Rotating reactors – A review

Accepted Manuscript Title: Rotating Reactors–A Review Author: Frans Visscher John van der Schaaf Xander Nijhuis Jaap Schouten PII: DOI: Reference:

S0263-8762(13)00298-0 http://dx.doi.org/doi:10.1016/j.cherd.2013.07.021 CHERD 1322

To appear in: Received date: Revised date: Accepted date:

15-4-2013 19-7-2013 22-7-2013

Please cite this article as: Visscher, F., van der Schaaf, J., Nijhuis, X., Schouten, J., Rotating Reactors–A Review, Chemical Engineering Research and Design (2013), http://dx.doi.org/10.1016/j.cherd.2013.07.021 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.

Research Highlights.doc

Research highlights

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The current state-of-the-art in the field of rotating reactors is presented. Rotating reactors are classified and discussed according to their geometry. Their main advantages and disadvantages are presented Additionally, the historical development of these reactors is presented.

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*Manuscript

Rotating Reactors – A Review Frans Visscher1, John van der Schaaf1, Xander Nijhuis1, Jaap Schouten*, 1

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Laboratory of Chemical Reactor Engineering, Department of Chemical Engineering and Chemistry, Eindhoven University of Technology, P.O. Box 513, 5600 MB Eindhoven, The Netherlands. Tel.: +31 40 247 2850; www.chem.tue.nl/scr E-mail address: [email protected]

Abstract This review-perspective paper describes the current state-of-the-art in the field of rotating reactors. The paper has a focus on rotating reactor technology with applications at lab scale, pilot scale and industrial scale. Rotating reactors are classified and discussed according to their geometry: stirred tanks, tubes, discs and miscellaneous reactors. Their operating characteristics, industrial applications, and their main advantages and disadvantages are discussed including power requirements, residence time distribution, reactor volume, gas-liquid mass transfer rate, and the micromixing time. Finally, the barriers for further industrial implementations are discussed.

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1. Introduction Mechanical agitation is applied to chemical reactors on pilot scale and on industrial scale with the aim to increase the mixing efficiency (Hemrajani and Tatterson, 2004). As a result, the reactor volume can be decreased by a factor 10-100, which is an urgent need in chemical industry; this need is commonly addressed as process intensification (Stankiewicz and Moulijn, 2000). Colin Ramshaw and co-workers were the first pioneers who envisioned that understanding the length and time scales relevant in plant design will lead to process intensification (Reay et al., 2008). The relevant length and time scales include the integral path of plant, reactor, fluid dynamics, catalyst and molecular level (Charpentier and McKenna, 2004; Lerou and Ng, 1996).

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The reactor volume reduction that originates from the application of mechanical agitation yields a potential cost benefit that acts as a driving force for the development of rotating reactors, especially in research dealing with the fine chemical and pharmaceutical industry (Anderson, 2012; Chaudhari and Mills, 2011). With a lower equipment volume the holdup of liquids, gases and solids in the reactor is smaller, which reduces the impact of potential escalation of dangerous situations. Additionally, the volume reduction allows for the application of expensive coating materials in the reactor, like platinum and tantalum.

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Mechanical agitation is the most common method that is applied in order to enhance mixing in industrial reactors. For this purpose a rotating element is added to the reactor with the purpose to increase the gas-liquid, liquid-liquid, and liquid-solid mass transfer rates as well as the heat transfer rate. As a result conversion and selectivity can be increased and the occurrence of hot-spots is prevented. Process intensification has been illustrated through the application of rotating reactors, for example through the application of thin-film spinning disc reactors with a high heat transfer coefficient (Aoune and Ramshaw, 1999; Zhang et al., 2010). Alternative methods to achieve the said intensification are the integration of reaction and separation in one unit, the switch from batch to continuous operated reactors and the application of microreactor technology. Microreactors allow for the reduction of byproduct formation and the increase of gas-liquid and liquid-liquid mass transfer rates (Hessel et al., 2009; Hessel, 2009). The main challenge for further implementation of this technology is the scale-up of the volumetric throughput, which can only be achieved by increasing the pressure drop over the microchannel or by parallel feeding of multiple channels (AlRawashdeh et al., 2012; Kashid et al., 2010). This review-perspective paper classifies rotating reactors and describes the current state-of-the-art in the field of rotating reactors which are applied on lab scale, pilot scale, and industrial scale. In this review a chemical reactor is defined as any device which is used to conduct a chemical reaction, more specifically, in which at least one molecular compound is transformed into another predetermined chemical compound (Trambouze and Euzen, 2002b). The discussed rotating reactors are applied for either single-phase systems (liquid or gas) or multi-phase systems (gas-liquid, gassolid, gas-liquid-solid, liquid-solid, or liquid-liquid mixtures). 2

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The reactor selection is narrowed down further by discarding reactors which are only used for the experimental determination of physical properties, and by discarding reactors solely used for the testing of catalytic properties or reaction kinetics (Doraiswamy and Tajbl, 1974; Dudukoviç et al., 2002; Kapteijn and Moulijn, 2008; Pavko et al., 1981; Weekman, 1974). Examples of these reactors are the Berty Stationary Catalyst Basket (Berty, 1974), the Mahoney-Robinson Spinning Catalyst Basket reactor (Mahoney, 1974), and the Carberry Spinning Catalyst Basket reactor (Carberry, 1964). These reactors are discarded because they are not developed with the aim of industrial implementation. Reactors in which mechanical energy is dissipated without the usage of rotation are also discarded (e.g. the oscillatory baffled reactors (Harvey et al., 2001), static mixers (Meijer et al., 2012), and reactors in which cavitation is ultrasonically induced cavitation (Rooze et al., 2013)).

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1.1. Reactor selection criteria Reactor selection is a challenging task in which multiple aspects need to be taken into account (Krishna and Sie, 1994). Reactors are developed dedicated to their application which can be the synthesis of nanoparticles (Dahl et al., 2007; Ng et al., 2012; Pask et al., 2012; Tai et al., 2007), exothermic reactions (Ogura et al., 2008), electrochemical reactions (Rivero et al., 2010), micromixing characteristics (Assirelli et al., 2002; Baldyga and Pohorecki, 1995; Bourne and Studer, 1992; Chu et al., 2007; Jiao et al., 2012; Rousseaux et al., 2001; Zhao et al., 2010), and others. Reactions can also be combined with separation steps in reactive separators (Stankiewicz, 2003). Specific aspects that cannot be neglected when reactors are compared, are the volumetric throughput, reactor volume, residence time distribution, effective catalyst loading, mass transfer rates, and pressure and temperature limitations of the reactor (Cybulski and Moulijn, 2005; Froment et al., 2011; Trambouze and Euzen, 2002b).

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Rotating reactors contain one or more rotating elements which may have various designs: impellers, tubes, or discs. The geometry of the rotating element and the rotational speed determine the power required for mechanical agitation. Energy dissipation in chemical reactors is given in W kgPRODUCT-1 or W m-3LIQUID (Baldyga and Pohorecki, 1995; Villermaux, 1988). The mixing intensity is mainly determined by the local energy dissipation and not by the power consumption of the auxiliary equipment (Laufhütte and Mersmann, 1987; Mukherjee and Wrenn, 2009). For mechanically agitated reactors the energy consumption by the motor that propels the rotating element is dominant over the energy consumed by pumps. The compressor duty significantly contributes to the total power consumption in the case of gas-liquid reactors. 1.2. Reactor classification In this review, first the development of rotating reactors over time is presented. Next the specific benefits and disadvantages of various rotating reactors are given. In this review rotating reactors are classified by four different groups according to their characteristic geometry: tanks, tubes, discs and miscellaneous rotating reactors (Figure 1). 3

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The development of various rotating reactors in time is schematically shown in Figure 2. In Figure 2 the year in which the first scientific or technical publication (peer-reviewed or patent) was published is used as allocation in time. A table with the relevant literature of each reactor is given in Table 1. The majority of rotating reactors has been developed after 1980. Exact sales information is not available, but most of the reactors developed since 1980 have not yet permanently penetrated the commercial market of industrial processes except for niche applications.

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2. Rotating reactors 2.1. Stirred tanks Agricola (1556) illustrated in his book “De Re Metallica” (On the Nature of Metals) how stirring reactors were used in the mining industry (Stankiewicz and Moulijn, 2000). Earlier references to vessels equipped with a stirrer date back to 77 A.D. where Pliny the Elder describes the leaching of metals and the purification of sulfur (Pliny the Elder, 1929). The volume of industrial used stirred tank reactors ranges from 2 10-3 to 3 102 m3. The reactors are either batch wise, semi-batch or continuously operated (Trambouze and Euzen, 2002a).

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The most basic form of a stirred tank reactor consists of a cylindrical tank with elliptical bottom, with one or more stirrers mounted on a central shaft. With increasing tank diameter, stirred tanks often exhibit poor mixing which is especially true for multiphase reactions in which the nonuniformity in mixing and mass transfer leads to significant variance in reaction rate and selectivity (Ståhl Wernersson and Trägårdh, 1999; Stitt, 2002). Some of the stirrers are shown schematically in Figure 3 (Hemrajani and Tatterson, 2004; Joshi et al., 1982). Both the power consumption in stirred tanks (Ascanio et al., 2004; Villermaux, 1988), and the various techniques to visualize the liquid flow behavior have been elaborated already earlier in a review (Mavros, 2001).

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The catalyst in a stirred tank can be either dispersed as a slurry (with a particle diameter below 1 10-3 m) or as a homogenous catalyst. The most important drawbacks of a heterogeneously dispersed catalyst is the separation of the catalyst from the reaction mixture, and the attrition and agglomeration of the catalyst particles. Heat transfer to or from stirred tank reactors can be obtained by jacketing the stirred tank or by using internal coils. Usage of structures inside the reactor allows for higher heat transfer rates, but increases also the risk of fouling, the non-uniformity of mixing intensity, and the time required for reactor cleaning (Kumaresan and Joshi, 2006). Various methods can be exploited in order to improve the mixing capability in stirred tanks. The application of vertical wall baffles mounted to the reactor wall is well known (Lu et al., 1997). A second method to enhance mixing capability in the stirred tank reactor is by applying a more sophisticated stirrer design, which can be categorized according to the induced direction of flow: up/down draft (disks, plates), radial (flat blade impeller, Rushton stirrer), axial (propeller, pitched blade turbine), or vortex (Kumaresan and Joshi, 2006; Stitt, 2002). 4

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Stirrer geometries can be, but are not limited to, propellers, turbines, anchors, or Archimedes screws (Hemrajani and Tatterson, 2004). An extensive recent review on typical impeller characteristics which are essential for further comparison of various impellers is given elsewhere and addresses the relevance of the power number, the flow number, the momentum number, and the Zwietering constant (Machado et al., 2012). Multiple stirrers on a single rotating shaft are needed when the aspect ratio, the ratio of the stirrer diameter over the tank diameter, exceeds 1.5. A last method to enhance mixing in stirred tanks is the application and improvement of a gas-distributing inlet, which will enhance the interfacial area and the gas-liquid mass transfer rate in the stirred tank. Four different stirrers are described in more detail here: Rushton stirrer, gas inducing stirrer, monolithic stirrer and the solid foam stirrer.

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2.1.1 Rushton stirrer The Rushton stirrer was developed around 1940 (Foust et al., 1944), and is a radial flow generating stirrer, which is equipped with six vertical (flat or curved) blades which are mounted on a disc (Figure 4). For the standard stirrer the blade length is equal to DI/4, the blade width is equal to DI/5. The disc diameters equals either 0.66DI or 0.75DI, in which DI is the impeller diameter (Hemrajani and Tatterson, 2004). The gas-liquid flow behavior in a Rushton stirred tank was studied using Laser-Doppler Anemometry (Wu et al., 1989) and Particle Image Velocimetry (Hill et al., 2000). The power characteristics of Rushton stirrer are related to physical properties of the liquid mixture and the geometry of the tank itself (Rushton et al., 1950a; Rushton et al., 1950b).

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2.1.2 Gas-inducing stirrer Often the per pass conversion of the gas phase is low when gas-liquid mass contacting is performed in a stirred tank, in that case it is beneficial to recycle the unreacted gas-phase back into the reactor. Dead end systems are than a solution in which expensive compression costs can be reduced: in these systems the remaining gas phase is forced into the free reactor volume from where it is recycled internally into the liquid mixture. A gas-inducing impeller enables efficient recycling of gas from the free reactor volume into the liquid-mixture. The critical impeller speed that is required for the start of gas induction follows from a balance between the velocity head generated by the impeller and the hydrostatic head above the impeller (Patwardhan and Joshi, 1998). Guidelines have been given about the desired geometry of gas-inducing impellers for achieving different design objectives such as heat transfer, mass transfer, mixing, solid suspension, froth flotation, and so forth. Non-intrusive electrical capacitance tomography (ECT) has been used to study the dispersed phase hold up, mixing times, and reaction metrics in a continuously operated stirred tank that was equipped with a gas-inducing stirrer (Bawadi et al., 2011). Application of a gas-inducing stirrer in a stirred tank gives an increase in the productivity (Bawadi et al., 2011). The influence of the stirrer diameter, the aspect ratio, the stirrer submergence from the liquid level, and the clearance between the stirrer and the tank bottom has been presented elsewhere (Saravanan et al., 1994). 5

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The gas−liquid mass transfer rate in stirred tanks that are equipped with gas-inducing impellers was measured, and can be described by a dimensionless correlation which contains the Froude number (gas-induction rate), the Reynolds number (turbulence intensity), and the Schmidt number (fluid properties) (Zieverink et al., 2006). The separation of the reaction mixture and the catalyst particles at the outlet of a stirred tank is often troublesome; there is therefore a tendency to immobilize the catalyst on the stirrer. In early attempts a lab-scale rotating basket was mounted in the reactor, such that kinetics could be measured without the presence of external mass transfer limitation in the gas phase (Carberry, 1964).

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2.1.3 Monolithic stirrer Mounting monolithic blocks on the stirrer shaft can be an attractive alternative for stirred tanks with dispersed catalyst particles (Figure 4). Most monoliths consist of one piece of ceramic material; within this piece a large number of parallel channels is present which extends over the entire length of the block. The concept of a monolithic stirrer was demonstrated in 1998 for liquid mixtures with a low viscosity (Albers et al., 1998). Because the catalyst is immobilized in the monolith there is no need for liquid-solid separation at the reactor outlet. Another advantage is the open structure of the monolithic block which results in a large geometrical area. The inside of the monolithic channels can be coated with a thin layer of either a conventional catalyst (Bennett et al., 1991) or a biocatalyst (De Lathouder et al., 2006). The monolith is characterized by its number of cells per square inch (Hoek, 2004b). With increasing cell density the catalyst layer thickness decreased, which proved to be beneficial for the performance of the monolithic stirrer reactor (Hoek et al., 2004). The volumetric gas-liquid mass transfer coefficient in a monolith increases from 0.015 s-1 at 200 RPM to 0.527 s-1 at 450 RPM when measured for the stirrer configuration with two monoliths in one plane (Hoek, 2004a). The cell density of the monoliths has no effect on the gas-liquid mass transfer rate. The volumetric gas-liquid mass transfer increases by a factor of three when a stirrer configuration consisting of four monoliths in one plane is used. The liquid-solid mass transfer rate increased with increasing stirrer speed (Hoek, 2004b). The biggest potential for industrial implementation would be the replacement of conventional slurry reactors that are nowadays used for multiphase processes, for example in fine chemical synthesis. 2.1.4 Foam based stirrer A recent innovation in stirrer design for stirred tanks is the solid foam based stirrer. Such stirrers contain a piece of open-celled solid foam which is made out of a reticulated structure of struts (Figure 5). Each strut has the function of a static mixer which splits and recombines the fluid stream that passes the strut. The solid foam combines a high surface area (160-8500 m2 m-3) with a high voidage (80-97%), which yields a high surface to volume ratio. As a result non-rotating foam packed columns exhibit a low pressure drop and high gas-liquid mass transfer rates (Stemmet et al., 2005; Stemmet et al., 2006). Due to the high surface area it is an excellent material for the deposition of catalysts (Ordomsky et al., 2012a; Ordomsky et al., 2012b; Wenmakers et al., 2008; Wenmakers et al., 2010). 6

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The rotating foam stirrer reactor yields higher gas-liquid and liquid-solid mass transfer rates than stirred tanks equipped with a Rushton stirrers or slurry bubble columns (Tschentscher et al., 2010a; Tschentscher et al., 2010b). Various foam structures have been applied including donuts, two-blades, and blocks (Leon et al., 2012a). Rotating foam stirrer reactors have promising applications for multiphase processes (Leon et al., 2012b). Mounting multiple foam blocks on a single horizontal shaft with baffles in between the consecutive foam blocks, yields plug flow behavior in this reactor, and accordingly a higher selectivity towards the desired product in selective reactions with unwanted consecutive reactions (Leon et al., 2013). Tomography measurements have shown that also liquids with a higher viscosity can be fed to such a reactor (Tschentscher et al., 2012). An important advantage for industrial implementation of this stirrer type is the very low pressure drop which ensures the good accessibility of the catalyst on the foam. For industrial implementation it is essential that the catalyst does not wash out, and thus that the appropriate catalyst coating technique is applied.

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2.2. High shear rotating tubes 2.2.1. Rotating packed bed Rotating Packed Bed (RPBs) have received considerable attention as a method of process intensification for gas-liquid mass transfer. Most publications mention that the first patent on RPBtechnology was filled by Ramshaw et al. in 1978 (Ramshaw and Mallinson, 1981). This patent described the application of such technology for distillation, absorption, and stripping (Ramshaw, 1983). An earlier patent application described this same technique about two decades earlier (Pilo, 1960). Rotating packed beds can be operated either co-currently, cross-currently or countercurrently. In the most basic form of a RPB, the rotor is an annular cylindrical packed bed which is housed in a cylindrical casing (Figure 6). In rotating packed beds the rotational speed ranges from 500 to 2500 RPM. The rotor can be made of various packing types, e.g. gas spheres, solid foam, and discs. Other rotor configurations which have been applied are disc plate packings (Jian et al., 1998), helical packings (Chen et al., 1995), multistaged spraying packings (Pan et al., 2006), and splitpackings (Chandra et al., 2005; Reddy et al., 2006). The gas phase and liquid phase are counter-currently contacted in a RPB. The gas phase is introduced near the casing, and flows radially inwards through the packing and exits the reactor through the shaft. The liquid phase is fed through a stationary feeding tube placed in the eye of the rotor, and touches the inner periphery of the packing as droplets, jets, or as a spray. For measurements at low gravity level (<60g) the liquids fills the voids of the packing, whereas at high gravity levels (>100g) the voids are only partial filled (Burns et al., 2003; Burns and Ramshaw, 1996). The residence time distribution in a RPB was determined experimentally, and showed that intense macromixing is obtained at the inner region of the packing (dR = 7 10-3 - 10 10-3) where the liquid impinges and deforms. The same study showed that the liquid volume in a packed bed under normal operation, does not exceed 5% of the total bed volume (Kenyvani and Gardner, 1989),. 7

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As a result the mean residence time of the liquid phase in the reactor is very short which limits the separation capacity. The micromixing efficiency of rotating packed beds is discussed in detail by Yang et. al. (Yang et al., 2005). There, the micromixing time is approximately evaluated to be about 10-4 seconds.

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Two in-depth reviews describe the functionality of the rotating packed bed and the relevant literature up to 2010 in great detail (Rao et al., 2004; Zhao et al., 2010), including liquid flow, gasliquid mass transfer, micromixing efficiency, stripping/absorption, and nano-particle preparation. The main conclusion is that adequate design of the rotor is essential for full exploitation of the centrifugal field. The height equivalent of a transfer unit is about 4-5 times lower in RPBs than in conventional extraction columns (Rao et al., 2004), and equals 1 10-2 – 2 10-2 m (Zhao et al., 2010). The intensification achieved so far falls short of the goal of 2 - 3 orders of magnitude volume reduction of the conventional columns volume (Rao et al., 2004).

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Relatively poor mass transfer performance of RPB technology was observed when used as bioreactor for the production of polyhydroxyalkanoate from a fermentation broth (Boodhoo et al., 2010). The apparent poor performance is attributed to the limitation of the gas and liquid flow rates that are a consequence of flooding characteristics. Also, the possible entrainment of air in the oxygen stripped end product could have influenced the reliability of these findings. The micromixing efficiency in the RPB was determined using the Villermaux-Dushman reaction system. In this study samples were drawn along the radial position. The micromixing efficiency was thus measured in different packing zones of the RPB. The micromixing time in the RPB can be as low as 10-4 s, which is one to three orders of magnitude smaller than in conventional packed beds. In the Impinging Stream Rotating Packed Bed (IS-RPB), two stainless capillary nozzles are located in-line with the rotational shaft. The two different liquids flow along the same axis in the opposite direction and collide, which causes a narrow zone in which a high turbulence intensity is created which leads to high micromixing (Qi et al., 2008), and excellent mass transfer efficiency (Jiao et al., 2012). Several implementations of rotating packed beds have been successfully obtained, including seawater treatment, HOCl production and nanoparticle preparation. However, industrial acceptance of this reactor type is yet to come, mainly due to concerns involving reliability and the energy consumption. More sound theoretical foundations are required, especially the liquid distribution at the entrance (in the centre) of packed bed is of main importance for the mixing intensity obtained in the reactor. This should be combined with long term operational data and economic evaluation of such data. 2.2.2. Rotating zigzag bed In the rotating zigzag bed (RZB) the packing from the rotating packed bed is replaced by a rotor which is a coaxial combination of a rotating disc with a stationary disk (Figure 7) (Li et al., 2012b; Li et al., 2013). Concentric circular sheets are mounted on the rotor and the stator, and act as rotating and stationary baffles, respectively. Typically the distance between the consecutive rings is below 15 10-3 m. 8

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The gas and liquid flow are counter currently contacted while flowing in a horizontal zigzag motion through the reactor. The RZB thus has in essence a similar geometry as conventional tray-distillation columns, but in the RZB the flow is oriented in a vertical plane and the gravitational field (1g) is replaced by a high centrifugal field (>500g) (Wang et al., 2011). When compared to a rotating packed bed reactor the diameter is increased by a factor two, whereas the length is reduced by a factor 16. The reactor volume thus decreased by a factor 5 (Li et al., 2013).

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The rotating zigzag reactor has a few beneficial advantages over common RPB technology: first, one of the dynamic seals can be replaced by a static seal, yielding a longer seal lifetime and thus less maintenance. Second, additional feed inlets can be applied through the stator which allows for continuous operation of the distillation process (Ji et al., 2008). Third, no additional liquid feed distributers are required: the zigzag flow enhances mixing between each consecutive ring. Fourth, the liquid hold-up of the bed is larger than in a conventional RPB, which leads to an increased residence time. Fifth, in theory multiple zigzag-rotors could be mounted on one rotational shaft (Wang et al., 2008).

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The number of theoretical plates increases with increasing rotor speed (Wang et al., 2008). The RZB may contain up to 20 theoretical plates per meter length, when operated at 1000 RPM and reflux rate of 600 10-3 m3 hr-1 (when measured with an inner rotor diameter 20 10-2 m, outer rotor diameter of 63 10-2, casing diameter of 80 10-2 m, casing height 55 10-2 m, measured for the separation of an ethanol-water mixture) (Wang et al., 2008). Seven industrial applications of RZB technology have been reported, with a maximum feed capacity of 48 tons day-1 (Wang et al., 2008). The overall pressure drop of the RZB is a combination of the pressure drops over the rotor, the casing and the gas outlet (Li et al., 2013), and increases with increasing rotor speed and the gas flow rate, and decreases with increasing liquid flow rate (Wang et al., 2011). Empirical correlations for the overall pressure drop over the RZB have been presented for a dry bed and a wetted bed (Li et al., 2013). The power required to propel the rotating zigzag bed is discussed in detail by (Agarwal et al., 2010) and (Li et al., 2012b). 2.2.3. Rotating fluidized bed Fluidized beds are widely applied in chemical industry and are used for a wide range of processes including olefin polymerization end detergent production. In a fluidized bed reactor the gas phase is forced through a bed of solid particles. As a result, at appropriate conditions the gas-solid mixture will exhibit fluid-like behavior, and can also be modeled as such (Deen et al., 2004). In fluidized beds, high catalyst loadings can be applied. Two types of rotating fluidized beds (RFB) have been applied with either a static (SG) or rotating geometry (RG) (Harish Kumar and Murthy, 2010). In the rotating fluidized bed with a rotating geometry the cylindrical enclosure rotates around its own axis of symmetry. As a result the intensity of the gravitational field depends on the rotational speed and the enclosure radius.

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Compared to conventional fluidized beds its main advantage is the wider range of operation (Harish Kumar and Murthy, 2010). The biggest disadvantage is that at extreme rotational speeds of the enclosure, particles will accumulate near the rim of the enclosure which leads to the channeling of the gas phase through the dead zone near the center of the enclosure (Kroger et al., 1980). The pressure drop and fluidization criteria have been presented by (Chen, 1987).

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The second type of rotating fluidized beds has a static geometry which makes that in essence this reactor is not a rotated reactor; however the fluid behavior has a large similarity with that of other equipment described in this review (De Wilde and de Broqueville, 2008a). This rotating fluidized bed with a static geometry is also known as the Gas-Solid Vortex Reactor (GSVR). In this reactor the gas phase is injected tangentially into the cylindrical reactor which contains the solid particles (Figure 8). Due to the transfer of momentum from the gas-phase to the particles, the particles will rotate in the reactor and generate a centrifugal force. High mass transfer rates are obtained in this reactor due to the counter play between the centrifugal force and the drag force applied by the inward-flowing gas. The solids experience a centrifugal force which is directed radially outwards, while the gas phase is forced to move radially inwards, which fluidizes the solids radially. The concept has been illustrated for low density polymer particles and high density aluminia particles, for various solids loadings (De Wilde and de Broqueville, 2007). The heat transfer rate is studied numerically by using computational fluid dynamics simulations of the particle bed temperature response to a step change in the fluidization gas temperature (de Broqueville and De Wilde, 2009).

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The functionality of the GSVR has been demonstrated for the pyrolysis of lignocellulosic biomass. The gas-solid heat transfer coefficient equals 650 W m-2 K-1, which is 3-5 times higher than in nonrotating fluidization reactors (Ashcraft et al., 2012; De Wilde and de Broqueville, 2008a). The gassolid mass transfer coefficient for the GSVR is about 10 times higher than for conventional riser technology (Ashcraft et al., 2013). The GSVR can be extended with a rotating chimney which leads to increase of the centrifugal force in the vicinity of the chimney, which on its turn allows to reduce solids losses via the chimney and simultaneously to build up a higher solids loading in the fluidization chamber at given gas flow rate and solids feeding rates (De Wilde and de Broqueville, 2008b). Rotating fluidized beds have a relatively large volume. In rotating fluidized beds this implies that the energy consumption is high when compared to other reactors. 2.2.4. Taylor-Couette reactor For reactors in which only the inner or both inner and the outer cylinders are rotating, a TaylorCouette type of flow is observed. Such flow offers the advantage of centrifugally accelerated settling, short residence times, low volume fractions, flexible phase ratios and a well-controlled inventory (Vedantam and Joshi, 2006). This class of reactors is also known as rotating annular rectors. The residence time distribution for a such a reactor can be described as near plug flow behavior for Taylor numbers (ratio of centrifugal to the viscous forces) above 60 where a laminar vortex flow regime is present (Pudjiono and Tavare, 1993). 10

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A detailed review on experimental correlations between the axial dispersion number and the Reynolds number is given elsewhere, together with a detailed CFD and RTD study of the fluid flow behavior (Vedantam et al., 2006). Normally a smooth shaped cylinder is used. Experiments with a ribbed rotor lead to an increase of the micromixing efficiency (Richter et al., 2008). Both the exothermic copper-catalyzed oxidation of isopropanol and the pyrolysis of acetone to ketene were studied (Cohen and Marom, 1983), and were executed under excellent control over the yield and selectivity.

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Rotating annular equipment also finds its application in the intensification of stage wise countercurrent liquid-liquid extraction (Schuur et al., 2012), and the continuous production of monodispersed nanoparticles (Ogihara et al., 1995). This reactor is quite similar to the rotating annular reactor, in which a gas-liquid or liquid-liquid mixture is forced through the radial duct between two coaxial cylinders, of which the inner cylinder is rotating (Lawrence et al., 2000).

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The radial distance between the two cylinders is small when compared to the radii of the cylinders and is in the range of 10 10-3 m, which is about 40 times higher when compared to the spinning tube-in-tube reactor. Typically the flow rate is in the range of 10 10-3 m3 hr-1 (Lawrence et al., 2000). As a result, the pressure drop in the rotating annular reactor is lower, which allows for a larger throughput.

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2.2.5. Spinning tube-in-tube reactor The last example of a reactor with tubular geometry is the spinning tube-in-tube reactor. The reactor patent for this reactor was filed in 2008 by Richard Holl. The reactor consists of a rotating cylinder (rotor) inside a stationary cylinder (stator), which are mounted at a concentric radial spacing between 0.25 10-3 and 0.44 10-3 m (Figure 9 from (Holl, 2010)).

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As a result of the small radial distance between the cylinders, the reactants inside the annular volume are exposed to elevated shear stress levels. The typical rotational speed of the rotor is between 3000 and 12000 RPM, with a reactor volume that varies from 1 10-5 to 1 10-3 m3. The reactor concept has been demonstrated for fine chemical production (Gonzalez and Ciszewski, 2008; Hampton et al., 2008). Due to the low reactor volume, and expected excellent heat transfer rates, this reactor is suited for highly exothermic reactions which require only a small amount of catalyst. 2.3. Low shear rotating tubes 2.3.1. Rotating tube reactor Several reactors have been developed which combine a tubular geometry with low shear conditions. The rotating tube reactor consists of a rotatable hollow cylinder, and has a rotational speed which is typically below 1000 RPM. The centrifugal force induced by the rotation results in thin liquid films inside the rotating rube, with a thickness between 700 10-6 and 1400 10-6 m. The technology has been applied in the transesterification of canola oil to biodiesel using a base catalyst. Using methanol and sodium hydroxide as catalysts, a conversion of 98% was achieved in residence times of 40 seconds (Lodha et al., 2012). 11

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The conversion obtained while using a residence time of 40 seconds is equal to the conversion obtained in a membrane reactor operated at a residence time of 6 hours, for a volumetric flow rate which is 150 times higher (Lodha et al., 2012). The mechanically simple design of this reactor and the low pressure drop over the cylinder, make that the reactor can be used for bulk processes in which mass transfer limitations is not limiting.

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2.3.2. Rotating tubular membrane reactor The rotating membrane bioreactor with a tubular geometry was patented in 1982 (Cowen et al., 1982). In this reactor the membrane is rotating around a shaft which is placed at the liquid level in the reactor. As a result the membrane is partially wetted but continuously refreshed during operation. The rotational speed of the membrane is most often below 10 RPM. A typical reactor volume equals 2 10-3 m3, and has a total effective filtration area of 0.043 m2 (Jiang et al., 2012). Rotation of the membrane allows for an increase of the permeate flux because cake formation is suppressed in microfiltration and concentration polarization is suppressed during ultrafiltration and reverse osmosis (Jaffrin, 2008).

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2.3.3. Rotating annular chromatographic reactor The rotating annual chromatographic reactor is a chemical reactor in which chromatographic separation takes place, and is used for liquid-solid or gas-solid systems (Cho et al., 1980a). The reactor consists of two rotatable cylinders which are rotating around one central shaft (Figure 11). The volume between the two cylinders is filled with the solid phase. Often ion-exchange resins are used which have the function of both catalyst and adsorbent. The adsorbent can also be used for the immobilization of biocatalysts (Sarmidi and Barker, 1993b).

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The solvent of the reacting molecules is distributed equally at the top of the reactor. The rotational speed of the cylinders, typically below 10 RPM, causes horizontal migration of the reacting molecules, whereas the fluid flow causes a vertical downward migration of the reacting molecules through the solid phase. Reaction products with a high affinity to the solid phase will have a large horizontal velocity component, whereas reaction products with a low affinity have a small horizontal velocity component (Sarmidi and Barker, 1993a). At different angular positions of the reactor, the separated products can be collected in high purity. The historical development is described in detail in a review elsewhere (Uretschläger and Jungbauer, 2002). Typically the reactor volume is in the range of 10 10-3 m3, whereas the volumetric flow rate equals 15 10-3 m3 hr-1 (Cho et al., 1980b). The pressure drop due to the packing inside the narrow concentric space limits the scalability in terms of volumetric throughput (Ströhlein et al., 2005).

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2.3.4. Rotating sorbent reactor The rotating sorbent reactor was developed in 2004 and is a combination of the rotating particle separator with a cyclonic reactor. The basic design consists of three stages. In the first stage, the solid phase is injected in a gas phase flow that contains an undesired component. In the second stage, the solid phase absorbs the undesired component from the gas phase, and in the third stage the solid phase is separated from the gas phase. Typical rotational speed of the separator part is higher than 12000 RPM. The volumetric flow rate equals 50 m3 hr-1 and the reactor volume equals 4 dm3. Industrial applications have not been reported in literature (Mondt et al., 2004).

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2.4. Low shear rotating disc reactors The rotating biological contactor (or rotating disc gas-liquid contactor) consists of a number of discs which are mounted on a central horizontal shaft (Figure 10). The stack of discs with a low inter-disc distance yields an equal geometry as a cylinder. The discs are partially immersed in the liquid (Zhevalkink et al., 1978). A liquid film is brought upwards over the surface of the discs, upon rotation of the discs. Further rotation will supply the liquid film to the bulk liquid. Typical rotation speeds are below 50 RPM. Due to the low rotational speed and the low power consumption of the motor the reactor is suited for a bulk process, like wastewater treatment.

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Several other low shear rotating disc reactors are applied on either lab scale or pilot scale. The rotating disc CVD reactor has solely applied to achieve a uniform layer thickness in chemical vapor deposition processes (Coltrin et al., 1989). More recent research describes a rotating disc photoelectrocatalytic reactor (Li et al., 2012a).

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2.5. High shear rotating disc reactor 2.5.1. Thin-film spinning disc The thin-film spinning disc reactor was commercially developed with the aim of intensifying gasliquid reactions (Jachuck et al., 1997), and finds its base in a patent from 1927 (Buhtz, 1927). In this reactor, the liquid phase is fed at the center of the rotating disc from where it flows radially outwards as a thin film over the disc. Extensive research has been done with respect to the gas-liquid mass transfer rate (Aoune and Ramshaw, 1999; Sisoev et al., 2006), liquid-solid mass transfer rate (Burns and Jachuck, 2005; Peev et al., 2007a), and the heat transfer rates (Harmand et al., 2013). The thinfilm spinning disc reactor has been used for polymerizations (Boodhoo and Jachuck, 2000), active pharmaceutical ingredient production (Oxley et al., 2000), crystallization (Pask et al., 2012; Tai et al., 2007), and nanoparticle synthesis (Chu et al., 2007; de Caprariis et al., 2012; Liu et al., 2012). The characteristics of the liquid film, i.e. the film thickness and radial velocity and residence time distribution, have been described numerically and experimentally (Mohammadi and Boodhoo, 2012). With increasing rotational disc speed the liquid film thickness decreases with increasing radius (Burns et al., 2003; Wood and Watts, 1973). 13

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The radial velocity increases with increasing rotational disc speed and increasing liquid flow rate, and decreases with increasing viscosity (Wood and Watts, 1973). Surface waves have been observed in the liquid film (Aoune and Ramshaw, 1999) which are classified as concentric waves or spiral waves (Sisoev et al., 2006), and are of major importance to the gas-liquid mass transfer rate (Peev et al., 2007b). Whereas in most studies a smooth disc is used as the rotor, also rotating discs with surface modifications have been used, which allow for a further increase of the gas-liquid mass transfer rate (Sisoev et al., 2006).

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Recently, two studies independently discussed the micromixing efficiency of the thin-film spinning disc reactor. Both studies concluded that this reactor type is a valuable intensified mixer (Boodhoo and Al-Hengari, 2012; Jacobsen and Hinrichsen, 2012), with a particular importance for the application in very fast reactions that need large heat dissipation (like nitrations, sulfonations, Darzens processes, crystallizations, and exothermic condensations) (Boodhoo, 2012). The scale-up of the thin-film spinning disc reactor is discussed in detail elsewhere (Boodhoo, 2012). The desired method of scale-up is a counter play of the residence time and the liquid film thickness. Both are determined by the rotational disc speed and the disc radius.

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2.5.2. Rotor-stator spinning disc reactor A high-shear multiphase rotating reactor that is developed as an improvement of the thin-film spinning disc reactor is the rotor-stator spinning disc reactor (van der Schaaf and Schouten, 2011). This reactor consists of a rotating disc (rotor) which is located between two stationary discs (stators). The axial distance between the rotor and the stators is typically in the range of 1.0 10-3 m (Figure 12). The rotational disc speed is most often around 1000 RPM but rotational speeds of 4500 RPM have been reported (Meeuwse et al., 2012). Pulse wise residence time distribution measurements have shown that the single phase fluid flow can be described by a combination of a plug flow volume and three ideally mixed volumes in series (Visscher et al., 2012c). The ideally mixed volumes in series originate from the boundary layer formation on the rotor and the stators (van Eeten et al., 2012; Visscher et al., 2012c). The gas-liquid mass transfer rate was measured as a function of rotational speed (Meeuwse et al., 2010b), rotor diameter and rotor-stator distance (Meeuwse et al., 2009), and feed location of the gas phase (Meeuwse et al., 2009). When the gas phase is fed through an inlet in the bottom stator located at the rim of the rotor, the gas bubble is sheared off due to the shear stress that is present in the cavity between the rotor and the stator. The gas bubble diameter decreases with increasing rotational disc speed typically below 1 10-3 m for rotational speeds above 1000 RPM (Meeuwse et al., 2010b), as a result gas-liquid mass transfer rates are reported of 0.43 m3L m3R s-1 at a gas flow rate of 7.3 10-6 m s-1.

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Typically the gas-hold up is below 5%. The volumetric mass transfer per unit volume of gas is 40 times higher than is reported for a bubble column. A catalytic coating was applied to the rotor (aL = 274 m2I m-3R ) in order to study the reaction and liquid-solid mass transfer in series. The volumetric liquid-solid mass transfer coefficient is one order of magnitude higher compare to values reported for packed beds (Meeuwse et al., 2010a). The volumetric mass transfer coefficient is increased by a factor 3 when the rotor radius is increased from 0.066 m to 0.135 m whereas the energy input increases by a factor 15 (Meeuwse et al., 2011). The preferred method to scale-up this reactor without changing the volumetric throughput is therefore not to increase the diameter but to increase the number of rotors-in-series (Meeuwse et al., 2012).

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The overall liquid-liquid mass transfer rate in the spinning disc was measured using an extraction system with water as the continuous phase and n-heptane as the dispersed phase, in combination with benzoic acid as the solute (Visscher et al., 2011). The liquid-liquid mass transfer rate equals 0.17 m3org m3R s-1 at 300 RPM, in this particular case both phases have plug flow behavior in the reactor. This value increases to 51.47 m3org m3R s-1 at 1600 RPM, in this case both liquids obey ideally mixed behavior. These mass transfer rates are at least 25 times higher compared to those in packed columns and at most 15 times higher compare to those measured in state-of-the-art microchannels (Visscher et al., 2012d). γ-ray tomography measurements have shown the volume fraction of the dispersed phase is closely related to the ratio of the dispersed phase volumetric flow rate over the total flow rate (Visscher et al., 2012b).

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A patented prototype of rotor-impeller-rotor spinning disc equipment has been reported in literature which exploits these high liquid-liquid mass transfer rate with the aim to intensify liquid-liquid contacting equipment (van der Schaaf et al., 2012; Visscher et al., 2012a). The height equivalent of a theoretical stage in the extractor equals 1.4 10-2 m (Visscher et al., 2013), which is a factor 10 higher when compared to other rotating equipment applied for liquid-liquid contacting. 2.6. Remaining reactor types 2.6.1. Shockwave power reactor The shockwave power reactor was patented in 1993 (Griggs, 1993), and consists of a spinning rotor that is baffled such that dead ended cavities are present at the rim of the rotor Figure 13 and Figure 14. The rotation creates locally a low pressure zone in the cavities which collapses under the emission of an energy wave in the surrounding liquid, called the shockwave. As a result gas-liquid mass transfer rates up to 5.2 s-1 have been obtained (Mancosky, 2013), which is about 25 times higher than in a mechanically agitated tank. The volumetric flow rate may vary from 227 10-3 m3 hr-1 to 1.14 m3 hr-1, with typical rotational speeds of the Shockwave power reactor range up to 3600 RPM. A challenge in the design of this reactor is the lifetime of the rotor and the cavities. The combination of cavitation with corrosive chemicals implies that the rotor needs to be coated with an extraordinary material to resist such harsh conditions. 15

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2.6.2. RAPTOR RAPTOR is an abbreviation of “Réacteur Agité Polyvalent à Transfert Optimisé Rectiligne”, or Reactor with Polyvalent Rectilinear Stirred Reactor with Optimized Transfer (Bahroun et al., 2010). The technology is developed by the AET group and was presented in 2007, and is shown in Figure 17 (Barillon et al., 2007). The exact geometry of the reactor is only scarcely disclosed in the literature (Bahroun et al., 2010). Rotational speeds of the inner part do not exceed 1500 RPM. The small reactor volume (smaller than 2.9 10-3 m³ (Milly et al., 2008)) allows for the application of temperatures up to 300oC and an operating pressure of 250 bar. The maximum flow rate is 0.15 m3 hr-1, with a residence time between 15 seconds and 10 minutes. The heat transfer area per reactor volume equals 150 m2m-3 which is at least 30 times higher than for an industrial scale stirred tank reactor which is batch wise operated (5 m2 m-3).

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2.6.3. Dynamically rotating axis micro reactor A unique combination of micro reactor technology and mechanical agitation is used in the dynamically rotating axis micro reactor. This reactor consists of two cylinders of which one is rotating. The inner cylinder has internal diameter of 8 10-3 m and the spacing between the two cylinders is 100 10-6 m (Figure 15, (Ogura et al., 2008)). The volumetric flow rate is typically in the range of 80 10-9 m3 s-1. The rotation speed of the outer cylinder is limited to 3600 RPM. Due to the high shear stress in the small radial gap between the 100 10-3 m long cylinders, high mass transfer rates are achieved for liquid-liquid applications, as has been demonstrated by generating emulsions of water and o-chlorobenzene, and the nitration of naphthalene (Ogura et al., 2013). Literature on this reactor type is too scarce to make a valid estimation of its feasibility.

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2.6.4. Coflore agitated cell reactor The Coflore Agitated Cell Reactor (ACR), is commercially available since 2006 and consists of two separate parts: a reactor block and a lateral moving plate (Jones et al., 2012). Each reactor block consists of a channel that connects ten consecutive interlinked reactor chambers. In each chamber a non-fixed mixing element is present (Figure 16). Upon moving the plate, the non-fixed mixing element enhances the dispersion in the reactor chambers. The moving plate causes the liquid in each reactor chamber to behave as ideally mixed flow; as a result plug flow behavior is mimicked in a single reactor block. The design of the non-fixed mixing element opens up the application for gas-liquid contacting as well as slurry handling. Catalysts can be loaded into the non-fixed mixing element. The reactor block volume may vary from 20 10-6 m3 to 100 10-6 m3, volumetric liquid flow rate equals up to 10 10-3 m3 s-1. Adjustments can be made to allow for counter-current liquid-liquid extraction over 6 stages in a single extractor block. Technically this reactor does not belong to this review since it’s shaken, not stirred. When the reactor is mounted on the moving plate the unit has a small footprint; accordingly the reactor is well applicable in laboratories and in fume hoods. 16

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The feasibility of various industrial applications have been reported by the manufacturer, including but limited to the Hoffman reaction, Suzuki reaction, polymerizations and Grignard reactions. The greatest benefit to conventional batch reactors is the easy scalability of the reactor. A challenge for further industrial implementation of this reactor will be the prevention of fouling, for processes in which solids build-up might occur.

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2.6.5. Rotating cone reactor The rotating cone reactor was developed at the University of Twente. In this reactor a solid phase is fed to the bottom inlet of a rotating cone, from where they travel spiral-wise upwards. The inner diameter of the cone has a maximum of 0.650 m. The rotational speed of the cone is 900 RPM. The reactor volume may range from 2 – 200 dm3. The combined feed rate of the gas and solid phase equals 13 kg hr-1. The first industrial test case for which this reactor is used was the pyrolysis of biomass (Wagenaar et al., 1994). More recently the flow behavior was presented (Leung, 2009), as well as the feasibility of the rotating cone reactor for the epoxidation of soybean oil (Chen et al., 2011).

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2.6.6. Rotating membrane reactor Rotating membranes find their application in ultrafiltration (Sarkar et al., 2012; Sarkar et al., 2009). A theoretical comparison of rotating and stationary membrane disk filters was published in 2000 (Serra and Wiesner, 2000). An in-depth review on this matter was published by Jaffrin (Jaffrin, 2008). To obtain sufficient surface area multidisc membrane units are mounted on a single shaft.

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Examples of equipment are the DMF (Pall Corp, USA), Dyno (Bokela, Germany), Optifilter (ABB Flootek, Finland) and the SpinTek (Huntington, USA) (Jaffrin, 2008). The membrane area per module may reach 150 m² (Bläse et al., 2006). Membrane diameters are typically below 1 m. The power required for rotation is below 30 kW (Jaffrin, 2008). The availability of ceramic membranes will facilitate the future development of rotating membrane reactors, its application for the production of galacto-oligosaccarides should an productivity increase by a factor two (Sen et al., 2012). 3. Summary and outlook This review-perspective paper describes the current state-of-the-art in the field of rotating reactors. An overview is given of the development of rotating reactors over time. The paper has a focus on rotating reactor technology with applications in lab scale, pilot scale and industrial chemical reaction engineering. Rotating reactors are classified according to their geometry. The reactors are classified as stirred tanks, tubes, discs and miscellaneous reactors. Their main advantages and disadvantages are presented, including the typical operational conditions (residence time, rotational speed, energy consumption). An overview is given in Table 1. The energy dissipation mentioned in Table 1 represent the maximum energy consumed by the motor(s) that propel the rotating element(s). This is not equal to the energy that is dissipated locally in the liquids. Where possible an accurate 17

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approximation of this difference is given. However, for many reactors the energy that is locally dissipated in the reactants is not well defined. As an indication the power to volume ratio could be studied. From such a ratio the Shockwave power reactor has high local energy dissipation, whereas a large scale stirred tank has a low power to volume ratio. Table 1 thus has the function of an orientation on the maximum power that is consumed by the reactors described.

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In academia there is a tendency to develop multifunctional modules suited for reactive extractions, one-pot synthesis, and juke-box like functionalities where multiple separations steps are combined in one compact mini-plant. The further development of such equipment requires that large scale equipment manufacturers, end-product consumers, prototype experts, and reactor scientists collaborate together from the start in research projects dealing with reactor development.

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A promising route to bridge the gap between lab-scale reactor development and industrial implementation of new, intensified reactors is the process of knowledge valorization. One method to achieve such an industrial implementation is the startup of spinoff companies (De Cleyn and Braet, 2009). Examples of such spinoff companies are Transatomic Power (molten-salt reactors), FLOWID (micro reactor technology), SOWARLA (wastewater treatment), and SPINID (rotorstator and rotor-rotor spinning technology). The presence of resource incubators at research institutes and universities will facilitate and speed up the development of such companies.

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Introducing new reactor concepts in chemical engineering is a lengthy process. Industry is reluctant to introduce novel chemical reactor types in existing processes when replacement is not essential, with the aim to minimize possible risks to plant performance. The introduction thus depends mainly on the development of new processes and new production plants.

AL DI Fr Re Sc dR

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A strong argument for the introduction of novel reactor types will come from the experience that is built up during public-private research projects in which both academic and industrial partners are collaborating. During such projects detailed information of chemical resistance, heat- and mass transfer performance, mechanical durability of rotating parts, and energy consumption can be collected.

4. List of symbols Surface area, m2 m-3 Impeller diameter, m Froude,Reynolds,Schmidt, Inner diameter, -

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5. References Agarwal, L., Pavani, V. et al., 2010, Process Intensification in HiGee Absorption and Distillation: Design Procedure and Applications, Ind. Eng. Chem. Res., 49(20), 10046-10058 Aiba, S., 1958, Flow patterns of liquids in agitated vessels, AlChE J., 4(4), 485-489 Al-Rawashdeh, M., Yu, F. et al., 2012, Numbered-up gas-liquid micro/milli channels reactor with modular flow distributor, Chem. Eng. J., 207208(0), 645-655 Albers, R. K. E., Houterman, M. J. J. et al., 1998, Novel monolithic stirred reactor, AlChE J., 44(11), 2459-2464 Anderson, N. G., 2012, Using continuous processes to increase production, Org. Process Res. Dev., 16(5), 852-869 Aoune, A. and Ramshaw, C., 1999, Process intensification: heat and mass transfer characteristics of liquid films on rotating discs, Int. J. of Heat and Mass Trans., 42(14), 2543-2556 Ascanio, G., Castro, B. et al., 2004, Measurement of power consumption in stirred vesselsA review, Chem. Eng. Res. Des., 82(9), 1282-1290 Ashcraft, R. W., Heynderickx, G. J. et al., 2012, Modeling fast biomass pyrolysis in a gassolid vortex reactor, Chem. Eng. J., 207-208(0), 195-208 Ashcraft, R. W., Kovacevic, J. et al., 2013, Assessment of a gassolid vortex reactor for SO2/NOx adsorption from Flue Gas, Ind. Eng. Chem. Res., 52(2), 861-875 Assirelli, M., Bujalski, W. et al., 2002, Study of micromixing in a stirred tank using a rushton turbine: comparison of feed positions and other mixing devices, Chem. Eng. Res. Des., 80(8), 855863 Bahroun, S., Li, S. et al., 2010, Control and optimization of a three-phase catalytic slurry intensified continuous chemical reactor, Journal of Process Control, 20(5), 664-675 Baldyga, J. and Pohorecki, R., 1995, Turbulent micromixing in chemical reactors -- a review, T. Chem. Eng. J. Biochem. Eng. J., 58(2), 183-195 Barillon, O., De Panthou, F., Marie, S., Trani, A., Falk, L., and Jenck, J., 1st European Process Intensifcation Conference,19-9-2007,Copenhagen, Denmark. Bawadi, Abdullahet al., Chapter Synthetic Liquids Production and RefiningAmerican Chemical Society, Bennett, C. J., Kolaczkowski, S. T. et al., 1991, Determination of heterogeneous reaction kinetics and reaction rates under mass transfer controlled conditions for a monolith reactor, Trans. Ins. Chem. Eng., B69(4), 209-220 Berty, J. M., 1974, Reactor for vapor-phase catalytic studies, Chem. Eng. Prog., 70(5), 78 Bläse, D., Feuerpeil, H. P., and Olapinski, H., EP1299177 B1, Rotating filter, Date of patent: 8/11/2006 Boodhoo, K. V. K., Chapter 3 in: Process Intensification Technologies for Green Chemistry: Engineering Solutions for Sustainable Chemical Processing, Eds. Boodhoo, K. V. K. and Harvey, A. (Wiley, ChiChester, UK). Boodhoo, K. V. K. and Al-Hengari, S. R., 2012, Micromixing Characteristics in a Small-Scale Spinning Disk Reactor, Chem. Eng. Tech., 35(7), 1229-1237 Boodhoo, K. V. K. and Jachuck, R. J., 2000, Process intensification: spinning disk reactor for styrene polymerisation, App. Therm. Eng., 20(12), 1127-1146 Boodhoo, K. V. K., Cartwright, C. D. et al., 2010, Development of a HIGEE bioreactor (HBR) for production of polyhydroxyalkanoate: Hydrodynamics, gasliquid mass transfer and fermentation studies, CEP:PI, 49(7), 748-758 19

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Bourne, J. R. and Studer, M., 1992, Fast reactions in rotor-stator mixers of different size, CEP:PI, 31(5), 285-296 Buhtz, E., US1629200, Method of and apparatus for carrying out chemical reactions or physical processes, Date of patent: 17/5/1927 Burns, J. R. and Jachuck, R. J. J., 2005, Determination of liquid-solid mass transfer coefficients for a spinning disc reactor using a limiting current technique, Int. J. of Heat and Mass Trans., 48(12), 2540-2547 Burns, J. R. and Ramshaw, C., 1996, Process intensification: Visual study of liquid maldistribution in rotating packed beds, Chem. Eng. Sci., 51(8), 1347-1352 Burns, J. R., Ramshaw, C. et al., 2003, Measurement of liquid film thickness and the determination of spin-up radius on a rotating disc using an electrical resistance technique, Chem. Eng. Sci., 58(11), 2245-2253 Carberry, J. J., 1964, Desiging laboratory catalytic reactors, Ind. Eng. Chem., 56(11), 39-46 Chandra, A., Goswami, P. S. et al., 2005, Characteristics of flow in a rotating packed bed (HIGEE) with split packing, Ind. Eng. Chem. Res., 44(11), 4051-4060 Charpentier, J. C. and McKenna, T. F., 2004, Managing complex systems: some trends for the future of chemical and process engineering, Chem. Eng. Sci., 59(8Çô9), 1617-1640 Chaudhari, R. V. and Mills, P. L., 2011, Multiphase catalysis and reaction engineering for emerging pharmaceutical processes, Chem. Eng. Sci., 59(22-23), 5337-5344 Chen, J., Jiang, J. et al., 2011, Catalytic cracking of soybean oil for production of renewable fuel using rotating cone reactor, Taiyangneng Xuebao/Acta Energiae Solaris Sinica, 32(3), 354357 Chen, Y. M., 1987, Fundamentals of a centrifugal fluidized bed, AlChE J., 33(5), 722-728 Chen, Z., Xiong, S. et al., 1995, Helical rotating absorber, J. Chem. Ind. Eng., 46, 388-392 Cho, B. K., Carr, J. et al., 1980a, A continuous chromatographic reactor, Chem. Eng. Sci., 35(12), 74-81 Cho, B. K., Carr, R. W. et al., 1980b, A new continuous flow reactor for simultaneous reaction and separation, Sep. Sci. Technol., 15(3), 679-696 Chu, G. W., Song, Y. H. et al., 2007, Micromixing efficiency of a novel rotor-stator reactor, Chem. Eng. J., 128(2-3), 191-196 Cohen, S. and Marom, D. M., 1983, Experimental and theoretical study of a rotating annular flow reactor, The. Chem. Eng. J., 27(2), 87-97 Coltrin, M. E., Kee, R. J. et al., 1989, A mathematical model of the fluid mechanics and gasphase chemistry in a rotating disk chemical vapor deposition reactor, J. Elec. Soc., 136(3), 819-829 Cowen, G., Morton-Berry, P., and Steel M.L., US4311570, Chemical process on the surface of a rotating body, Date of patent: 19/1/1982 Cybulski, A. and Moulijn, J., Structured catalysts and reactors, 2nd, 2005 (Taylor & Francis, Dahl, J. A., Maddux, B. L. S. et al., 2007, Toward greener nanosynthesis, Chemical Reviews, 107(6), 2228-2269 de Broqueville, A. and De Wilde, J., 2009, Numerical investigation of gas-solid heat transfer in rotating fluidized beds in a static geometry, Chem. Eng. Sci., 64(6), 1232-1248 de Caprariis, B., Di Rita, M. et al., 2012, Reaction-precipitation by a spinning disc reactor: Influence of hydrodynamics on nanoparticles production, Chem. Eng. Sci., 76(0), 73-80 De Cleyn, S. H. and Braet, J., 2009, Research valorisation through spin-off ventures: integration of existing concepts and typologies, World Review of Entrepreneurship, Management and Sustainable Development, 5(4), 325-352 20

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De Lathouder, K. M., Bakker, J. J. W. et al., 2006, Structured reactors for enzyme immobilization: A monolithic stirrer reactor for application in organic media, Chem. Eng. Res. Des., 84(5), 390398 De Wilde, J. and de Broqueville, A., 2007, Rotating fluidized beds in a static geometry: Experimental proof of concept, AlChE J., 53(4), 793-810 De Wilde, J. and de Broqueville, A., 2008a, Experimental investigation of a rotating fluidized bed in a static geometry, Powder Technology, 183(3), 426-435 De Wilde, J. and de Broqueville, A., 2008b, Experimental study of fluidization of 1G-Geldart D-type particles in a rotating fluidized bed with a rotating chimney, AlChE J., 54(8), 2029-2044 Deen, N. G., van Sint Annaland, M. et al., 2004, Multi-scale modeling of dispersed gas-liquid twophase flow, Chem. Eng. Sci., 59(8Çô9), 1853-1861 Doraiswamy, L. K. and Tajbl, D. G., 1974, Laboratory catalytic reactors, Cat. Rev., 10(1), 177-219 Dudukoviç, M. P., Larachi, F. et al., 2002, Multiphase catalytic reactors: a perspective on current knowledge and future trends, Cat. Rev., 44(1), 123-246 Foust, H. C., Mack, D. E. et al., 1944, Gas-liquid contacting by mixers, Ind. Eng. Chem., 36(6), 517522 Froment, G. F., DeWilde, J., and Bischoff, K., Chemical Reactor anaylysis and design, 3rd edition, 2011 (Wiley, Hoboken, NJ, USA). Gonzalez, M. A. and Ciszewski, J. T., 2008, High conversion, solvent free, continuous synthesis of imidazolium ionic liquids in spinning tube-in-tube reactors, Org. Process Res. Dev., 13(1), 64-66 Griggs, J. L., US5188090, Apparatus for heating fluids, Date of patent: 23/2/1993 Hampton, P. D., Whealon, M. D. et al., 2008, Continuous organic synthesis in a spinning tube-intube Reactor: TEMPO-catalyzed oxidation of alcohols by hypochlorite, Org. Process Res. Dev., 12(5), 946-949 Harish Kumar, S. and Murthy, D. V. R., 2010, Minimum superficial fluid velocity in a gassolid swirled fluidized bed, CEP:PI, 49(10), 1095-1100 Harmand, S., Pelle, J. et al., 2013, Review of fluid flow and convective heat transfer within rotating disk cavities with impinging jet, Int. J. Therm. Sci., 67, 1-30 Harvey, A. P., Mackley, M. R. et al., 2001, Operation and optimization of an oscillatory flow continuous reactor, Ind. Eng. Chem. Res., 40(23), 5371-5377 Hemrajani, R. R. and Tatterson, G. B., Chapter 6 in: Handbook of Industrial Mixing - Science and Practice, Eds. Paul, E. L., Atiemo-Obeng, V. A., and Kresta, S. M. (John Wiley & Sons, Hoboken, New Jersey). Hessel, V., 2009, Novel process windows - Gate to maximizing process intensification via flow chemistry, Chem. Eng. Tech., 32(11), 1655-1681 Hessel, V., Renken, A., Schouten, J. C., and Yoshida, J., Micro process engineering: A comprehensive handbook, 2009 (Wiley, Weinheim, Germany). Hill, D. F., Sharp, K. V. et al., 2000, Stereoscopic particle image velocimetry measurements of the flow around a Rushton turbine, Exp. Fluids., 29(5), 478-485 Hoek, I., Chapter 2 in: Towards the catalytic application of a monolithic stirrer reactor, PhD-thesis, TU Delft, The Netherlands Hoek, I., Chapter 3 in: Towards the catalytic application of a monolithic stirrer reactor, PhD-thesis, TU Delft, The Netherlands Hoek, I., Nijhuis, T. A. et al., 2004, Performance of the monolithic stirrer reactor: applicability in multi-phase processes, Chem. Eng. Sci., 59(2223), 4975-4981 21

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Ogura, T, Ohta, T, Takahashi, Y, and Mae, K., World Conference of Chemical Engineering - 8,23-82009,Montreal, Quebec, Canada. Ogura, T., Ohta, T., Takahashi, Y., and Mae, K., AIChE Spring meeting & IMRET-10,2008,New Orleans,Louisiana, USA. Ordomsky, V. V., Schouten, J. C. et al., 2012a, Foam supported sulfonated polystyrene as a new acidic material for catalytic reactions, Chem. Eng. J., 207208(0), 218-225 Ordomsky, V. V., Schouten, J. C. et al., 2012b, Zirconium phosphate coating on aluminum foams by electrophoretic deposition for acidic catalysis, ChemCatChem, 4(1), 129-133 Oxley, P., Brechtelsbauer, C. et al., 2000, Evaluation of spinning disk reactor technology for the manufacture of pharmaceuticals, Ind. Eng. Chem. Res., 39(7), 2175-2182 Pan, Z. Q., Zhang, Y. J. et al., 2006, Experimental investigation into mass transfer between liquid and gas in multi-staged spraying rotating packed bed, J. South China Univ. Technol., 34(3), 67-71 Pask, S. D., Nuyken, O. et al., 2012, The spinning disk reactor: an example of a process intensification technology for polymers and particles, Polym. Chem., 3(10), 2698-2707 Patwardhan, A. W. and Joshi, J. B., 1998, Design of gas-inducing reactors, Ind. Eng. Chem. Res., 38(1), 49-80 Pavko, A., Misic, D. M. et al., 1981, Kinetics in three-phase reactors, The. Chem. Eng. J., 21(2), 149154 Peev, G., Nikolova, A. et al., 2007a, Solid dissolution in a thin liquid film on a horizontal rotating disk, Heat Mass Transfer, 43(4), 397-403 Peev, G., Peshev, D. et al., 2007b, Gas absorption in a thin liquid film flow on a horizontal rotating disk, Heat Mass Transfer, 43(8), 843-848 Pilo, C. W., US2941872, Apparatus for intimate contacting of two fluids, Date of patent: 21/6/1960 Pliny the Elder, Naturalis Historia : The Elder Pliny's chapters on chemical subjects, Translated by Bailey, K.C., 1929 (Arnold, London, UK). Pudjiono, P. I. and Tavare, N. S., 1993, Residence time distribution analysis from a continuous couette flow device around critical taylor number, Can. J. Chem. Eng., 71(2), 312-318 Qi, G. S., Liu, Y. Z. et al., 2008, Experimental research on extraction of acetic acid from dilute solution by chemical complexation impinging stream-rotating packed bed, Xiandai Huagong/Modern Chemical Industry, 28(11), 65-67 Ramshaw, C., 1983, HIGEE distillation-an example of process intensification, Chem. Eng., 389, 1314 Ramshaw, C. and Mallinson, R. H., US4283255, Mass transfer process, Date of patent: 11/8/1981 Rao, D. P., Bhowal, A. et al., 2004, Process intensification in rotating packed beds (HIGEE): An Appraisal, Ind. Eng. Chem. Res., 43(4), 1150-1162 Reay, D., Ramshaw, C., and Harvey, A., Process intensification, engineering for efficiency, sustainability and flexibility, 1, 2008 (Butterworth-Heinemann, Oxford). Reddy, K. J., Gupta, A. et al., 2006, Process Intensification in a HIGEE with Split Packing, Ind. Eng. Chem. Res., 45(12), 4270-4277 Richter, O., Hoffmann, H. et al., 2008, Effect of the rotor shape on the mixing characteristics of a continuous flow Taylor-vortex reactor, Chem. Eng. Sci., 63(13), 3504-3513 Rivero, E. P., Granados, P. et al., 2010, Mass transfer modeling and simulation at a rotating cylinder electrode (RCE) reactor under turbulent flow for copper recovery, Chem. Eng. Sci., 65(10), 3042-3049 24

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Rooze, J., Rebrov, E. V. et al., 2013, Dissolved gas and ultrasonic cavitation - A review, Ultrason. Sonochem., 20(1), 1-11 Rosales Trujillo, W. and De Wilde, J., 2012, Fluid catalytic cracking in a rotating fluidized bed in a static geometry: a CFD analysis accounting for the distribution of the catalyst coke content, Powder Technology, 221(0), 36-46 Rousseaux, J. M., Muhr, H. et al., 2001, Mixing and micromixing times in the forced vortex region of unbaffled mixing devices, Can. J. Chem. Eng., 79(5), 697-707 Rushton, J. H., Costich, E. W. et al., 1950a, Power characteristics of mixing impellers, Part 1, Chem. Eng. Progr. , Part III, 46(8), 395-404 Rushton, J. H., Costich, E. W. et al., 1950b, Power characteristics of mixing impellers, Part 2, Chem. Eng. Progr. , Part III, 46(9), 467-476 Saravanan, K., Mundale, V. D. et al., 1994, Gas inducing type mechanically agitated contactors, Ind. Eng. Chem. Res., 33(9), 2226-2241 Sarkar, A., Moulik, S. et al., 2012, Performance characterization and CFD analysis of a novel shear enhanced membrane module in ultrafiltration of Bovine Serum Albumin (BSA), Desalination, 292(0), 53-63 Sarkar, P., Ghosh, S. et al., 2009, Effect of different operating parameters on the recovery of proteins from casein whey using a rotating disc membrane ultrafiltration cell, Desalination, 249(1), 5-11 Sarmidi, M. R. and Barker, P. E., 1993a, Saccharification of modified starch to maltose in a continuous rotating annular chromatograph (CRAC), J. Chem. Tech. Bio., 57(3), 229-235 Sarmidi, M. R. and Barker, P. E., 1993b, Simultaneous biochemical reaction and separation in a rotating annular chromatograph, Chem. Eng. Sci., 48(14), 2615-2623 Schuur, B., Kraai, G. N. et al., 2012, Hydrodynamic features of centrifugal contactor separators: Experimental studies on liquid hold-up, residence time distribution, phase behavior and drop size distributions, CEP:PI, 55(0), 8-19 Sen, D., Sarkar, A. et al., 2012, Batch Hydrolysis and Rotating Disk Membrane Bioreactor for the Production of Galacto-oligosaccharides: A Comparative Study, Ind. Eng. Chem. Res., 51(32), 10671-10681 Serra, C. A. and Wiesner, M. R., 2000, A comparison of rotating and stationary membrane disk filters using computational fluid dynamics, Journal of Membrane Science, 165(1), 19-29 Sisoev, G. M., Matar, O. K. et al., 2006, The flow of thin liquid films over spinning discs, Can. J. Chem. Eng., 84(6), 625-642 Ståhl Wernersson, E. and Trägårdh, C., 1999, Scale-up of Rushton turbine-agitated tanks, Chem. Eng. Sci., 54(19), 4245-4256 Stankiewicz, A. I. and Moulijn, J. A., 2000, Process intensification: Transforming chemical engineering, Chem. Eng. Prog., 96(1), 22-33 Stankiewicz, A., 2003, Reactive separations for process intensification: an industrial perspective, CEP:PI, 42(3), 137-144 Stemmet, C. P., Jongmans, J. N. et al., 2005, Hydrodynamics of gasliquid counter-current flow in solid foam packings, Chem. Eng. Sci., 60(22), 6422-6429 Stemmet, C. P., van der Schaaf, J. et al., 2006, Solid foam packings for multiphase reactors: modelling of liquid holdup and mass transfer, Chem. Eng. Res. Des., 84(12), 1134-1141 Stitt, E. H., 2002, Alternative multiphase reactors for fine chemicals: A world beyond stirred tanks?, Chem. Eng. J., 90(1-2), 47-60 25

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Ströhleinet al., Chapter Integrated Chemical Processes, Eds. Sundmacher, K., Kienle, A., and SeidelMorgenstern, A. (Wiley, Tai, C. Y., Tai, C. T. et al., 2007, Synthesis of magnesium hydroxide and oxide nanoparticles using a spinning disk reactor, Ind. Eng. Chem. Res., 46(17), 5536-5541 Trambouze, P. and Euzen, J. P, Chapter 4 in: Chemical Reactors, From design to operationInstitute Français du pétrole publications, Paris, Fr). Trambouze, P. and Euzen, J.-P., Chemical Reactors, Technip, 2002b (Insititut Français du pétrole publications, Paris, Fr.). Tschentscher, R., Nijhuis, T. A. et al., 2010a, Gasliquid mass transfer in rotating solid foam reactors, Chem. Eng. Sci., 65(1), 472-479 Tschentscher, R., Spijkers, R. J. P. et al., 2010b, Liquidsolid mass transfer in agitated slurry reactors and rotating solid foam reactors, Ind. Eng. Chem. Res., 49(21), 10758-10766 Tschentscher, R., Schubert, M. et al., 2012, Gas holdup of rotating foam reactors measured by tomography - effect of solid foam pore size and liquid viscosity, AlChE J., DOI: 10.1002/aic.13787 Uretschläger, A. and Jungbauer, A., 2002, Preparative continuous annular chromatography (P-CAC), a review, Bioproc. Biosys. Eng., 25(2), 129-140 van der Schaaf, J. and Schouten, J. C., 2011, High-gravity and high-shear gas-liquid contactors for the chemical process industry, Curr. Opin. Chem. Eng., 1(1), 84-88 van der Schaaf, J., Visscher, F., Bindraban, D., and Schouten, J. C., WO/2012/150226 A1, Device for multi phase and single phase contacting, Date of patent: 30/4/2012 van Eeten, K. M. P., van der Schaaf, J. et al., 2012, Boundary layer development in the flow field between a rotating and a stationary disk, Physics of Fluids, 24(3), 033601-033618 Vedantam, S. and Joshi, J. B., 2006, Annular centrifugal contactorsA review, Chem. Eng. Res. Des., 84(7), 522-542 Vedantam, S., Joshi, J. B. et al., 2006, Three-Dimensional CFD Simulation of Stratified Two-Fluid Taylor-Couette Flow, Can. J. Chem. Eng., 84(3), 279-288 Villermaux, J., 1988, The role of energy dissipation in contacting and mixing devices, Chem. Eng. Tech., 11(1), 276-287 Visscher, F., Nijhuis, R. T. R. et al., 2012a, Liquid-liquid flow in an impeller-stator spinning disc reactor, CEP:PI, DOI:10.1016/j.cep.2013.01.015 Visscher, F., Saffarionpour, S. et al., 2013, Counter-current liquid-liquid contacting in an spinning disc reactor, Unpublished Work Visscher, F., Bieberle, A. et al., 2012b, Water and n-heptane volume fractions in a rotor-stator spinning disc reactor, Ind. Eng. Chem. Res., 51(51), 16670-16676 Visscher, F., de Hullu, J. et al., 2012c, Residence time distribution in a single phase rotor-stator spinning disc reactor, AlChE J., 59(7), 2686-2693 Visscher, F., Gaakeer, W. A. et al., 2011, Liquid-liquid extraction systems of benzoic acid in water and heptane, methylbenzene, or trichloroethylene as cosolvent, J. Chem. Eng. Data., 56(9), 3630-3636 Visscher, F., van der Schaaf, J. et al., 2012d, Liquid-liquid mass transfer in a rotorstator spinning disc reactor, Chem. Eng. J., 185-186(0), 267-273 Wagenaar, B. M., Prins, W. et al., 1994, Pyrolysis of biomass in the rotating cone reactor: modelling and experimental justification, Chem. Eng. Sci., 49(24, Part 2), 5109-5126 Wang, G. Q., Xu, Z. C. et al., 2011, Progress on Higee distillation--Introduction to a new device and its industrial applications, Chem. Eng. Res. Des., 89(8), 1434-1442 26

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Wang, G. Q., Xu, Z. C. et al., 2008, Performance of a rotating zigzag bedA new HIGEE, CEP:PI, 47(12), 2131-2139 Weekman, V. W., 1974, Laboratory reactors and their limitations, AlChE J., 20(5), 833-840 Wenmakers, P. W. A. M., van der Schaaf, J. et al., 2008, "Hairy Foam": carbon nanofibers grown on solid carbon foam. A fully accessible, high surface area, graphitic catalyst support, J. Mater. Chem., 18(21), 2426-2436 Wenmakers, P. W. A. M., van der Schaaf, J. et al., 2010, Comparative modeling study on the performance of solid foam as a structured catalyst support in multiphase reactors, Ind. Eng. Chem. Res., 49(11), 5353-5366 Wood, R. M. and Watts, B. E., 1973, The flow, heat and mass transfer charactersitsics of liquid films on rotating discs, Chem. Eng. Res. Des., 51(a), 315-322 Wu, H., Patterson, G. K. et al., 1989, Distribution of turbulence energy dissipation rates in a Rushton turbine stirred mixer, Exp. Fluids., 8(3-4), 153-160 Yang, H. J., Chu, G. W. et al., 2005, Micromixing efficiency in a rotating packed bed: Experiments and simulation, Ind. Eng. Chem. Res., 44(20), 7730-7737 Zhang, D., Zhang, P. Y. et al., 2010, Application of HIGEE process intensification technology in synthesis of petroleum sulfonate surfactant, CEP:PI, 49(5), 508-513 Zhao, H., Shao, L. et al., 2010, High-gravity process intensification technology and application, Chem. Eng. J., 156(3), 588-593 Zhevalkink, J. A., Kelderman, P. et al., 1978, Liquid film thickness in a rotating disc gas-liquid contactor, Water Res., 12(8), 577-581 Zieverink, M. M. P., Kreutzer, M. T. et al., 2006, Gasliquid mass transfer in benchscale stirred tanks - fluid properties and critical impeller speed for gas induction, Ind. Eng. Chem. Res., 45(13), 4574-4581

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6. Figures

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Figure 1. Classification of rotating reactors based on the geometry.

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Figure 2. The development of rotating reactors is schematically shown as a function of time. The year in which the first scientific publication (peer-reviewed or patent) is published is used as allocation in time.

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Figure 3. Different liquid flow direction for various stirrers. Stirrer A shows the flow profile generated by a radial flow impeller (Rushton stirrer), Stirrer B for a three blade propeller (Aiba, 1958). Stirrer C for an axial flow turbine in which liquid is progressively sucked in (axial downwards) near the center and is forced radially outwards. An extensive list of different stirrers is presented by (Hemrajani and Tatterson, 2004) and (Joshi et al., 1982).

Figure 4. The top view and the side view of a Rushton stirrer (left) and a side view of a monolithic stirred tank reactor (right). 29

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Figure 5. Schematic 3D-view of the foam stirrer tank. The stirrer is equipped with two equal foam blocks on which a high catalyst loading can be coated. Courtesy to (Tschentscher et al., 2010a)

Figure 6. Schematic side view of a rotating packed bed in which the vapor phase (dotted lines) and the liquid phase (solid line) are contacted counter-currently over the rotating packing (grey).

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Figure 7. Schematic side view of the rotating zigzag bed. In this reactor the packing consists of a rotating disc (grey) and a stator (black). Two series of concentric circular sheets are fixed on the rotating and the stationary disc. The circular sheets on the rotor are perforated.

Figure 8. A schematic representation of the top view of a rotating fluidized bed in a static geometry. The gas phase is injection through the tangential inlets. Due to the rotation of the injected gas phase, the solids present in the reactor will start to fluidize (Courtesy to (Rosales Trujillo and De Wilde, 2012). 31

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Figure 9. The spinning tube-in-tube reactor. The temperature can be controlled by a heat exchanger which is located around the process volume (green). Reprinted with permission from (Gonzalez and Ciszewski, 2008). Copyright 2009, American Chemical Society.

Figure 10. A schematic side view and front view of the rotating biological contactor. A stack of discs (black) is partially immersed in the liquid (grey). Rotation of the stack of discs increases the liquid film renewal at the disc surface. 32

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Figure 11. A schematic side view of the rotating annular chromatographic reactor (Sarmidi and Barker, 1993a; Stankiewicz, 2003). Component C is added at the stationary inlet. While flowing downwards, the components A, B and C are selectively recovered over the rim of the rotating cylinders.

Figure 12. The side view of the rotor-stator spinning disc reactor. The axial clearance between the rotor and the stator is in the order of millimeters. Rotation of the rotor induces a velocity gradient over the height between the rotor and the stator. As a result a shear force acts on the droplets, bubbles or particles between the rotor and the stator. This results in high liquid-liquid, liquid-solid, and gas-liquid mass transfer rates (Visscher et al., 2012c). 33

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Figure 13. Shockwave power reactor (Mancosky, 2013).

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Figure 14. Schematic view of the rotor inside the shockwave power reactor (Mancosky, 2013). Upon rotation of the rotor (light grey) liquid is forced radially outwards, thereby creating cavitation at the radial inner position of the cavities inside the rotor.

Figure 15. Dynamically rotating axis micro reactor. Courtesy to (Ogura et al., 2008).

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Figure 16. The reactorblock of the Coflore reactor. Ten consecutive mixing chambers are interconnected. Once ideally mixed flow behavior is achieved in each mixing chamber, plug flow behavior is mimicked over the reactor block. Courtesy to (Jones et al., 2012).

Figure 17. The experimental set-up of the RAPTOR. The reactor exhibits high heat and mass transfer rates and can handle flow rates with a maximum of 0.15 m3 hr-1. Courtesy to: (Barillon et al., 2007) 35

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Table 1. Rotating reactor comparison. Typical values are given of the volumetric throughput (Q, [m³ hr-1]), reactor volume (V, [m³]), residence time (τ), power consumption (E, [Wm-³]), maximum rotation speed (N, [RPM]), gas-liquid mass transfer rate (GLMT, [s-1]), overall heat transfer coefficient (U, [Wm-2K-1]), and micromixing time (tm, [s]). Phase

Q

V

τ

E

N

GLMT

U

tm

GLS GLS GL GLS GLS GL GS GL GL GL GL GLS GLS GLS GL GL GL GL

10-3 – 101 10-3 – 101 – 10-2 10-2 10-4 – 10-6 10-7 10-5 10-3 10-6 -3 10 – 10-2 10-3 – 10-1 10-1 – 101 10-5 – 10-3 10-7 10-5

10-3 – 102 10-3 – 102 10-3 10-3 10-3 10-1 10-3 10-6 -6 10 – 10-3 10-4 10-2 10-2 10-3 -4 10 – 10-2 10-3 10-6 10-5

1 min – 5 hr 1 min – 5 hr 1-100 min 1 s – 10 min 1 s – 10 min 1 s – 10 min 1 s – 10 min 1 s – 10 min 1 s – 3 min 1 – 15 min 1 s – 10 min 1 s – 60 min 1 s – 10 min 1 s – 10 min 10 s–10 min 1s 1 s – 10 min

1 104 1 104 2 103 6 103 5 104 3 104 – 3 104 – – 1 103 – 5 103 5 106 – 3 106 – –

1500 1500 800 600 2500 1400 500 1000 12000 870 30 20 5000 4500 3600 1500 3600 –

0.01 – 2 0.01 – 2 0.8 0.2 12 12 – 0.1 – – – – 6 2 5.2 – – –

103 103 103 103 105 – 103 – 104 102 – – 104 105 104 104 – –

10-2–10-1 – – 10-4 – – 10-3-101 – – – – 10-2–10-1 10-4–10-2 – – – –

Relevant reference (Foust et al., 1944) (Arbiter and Harris, 1962) (Albers et al., 1998) (Tschentscher et al., 2010a) (Pilo, 1960) (Ji et al., 2008) (Kroger et al., 1980) (Pudjiono and Tavare, 1993) (Hampton et al., 2008) (Lodha and Jachuck, 2007) (Cowen et al., 1982) (Sarmidi and Barker, 1993a) (Buhtz, 1927), (Boodhoo, 2012) (Meeuwse et al., 2010b) (Mancosky, 2013) (Barillon et al., 2007) (Ogura et al., 2008) (Jones et al., 2012)

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ed

M an

Reactor Rushton stirrer Gas-inducing stirrer Monolithic stirrer Solid foam stirrer Rot. packed bed Rot. zigzag bed Rot. fluidized bed Taylor-Couette flow Spinning tube-in-tube Rot. tube Rot. Tub. Mem. Rot. Ann. Chrom. Thin film SDR Rotor-stator SDR Shockwave power RAPTOR Dyn. rot. axis micro Coflore agitated cell

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