Development of a CFD model for the simulation of a novel multiphase counter-current loop reactor

Accepted Manuscript Development of a CFD Model for the simulation of a novel multiphase countercurrent loop reactor Andreas Bednarz, Benedikt Weber, Andreas Jupke PII: DOI: Reference:

S0009-2509(16)30713-8 http://dx.doi.org/10.1016/j.ces.2016.12.048 CES 13320

To appear in:

Chemical Engineering Science

Received Date: Revised Date: Accepted Date:

31 August 2016 22 November 2016 19 December 2016

Please cite this article as: A. Bednarz, B. Weber, A. Jupke, Development of a CFD Model for the simulation of a novel multiphase counter-current loop reactor, Chemical Engineering Science (2016), doi: http://dx.doi.org/ 10.1016/j.ces.2016.12.048

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Development of a CFD Model for the simulation of a novel multiphase counter-current loop reactor Andreas Bednarz1*, Benedikt Weber1*, Andreas Jupke1 1) AVT – Fluid Process Engineering, RWTH Aachen University, Wüllnerstr. 5, D-52062 Aachen, Germany, Tel.: +49 241 80-95490, Fax: +49 241 80-92332, www.avt.rwthaachen.de, e-mail: [email protected] *) The first two authors equally contributed to this manuscript

Address correspondence to Andreas Jupke, AVT – Fluid Process Engineering, RWTH Aachen University, Wüllnerstr. 5, D-52062 Aachen, Germany, Tel.: +49 241 80-95490, Fax: +49 241 80-92332, www.avt.rwth-aachen.de, e-mail: [email protected]

1

Abstract Many bio-based processes exhibit low yields resulting from inhibitions. This innovative concept for a multiphase-loop reactor can overcome these limitations with a simultaneous aeration and extraction of valuable compounds in one unit operation. One dispersed phase is used to initiate a loop flow in the reactor while the other phase is dispersed in the downcomer. Therefore, the second dispersed phase rises in the countercurrent. The multiphase flow was simulated using a 2D axis-symmetric Euler-Euler approach in CFD. Besides friction force and gravitation the turbulent dispersion force was considered applying the standard k-ε turbulence model. The simulation of the reactor was validated in pilot-plant experiments. A comparison of the holdup of the dispersed phases and a visual evaluation of the flow field presented very good agreements between the experiments and simulations. Water, Shellsol T and synthetic air were used in the experiments. An improvement of the reactor design was possible based on additional simulation studies.

Keywords Biocompatibility, Loop reactor, Multiphase, Counter-current, CFD

2

1. Introduction: The expected climate change induces the demand of a change in the feedstock for chemical processes. The development of processes based on regenerative resources has been an issue in the research of many scientists in the recent years. Especially the production of bio-based intermediates for the utilization of petro-based fuels and monomers for polymer production has been investigated (Schneider and Wendisch, 2011). Besides the shift from fossil resources to bio-based and regenerative resources the development of biotechnological processes is being investigated (Maass et al., 2002; Kurzrock and Weuster-Botz, 2010; Schneider and Wendisch, 2011; Buschke, 2013; Klement and Büchs, 2013; Polen et al., 2013). Biotechnological processes have certain advantages concerning the selectivity of the transformation and the processes are usually conducted at ambient temperatures reducing the demand in energy for heating. Challenges arise concerning the common low concentrations of the target compound often resulting from product inhibition or even product toxicity. Until now most research was focused either on the biotransformation or on the separation of the product from a filtered fermentation broth. The investigation of a cell-afflicted in-situ extraction has only been an issue for a small group of researchers (Job and Blass, 1990, 1994). The coupling of a biotechnological production and separation in one unit has been investigated concerning for example an in-situ diafiltration (Meier et al., 2014), an airlift reactor with a membrane separation in the downcomer (Mihaľ et al., 2013) and an innovative four phase loop reactor (Sajc et al., 1995). In this study, an alternative reactor setup is presented for an airlift reactor with an integrated counter-current liquid-liquid extraction in the downcomer. The reactor is applied for patent at the European Patent Office (Bednarz et al., 2016). In this reactor a gaseous phase is dispersed in the upcomer of the reactor inducing the loop flow of the continuous aqueous phase. The extractive organic phase is dispersed in the downcomer and is capable of extracting a value 3

compound out of the continuous phase. The principle of the reactor is depicted in Figure 1 for a setup as airlift reactor. Here the gaseous phase is dispersed in the annular gap between the inner and the outer cylinder and causes the loop flow. The organic phase is dispersed in the inner cylinder and rises in a counter-current manner to the continuous phase. Since an oxygen/solvent atmosphere depending on the chemicals used in the reactor can occur appropriate safety precautions have to be taken.

Synthetic air

Circulating aqueous phase Organic phase

Air disperser Organic disperser

Figure 1: Principal flowsheet of the innovative reactor concept For the development of different multiphase reactors and separation apparatus the understanding of the hydrodynamics is of particular importance. The velocity fields at low dispersed phase fractions can be analyzed by expensive experimental setups such as Particle Image Velocimetry (PIV), Laser Doppler Anemometry (LDA) and Laser Doppler Velocimetry (LDV). But at higher fractions the dispersed phase hides the flow of the continuous phase and complicates the measurements. In the recent year’s computational fluid dynamics (CFD) has been advanced and widely used for the prediction of flow fields for single phase and two phase contacting apparatus. Different authors showed by comparison of 4

simulation results with experiments like PIV and LDA that CFD is suitable for the description of the flow field in rotating disc contactors (RDC) (Rieger et al., 1996; Fei et al., 2000; Drumm and Bart, 2006; Drumm et al., 2011; Chen et al., 2014), the Taylor-Couette Disc Contactor (Aksamija et al., 2015), pulsed columns (Bujalski et al., 2006; Amokrane et al., 2014), bubble columns (Bhole et al., 2008; Hlawitschka et al., 2013) and bioreactors (Coughtrie et al., 2013), mostly capturing the overall flow profile. The deviations lie in the exact shape of vortexes and in the position of dead zones (Drumm et al., 2011; Chen et al., 2014). Only few publications are found that deal with multiphase CFD simulations with more than two phases. Li et al. (2012) developed an simulation model for a centrifuge (gas-liquidliquid). They use a coupled model consisting of an Euler-Euler and a population-balance model. The experimental throughput of the centrifuge could be simulated by the model. A cracking reaction of a three-phase mixture (gas-liquid-solid) in an airlift-loop reactor was analyzed by Chang et al. (2012). Reaction and heat transfer were included into the model besides the hydrodynamics. However, there has been no experimental validation. A CFD simulation of a three-phase separator in industrial scale was realized by Kharoua et al. (2013). Their grid contained about 8.5 million elements but still the element size was around 5 cm. The authors noted that the grid is too coarse but for an industrial-scale separator the already high computing effort would be even higher. They also coupled the Euler-Euler approach with population balances. All the above mentioned groups used the Euler-Euler methodology. In contrast Zhang and Ahmadi (2005) incorporated a Euler-Lagrange model to simulate a gasliquid-solid stream in a bubble column. As the literature overview shows there has not been a three phase CFD simulation so far that is comparable to the here introduced novel loop reactor. But the reactor consists of parts comparable with conventional apparatus. The reactor is composed of a modified airlift reactor 5

with an extraction part in the center of the reactor. For these reactors there have been plenty publications concerning CFD models. The standard approach in a multiphase CFD simulation is the Euler-Euler approach. There each bubble or droplet is not simulated individually but is accounted for by volume fractions in each element of the grid. Euler-Euler models have been successfully used for different extraction columns such as RDC (Vikhansky and Kraft, 2004; Drumm et al., 2009; Drumm et al., 2010; Drumm et al., 2011; Aksamija et al., 2012; Jildeh et al., 2012; Chen et al., 2014; Aksamija et al., 2015) and Kühni (Bart et al., 2011; Jildeh et al., 2014). Another possibility offers the Euler-Lagrange approach. Here the dispersed phase is tracked directly. This leads to high computational time when simulating higher holdups. This concept is used to analyze different setups of bubble columns and bubble interaction (Bai et al., 2012; Jain et al., 2013; Roghair et al., 2013; Jain et al., 2014). In recent years the Euler-Euler simulation model of the RDC was improved by means of adding a population balance model to account for coalescence and breakage of the droplets (Drumm et al., 2009; Drumm et al., 2010; Chen et al., 2014). These simulations give better results than the simulation without population balances. Though it is mentioned by Bart et al. (2011) who coupled the population balances with the Euler-Euler approach for a Kühni compartment that model parameters have to be fitted to experiments. The here proposed novel reactor does not only consist of an extraction part but also includes an oxygen supply for the microorganisms. The reactor design is similar to an internal airlift reactor. Advantages of an airlift reactor are a good gas-liquid mass transfer and less shear stress compared to conventional aerated stirrer vessels (Contreras et al., 1999; Ruitenberg et al., 2001). Airlift reactors are often simulated with the Euler-Euler approach (Oey et al., 2003; Talvy et al., 2005; Talvy et al., 2007; Xu et al., 2011; Ghasemi and Hosseini, 2012; Coughtrie et al., 2013; Rzehak et al., 2015). In these models the drag force is always included. The turbulent distribution is only considered by Oey et al. (2003), Talvy et al. (2007) and (Rzehak 6

et al., 2015). The coupling with population balances was integrated into models of bubble columns (Chen et al., 2005; Wang and Wang, 2007; Bhole et al., 2008; Hlawitschka et al., 2013). It was stated that the simulation results are in better agreement with the experiments compared to a single Euler-Euler simulation. The aim of this study is to present a novel multiphase loop reactor. The reactor is developed by incorporating a CFD simulation. A priori experiments of a first reactor design were used to prove the validity of the CFD model. The study begins with a presentation of the experimental set up and the simulation model. Afterwards the comparison of experiments and simulation is presented and an improvement of the reactor based on additional simulations. At the end a short conclusion will be given. 2. Materials and methods 2.1 Chemicals Water, distilled in a MonoDest 3000 (Lenz Glas Instrumente, Wertheim, Germany), and a typical organic solvent, represented by Shellsol T (Technical grade, CAS: 64741-65-7), also referred to as kerosene, from Shell Chemicals, Rotterdam, Netherlands were chosen to investigate the prototype of the reactor and to validate the results from the simulations. Sodium chloride (Analytical grade, CAS: 7647-14-5) supplied by Merck Millipore, Billerica, USA was added to the aqueous phase in a concentration of

to ensure consistent phase

properties (Soika and Pfennig, 2005). As gaseous phase synthetic air from Praxair, Danbury, USA consisting of

oxygen and nitrogen was chosen.

2.2 Experimental setup The experimental setup of this study is based on an outer cylinder and an inner cylinder. The dimensions of the experimental setup that was used in the first simulations are given in Fig. 2 and Tab. 1. The lab setup only differs in two points: The fixation of the internals and the dead

7

volume at the bottom of the reactor are neglected for simplification. At the top of the inner cylinder a baffle plate is installed to prevent the gaseous phase from entering the inner cylinder. Above the inner cylinder a capturing cylinder is positioned to separate the two dispersed phases. The disperser of the gaseous phase consists of 30 evenly distributed holes of 0.6 mm diameter in a ring. The disperser of the organic phase consists of a filter plate where the droplets of the organic phase are evenly distributed. The organic phase was delivered by a dosing pump DULCOMETER D_4a from ProMinent Dosiertechnik, Heidelberg, Germany. To investigate the flow field inside the reactor in some experiments

PA 6.6 spheres

were added to the reactor and their course was visually tracked to qualitatively evaluate the flow field. a b

e

c

k

g

f

d

j h

aa

x

n

s

l

y z bb

m

i

cc t

o

r

w v u

p q

Figure 2: Experimental setup of developed multiphase-loop reactor Table 1: Dimensions of experimental setup of developed multiphase-loop reactor in Fig. 2 abbr. [mm] abbr. [mm] abbr. [mm] a

100

k

0

u

30.3

b

50

l

10

v

25.5

8

c

10

m

15

w

4

d

14

n

90

x

70

e

5

o

7

y

5

f

50

p

10

z

35

g

104

q

7

aa

1.5

h

300

r

3

bb

10

i

409

s

1

cc

70

j

45°

t

3

2.3 Experimental properties The aeration rate was determined based on an approximated required Oxygen Transfer Rate (OTR) of

which is a sound magnitude for common microorganisms like

Escherichia coli (Hansen et al., 2012; Rahmen et al., 2015). The OTR inside an airlift reactor has already been investigated (Kawalec-Pietrenko and Holowacz, 1998). Resulting from the reactor design and including a safety factor a maximum mass flow of

pure oxygen

was determined. For the preliminary examinations in the pilot plant synthetic air was utilized as gaseous phase due to costs and safety issues. To define the necessary mass flow of the organic solvent a typical biotechnological production was utilized. A production rate of the microorganism was assumed to account for with an inhibiting concentration of 100 and presumed equilibrium at the outlet a mass flow of

. With a partition coefficient of was determined for the

organic phase. 2.4 Simulation Model In a first approach simulations are carried out to determine the best position of the capturing cylinder for the organic phase. Then the geometry of the experimental setup was simulated and compared to experimental results. Afterwards the geometry was adjusted by the means of 9



the size of the internal cylinder and simultaneously changing the baffle plate and



the size of the organic phase disperser.

The modified geometry of the apparatus is then computed and checked if the flow pattern has improved. Different operation points were tested by varying the inlet mass flow of air between

and

. The inlet flow of kerosene is mainly set constant at

But in some simulations it is increased to

.

to check the influence on the performance

of the reactor. At the end the geometry given in Fig. 3 turned out to be the most promising.

110 5

Ø200 Ø100 10 7

64.6

Ø10

0 Ø1

70

410

300

° 20

3 15 Ø160 Ø180

Figure 3: Dimensions [mm] of the developed multiphase-loop reactor 2.4.1 Three-fluid model The simulations were conducted with the CFD software Fluent 15.0 from ANSYS, Canonsburg, USA. For the simulation an Euler-Euler approach is applied. Even if the phase interface cannot be resolved with this technique it incorporates a momentum equation for each phase. The aim of the simulations is to prove the functionality of the reactor design. Since the functionality is mainly affected by the fluiddynamics in the reactor the mass transfer between the phases is neglected at this stage. As the drop and bubble distribution is not of interest in 10

this stage of investigation population pheomena will not be considered either. In the momentum equation only interphase exchange forces as gravity, drag and turbulent dispersion force are included. Other forces such as virtual mass have a minor effect (Talvy et al., 2007). Consequently they are not included. The mass and momentum balance can be written for the phase

as following: (1) (2)

Where

expresses the volume fraction,

by all phases,

the stress tensor,

the density,

the gravity,

the velocity,

the pressure shared

the dragforce and

the turbulent

dispersion force. The sum of all volume fractions has to be one: (3) The model of Schiller and Neumann (1935) has been applied for simulations of extraction columns (Bardin-Monnier et al., 2003; Vikhansky and Kraft, 2004; Drumm et al., 2010; Chen et al., 2014; Attarakih et al., 2015; Chen et al., 2016) as well as bubbly flows (Xu et al., 2011; Ghasemi and Hosseini, 2012; Elqotbi et al., 2013; Morchain et al., 2014). Therefore, even though it was developed for solid particles for the first development stage of the multiphaseloop reactor it is employed for the drag modeling as follows: (4) (5) (6) (7) (8) 11

Here

represents the interphase exchange coefficient,

the viscosity of the continuous phase, drag coefficient and

the diameter of the particles,

the velocity of phase ,

the drag function,

the

the relative Reynolds number.

In lots of CFD publications the turbulent dispersion is neglected (Drumm et al., 2008, Drumm et al., 2010, 2010; Drumm et al., 2011; Xu et al., 2011; Ghasemi and Hosseini, 2012; Coughtrie et al., 2013; Chen et al., 2014). In the presented CFD model turbulent dispersion has to be considered. Without this model the dispersed phase would not spread in the simulations as observed in the experiments. Since Talvy et al., 2007 used the model of Simonin and Viollet (1990) in an airlift reactor it is included here and looks like subsequent (9) where

is a constant,

of turbulent quantities and

is a fluid-particle turbulent dispersion term that is calculated out the dispersion Prandtl number. Since constant

was varied

between 0 (no turbulent dispersion), 0.2 and 0.4 to fit experiments the dispersion Prandtl number can be set to its standard value of 0.75. The comparison between simulations and experiments showed that a quantity of 0.2 is adequate for both dispersed phases. The turbulence quantities are calculated with the standard k-ε turbulence model using the averaged quantities of the mixture. The k-ω SST turbulence model suggested by Coughtrie et al. (2013) was tested and resulted in unrealistic results. In these simulations the gas phase takes a curved path in the riser as periodic vortexes in the continuous phase are calculated. This could not be observed during the conducted experiments. 2.4.2 Mesh and boundary conditions A 2D axis-symmetric model of the reactor was used in the simulation. Geometry and mesh were generated with DesignModeler of ANSYS, Canonsburg, USA and ANSYS Meshing. 3D models are known to be more accurate but at the same time these simulations need much 12

more computational resources. Still Xu et al. (2011), Ghasemi and Hosseini (2012) and Coughtrie et al. (2013) use a 2D geometry with rotational symmetry for internal airlift reactors. This assumption is applicable since the main flow in airlift reactors and in the here developed loop reactor is primarily axis-symmetrically. In addition, Coughtrie et al. (2013) achieves a good agreement between simulation and experiments monitored by PIV analysis. For sensitivity analysis four different mesh sizes (number of nodes

,

respectively element size in

,

and ) were

analyzed concerning gas hold-up, gas mass-flow at the top of the liquid surface and water mass-flow at two different interfaces in the reactor. The simulation was performed only with aeration since air bubbles rise faster than the extractive organic phase. The second finest mesh with an element size between

and

is chosen because the results of holdup and mass

flow differ from the finest mesh with a maximum deviation of

.

For the boundary condition of the wall the standard wall function with a no slip condition is used. The dispersed phases enter the loop reactor at the top of the associated disperser as given in Fig. 3. Therefore, velocity inlet boundary contitions are applied. The phases in the simulation are the same as in the experiments (air, water and kerosene). It is assumed that there is no population behaviour like coalescence or breakage of drops or bubbles. Thus the simulations are performed with a monomodal distribution of spherical air bubbles with a diameter of (2011) analysed a diameter of

and of spherical kerosene droplets with a diameter of

. Šimčík et al.

airlift reactor with a high speed camera and stated that a bubble is adequate. During the performed experiments the bubble size was

visually obtained and a size of

gives a good approximation of the bubble size

distribution. Though the drag model of Schiller and Neumann (1935) does not account for deformed bubbles it was successfully used for the CFD simulation of extraction columns and 13

bubbly flows as discussed in section 2.4.1. Thus for the development of the reactor this assumption is accepted. The bubbles enter the computational domain with a velocity of . The mass flow is adjusted by the porosity (between area so that a mass flow between

and

enters the domain with a velocity of which corresponds to a mass flow of

) of the inlet

of air is achieved. The kerosene or

or

and

(porosity of inlet area

)

. At the top of the reactor pressure

outlet boundary conditions are used to account for the exiting air. Since during the simulations only a little amount of kerosene accumulates at the top of the water surface the organic does not need an outlet boundary condition. To sum up for the testing of the different geometries and operation points 39 CFD simulations have been performed. 2.4.3 Discretization schemes and solver For the solution of the Navier-Stokes equation system the pressure-based solver is used. Thus the pressure equation is solved to achieve mass conservation. The task is solved in a transient manner and the pressure-velocity coupling is performed with the coupled algorithm. The time-step size was set to

. Fluent uses a variable called “Flow Courant Number” for

the stabilization of the convergence behavior. A quantity of

worked best for the

simulations. For the discretization of the convection terms of momentum and turbulence an upwind scheme is used. The volume fractions are discretized by the QUICK scheme and the gradients with a least-square cell based method. For the temporal discretization a first order implicit scheme is applied. During the simulation the residuals lie between 10-4 and 10-5. The physical convergence is accomplished by monitoring the volume fraction of kerosene and the mass flow of air, water and kerosene at different control areas in the reactor such as bottom, middle and top. The data are recorded every tenth time step. The volume fraction of air is recorded every fifth time step. 14

For the simulation of the three-phase system following procedure was applied. First, the reactor was operated for

with water only and an inlet stream of air. In this time the inner

circulation of the water was formed by the loop-inducing gaseous phase. Subsequently, the inlet of the organic phase is added and the simulation continued for another

. For all

simulations this procedure was sufficient to achieve steady state in the reactor. 3. Results & Discussion 3.1 Proof of Concept The accuracy of the simulation is determined qualitatively (phase distribution) as well as quantitatively (gas hold-up) in comparison to the experimental study. In Fig. 4 the simulation of the experimental setup with a mass flow of

air and

kerosene is presented.

The volume fraction of the dispersed phases is shown in levels of grey. The mass flow of air is sufficient for inducing the loop flow. With this setup no bubbles of the gaseous phase enter the inner part of the downcomer (Fig. 4 – left). But an accumulation of the gas phase at the baffle plate is obtained in the experiment and simulation. At the edge of the baffle plate and underneath the capturing cylinder an accumulation of air bubbles occurs. During the start-up process of the reactor the first bubbles are entrained into the capturing cylinder and accumulate at both spaces. Afterwards the water velocity is higher than the terminal velocity of the bubbles so that in the simulation air is trapped at this location. Looking at the right side of Fig. 4 part of the dispersed phase is carried into the upcomer (the corresponding video of the

simulation

is

found

in

simulation_volumefraction_kerosene_massflowrate_0.6g_per_s.avi). This is caused by the high velocity of the continuous aqueous phase in the zone between organic disperser and inner cylinder. From the organic phase disperser, the other part of the organic phase rises

15

upwards. When it reaches the top of the reactor it is led into the center of the inner cylinder because of the inlet flow of the aqueous phase.

αair [-]

αkerosene [-]

1.0

0.020

0.9

0.018

0.8

0.016

0.7

0.014

0.6

0.012

0.5

0.010

0.4

0.008

0.3

0.006

0.2

0.004

0.1

0.002

0.0

0.000

Figure 4: Contour with volume fraction of air (left) and kerosene (right) of the experimental geometry with an inlet mass flow of 0.52 g/s air and 0.6 g/s kerosene after 50 s The flow field of the continuous phase is depicted in Fig. 5. In the simulations the flow field can be roughly divided into three areas. In the upcomer an even distribution is present with low velocities near the walls and a high velocity in the gap where the continuous phase is accelerated and rises in conjunction with the dispersed gaseous phase. The second area consists of the outer part of the inner cylinder, where the velocity is highest in the inlet at the baffle plate and slowly decreases downwards while dividing into a larger spatial area until the organic disperser is reached. The third area includes the upper inner part of the inner cylinder. In this area an additional vortex is observed in the simulations of all air flow-rates as well as during the corresponding experiments. This vortex results from the flow of the continuous 16

phase entering the downcomer through the gap between the baffle plate and the capturing cylinder.

vwater [m/s]

0.50 0.45 0.40 0.35 0.30 0.25 0.20 0.15 0.10 0.05 0.00 Figure 5: Contour and vectors with water velocity of the experimental geometry with an inlet mass flow of 0.52 g/s air and 0.6 g/s kerosene after 50 s The experimental study of the reactor is given in Fig. 6. A lower air flow was tested in Fig. 6 a) and b) than in the presented simulations in Fig. 4 and 5 but due to the reactor set up at higher air flows bubbles hide the organic droplets and particles in the inner cylinder. Therefore, the inner flow is hardly visible on a frame of the videos. Since the principle behavior remains the same some results with less aeration are presented. The flow field of the continuous phase was observed by the means of adding spheres to the reactor. In Fig. 6 a) the qualitative track of three spheres captured with a high-speed camera is shown. The camera 17

shows the upper three-fourths of the reactor. The PA 6.6 spheres enter the inner cylinder from the top coming from the riser. Then they are captured in the vortex which is also seen in the simulations and rise into the capturing cylinder. The path of the organic phase during an experiment with an air flow of

is presented in Fig. 6 b). Here the inner cylinder

above the organic disperser is shown. After injection the drops are nearly in the middle of the cylinder. During their rise they move closer together but more to the right side of the reactor. The concentration of the drops in the center of the reactor was represented in the simulations (Fig. 4 right). The shift to the right is only found in the experiment with low volumetric flow rate of the gaseous phase. At this operation point the air is mainly dispersed on the left side of the reactor. Consequently, the vortex loses its symmetry and the drops are dragged to the right. The entrainment of drops is shown in Fig. 6 c). The record of the experiment is given in experiment_entrained_drops.avi. The drops are partly dragged underneath the ring that holds the inner cylinder and then enter the riser. Since more drops were dispersed at the left side of the disperser on this side more drops were entrained. The entrainment was also observed in the simulations in Fig. 4 right. Summing up the principal behavior of the drops and of the water flow in the reactor could be simulated with the chosen CFD model.

18

Capturing Cylinder a)

b)

c)

Figure 6: Experimental study in a pilot plant of the novel multiphase-loop reactor. The particles are marked with circles. a) Qualitative path of 4 mm PA 6.6 spheres with an inlet mass flow of 0.26 g/s air; b) Movement of drops in the inner cylinder with an inlet mass flow of 0.13 g/s air and 0.6 g/s kerosene; c) Entrained drops into the riser with an inlet mass flow of 0.52 g/s air and 0.6 g/s kerosene Besides the visual evaluation and validation of the flow field the holdup of the gaseous phase was determined as a quantitative indicator for comparison with the simulations. The holdup of the gaseous phase is determined based on the shift in height of the continuous phase during operation and after stopping both disperse phase flows. The determination of the holdup of the 19

organic phase was not possible resulting from the low volumetric flow rates of the organic phase. In Fig. 7 the holdup of the gaseous phase is shown depending on the volumetric aeration rate. At lower aeration rates slightly lower values for the holdup are obtained compared to the simulation while at higher aeration rates the deviations decrease. 6

5

holdup in %

4

continuous phase: organic phase: organic phase flow: gaseous phase:

water Shellsol T 0.6 g/s synthetic air

simulation experiment

3

2

1

0 0.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

aeration rate in m³/h

Figure 7: Comparison of experimental and simulated hold-up of air In Fig. 8 the entrained organic phase into the riser is plotted over the flow rate of air and compared to entrained drops during the experiments as given in Fig. 6 c. The entrained drops were recorded with a high speed camera and counted in addition. The quite large error bars occur since the organic phase disperser hides part of the drops. Similar to the results in Fig. 7 at high flow rates simulation and experimental results fit well. Only at an aeration rate of experiments and simulations differ. The reason for the deviation is probably the non-even and pulsed dispersion of the organic phase and in homogeneities of the gas phase dispersion at lower aeration rates.

20

45

entrained dispersed phase in %

40

simulation experiments

35 30 25 20 15 10 5 0 0.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

aeration rate in m³/h

Figure 8: Comparison of experimental and simulated entrainment of dispersed phase Overall, the simulation and the experiments (Fig. 7 and 8) are in good agreement and validate the simulation in addition to the presented qualitative visual observation of the flow field in the reactor. 3.2 Simulation-based improvement of reactor design On the basis of the validated process model a study focusing on the improvement of the design is performed. The results for the realized experimental setup reveal that the disperser of the organic phase is not suitable for the operation of the reactor. About

of the

entering kerosene is entrained into the area of ascending air bubbles due to a high downward directed velocity of the continuous phase (air inlet of

) next to the disperser. Thus the

requirement that both dispersed phases can be removed separately and do not get in contact with each other is violated under these conditions. As a consequence, instead of the funnel like disperser shape (Fig. 2) a ring disperser is recommended and installed. In the first approach the disperser has a diameter of is arranged in a circle with a diameter of

and

. The advantage is that the continuous

aqueous phase can flow downwards alongside and through the disperser ring. Thus, the surrounding water velocity should be less. In addition, the inner cylinder is enlarged in two steps from

initially to

for the final design so that the interface in the region 21

of ascending air is reduced and the interface in the inner cylinder increases. This results in a lower velocity of the continuous phase in the inner cylinder. The simulation (mass flow inlet air

and kerosene

) shows that still a part of the fed kerosene is entrained

into the area of aeration. At the beginning of the input of the organic phase the ascending kerosene can rise. After about

kerosene is dragged towards the wall of the inner cylinder

and then swept away by the continuous phase flowing downwards. This phenomenon occurs periodically. Therefore, with this reactor design no stationary operation point can be met under these circumstances. By reducing the size of the disperser to a diameter of

kerosene can rise without being

entrained. This results in the previously described geometry in Fig. 3. The simulation results of the developed reactor geometry are described in the following. 3.2.1 Circulating continuous phase In the reactor the continuous aqueous phase circulates to confirm a good oxygen supply of the microorganisms and to achieve a counter-current flow for the extraction. The velocity profile and the flow field evolving in the reactor are shown in Fig. 9. In the gap between inner cylinder and outer wall the continuous phase flows upwards since it is dragged by the air. At the upper part of the gap the water flow makes a turn and enters the inner cylinder. Due to the jet into the inner cylinder the water flow shapes a vortex directly underneath the capturing cylinder. Thus, in the upper central part of the inner cylinder the water flows upwards resulting in the already presented and experimentally confirmed internal vortex. Further down in the cylinder the flow of the continuous aqueous phase is directed downwards. In the whole inner cylinder the main downwards directed flow is next to the cylinder wall. The arising vortex and concentration of flow next to the wall leads a decreased performance of the reactor. On the one hand a part of the aqueous phase entering the inner cylinder is temporary 22

captured in the vortex. So this part of the continuous phase exhibits less oxygen supply since it gets less contact with the oxygen containing dispersed gaseous phase. On the other hand, there is no complete counter-current flow for the extraction and the flow is smallest next to the disperser of kerosene. The separation performance is going to be lower as the residence time is reduced compared to complete and unvarying counter-current flow.

v water [m/s] 0.55 0.50 0.45 0.40 0.35 0.30 0.25 0.20 0.15 0.10 0.05 0.00

Figure 9: Contour and vectors with water velocity of develop reactor geometry with an inlet mass flow of 0.52 g/s air and 0.6 g/s kerosene after 50 s 3.2.2 Distribution of gaseous phase One important characteristic is the distribution of the oxygen containing gaseous phase in the reactor. The contour of the volume fraction of air in Fig. 10 shows that the air rises in the gap and is then dragged with the water around the baffle plate. Next to the baffle plate the air 23

accumulates. The rest leaves the reactor close to the wall of the capturing cylinder. A little amount of air accumulates underneath the capturing cylinder. This has the disadvantage that some air is entering the capturing cylinder which is intended to collect only the dispersed organic phase. This could result in an anewed dispersion of the coherent organic phase in the top of the capturing cylinder. To prevent this amongst others three options exist. One is to place the capturing cylinder a little deeper. The other is to chamfer the downside of the cylinder. In addition, the baffle plate could be modified so that the fluids do not have to take a 180° turn. This would also reduce energy input at the top of the baffle plate so that the formation of small bubbles is reduced. But all in all the functionality of the reactor according to the distribution of air is achieved.

αair [-]

1.0

αkerosene [-] 0.020

0.9

0.018

0.8

0.016

0.7

0.014

0.6

0.012

0.5

0.010

0.4

0.008

0.3

0.006

0.2

0.004

0.1

0.002

0.0

0.000

Figure 10: Contour with volume fraction of air (left) and kerosene (right) of develop reactor geometry with an inlet mass flow of 0.52 g/s air and 0.6 g/s kerosene after 50 s 3.2.3 Distribution of kerosene 24

On the right hand side of Fig. 10 the contour of the volume fraction of kerosene is depicted. The organic phase is rising and spreads in the lower part of the inner cylinder. No part of the entering fluid is entrained by the continuous aqueous phase and the complete phase is collected in the capturing cylinder. When the dispersed phase reaches the region with the vortex it accumulates in the center. Because the continuous phase changes flow direction the organic phase is accelerated and the fraction of the organic phase decreases in the upper part of the cylinder. This causes a possible loss in separation performance of the extraction. The simulations were also performed with a much higher mass flow of dispersed organic phase. Instead of

the mass flow was adjusted to

. The results show only

slight differences to the above presented and the overall behavior remained the same. 4. Conclusion In this study a successful development of an innovative reactor concept was presented. This reactor design enables an aeration of a fermentation process combined with an integrated extraction. Thus it is possible to produce and simultaneously separate toxic or inhibiting products from aerobe biotransformations in one apparatus. The proposed reactor setup was put into operation in pilot-plant scale where part of the dispersed organic phase was entrained into the outer cylinder gap. To analyze the flow CFD simulations were conducted which were validated in the pilot-plant setup. The comparison of experiments to the chosen CFD model show good agreement. Thus the model is suitable for the presented reactor development. For a more accurate description of the flow more detailed models have to be used as for example given in Rzehak et al., 2015. Therefore, based on these simulations the design of the reactor was adjusted and improved so that no entrainment occurs. In further investigations the opposite circulation direction of the continuous phase will be tested in the reactor. In this setup the loop-inducing gaseous phase will be dispersed in the inner cylinder. Consequently, the arrangement of the other parts has to be adapted. This set up 25

has the advantage that the vortex in the upper part of the downcomer would evolve smaller in the outer gap. This would result in a nearly complete counter-current flow and an increase of the possible separation performance. As a next step the fluid dynamic and the mass transfer of this enhanced reactor setup will be examined. Summing up, the new and innovative reactor setup enables the realization of a wide range of possible processes. Especially biotransformations with a product inhibition at low concentrations and necessary aeration appear to be a promising field for utilization and will be investigated.

Acknowledgments The authors are very grateful for innumerous contributions to the case by Markus Schmidt, Bettina Rüngeler and Markus Lückge (All AVT-Fluid Process Engineering of RWTH Aachen University). They are thankful for the experimental assistance of various student assistants. We would also like to thanks the IT-Center of the RWTH Aachen University for the computational resources.

This research did not receive any specific grant from funding agencies in the public, commercial, or not-for-profit sectors.

Symbols [-]

volume fraction of phase q

[-]

volume fraction of phase p

[kg/m³]

density of phase q

[Pa s]

viscosity of phase q 26

[kg/(m² s²)]

stress tensor

[-]

drag coefficient

[-]

constant for turbulent dispersion

[m]

diameter of particle of phase q

[m/s²]

fluid-particulate dispersion tensor

[-]

drag function

[N/m³]

drag force

[N/m³]

turbulent dispersion force

[m/s²]

gravitation

[kg/(m³s)]

interphase exchange coefficient

[Pa]

pressure

[-]

dispersion Prandtl number (definition:

)

[m/s]

velocity of phase p

[m/s]

velocity of phase q

[-]

relative Reynolds number

Abbreviations CFD

Computational Fluid Dynamics

LDA

Laser Doppler Anemometry

LDV

Laser Doppler Velocimetry

PIV

Particle Image Velocimetry

RDC

Rotating Disc Contactor

27

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Graphical abstract

35

Highlights

-

Innovative multiphase loop reactor for in-situ extraction and aeration

-

Experimental validation of multiphase CFD simulations

-

Loop reactor with countercurrent extraction

36