kLS [m/s] Liquid–solid mass transfer coefficient kGLaGL [1/s] Volumetric mass transfer coefficient through the gas–liquid interphase per unit flow reactor volume kGLSaGLS [1/s] Equivalent volumetric mass transfer coefficient per unit flow reactor volume for the parallel and consecutive steps of gas–liquid, liquid–solid and gas–solid mass transfer kGL þ LSaGL þ LS [1/s] Equivalent volumetric mass transfer coefficient per unit flow reactor volume for the consecutive steps of gas–liquid and liquid-solid mass transfer kintr,v [1/s] Intrinsic rate constant per active catalyst volume LB,CMC [m] Gas bubble length in the composite minichannel LB,EMC [m] Gas bubble length in the empty minichannel Lchar [m] Characteristic length LR [m] Reactor length LS,CMC [m] Liquid slug length in the composite minichannel LS,EMC [m] Liquid slug length in the empty minichannel _ L [kg/s]Feed liquid mass per time m P [Pa] Pressure P [W] Power Pm CMC [m] Mean wetted perimeter of the composite mini channel R2 Coefficient of determination with respect to the overall volumetric mass transfer coefficient per unit flow reacP ð½kGLS aGLS Exp,i ½kGLS aGLS Model,i Þ2 i tor volume R2 ¼ P 2 i

ð½kGLS aGLS Exp,i kGLS aGLS Þ

rm [mol=ðs g Pd Þ] Reaction rate per palladium mass rapp,v [mol=ðs m3 Þ] Apparent reaction rate in the composite minichannel reactor per catalyst volume ReS,G Reynolds number in the composite channel for the gas phase ( ¼ rG uTP,is dS =mG Þ ReS,L Reynolds number in the composite channel for the liquid phase ( ¼ rL uTP,is dS =mL Þ ReTP,is Reynolds number in the composite channel ( ¼ rL uTP,is dh,CMC =mL Þ ScL,H2 Schmidt number of hydrogen in the gas phase ( ¼ mG =ðDH2,G rG ÞÞ ScL,H2 Schmidt number of hydrogen in the liquid phase ( ¼ mL =ðDH2,L rL ÞÞ ShG,H2 Sherwood number in the gas phase ð ¼ kGS dS =DG,H2 Þ ShL,H2 Sherwood number in the liquid phase ( ¼kLSdS/DL,H2) TR [K] Reaction temperature uB,CMC [m/s] Gas bubble velocity in the composite minichannel uB,EMC [m/s] Gas bubble velocity in the empty minichannel uL,s [m/s]Superficial liquid velocity (corresponds to the unpacked channel) uG,s [m/s] Superficial gas velocity (corresponds to the unpacked channel) uTP,is [m/s] Interstitial two-phase velocity (corresponds to the packed minichannel) uTP,s [m/s] Superficial two-phase velocity (corresponds to the unpacked channel) V [m3] Volume

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V_ G , V_ L [m3/s] Gas and liquid volumetric flow rates, respectively VR [m3] Reactor volume out in XAMS Alpha-methylstyrene conversion ð ¼ ½cin AMS cAMS =cAMS Þ bG,in Gas holdup at the reactor inlet bG,m Mean gas holdup inside the flow channel bL,m Mean liquid holdup inside the flow channel bG,UC Gas holdup inside one unit cell Dp [Pa] Pressure drop Dpdyn,R [Pa] Dynamic pressure drop inside the reactor DpL,feed [Pa] Pressure drop inside the liquid feeding eBed [m3free =m3Bed ] Bed porosity ZCat Catalyst effectiveness factor m [Pa s] Dynamic viscosity p Mathematical constant (3.14159y) r [kg=m3 ] Density f Thiele modulus

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