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chemical engineering research and design 9 1 ( 2 0 1 3 ) 1923–1940

Contents lists available at ScienceDirect

Chemical Engineering Research and Design

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eview

otating reactors – A review

. Visscher, J. van der Schaaf, T.A. Nijhuis, J.C. Schouten ∗

aboratory of Chemical Reactor Engineering, Department of Chemical Engineering and Chemistry, Eindhoven University of Technology,.O. Box 513, 5600 MB Eindhoven, The Netherlands

a b s t r a c t

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.

© 2013 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.

Keywords: Reactor; Development; Gravity; Multiphase; Intensification

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ontents

1. Introduction . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

1.1. Reactor selection criteria . . . . . . . . . . . . . . . . . . . . . . . . . . .1.2. Reactor classification . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

2. Rotating reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .

2.1. Stirred tanks . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .2.1.1. Rushton stirrer . . . . . . . . . . . . . . . . . . . . . . . . . . . . .2.1.2. Gas-inducing stirrer . . . . . . . . . . . . . . . . . . . . . . .2.1.3. Monolithic stirrer . . . . . . . . . . . . . . . . . . . . . . . . . .2.1.4. Foam based stirrer . . . . . . . . . . . . . . . . . . . . . . . . .

2.2. High shear rotating tubes . . . . . . . . . . . . . . . . . . . . . . . . . .2.2.1. Rotating packed bed . . . . . . . . . . . . . . . . . . . . . . .2.2.2. Rotating zigzag bed . . . . . . . . . . . . . . . . . . . . . . . .2.2.3. Rotating fluidized bed . . . . . . . . . . . . . . . . . . . . .2.2.4. Taylor–Couette reactor. . . . . . . . . . . . . . . . . . . .

2.2.5. Spinning tube-in-tube reactor. . . . . . . . . . . .

2.3. Low shear rotating tubes . . . . . . . . . . . . . . . . . . . . . . . . . . .2.3.1. Rotating tube reactor . . . . . . . . . . . . . . . . . . . . . .2.3.2. Rotating tubular membrane reactor . . . . . .2.3.3. Rotating annular chromatographic reacto
2.3.4. Rotating sorbent reactor . . . . . . . . . . . . . . . . . . . . . . . .
∗ Corresponding author. Tel.: +31 40 247 2850.E-mail address: [email protected] (J.C. Schouten).URL: http://www.chem.tue.nl/scr (J.C. Schouten).Received 15 April 2013; Received in revised form 19 July 2013; Accepte

263-8762/$ – see front matter © 2013 The Institution of Chemical Engittp://dx.doi.org/10.1016/j.cherd.2013.07.021

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d 22 July 2013neers. Published by Elsevier B.V. All rights reserved.

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1924 chemical engineering research and design 9 1 ( 2 0 1 3 ) 1923–1940

2.4. Low shear rotating disc reactors . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19322.5. High shear rotating disc reactor. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1933

2.5.1. Thin-film spinning disc. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19332.5.2. Rotor–stator spinning disc reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1933

2.6. Remaining reactor types. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19342.6.1. Shockwave power reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19342.6.2. RAPTOR. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19342.6.3. Dynamically rotating axis micro reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19342.6.4. Coflore agitated cell reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19352.6.5. Rotating cone reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 19352.6.6. Rotating membrane reactor . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1935

3. Summary and outlook . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1935References. . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . 1936

List of symbols

AL surface area (m2 m−3)DI impeller diameter (m)dR inner diameter (–)E power consumption (W m−3)N maximum rotation speed (RPM)Q volumetric throughput (m3 h−1)� residence timetm micromixing time (s)U overall heat transfer coefficient (W m−2 K−1)V reactor volume (m3)

Reactors are developed dedicated to their application which

1. Introduction

Mechanical agitation is applied to chemical reactors on pilotscale and on industrial scale with the aim to increase the mix-ing efficiency (Hemrajani and Tatterson, 2004). As a result, thereactor volume can be decreased by a factor 10–100, whichis an urgent need in chemical industry; this need is com-monly addressed as process intensification (Stankiewicz andMoulijn, 2000). Ramshaw and co-workers were the first pio-neers who envisioned that understanding the length and timescales relevant in plant design will lead to process intensifica-tion (Reay et al., 2008). The relevant length and time scalesinclude the integral path of plant, reactor, fluid dynamics, cat-alyst and molecular level (Charpentier and McKenna, 2004;Lerou and Ng, 1996).

The reactor volume reduction that originates from theapplication of mechanical agitation yields a potential cost ben-efit that acts as a driving force for the development of rotatingreactors, especially in research dealing with the fine chemicaland pharmaceutical industry (Anderson, 2012; Chaudhari andMills, 2011). With a lower equipment volume the holdup of liq-uids, gases and solids in the reactor is smaller, which reducesthe impact of potential escalation of dangerous situations.Additionally, the volume reduction allows for the applicationof expensive coating materials in the reactor, like platinumand tantalum.

Mechanical agitation is the most common method thatis applied in order to enhance mixing in industrial reactors.For this purpose a rotating element is added to the reactorwith the purpose to increase the gas–liquid, liquid–liquid, andliquid–solid mass transfer rates as well as the heat transferrate. As a result conversion and selectivity can be increasedand 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). Alternativemethods to achieve the said intensification are the integra-tion of reaction and separation in one unit, the switch frombatch to continuous operated reactors and the application ofmicroreactor technology. Microreactors allow for the reduc-tion of byproduct formation and the increase of gas–liquidand liquid–liquid mass transfer rates (Hessel et al., 2009;Hessel, 2009). The main challenge for further implementationof this technology is the scale-up of the volumetric through-put, which can only be achieved by increasing the pressuredrop over the microchannel or by parallel feeding of multiplechannels (Al-Rawashdeh et al., 2012; Kashid et al., 2010).

This review-perspective paper classifies rotating reactorsand describes the current state-of-the-art in the field ofrotating reactors which are applied on lab scale, pilot scale,and industrial scale. In this review a chemical reactor isdefined as any device which is used to conduct a chemicalreaction, more specifically, in which at least one molecu-lar compound is transformed into another predeterminedchemical compound (Trambouze and Euzen, 2002b). Thediscussed rotating reactors are applied for either single-phasesystems (liquid or gas) or multi-phase systems (gas–liquid,gas–solid, gas–liquid–solid, liquid–solid, or liquid–liquidmixtures).

The reactor selection is narrowed down further by discard-ing reactors which are only used for the experimental determi-nation of physical properties, and by discarding reactors solelyused for the testing of catalytic properties or reaction kinetics(Doraiswamy and Tajbl, 1974; Dudukovic et al., 2002; Kapteijnand Moulijn, 2008; Pavko et al., 1981; Weekman, 1974). Exam-ples of these reactors are the Berty Stationary Catalyst Basket(Berty, 1974), the Mahoney–Robinson Spinning Catalyst Basketreactor (Mahoney, 1974), and the Carberry Spinning CatalystBasket reactor (Carberry, 1964). These reactors are discardedbecause they are not developed with the aim of industrialimplementation. Reactors in which mechanical energy is dis-sipated without the usage of rotation are also discarded (e.g.the oscillatory baffled reactors, (Harvey et al., 2001), static mix-ers, (Meijer et al., 2012), and reactors in which cavitation isultrasonically induced cavitation, (Rooze et al., 2013)).

1.1. Reactor selection criteria

Reactor selection is a challenging task in which multipleaspects need to be taken into account (Krishna and Sie, 1994).

can be the synthesis of nanoparticles (Dahl et al., 2007; Ng

chemical engineering research and design 9 1 ( 2 0 1 3 ) 1923–1940 1925

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Fig. 2 – The development of rotating reactors isschematically shown as a function of time. The year inwhich the first scientific publication (peer-reviewed or

t al., 2012; Pask et al., 2012; Tai et al., 2007), exothermiceactions (Ogura et al., 2008), electrochemical reactionsRivero et al., 2010), micromixing characteristics (Assirellit al., 2002; Baldyga and Pohorecki, 1995; Bourne and Studer,992; Chu et al., 2007; Jiao et al., 2012; Rousseaux et al., 2001;hao et al., 2010), and others. Reactions can also be combinedith separation steps in reactive separators (Stankiewicz,

003). Specific aspects that cannot be neglected when reactorsre compared, are the volumetric throughput, reactor volume,esidence time distribution, effective catalyst loading, massransfer rates, and pressure and temperature limitations ofhe reactor (Cybulski and Moulijn, 2005; Froment et al., 2011;rambouze and Euzen, 2002b).

Rotating reactors contain one or more rotating elementshich may have various designs: impellers, tubes, or discs.he geometry of the rotating element and the rotationalpeed determine the power required for mechanical agi-ation. Energy dissipation in chemical reactors is given in

kg−1PRODUCT or W m−3

LIQUID (Baldyga and Pohorecki, 1995;illermaux, 1988). The mixing intensity is mainly determinedy the local energy dissipation and not by the power consump-ion of the auxiliary equipment (Laufhütte and Mersmann,987; Mukherjee and Wrenn, 2009). For mechanically agitatedeactors the energy consumption by the motor that propelshe rotating element is dominant over the energy consumedy pumps. The compressor duty significantly contributes tohe total power consumption in the case of gas–liquid reactors.

.2. Reactor classification

n this review, first the development of rotating reactors overime is presented. Next the specific benefits and disadvantagesf various rotating reactors are given. In this review rotatingeactors are classified by four different groups according toheir characteristic geometry: tanks, tubes, discs and miscel-aneous rotating reactors (Fig. 1).

The development of various rotating reactors in time ischematically shown in Fig. 2. In Fig. 2 the year in whichhe first scientific or technical publication (peer-reviewed oratent) was published is used as allocation in time. A tableith the relevant literature of each reactor is given in Table 1.he majority of rotating reactors has been developed after980. Exact sales information is not available, but most of theeactors developed since 1980 have not yet permanently pen-
trated the commercial market of industrial processes exceptor niche applications.
Fig. 1 – Classification of rotating re

patent) is published is used as allocation in time.

2. Rotating reactors

2.1. Stirred tanks

Agricola (1556) illustrated in his book “De Re Metallica”(On the Nature of Metals, http://en.wikipedia.org/wiki/De remetallica) how stirring reactors were used in the mining

industry (Stankiewicz and Moulijn, 2000). Earlier referencesto vessels equipped with a stirrer date back to 77 A.D. wherePliny the Elder describes the leaching of metals and the purifi-cation of sulfur (Pliny the Elder, 1929). The volume of industrialused stirred tank reactors ranges from 2 × 10−3 to 3 × 102 m3.The reactors are either batch wise, semi-batch or continuouslyoperated (Trambouze and Euzen, 2002a).

The most basic form of a stirred tank reactor consists ofa cylindrical tank with elliptical bottom, with one or morestirrers mounted on a central shaft. With increasing tankdiameter, stirred tanks often exhibit poor mixing which isespecially true for multiphase reactions in which the non-uniformity in mixing and mass transfer leads to significantvariance in reaction rate and selectivity (Ståhl Wernerssonand Trägårdh, 1999; Stitt, 2002). Some of the stirrers areshown schematically in Fig. 3 (Hemrajani and Tatterson, 2004;
Joshi et al., 1982). Both the power consumption in stirredtanks (Ascanio et al., 2004; Villermaux, 1988), and the various
actors based on the geometry.

http://en.wikipedia.org/wiki/De_re_metallica

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

Reactor Phase Q V � E N GLMT U tm Relevant reference

Rushton stirrer GLS 10−3 to 101 10−3 to 102 1 min–5 h 1 × 104 1500 0.01–2 103 10−2 to 10−1 Foust et al. (1944)Gas-inducing stirrer GLS 10−3 to 101 10−3 to 102 1 min–5 h 1 × 104 1500 0.01–2 103 Arbiter and Harris (1962)Monolithic stirrer GL – 10−3 1–100 min 2 × 103 800 0.8 103 – Albers et al. (1998)Solid foam stirrer GLS 10−2 10−3 1 s–10 min 6 × 103 600 0.2 103 – Tschentscher et al. (2010a)Rot. packed bed GLS 10−2 10−3 1 s–10 min 5 × 104 2500 12 105 10−4 Pilo (1960)Rot. zigzag bed GL 10−4 10−1 1 s–10 min 3 × 104 1400 12 – – Ji et al. (2008)Rot. fluidized bed GS – 10−3 1 s–10 min – 500 – 103 – Kroger et al. (1980)Taylor–Couette flow GL 10−6 10−6 1 s–10 min 3 × 104 1000 0.1 – 10−3 to 101 Pudjiono and Tavare (1993)Spinning tube-in-tube GL 10−7 10−6 to 10−3 1 s–3 min – 12,000 – 104 – Hampton et al. (2008)Rot. tube GL 10−5 10−4 1–15 min – 870 – 102 – Lodha and Jachuck (2007)Rot. tub. mem. GL 10−3 10−2 1 s–10 min 1 × 103 30 – – – Cowen et al. (1982)Rot. ann. chrom. GLS 10−6 10−2 1 s–60 min – 20 – – – Sarmidi and Barker (1993a)Thin film SDR GLS 10−3 to 10−2 10−3 1 s–10 min 5 × 103 5000 6 104 10−2 to 10−1 Buhtz (1927) and Boodhoo and

Al-Hengari (2012)Rotor–stator SDR GLS 10−3 to 10−1 10−4 to 10−2 1 s–10 min 5 × 106 4500 2 105 10−4 to 10−2 Meeuwse et al. (2010b)Shockwave power GL 10−1 to 101 – – – 3600 5.2 104 – Mancosky (2013)RAPTOR GL 10−5 to 10−3 10−3 10 s–10 min 3 × 106 1500 – 104 – Barillon et al. (2007)Dyn. rot. axis micro GL 10−7 10−6 1 s – 3600 – – – Ogura et al. (2008)Coflore agitated cell GL 10−5 10−5 1 s–10 min – – – – – Jones et al. (2012)


Fig. 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 isprogressively sucked in (axial downwards) near the center and is forced radially outwards. An extensive list of differents Josh

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tirrers is presented by Hemrajani and Tatterson (2004) and

echniques to visualize the liquid flow behavior have beenlaborated already earlier in a review (Mavros, 2001).

The catalyst in a stirred tank can be either disperseds a slurry (with a particle diameter below 1 × 10−3 m) ors a homogenous catalyst. The most important drawbacksf a heterogeneously dispersed catalyst are the separationf the catalyst from the reaction mixture, and the attritionnd agglomeration of the catalyst particles. Heat transfer tor from stirred tank reactors can be obtained by jacketinghe stirred tank or by using internal coils. Usage of struc-ures inside the reactor allows for higher heat transfer rates,ut increases also the risk of fouling, the non-uniformity ofixing intensity, and the time required for reactor clean-

ng (Kumaresan and Joshi, 2006). Various methods can bexploited in order to improve the mixing capability in stirredanks.

The application of vertical wall baffles mounted to the reac-or wall is well known (Lu et al., 1997). A second method tonhance mixing capability in the stirred tank reactor is bypplying a more sophisticated stirrer design, which can be cat-gorized according to the induced direction of flow: up/downraft (discs, plates), radial (flat blade impeller, Rushton stirrer),xial (propeller, pitched blade turbine), or vortex (Kumaresannd Joshi, 2006; Stitt, 2002).

Stirrer geometries can be, but are not limited to, pro-ellers, turbines, anchors, or Archimedes screws (Hemrajanind Tatterson, 2004). An extensive recent review on typicalmpeller characteristics which are essential for further com-arison of various impellers is given elsewhere and addresseshe relevance of the power number, the flow number, the

omentum number, and the Zwietering constant (Machadot al., 2012). Multiple stirrers on a single rotating shaft areeeded when the aspect ratio, the ratio of the stirrer diameterver the tank diameter, exceeds 1.5. A last method to enhanceixing in stirred tanks is the application and improvement

f a gas-distributing inlet, which will enhance the interfacialrea and the gas–liquid mass transfer rate in the stirred tank.our different stirrers are described in more detail here: Rush-on stirrer, gas inducing stirrer, monolithic stirrer and the solidoam stirrer.

.1.1. Rushton stirrerhe Rushton stirrer was developed around 1940 (Foust et al.,944), and is a radial flow generating stirrer, which is equipped
ith six vertical (flat or curved) blades which are mountedn a disc (Fig. 4). For the standard stirrer the blade length is
i et al. (1982).

equal to DI/4, the blade width is equal to DI/5. The disc diam-eters equals either 0.66DI or 0.75DI, in which DI is the impellerdiameter (Hemrajani and Tatterson, 2004). The gas–liquid flowbehavior in a Rushton stirred tank was studied using Laser-Doppler Anemometry (Wu et al., 1989) and Particle ImageVelocimetry (Hill et al., 2000). The power characteristics ofRushton stirrer are related to physical properties of the liquidmixture and the geometry of the tank itself (Rushton et al.,1950a,b).

2.1.2. Gas-inducing stirrerOften the per pass conversion of the gas phase is low whengas–liquid mass contacting is performed in a stirred tank, inthat case it is beneficial to recycle the unreacted gas-phaseback into the reactor. Dead end systems are than a solution inwhich expensive compression costs can be reduced: in thesesystems the remaining gas phase is forced into the free reac-tor volume from where it is recycled internally into the liquidmixture. A gas-inducing impeller enables efficient recyclingof gas from the free reactor volume into the liquid-mixture.The critical impeller speed that is required for the start of gasinduction follows from a balance between the velocity headgenerated by the impeller and the hydrostatic head above theimpeller (Patwardhan and Joshi, 1998). Guidelines have beengiven about the desired geometry of gas-inducing impellersfor achieving different design objectives such as heat transfer,mass transfer, mixing, solid suspension and froth flotation.

Non-intrusive electrical capacitance tomography (ECT) hasbeen used to study the dispersed phase hold up, mixing times,and reaction metrics in a continuously operated stirred tankthat was equipped with a gas-inducing stirrer (Bawadi et al.,2011). Application of a gas-inducing stirrer in a stirred tankgives an increase in the productivity (Bawadi et al., 2011). Theinfluence of the stirrer diameter, the aspect ratio, the stirrersubmergence from the liquid level, and the clearance betweenthe stirrer and the tank bottom has been presented elsewhere(Saravanan et al., 1994).

The gas–liquid mass transfer rate in stirred tanks that areequipped with gas-inducing impellers was measured, and canbe described by a dimensionless correlation which containsthe Froude number (gas-induction rate), the Reynolds number(turbulence intensity), and the Schmidt number (fluid prop-erties) (Zieverink et al., 2006). The separation of the reactionmixture and the catalyst particles at the outlet of a stirred tank
is often troublesome; there is therefore a tendency to immo-bilize the catalyst on the stirrer. In early attempts a lab-scale


(left)

Tomography measurements have shown that also liquids witha higher viscosity can be fed to such a reactor (Tschentscher

Fig. 5 – Schematic 3D-view of the foam stirrer tank. Thestirrer is equipped with two equal foam blocks on which a

Fig. 4 – The top view and the side view of a Rushton stirrer

rotating basket was mounted in the reactor, such that kinet-ics could be measured without the presence of external masstransfer limitation in the gas phase (Carberry, 1964).

2.1.3. Monolithic stirrerMounting monolithic blocks on the stirrer shaft can be anattractive alternative for stirred tanks with dispersed cata-lyst particles (Fig. 4). Most monoliths consist of one piece ofceramic material; within this piece a large number of parallelchannels is present which extends over the entire length of theblock. The concept of a monolithic stirrer was demonstratedin 1998 for liquid mixtures with a low viscosity (Albers et al.,1998). Because the catalyst is immobilized in the monoliththere is no need for liquid–solid separation at the reactor out-let. Another advantage is the open structure of the monolithicblock which results in a large geometrical area. The inside ofthe monolithic channels can be coated with a thin layer ofeither a conventional catalyst (Bennett et al., 1991) or a biocat-alyst (De Lathouder et al., 2006). The monolith is characterizedby its number of cells per square inch (Hoek, 2004b). Withincreasing cell density the catalyst layer thickness decreased,which proved to be beneficial for the performance of themonolithic stirrer reactor (Hoek et al., 2004). The volumetricgas–liquid mass transfer coefficient in a monolith increasesfrom 0.015 s−1 at 200 RPM to 0.527 s−1 at 450 RPM when mea-sured for the stirrer configuration with two monoliths in oneplane (Hoek, 2004a). The cell density of the monoliths has noeffect on the gas–liquid mass transfer rate. The volumetricgas–liquid mass transfer increases by a factor of three whena stirrer configuration consisting of four monoliths in oneplane is used. The liquid–solid mass transfer rate increasedwith increasing stirrer speed (Hoek, 2004b). The biggestpotential for industrial implementation would be the replace-ment of conventional slurry reactors that are nowadaysused for multiphase processes, for example in fine chemicalsynthesis.

2.1.4. Foam based stirrerA recent innovation in stirrer design for stirred tanks is thesolid foam based stirrer. Such stirrers contain a piece ofopen-celled solid foam which is made out of a reticulatedstructure of struts (Fig. 5). Each strut has the function of astatic mixer which splits and recombines the fluid stream
that passes the strut. The solid foam combines a high surfacearea (160–8500 m2 m−3) with a high voidage (80–97%), which
and a side view of a monolithic stirred tank reactor (right).

yields a high surface to volume ratio. As a result non-rotatingfoam packed columns exhibit a low pressure drop and highgas–liquid mass transfer rates (Stemmet et al., 2005, 2006). Dueto the high surface area it is an excellent material for the depo-sition of catalysts (Ordomsky et al., 2012a,b; Wenmakers et al.,2008, 2010).

The rotating foam stirrer reactor yields higher gas–liquidand liquid–solid mass transfer rates than stirred tanksequipped with a Rushton stirrers or slurry bubble columns(Tschentscher et al., 2010a,b). Various foam structures havebeen applied including donuts, two-blades, and blocks (Leonet al., 2012a). Rotating foam stirrer reactors have promisingapplications for multiphase processes (Leon et al., 2012b).Mounting multiple foam blocks on a single horizontal shaftwith baffles in between the consecutive foam blocks, yieldsplug flow behavior in this reactor, and accordingly a higherselectivity toward the desired product in selective reactionswith unwanted consecutive reactions (Leon et al., 2013).

high catalyst loading can be coated.Courtesy to Tschentscher et al. (2010a).


Fig. 6 – Schematic side view of a rotating packed bed inwhich the vapor phase (dotted lines) and the liquid phase(solid line) are contacted counter-currently over the rotatingpacking (gray).

eteint

2

2Rtmo(n1nbcadsvOp1s

tarstotiR

mi(dat

is oriented in a vertical plane and the gravitational field (1 g) is

t al., 2012). An important advantage for industrial implemen-ation of this stirrer type is the very low pressure drop whichnsures the good accessibility of the catalyst on the foam. Forndustrial implementation it is essential that the catalyst doesot wash out, and thus that the appropriate catalyst coatingechnique is applied.

.2. High shear rotating tubes

.2.1. Rotating packed bedotating packed beds (RPBs) have received considerable atten-ion as a method of process intensification for gas–liquid

ass transfer. Most publications mention that the first patentn RPB-technology was filled by Ramshaw and Mallinson

1981). This patent described the application of such tech-ology for distillation, absorption, and stripping (Ramshaw,983). An earlier patent application described this same tech-ique about two decades earlier (Pilo, 1960). Rotating packededs can be operated either co-currently, cross-currently orounter-currently. In the most basic form of a RPB, the rotor isn annular cylindrical packed bed which is housed in a cylin-rical casing (Fig. 6). In rotating packed beds the rotationalpeed ranges from 500 to 2500 RPM. The rotor can be made ofarious packing types, e.g. gas spheres, solid foam, and discs.ther rotor configurations which have been applied are disclate packings (Jian et al., 1998), helical packings (Chen et al.,995), multistaged spraying packings (Pan et al., 2006), andplit-packings (Chandra et al., 2005; Reddy et al., 2006).

The gas phase and liquid phase are counter-currently con-acted in a RPB. The gas phase is introduced near the casing,nd flows radially inwards through the packing and exits theeactor through the shaft. The liquid phase is fed through atationary feeding tube placed in the eye of the rotor, andouches the inner periphery of the packing as droplets, jets,r as a spray. For measurements at low gravity level (<60 g)he liquids fills the voids of the packing, whereas at high grav-ty levels (>100 g) the voids are only partial filled (Burns andamshaw, 1996; Burns et al., 2003).

The residence time distribution in a RPB was deter-ined experimentally, and showed that intense macromix-

ng is obtained at the inner region of the packingdR = 7 × 10−3–10 × 10−3) where the liquid impinges andeforms. The same study showed that the liquid volume in

packed bed under normal operation, does not exceed 5% of
he total bed volume (Kenyvani and Gardner, 1989).
As a result the mean residence time of the liquid phase inthe reactor is very short which limits the separation capacity.The micromixing efficiency of rotating packed beds is dis-cussed in detail by Yang et al. (2005). There, the micromixingtime is approximately evaluated to be about 10−4 s.

Two in-depth reviews describe the functionality of therotating packed bed and the relevant literature up to 2010 ingreat detail (Rao et al., 2004; Zhao et al., 2010), including liquidflow, gas–liquid mass transfer, micromixing efficiency, strip-ping/absorption, and nano-particle preparation. The mainconclusion is that adequate design of the rotor is essential forfull exploitation of the centrifugal field. The height equiva-lent of a transfer unit is about 4–5 times lower in RPBs thanin conventional extraction columns (Rao et al., 2004), andequals 1 × 10−2–2 × 10−2 m (Zhao et al., 2010). The intensifi-cation achieved so far falls short of the goal of 2–3 ordersof magnitude volume reduction of the conventional columnsvolume (Rao et al., 2004).

Relatively poor mass transfer performance of RPB technol-ogy was observed when used as bioreactor for the productionof polyhydroxyalkanoate from a fermentation broth (Boodhooet al., 2010). The apparent poor performance is attributedto the limitation of the gas and liquid flow rates that area consequence of flooding characteristics. Also, the possi-ble entrainment of air in the oxygen stripped end productcould have influenced the reliability of these findings. Themicromixing efficiency in the RPB was determined using theVillermaux–Dushman reaction system. In this study sampleswere drawn along the radial position. The micromixing effi-ciency was thus measured in different packing zones of theRPB. The micromixing time in the RPB can be as low as 10−4 s,which is one to three orders of magnitude smaller than inconventional packed beds. In the Impinging Stream RotatingPacked Bed (IS-RPB), two stainless capillary nozzles are locatedin-line with the rotational shaft. The two different liquidsflow along the same axis in the opposite direction and col-lide, which causes a narrow zone in which a high turbulenceintensity 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 beensuccessfully obtained, including seawater treatment, HOClproduction and nanoparticle preparation. However, industrialacceptance of this reactor type is yet to come, mainly due toconcerns involving reliability and the energy consumption.More sound theoretical foundations are required, especiallythe liquid distribution at the entrance (in the center) of packedbed is of main importance for the mixing intensity obtainedin the reactor. This should be combined with long term oper-ational data and economic evaluation of such data.

2.2.2. Rotating zigzag bedIn the rotating zigzag bed (RZB) the packing from the rotatingpacked bed is replaced by a rotor which is a coaxial combina-tion of a rotating disc with a stationary disc (Fig. 7) (Li et al.,2012, 2013). Concentric circular sheets are mounted on therotor and the stator, and act as rotating and stationary baffles,respectively. Typically the distance between the consecutiverings is below 15 × 10−3 m.

The gas and liquid flow are counter-currently contactedwhile flowing in a horizontal zigzag motion through the reac-tor. The RZB thus has in essence a similar geometry asconventional tray-distillation columns, but in the RZB the flow

replaced by a high centrifugal field (>500 g) (Wang et al., 2011).


Fig. 7 – Schematic side view of the rotating zigzag bed. Inthis reactor the packing consists of a rotating disc (gray)and a stator (black). Two series of concentric circular sheetsare fixed on the rotating and the stationary disc. Thecircular sheets on the rotor are perforated.

Fig. 8 – A schematic representation of the top view of arotating fluidized bed in a static geometry. The gas phase isinjection through the tangential inlets. Due to the rotationof the injected gas phase, the solids present in the reactorwill start to fluidize.Courtesy to Rosales Trujillo and De Wilde (2012).

The solids experience a centrifugal force which is directed

When compared to a rotating packed bed reactor the diameteris increased by a factor two, whereas the length is reduced bya factor 16. The reactor volume thus decreased by a factor 5(Li et al., 2013).

The rotating zigzag reactor has a few beneficial advantagesover common RPB technology: first, one of the dynamic sealscan be replaced by a static seal, yielding a longer seal life-time and thus less maintenance. Second, additional feed inletscan be applied through the stator which allows for continu-ous operation of the distillation process (Ji et al., 2008). Third,no additional liquid feed distributers are required: the zigzagflow enhances mixing between each consecutive ring. Fourth,the liquid hold-up of the bed is larger than in a conventionalRPB, which leads to an increased residence time. Fifth, in the-ory multiple zigzag-rotors could be mounted on one rotationalshaft (Wang et al., 2008).

The number of theoretical plates increases with increas-ing rotor speed (Wang et al., 2008). The RZB may contain upto 20 theoretical plates per meter length, when operated at1000 RPM and reflux rate of 600 × 10−3 m3 h−1 (when measuredwith an inner rotor diameter 20 × 10−2 m, outer rotor diame-ter of 63 × 10−2, casing diameter of 80 × 10−2 m, casing height55 × 10−2 m, measured for the separation of an ethanol-watermixture) (Wang et al., 2008). Seven industrial applications ofRZB technology have been reported, with a maximum feedcapacity of 48 tons day−1 (Wang et al., 2008). The overall pres-sure drop of the RZB is a combination of the pressure dropsover the rotor, the casing and the gas outlet (Li et al., 2013), andincreases 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 overthe RZB have been presented for a dry bed and a wetted bed (Liet al., 2013). The power required to propel the rotating zigzagbed is discussed in detail by Agarwal et al. (2010) and Li et al.(2012).

2.2.3. Rotating fluidized bedFluidized beds are widely applied in chemical industry and areused for a wide range of processes including olefin polymer-ization end detergent production. In a fluidized bed reactorthe gas phase is forced through a bed of solid particles. Asa 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 canbe applied.

Two types of rotating fluidized beds (RFBs) have beenapplied with either a static (SG) or rotating geometry (RG)(Harish Kumar and Murthy, 2010). In the rotating fluidizedbed with a rotating geometry the cylindrical enclosure rotatesaround its own axis of symmetry. As a result the intensity ofthe gravitational field depends on the rotational speed and theenclosure radius.

Compared to conventional fluidized beds its main advan-tage is the wider range of operation (Harish Kumar and Murthy,2010). The biggest disadvantage is that at extreme rotationalspeeds of the enclosure, particles will accumulate near therim of the enclosure which leads to the channeling of the gasphase through the dead zone near the center of the enclosure(Kroger et al., 1980). The pressure drop and fluidization criteriahave been presented by (Chen, 1987).

The second type of rotating fluidized beds has a staticgeometry which makes that in essence this reactor is not arotated reactor; however the fluid behavior has a large simi-larity with that of other equipment described in this review (DeWilde and de Broqueville, 2008a). This rotating fluidized bedwith a static geometry is also known as the Gas–Solid VortexReactor (GSVR). In this reactor the gas phase is injected tan-gentially into the cylindrical reactor which contains the solidparticles (Fig. 8). Due to the transfer of momentum from thegas-phase to the particles, the particles will rotate in the reac-tor and generate a centrifugal force. High mass transfer ratesare obtained in this reactor due to the counter play betweenthe centrifugal force and the drag force applied by the inward-flowing gas.

radially outwards, while the gas phase is forced to move


rch(ssca

tthemtTlctlaRih

2Ftflgftadflvi

teflst2p(c

ted1lfd2

wotial

Fig. 9 – The spinning tube-in-tube reactor. The temperaturecan be controlled by a heat exchanger which is locatedaround the process volume (green). Reprinted withpermission from (Gonzalez and Ciszewski, 2008). Copyright2009, American Chemical Society. (For interpretation of thereferences to color in this figure legend, the reader isreferred to the web version of the article.)

adially inwards, which fluidizes the solids radially. The con-ept has been illustrated for low density polymer particles andigh density aluminia particles, for various solids loadings

De Wilde and de Broqueville, 2007). The heat transfer rate istudied numerically by using computational fluid dynamicsimulations of the particle bed temperature response to a stephange in the fluidization gas temperature (de Broquevillend De Wilde, 2009).

The functionality of the GSVR has been demonstrated forhe pyrolysis of lignocellulosic biomass. The gas–solid heatransfer coefficient equals 650 W m−2 K−1, which is 3–5 timesigher than in non-rotating fluidization reactors (Ashcraftt al., 2012; De Wilde and de Broqueville, 2008a). The gas–solidass transfer coefficient for the GSVR is about 10 times higher

han for conventional riser technology (Ashcraft et al., 2013).he GSVR can be extended with a rotating chimney which

eads to increase of the centrifugal force in the vicinity of thehimney, which on its turn allows to reduce solids losses viahe chimney and simultaneously to build up a higher solidsoading in the fluidization chamber at given gas flow ratend solids feeding rates (De Wilde and de Broqueville, 2008b).otating fluidized beds have a relatively large volume. In rotat-

ng fluidized beds this implies that the energy consumption isigh when compared to other reactors.

.2.4. Taylor–Couette reactoror reactors in which only the inner or both inner andhe outer cylinders are rotating, a Taylor–Couette type ofow is observed. Such flow offers the advantage of centrifu-ally accelerated settling, short residence times, low volumeractions, flexible phase ratios and a well-controlled inven-ory (Vedantam and Joshi, 2006). This class of reactors islso known as rotating annular rectors. The residence timeistribution for a such a reactor can be described as near plugow behavior for Taylor numbers (ratio of centrifugal to theiscous forces) above 60 where a laminar vortex flow regimes present (Pudjiono and Tavare, 1993).

A detailed review on experimental correlations betweenhe axial dispersion number and the Reynolds number is givenlsewhere, together with a detailed CFD and RTD study of theuid flow behavior (Vedantam et al., 2006). Normally a smoothhaped cylinder is used. Experiments with a ribbed rotor leado an increase of the micromixing efficiency (Richter et al.,008). Both the exothermic copper-catalyzed oxidation of iso-ropanol and the pyrolysis of acetone to ketene were studied

Cohen and Marom, 1983), and were executed under excellentontrol over the yield and selectivity.

Rotating annular equipment also finds its application inhe intensification of stage wise counter-current liquid–liquidxtraction (Schuur et al., 2012), and the continuous pro-uction of monodispersed nanoparticles (Ogihara et al.,995). This reactor is quite similar to the rotating annu-ar reactor, in which a gas–liquid or liquid–liquid mixture isorced through the radial duct between two coaxial cylin-ers, of which the inner cylinder is rotating (Lawrence et al.,000).

The radial distance between the two cylinders is smallhen compared to the radii of the cylinders and is in the rangef 10 × 10−3 m, which is about 40 times higher when comparedo the spinning tube-in-tube reactor. Typically the flow rates in the range of 10 × 10−3 m3 h−1 (Lawrence et al., 2000). As

result, the pressure drop in the rotating annular reactor isower, which allows for a larger throughput.

2.2.5. Spinning tube-in-tube reactorThe last example of a reactor with tubular geometry is thespinning tube-in-tube reactor. The reactor patent for this reac-tor was filed in 2008 by Richard Holl (Holl, 2010). The reactorconsists of a rotating cylinder (rotor) inside a stationary cylin-der (stator), which are mounted at a concentric radial spacingbetween 0.25 × 10−3 and 0.44 × 10−3 m (Fig. 9).

As a result of the small radial distance between the cylin-ders, the reactants inside the annular volume are exposedto elevated shear stress levels. The typical rotational speedof the rotor is between 3000 and 12,000 RPM, with a reactorvolume that varies from 1 × 10−5 to 1 × 10−3 m3. The reactorconcept has been demonstrated for fine chemical production(Gonzalez and Ciszewski, 2008; Hampton et al., 2008). Due tothe low reactor volume, and expected excellent heat trans-fer rates, this reactor is suited for highly exothermic reactionswhich require only a small amount of catalyst.

2.3. Low shear rotating tubes

2.3.1. Rotating tube reactorSeveral reactors have been developed which combine a tubulargeometry with low shear conditions. The rotating tube reac-tor consists of a rotatable hollow cylinder, and has a rotationalspeed which is typically below 1000 RPM. The centrifugal forceinduced by the rotation results in thin liquid films insidethe rotating rube, with a thickness between 700 × 10−6 and1400 × 10−6 m. The technology has been applied in the trans-esterification of canola oil to biodiesel using a base catalyst.Using methanol and sodium hydroxide as catalysts, a conver-sion of 98% was achieved in residence times of 40 s (Lodhaet al., 2012).

The conversion obtained while using a residence time of40 s is equal to the conversion obtained in a membrane reactoroperated at a residence time of 6 h, for a volumetric flow ratewhich is 150 times higher (Lodha et al., 2012). The mechani-cally simple design of this reactor and the low pressure dropover the cylinder, make that the reactor can be used for bulkprocesses in which mass transfer limitations is not limiting.


Fig. 10 – A schematic side view and front view of the rotating biological contactor. A stack of discs (black) is partiallyincre
immersed in the liquid (gray). Rotation of the stack of discs
2.3.2. Rotating tubular membrane reactorThe rotating membrane bioreactor with a tubular geometrywas patented in 1982 (Cowen et al., 1982). In this reactor themembrane is rotating around a shaft which is placed at theliquid level in the reactor. As a result the membrane is par-tially wetted but continuously refreshed during operation. Therotational speed of the membrane is most often below 10 RPM.A typical reactor volume equals 2 × 10−3 m3, and has a totaleffective filtration area of 0.043 m2 (Jiang et al., 2012). Rotationof the membrane allows for an increase of the permeate fluxbecause cake formation is suppressed in microfiltration andconcentration polarization is suppressed during ultrafiltrationand reverse osmosis (Jaffrin, 2008).

2.3.3. Rotating annular chromatographic reactorThe rotating annual chromatographic reactor is a chemicalreactor in which chromatographic separation takes place, andis used for liquid–solid or gas–solid systems (Cho et al., 1980a).The reactor consists of two rotatable cylinders which are rotat-
ing around one central shaft (Fig. 11). The volume between the
Fig. 11 – A schematic side view of the rotating annularchromatographic reactor (Sarmidi and Barker, 1993a;Stankiewicz, 2003). Component C is added at the stationaryinlet. While flowing downwards, the components A–C areselectively recovered over the rim of the rotating cylinders.

ases the liquid film renewal at the disc surface.

two cylinders is filled with the solid phase. Often ion-exchangeresins are used which have the function of both catalyst andadsorbent. The adsorbent can also be used for the immobi-lization of biocatalysts (Sarmidi and Barker, 1993b).

The solvent of the reacting molecules is distributed equallyat the top of the reactor. The rotational speed of the cylin-ders, typically below 10 RPM, causes horizontal migration ofthe reacting molecules, whereas the fluid flow causes a ver-tical downward migration of the reacting molecules throughthe solid phase. Reaction products with a high affinity to thesolid phase will have a large horizontal velocity component,whereas reaction products with a low affinity have a smallhorizontal velocity component (Sarmidi and Barker, 1993a).

At different angular positions of the reactor, the sepa-rated products can be collected in high purity. The historicaldevelopment is described in detail in a review elsewhere(Uretschläger and Jungbauer, 2002). Typically the reactor vol-ume is in the range of 10 × 10−3 m3, whereas the volumetricflow rate equals 15 × 10−3 m3 h−1 (Cho et al., 1980b). The pres-sure drop due to the packing inside the narrow concentricspace limits the scalability in terms of volumetric throughput(Ströhlein et al., 2007).

2.3.4. Rotating sorbent reactorThe rotating sorbent reactor was developed in 2004 and is acombination of the rotating particle separator with a cyclonicreactor. The basic design consists of three stages. In the firststage, the solid phase is injected in a gas phase flow that con-tains an undesired component. In the second stage, the solidphase absorbs the undesired component from the gas phase,and in the third stage the solid phase is separated from the gasphase. Typical rotational speed of the separator part is higherthan 12,000 RPM. The volumetric flow rate equals 50 m3 h−1

and the reactor volume equals 4 × 10−3 m3. Industrial applica-tions have not been reported in literature (Mondt et al., 2004).

2.4. Low shear rotating disc reactors

The rotating biological contactor (or rotating disc gas–liquidcontactor) consists of a number of discs which are mounted ona central horizontal shaft (Fig. 10). The stack of discs with a lowinter-disc distance yields an equal geometry as a cylinder. Thediscs are partially immersed in the liquid (Zhevalkink et al.,1978). A liquid film is brought upwards over the surface of thediscs, upon rotation of the discs. Further rotation will supply
the liquid film to the bulk liquid. Typical rotation speeds arebelow 50 RPM. Due to the low rotational speed and the low


pb

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Fig. 12 – The side view of the rotor–stator spinning discreactor. The axial clearance between the rotor and thestator is in the order of millimeters. Rotation of the rotorinduces a velocity gradient over the height between therotor and the stator. As a result a shear force acts on thedroplets, bubbles or particles between the rotor and thestator. This results in high liquid–liquid, liquid–solid, and

ower consumption of the motor the reactor is suited for aulk process, like wastewater treatment.

Several other low shear rotating disc reactors are appliedn either lab scale or pilot scale. The rotating disc CVD reac-or has solely applied to achieve a uniform layer thickness inhemical vapor deposition processes (Coltrin et al., 1989). Moreecent research describes a rotating disc photoelectrocatalyticeactor (Li et al., 2012).

.5. High shear rotating disc reactor

.5.1. Thin-film spinning dische thin-film spinning disc reactor was commercially devel-ped with the aim of intensifying gas–liquid reactions (Jachuckt al., 1997), and finds its base in a patent from 1927 (Buhtz,927). In this reactor, the liquid phase is fed at the centerf the rotating disc from where it flows radially outwards as

thin film over the disc. Extensive research has been doneith respect to the gas–liquid mass transfer rate (Aoune andamshaw, 1999; Sisoev et al., 2006), liquid–solid mass trans-

er rate (Burns and Jachuck, 2005; Peev et al., 2007a), and theeat transfer rates (Harmand et al., 2013). The thin-film spin-ing disc reactor has been used for polymerizations (Boodhoond Jachuck, 2000), active pharmaceutical ingredient produc-ion (Oxley et al., 2000), crystallization (Pask et al., 2012; Tait al., 2007), and nanoparticle synthesis (Chu et al., 2007; deaprariis et al., 2012; Liu et al., 2012).

The characteristics of the liquid film, i.e. the film thick-ess and radial velocity and residence time distribution, haveeen described numerically and experimentally (Mohammadind Boodhoo, 2012). With increasing rotational disc speed theiquid film thickness decreases with increasing radius (Burnst al., 2003; Wood and Watts, 1973).

The radial velocity increases with increasing rotational discpeed and increasing liquid flow rate, and decreases withncreasing viscosity (Wood and Watts, 1973). Surface wavesave been observed in the liquid film (Aoune and Ramshaw,999) which are classified as concentric waves or spiral wavesSisoev et al., 2006), and are of major importance to theas–liquid mass transfer rate (Peev et al., 2007b). Whereas inost studies a smooth disc is used as the rotor, also rotat-

ng discs with surface modifications have been used, whichllow for a further increase of the gas–liquid mass transferate (Sisoev et al., 2006).

Recently, two studies independently discussed theicromixing efficiency of the thin-film spinning disc reactor.

oth studies concluded that this reactor type is a valuablentensified mixer (Boodhoo and Al-Hengari, 2012; Jacobsennd Hinrichsen, 2012), with a particular importance forhe application in very fast reactions that need large heatissipation (like nitrations, sulfonations, Darzens processes,rystallizations, and exothermic condensations) (Boodhoo,012). The scale-up of the thin-film spinning disc reactor isiscussed in detail elsewhere (Boodhoo, 2012). The desiredethod of scale-up is a counter play of the residence time

nd the liquid film thickness. Both are determined by theotational disc speed and the disc radius.

.5.2. Rotor–stator spinning disc reactor high-shear multiphase rotating reactor that is developeds an improvement of the thin-film spinning disc reactor ishe rotor–stator spinning disc reactor (van der Schaaf and
chouten, 2011). This reactor consists of a rotating disc (rotor)hich is located between two stationary discs (stators). The
gas–liquid mass transfer rates (Visscher et al., 2012c).

axial distance between the rotor and the stators is typically inthe range of 1.0 × 10−3 m (Fig. 12). The rotational disc speed ismost often around 1000 RPM but rotational speeds of 4500 RPMhave been reported (Meeuwse et al., 2012). Pulse wise resi-dence time distribution measurements have shown that thesingle phase fluid flow can be described by a combination ofa plug flow volume and three ideally mixed volumes in series(Visscher et al., 2012c). The ideally mixed volumes in seriesoriginate from the boundary layer formation on the rotor andthe stators (van Eeten et al., 2012; Visscher et al., 2012c).

The gas–liquid mass transfer rate was measured as a func-tion of rotational speed (Meeuwse et al., 2010b), rotor diameterand rotor–stator distance (Meeuwse et al., 2009), and feed loca-tion of the gas phase (Meeuwse et al., 2009). When the gasphase is fed through an inlet in the bottom stator located atthe rim of the rotor, the gas bubble is sheared off due to theshear stress that is present in the cavity between the rotor andthe stator. The gas bubble diameter decreases with increasingrotational disc speed typically below 1 × 10−3 m for rotationalspeeds above 1000 RPM (Meeuwse et al., 2010b), as a resultgas–liquid mass transfer rates are reported of 0.43 m3

L m3R s−1

at a gas flow rate of 7.3 × 10−6 m s−1.Typically the gas-hold up is below 5%. The volumetric mass

transfer per unit volume of gas is 40 times higher than isreported for a bubble column. A catalytic coating was appliedto the rotor (aL = 274 m2

I m−3R ) in order to study the reac-

tion and liquid–solid mass transfer in series. The volumetricliquid–solid mass transfer coefficient is one order of mag-nitude higher compare to values reported for packed beds(Meeuwse et al., 2010a). The volumetric mass transfer coef-ficient is increased by a factor 3 when the rotor radius isincreased from 0.066 m to 0.135 m whereas the energy inputincreases by a factor 15 (Meeuwse et al., 2011). The preferredmethod to scale-up this reactor without changing the volu-metric throughput is therefore not to increase the diameterbut to increase the number of rotors-in-series (Meeuwse et al.,2012).

The overall liquid–liquid mass transfer rate in the spin-ning disc was measured using an extraction system withwater as the continuous phase and n-heptane as the dis-persed phase, in combination with benzoic acid as the solute(Visscher et al., 2011). The liquid–liquid mass transfer rateequals 0.17 m3

org m3R s−1 at 300 RPM, in this particular case

both phases have plug flow behavior in the reactor. Thisvalue increases to 51.47 m3

org m3R s−1 at 1600 RPM, in this case


Fig. 13 – Shockwave power reactor (Mancosky, 2013). Fig. 15 – Dynamically rotating axis micro reactor.

A unique combination of micro reactor technology andmechanical agitation is used in the dynamically rotating

both liquids obey ideally mixed behavior. These mass trans-fer rates are at least 25 times higher compared to those inpacked columns and at most 15 times higher compare to thosemeasured in state-of-the-art microchannels (Visscher et al.,2012d). �-Ray tomography measurements have shown the vol-ume fraction of the dispersed phase is closely related to theratio of the dispersed phase volumetric flow rate over the totalflow rate (Visscher et al., 2012b).

A patented prototype of rotor–impeller–rotor spinning discequipment has been reported in literature which exploitsthese high liquid–liquid mass transfer rate with the aim tointensify liquid–liquid contacting equipment (van der Schaafet al., 2012; Visscher et al., 2012a). The height equivalent of atheoretical stage in the extractor equals 1.4 × 10−2 m (Visscheret al., 2013), which is a factor 10 higher when compared toother rotating equipment applied for liquid–liquid contacting.

2.6. Remaining reactor types

2.6.1. Shockwave power reactorThe shockwave power reactor was patented in 1993 (Griggs,1993), and consists of a spinning rotor that is baffled suchthat dead ended cavities are present at the rim of the rotorFigs. 13 and 14. The rotation creates locally a low pressurezone in the cavities which collapses under the emission ofan energy wave in the surrounding liquid, called the shock-wave. As a result gas–liquid mass transfer rates up to 5.2 s−1

have been obtained (Mancosky, 2013), which is about 25 timeshigher than in a mechanically agitated tank. The volumetricflow rate may vary from 0.23 m3 h−1 to 1.14 m3 h−1, with typ-ical rotational speeds of the Shockwave power reactor rangeup to 3600 RPM. A challenge in the design of this reactor is thelifetime of the rotor and the cavities. The combination of cavi-tation with corrosive chemicals implies that the rotor needs tobe coated with an extraordinary material to resist such harshconditions.

Fig. 14 – Schematic view of the rotor inside the shockwavepower reactor (Mancosky, 2013). Upon rotation of the rotor(light gray) liquid is forced radially outwards, therebycreating cavitation at the radial inner position of thecavities inside the rotor.

Courtesy to Ogura et al. (2008).

2.6.2. RAPTORRAPTOR is an abbreviation of “Réacteur Agité Polyvalent àTransfert Optimisé Rectiligne”, or Reactor with PolyvalentRectilinear Stirred Reactor with Optimized Transfer (Bahrounet al., 2010).

The technology is developed by the AET group and was pre-sented in 2007, and is shown in Fig. 17 (Barillon et al., 2007).The exact geometry of the reactor is only scarcely disclosedin the literature (Bahroun et al., 2010). Rotational speeds ofthe inner part do not exceed 1500 RPM. The small reactor vol-ume (smaller than 2.9 × 10−3 m3, Milly et al., 2008) allows forthe application of temperatures up to 300 ◦C and an operat-ing pressure of 250 bar. The maximum flow rate is 0.15 m3 h−1,with a residence time between 15 s and 10 min. The heat trans-fer area per reactor volume equals 150 m2 m−3 which is at least30 times higher than for an industrial scale stirred tank reactorwhich is batch wise operated (5 m2 m−3).

2.6.3. Dynamically rotating axis micro reactor

Fig. 16 – The reactorblock of the Coflore reactor. Tenconsecutive mixing chambers are interconnected. Onceideally mixed flow behavior is achieved in each mixingchamber, plug flow behavior is mimicked over the reactorblock.Courtesy to Jones et al. (2012).


Fig. 17 – The experimental set-up of the RAPTOR. Thereactor exhibits high heat and mass transfer rates and canhandle flow rates with a maximum of 0.15 m3 h−1.Courtesy to Barillon et al. (2007).

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xis micro reactor. This reactor consists of two cylinders ofhich one is rotating. The inner cylinder has internal diame-

er of 8 × 10−3 m and the spacing between the two cylinderss 100 × 10−6 m (Fig. 15, Ogura et al., 2008). The volumet-ic flow rate is typically in the range of 80 × 10−9 m3 s−1.he rotation speed of the outer cylinder is limited to600 RPM. Due to the high shear stress in the small radialap between the 100 × 10−3 m long cylinders, high mass trans-er rates are achieved for liquid–liquid applications, as haseen demonstrated by generating emulsions of water and o-hlorobenzene, and the nitration of naphthalene (Ogura et al.,013). Literature on this reactor type is too scarce to make aalid estimation of its feasibility.

.6.4. Coflore agitated cell reactorhe coflore agitated cell reactor (ACR) is commercially avail-ble since 2006 and consists of two separate parts: a reactorlock and a lateral moving plate (Jones et al., 2012). Each reac-or block consists of a channel that connects ten consecutiventerlinked reactor chambers. In each chamber a non-fixed

ixing element is present (Fig. 16). Upon moving the plate,he non-fixed mixing element enhances the dispersion inhe reactor chambers. The moving plate causes the liquidn each reactor chamber to behave as ideally mixed flow; as

result plug flow behavior is mimicked in a single reactorlock.

The design of the non-fixed mixing element opens uphe application for gas–liquid contacting as well as slurryandling. Catalysts can be loaded into the non-fixed mixinglement. The reactor block volume may vary from 20 × 10−6 m3

o 100 × 10−6 m3, volumetric liquid flow rate equals up to0 × 10−3 m3 s−1. Adjustments can be made to allow forounter-current liquid–liquid extraction over 6 stages in a sin-le extractor block. Technically this reactor does not belong tohis review since it is shaken, not stirred. When the reactor is

ounted on the moving plate the unit has a small footprint;ccordingly the reactor is well applicable in laboratories andn fume hoods.

The feasibility of various industrial applications have beeneported by the manufacturer, including but limited to theoffman reaction, Suzuki reaction, polymerizations and Grig-ard reactions. The greatest benefit to conventional batcheactors is the easy scalability of the reactor. A challenge forurther industrial implementation of this reactor will be therevention of fouling, for processes in which solids build-up
ight occur.
2.6.5. Rotating cone reactorThe rotating cone reactor was developed at the University ofTwente. In this reactor a solid phase is fed to the bottom inletof a rotating cone, from where they travel spiral-wise upwards.The inner diameter of the cone has a maximum of 0.650 m. Therotational speed of the cone is 900 RPM. The reactor volumemay range from 2 to 200 dm3. The combined feed rate of thegas and solid phase equals 13 kg h−1. The first industrial testcase for which this reactor is used was the pyrolysis of biomass(Wagenaar et al., 1994). More recently the flow behavior waspresented (Leung et al., 2012), as well as the feasibility of therotating cone reactor for the epoxidation of soybean oil (Chenet al., 2011).

2.6.6. Rotating membrane reactorRotating membranes find their application in ultrafiltration(Sarkar et al., 2009, 2012). A theoretical comparison of rotat-ing and stationary membrane disc filters was published in2000 (Serra and Wiesner, 2000). An in-depth review on thismatter was published by Jaffrin (2008). To obtain sufficient sur-face area multidisc membrane units are mounted on a singleshaft.

Examples of equipment are the DMF (Pall Corp., USA), Dyno(Bokela, Germany), Optifilter (ABB Flootek, Finland) and theSpinTek (Huntington, USA) (Jaffrin, 2008). The membrane areaper module may reach 150 m2 (Bläse et al., 2006). Membranediameters are typically below 1 m. The power required for rota-tion is below 30 kW (Jaffrin, 2008). The availability of ceramicmembranes will facilitate the future development of rotat-ing membrane reactors, its application for the production ofgalacto-oligosaccarides should an productivity increase by afactor 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 givenof the development of rotating reactors over time. The paperhas a focus on rotating reactor technology with applicationsin lab scale, pilot scale and industrial chemical reaction engi-neering. Rotating reactors are classified according to theirgeometry. The reactors are classified as stirred tanks, tubes,discs and miscellaneous reactors. Their main advantages anddisadvantages are presented, including the typical operationalconditions (residence time, rotational speed, energy consump-tion). An overview is given in Table 1. The energy dissipationmentioned in Table 1 represent the maximum energy con-sumed by the motor(s) that propel the rotating element(s).This is not equal to the energy that is dissipated locally inthe liquids. Where possible an accurate approximation of thisdifference is given. However, for many reactors the energythat 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 localenergy dissipation, whereas a large scale stirred tank has alow power to volume ratio. Table 1 thus has the function of anorientation on the maximum power that is consumed by thereactors described.

In academia there is a tendency to develop multifunctionalmodules suited for reactive extractions, one-pot synthesis,and juke-box like functionalities where multiple separations
steps are combined in one compact mini-plant. The furtherdevelopment of such equipment requires that large scale


of a novel rotor–stator reactor. Chem. Eng. J. 128 (2–3),

equipment manufacturers, end-product consumers, proto-type experts, and reactor scientists collaborate together fromthe start in research projects dealing with reactor develop-ment.

A promising route to bridge the gap between lab-scalereactor development and industrial implementation of new,intensified reactors is the process of knowledge valorization.One method to achieve such an industrial implementation isthe 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 (rotor–statorand rotor–rotor spinning technology). The presence ofresource incubators at research institutes and universities willfacilitate and speed up the development of such companies.

Introducing new reactor concepts in chemical engineeringis a lengthy process. Industry is reluctant to introduce novelchemical reactor types in existing processes when replace-ment is not essential, with the aim to minimize possible risksto plant performance. The introduction thus depends mainlyon the development of new processes and new productionplants.

A strong argument for the introduction of novel reactortypes will come from the experience that is built up dur-ing public-private research projects in which both academicand industrial partners are collaborating. During such projectsdetailed information of chemical resistance, heat- and masstransfer performance, mechanical durability of rotating parts,and energy consumption can be collected.

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