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Hydrodynamics and mass transfer in multiphase monolithic reactors with different distributors: An experimental and modeling study
Accepted Manuscript Title: Hydrodynamics and mass transfer in multiphase monolithic reactors with different distributors: An experimental and modeling study Authors: Yanyan Guo, Chengna Dai, Zhigang Lei PII: DOI: Reference:
S0255-2701(17)30896-6 https://doi.org/10.1016/j.cep.2018.01.026 CEP 7181
To appear in:
Chemical Engineering and Processing
Received date: Revised date: Accepted date:
6-9-2017 27-1-2018 29-1-2018
Please cite this article as: Yanyan Guo, Chengna Dai, Zhigang Lei, Hydrodynamics and mass transfer in multiphase monolithic reactors with different distributors: An experimental and modeling study, Chemical Engineering and Processing https://doi.org/10.1016/j.cep.2018.01.026 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
Hydrodynamics and mass transfer in multiphase monolithic reactors with different distributors: An experimental and modeling study
Yanyan Guo, Chengna Dai, Zhigang Lei* State Key Laboratory of Chemical Resource Engineering, Beijing Key Laboratory of Energy
Corresponding author. E-mail address: [email protected] (Z. Lei).
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*
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Environmental Catalysis, Beijing University of Chemical Technology, Box 266, Beijing, 100029, China
Graphical abstract
The hydrodynamic and mass transfer performances in monolith bed with different
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distributors were studied by the combination of experiments and CFD simulation.
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Highlights o The gas-liquid distribution of monolith bed with different distributors was studied; The hydrodynamic and mass transfer experiments and CFD simulation were carried
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out;
Packed bed distributor was superior to the foam and porcelain ring distributors.
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Abstract This work focuses on the hydrodynamic and mass transfer performances in
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multiphase monolithic reactor with different distributors by the combination of experiments
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and CFD simulation. The experimental results showed that the pressure drop in the monolithic bed with packed bed distributor is lower than that with other distributors, and that
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the liquid holdup follows the trend of packed bed distributor > foam distributor > porcelain
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ring distributor. Moreover, the gas-liquid mass transfer coefficients increase with the increase of superficial velocities of both gas and liquid phases, and mainly depend on the superficial
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liquid velocity. The CFD simulation showed that the gas-liquid phase distribution is mainly dependent on the superficial velocity of liquid phase, and the changes of superficial gas
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velocity don’t have an obvious effect on the gas-liquid phase distribution. Furthermore, the experimental and simulated results showed that the mass transfer coefficients in the monolithic bed with packed bed distributor is much higher than those with foam distributor and porcelain ring distributor.
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Keywords: liquid holdup; pressure drop; mass transfer; gas-liquid phase distribution; CFD
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simulation; distributor
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1. Introduction In recent years, the multiphase monolithic reactors have received wide attention and are regarded as an intensification replacement of conventional multiphase reactors (e.g., trickle beds and slurry bubble columns) in the catalysis community, due to the low pressure drop and
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high mass transfer coefficients [1-10]. Gao et al. [11] prepared a novel cordierite monolithic catalyst for CO coupling to oxalate, and the result showed that the Pd efficiency of the
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monolithic catalyst (274 h-1) is much higher than that of a traditional pellet catalyst in packed bed reactor (46 h-1), which is attributed to the high Pd dispersion on the surface of the
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monolithic catalysts. Guo et al. [12] demonstrated that the monolithic reactor has higher
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propylene selectivity and methanol conversion than the random packed bed for the
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methanol-to-propylene process. Zamaniyan et al. [13] reported that tube fitted bulk
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monolithic reactor has superior advantages over the conventional packed reactor, especially
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with the low pressure drop and high effective usage of the catalyst in the natural gas reforming process. However, the channels of monolithic reactors are independent, and there
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is no redistribution for gas and liquid in the monolithic bed. Thus, the gas-liquid distributor at
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the entrance of the reactor plays an important and crucial role in the multiphase monolithic reactor performance, and a uniform gas-liquid phase distribution among monolithic channels is essential for reactor scale-up.
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The main function of the distributor is to disperse the gas and liquid phases. It has a
crucial effect on the shape of liquid drop and the lengths of bubble and liquid slug in the channels [14,15]. Many efforts have been done on the effect of distributor design on the gas-liquid phase distribution, gas and liquid holdups, and gas-liquid mass transfer by
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experiments and CFD simulation. Satterfield and Ozel investigated the relation between different distributors and pressure drop in monolithic beds, and found that the design of suitable distributors plays a vital role on the bed pressure drop [16]. Xu et al. [17] investigated the pressure drop and liquid holdup in multiphase monolithic reactor with
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different distributors, with the result that the packed bed distributor has better gas-liquid distribution capability and the design of distributor has a crucial influence on the liquid
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holdup, pressure drop, and mass transfer. Li et al. [18] investigated the effect of gas
distributors on the gas holdup and mixing characteristics by CFD simulation, and found that
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the gas distributor is conductive to improving the gas holdup. Roy and Al-Dahhan [19]
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investigated the liquid phase distribution in a monolithic bed with three different distributors
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(i.e., nozzle, foam, and showerhead) under Taylor flow regime, nozzle being the most
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suitable distributor for the monolithic bed. However, the studies on the effect of distributor
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types (i.e., porcelain ring distributor, foam distributor, and packed bed distributor) on the hydrodynamic (i.e., gas-liquid phase distribution, pressure drop, and liquid holdup) and mass
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transfer performances are still very limited.
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Thus, this work is on addressing on the dynamic intensification of monolithic bed with different distributors from the following three aspects: (i) determining the optimal type of distributor by measuring the liquid holdup, pressure drop, and mass transfer in the monolithic
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bed with different distributors and different channel sizes; (ii) establishing the CFD model to simulate the gas-liquid phase distribution at the entrance of monolithic bed with different distributors and to compared with the experimental results; and (iii) using the CFD model based on a Taylor unit to obtain the mass transfer coefficients in the monolithic bed with
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different distributors, and to compare between the experimental and simulated mass transfer coefficients. 2. Experimental section 2.1. Experimental setup
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The cold model experimental setup mainly consisted of a monolithic reactor (D = 100 mm, and H = 600 mm) and the measurement system for liquid holdup, pressure drop and
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gas-liquid mass transfer. The cold model setup was carried out in the continuous co-current
down-flow mode, with air and water introduced at the top through the distributor. In this
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work, three distributors (i.e., porcelain ring distributor, foam distributor, and packed bed
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distributor) were tested comparatively on the liquid holdup, pressure drop and gas-liquid
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mass transfer. The packed bed distributor (200 mm in height) consisted of glass beads with
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the diameter of 1 mm, and was positioned straightly on the top of monolithic bed, with a
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close wire netting to prevent glass beads entering into the channels of monolithic bed. For porcelain ring distributor the diameter, average length, and height of the porcelain ring are 2
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mm, 3 mm, and 200 mm, respectively. The foam distributor consisted of two pieces of foam
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packing (50 mm in height and 100 mm in diameter). Both porcelain ring distributor and foam distributor were positioned straightly on the top of monolithic bed. To make sure that liquid phase spreads uniformly over the packed bed distributor, a
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plate (100 mm in diameter) containing 210 holes (2 mm in diameter) was put on the top of packed bed distributor. Below the distributors, four blocks of cordierite monoliths ( = 95 mm; H = 150 mm; and diameters of channel = 1.1 mm or 1.3 mm) were stacked as the column bed, and sealed inside the column by Teflon tapes to prevent by-pass flows of gas or
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liquid near the column wall. 2.2. Experimental procedure Hydrodynamics and mass transfer coefficients were measured with the superficial gas and liquid velocities in the range of 0.0508–0.2344 m·s-1 and 0.0392–0.1019 m·s-1,
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respectively. According to the flow regime map for a capillary two-phase flow developed by Mishima and Hibiki [20], the flow investigated in this work is within the Taylor flow regime.
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Pressure drop between the top and the bottom of monolithic bed was measured by the differential pressure transducer (MDM4951 DP3E 0.5%, MICRO SENSOR Co. Ltd.,
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China), which is capable for pressure measurements in the range from 0 to 6 kPa.
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The liquid holdup measurement was performed following two steps. First, the liquid
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holdup in the reactor as a whole, including the contributions of distributor, separator and
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monolithic bed, was measured by collecting and weighting the liquid phase after the liquid
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and gas feedings were shut down simultaneously. Then, a blank experiment, where only one block of monolith was stacked within the reactor in place of four blocks of monoliths in the
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previous case, was carried out in the similar manner. Thus, a net weight value of liquid phase
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was obtained by the subtraction of the weight obtained in the blank experiment from that obtained in previous case. That is, the net weight is the part of liquid retaining in the blocks of monolith, excluding the contributions of distributor and separator. The liquid holdup was
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measured three times in each experiment, and the repetition was found to be within 5%. The mass transfer coefficients (kLa) were obtained following two steps similar as the liquid holdup measurement. First, the kLa coefficients in the reactor as a whole, including the contributions of distributor, separator and monolithic bed, were measured by stationary
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oxygen absorption technique, in which the amount of oxygen dissolved in water was measured with Luminescent dissolve oxygen (LDO) probe (Hach, model sc100TM). Then, a blank experiment, where two blocks of monolith is in place of four blocks of monoliths, was carried out in the similar manner. The net kLa was obtained by substracting the kLa achieved
e
e '
(1)
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QL / A CL,O2 kL a = ln C Z L,O2
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in the blank experiment from that obtained in the previous case.
where QL (m3·s-1) is the volumetric liquid flow rate, A (m2) and Z (m) are the
L,O2 e
(mol·L-1) and
C L,O2
' e
(mol·L-1) are the oxygen concentrations at the exit of the
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cross-section area of monolithic bed and the height of two blocks of monolith, respectively.
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column packed with four blocks and two blocks of monoliths, respectively. The details on the
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derivation of Eq. (1) can be obtained in previous publication [21].
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The obtained kLa(T) values are converted to 20 C by the following equation [22]: kL a (20) = kL a θ 20T
(2)
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where θ ( θ = 1.02) and T (C) are the temperature coefficient and experimental
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temperature, respectively. All the kLa as mentioned below are the corrected ones kLa(20). 3. CFD simulation
3.1. Modeling of hydrodynamic performance
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3.1.1. Model description The computational domain is shown in Fig. 2(a). The distributor was positioned straightly on the top of monolithic bed. The gas phase was introduced from the side of the column, and the liquid inlet was positioned at the top. The hydraulic diameters of gas and
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liquid inlets are 25 mm and 27 mm, respectively. 3.1.2. Governing equations The Eulerian-Eulerian approach was selected to describe the gas-liquid two-phase flow using the Fluent software (version 14.5), and a steady state model with standard k–ε model to
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describe the turbulence viscosity was used. The governing equations are given as follows:
(hi i ui ) 0 Mass balance for species:
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iui hi w i Di w
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Continuity equation:
(3)
(4)
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where ρi (kg·m-3), hi and ui (m·s-1) are density, volume fraction and mean velocity of phase i
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(liquid or gas), respectively; w is the mass fraction; and Di (m·s-2) is the diffusion coefficient in phase i.
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hL hG 1
(5)
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where hL and hG are the gas and liquid volume fractions as calculated by the Anderson and Jackson method [23].
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Momentum balance equation:
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2 T hi uu hi eff uI hiP FLG hi ( FLS g ) 3
eff t
(6) (7)
where μ, μt and μeff (kg·m-1·s-1) are the corresponding molecular, turbulent and effective viscosities, respectively; and μt was obtained from the k–ε model equations; I is the unit tensor; FLG (N) is the interface drag force between gas and liquid phases, which was
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calculated by the Schiller–Naumann method [24]; FLS (N) is the flow resistance created by solid phase; and P (Pa) is system pressure. The homogeneous and isotropic porous media model was used for distributors (i.e., porcelain ring distributor, foam distributor, and packed bed distributor), and a source term (Si)
Si
a
ui C2
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is added at the right side in the momentum balance equation:
1 u ui 2
(8)
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where a (m2) and C2 (m-1) are the permeability and inertial resistance coefficient of porous
d P2 3 a 150 (1 )2
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3.5 (1 ) dP 2
(9)
(10)
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C2
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media, respectively.
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where dp = 1.0 mm, 4.0 mm and 2.0 mm are the mean diameters of glass beads particles,
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porcelain ring particles and foam particles in the distributors, respectively; and δ = 0.375, 0.78 and 0.85 are the void fractions of glass beads particles, porcelain ring particles and foam
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particles, respectively.
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3.1.3. k–ε model
In the k–ε model, the turbulence viscosity can be calculated by
t Cμ
k
(11)
2
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The turbulent kinetic energy k equation is uk t k
2 T k eff uI 3
The turbulent dissipation rate ε equation is
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(12)
u t ε
2 2 T ε C u u uI C 1 eff 2 3 k k
(13)
where the default values were used ( Cμ 0.09 , k 1.0 , ε 1.3 , C1 1.44 , and
C2 1.92 ). I is the turbulent intensity and can be calculated by 1 u' 0.16Re 8 uavg
(14)
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I
where u ' (m·s-1) and uavg (m·s-1) are the root-mean-square of the velocity and the mean
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flow velocity, respectively; Re is Reynolds number. 3.1.4. Boundary conditions
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The liquid and gas in the inlet were set to be uG = uG,inlet and uL = uL,inlet in the direction
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normal to boundary. The liquid and gas at the outlet were set to be pressure-out boundary
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condition. The walls of monolithic channel were set to be no-slip boundary condition. The
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interior of distributor in the simulation was set as porous media.
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The geometrical structure was dealt with the structured grid generation technique, as shown in Figs. 2(b) and (c). The grid independency was validated to ensure the numerical
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accuracy with different grid numbers, and the final refined grids number is up to 430,000. In
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the simulation the first-order upwind differencing scheme was chosen for the solution of all differential equations, and the SIMPLE (Semi-Implicit Method for Pressure Linked Equations) algorithm was selected for pressure-velocity coupling. The residual of 10-3 was set
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for the convergence criterion. 3.2. Modeling of mass transfer in Taylor flow The Taylor unit model consists of a gas bubble and two halves of liquid slugs on both ends of the bubble. The Taylor gas bubble is considered to be a “void” space, which acts as a
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free surface with the surrounding liquid, as shown in Fig. 3(a). Mass transfer was simulated at 25 C and 101 kPa, and the diameter of monolithic channel was 1.1 mm. The calculation procedure of liquid phase mass transfer coefficients is divided into two steps. The velocity field of liquid flow was first calculated until it was stable. Then, the mass transfer was
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calculated, and the concentration of air at the surface of gas bubble is set to be 1. The volume-averaged mass and momentum conservation equations are given as follows.
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Continuity equation:
(u L ) 0
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u L ( L u L u L L (u L (u )T ) P L g t
(16)
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Momentum balance equation:
(15)
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Mass transfer convection diffusion equation:
(17)
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( LC ) ( L vLC DC ) 0 t
where D (= 0.260 cm2·s-1) is the diffusion coefficient of air in water at 25 C and 101 kPa; C
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is the concentration of the air; and ρL (= 1000 kg·m-3) and vL are the density and volume of
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water, respectively.
The periodic boundary conditions for liquid and gas at the inlet and outlet were set (ptop
= pbottom, uL,top uL,bottom , and Ctop = Cbottom). In order to correct for the periodic boundary
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condition, the source term of hydrostatic pressure drop ρLg is added to the momentum balance equation (16) in the y direction. Simulations were carried out in a reference frame, where the system moves upward with the bubble velocity uB and the gas bubble remains stationary. No-slip boundary condition was applied for the wall of monolithic channels, the
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velocity of which was set to be uwall = –uB. Free-slip was set for the gas bubble surface: du⊥ /dn = 0 (u⊥ is the velocity component in the direction of bubble surface, and n is the normal direction to bubble surface). The geometrical structure was meshed with structured grids. The grid size less than 1
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μm was used close to the surface, with an exponential increase away from the surface as shown in Figs. 3(b) and (c). The grid independency was validated to ensure the numerical
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accuracy with different grid numbers and sizes, and the final refined grids number was 20,000. The third-order MUSCL (Monotonic Upstream-Centered Scheme for Conservation
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Laws) differencing scheme was used to deal with the momentum and mass transfer
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convection diffusion equations. The SIMPLE (Semi-Implicit Method for Pressure Linked
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Equations) algorithm was used to deal with the pressure-velocity coupling. The residuals less
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than 10-5 were chosen for the convergence criterion, and 1×10-5 s was taken as the time step.
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The average liquid velocity (uL,domain) in computational domain can be calculated by
vu
top i i
v
(18)
top i
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uL ,domain
where vi (m3) and ui (m·s-1) are the volume and the vertical velocity of cell i, respectively.
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The superficial liquid velocity (UL) can be calculated by U L (uL,domain uwall )(1 G )
(19)
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where εG is gas holdup and calculated by
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bubble volume LUC d h2 / 4
where LUC (m) and dh (m) are unit cell length and hydraulic diameter of the unit. The superficial gas velocity (UG) is written as
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(20)
U G GU B
(21)
In this work, the mass transfer coefficient in liquid phase (kLa) is calculated by kLa
(CL,system,t t CL,system,t ) / t
(22)
CL,s CL,system,t
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where CL,s (= 1) is the concentration of air in water phase at bubble surface. CL,system,t and CL,system,t+Δt are the average liquid concentrations in computational domain at two consecutive
CL,system,t
top viCL ,i top vi
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time steps (time t and t+Δt), and can be derived from
(23)
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To validate the Taylor unit model established in this work, the mass transfer coefficients
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(kLa) for methane dissolved in water were first simulated. As shown in Fig. 4, the simulated
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4.1. Gas-liquid phase distribution
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4. Results and discussion
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values agree well with literature data [25], confirming the reliability of our Taylor unit model.
Figs. 5-7 show the effect of superficial velocities of gas and liquid phases on the
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gas-liquid phase distribution in the monolithic bed with packed bed distributor, foam
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distributor, and porcelain ring distributor, respectively. It can be seen that the gas-liquid phase distribution is mainly dependent on the superficial liquid velocity, and the increase of superficial gas velocity doesn’t have an obvious effect on the distribution. That is, the
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gas-liquid phase distribution isn’t affected appreciably by the superficial gas velocity. Compared to foam distributor and porcelain ring distributor, the gas-liquid phase distribution in the monolithic bed with packed bed distributor is more uniform, which is attributed to the uniformity of glass beads in packed bed distributor.
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4.2. Liquid holdup Figs. 8 and 9 show the effect of superficial gas velocity and superficial liquid velocity on the liquid holdup with different distributors and channel densities as obtained from experiments. With the increase of superficial gas velocity, the liquid holdup in monolithic bed
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decreases; while with the increase of superficial liquid velocity, the liquid holdup increases. Evidently, the liquid holdup is affected by the superficial liquid velocity, indicating that the
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gas-liquid phase distribution is mainly influenced by the superficial liquid velocity. On the other hand, under the same superficial gas and liquid velocities, the liquid holdup in the 600
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cpsi monolithic bed is higher than that in the 400 cpsi monolithic bed for the three
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distributors, which is attributed to the larger surface area leading to better uniform dispersion
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of gas and liquid phases in the 600 cpsi monolithic bed. Among the three distributors
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investigated in this work, the porcelain ring distributor has the lowest liquid holdup, while the
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packed bed distributor has the highest liquid holdup for both 400 cpsi and 600 cpsi monolithic beds under the same operating conditions, due to the worst distribution
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characteristics for porcelain ring distributor and the uniformity of glass beads in packed bed
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distributor. The liquid holdup follows the trend of packed bed distributor > foam distributor > porcelain ring distributor. 4.3. Pressure drop
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The pressure drop reflects the energy loss, which is caused by the fluid flowing through
the monolithic bed. It directly relates to the momentum transfer, and is crucial to determine the energy dissipation. Figs. 10 and 11 show the influence of superficial gas and liquid velocities on the experimental measured total pressure drop with different distributors. The
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results show that with the increase of the superficial gas and liquid velocities, pressure drop increases. In addition, pressure drop follows the trend of porcelain ring distributor > foam distributor > packed bed distributor. The uniformity of packed bed distributor leads to a decrease in pressure drop. Thus, it is advisable to select the packed bed distributor from the
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aspect of pressure drop. 4.4. Liquid phase mass transfer
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The porcelain ring distributor exhibits the largest pressure drop with the highest energy loss among the three types of distributors investigated. Therefore, the mass transfer
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performance of other two types of distributors (i.e., packed bed distributor and foam
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distributor) was further studied here. The experimental mass transfer coefficients in the 400
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cpsi and 600 cpsi monolithic beds with packed bed distributor and foam distributor under
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different superficial gas and liquid velocities are illustrated in Fig. 12. It can be seen that the
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mass transfer coefficients in the 600 cpsi monolithic bed are higher than those in the 400 cpsi monolithic bed for both foam distributor and packed bed distributor, which is attributed to
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larger surface area in the 600 cpsi monolithic bed leading to better gas-liquid distribution.
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When the liquid velocity is low, mass transfer coefficients increase slowly with the increase of superficial gas velocity. However, when the superficial liquid velocity is high, mass transfer coefficients increase quickly with the increase of superficial gas velocity. It seems
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that mass transfer is mainly affected by the superficial liquid velocity, which should play a vital role on the gas-liquid phase distribution. Significant difference of mass transfer coefficients between packed bed distributor and foam distributor at high liquid velocity region arises. This indicates that distribution performance is particularly important at high
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liquid velocity region. Compared to foam distributor, the mass transfer with packed bed distributor is enhanced. Therefore, it is advisable to select the packed bed distributor when considering both pressure drop and mass transfer together. Fig. 13 shows the experimental mass transfer coefficients in the 600 cpsi monolithic bed
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with packed bed distributor under different superficial gas and liquid velocities, along with the simulated results by CFD model. It can be seen that both simulated and experimental
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results exhibit the same trend. That is, the mass transfer coefficients increases with the increase of superficial gas and liquid velocities. The simulated mass transfer coefficients are
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higher than experimental data, but at high liquid velocity region the discrepancy becomes
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much smaller, which can be attributed to the poor gas and liquid distribution at low liquid
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velocity region.
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5. Conclusion
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Three types of distributors, i.e., packed bed distributor, foam distributor, and porcelain ring distributor, were adopted in this work to investigate the hydrodynamic (i.e., gas-liquid
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phase distribution, pressure drop, and liquid holdup) and mass transfer performances in
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monolithic bed by the combination of experiments and CFD simulation. It was found that the gas-liquid phase distribution in the monolithic bed with packed bed distributor is the most uniform, and the superficial liquid velocity plays an important role on the gas-liquid phase
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distribution. The liquid holdup follows the trend of packed bed distributor > foam distributor > porcelain ring distributor. The pressure drop follows the trend of porcelain ring distributor > foam distributor > packed bed distributor. Therefore, one advantage of using packed bed distributor is brought out in decreasing pressure drop, which means that energy
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consumption can be saved. Another advantage is the higher mass transfer coefficients than those of other distributors. It is known that the distributor design is crucial to the reactor hydrodynamic and mass transfer performances. The packed bed distributor is considered as the most suitable when considering the holdup, pressure drop and mass transfer performances
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together. The good hydrodynamic and mass transfer performances of packed bed distributor can be attributed to the uniform gas-liquid phase distribution. Moreover, when the superficial
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gas and liquid velocities are high, the experimental data agree well with the simulated results.
This may be attributed to the better gas-liquid distribution at high liquid velocity region,
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which is close to the model assumption made in CFD simulation.
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Acknowledgments
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This work is finally supported by the National Natural Science Foundation of China under Grants
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(Nos. 21476009 and U1462104), and State Key Laboratory of Heavy Oil Processing (SKLHOP201502).
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Nomenclature kL = liquid-phase mass transfer coefficient [m·s-1] u = velocity [m·s-1] k = turbulent kinetic energy [J·kg-1] a = specific geometrical area [m-1]
UL = superficial liquid velocity referred to the column cross section [m·s-1] UG = superficial gas velocity referred to the column cross section [m·s-1]
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UB = bubble velocity [m·s-1] S = source term [Pa·m-1] uL,domain = the average liquid velocity [m·s-1]
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= turbulent dissipation rate [J·kg-1·s-1]
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ρ = density [kg·m-3]
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σ = surface tension [N·m-1]
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δf = the thickness of liquid film [mm]
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LB = the length of bubble [mm]
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g = acceleration due to gravity [m·s-2]
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[12] W. Guo, W. Xiao, M. Luo. Comparison among monolithic and randomly packed reactors for the methanol-to-propylene process, Chemical Engineering Journal S207–208 (2012) 734–745. [13] A. Zamaniyan, A.A. Khodadadi, Y. Mortazavi, H. Manafi, Comparative model analysis of the
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performance of tube fitted bulk monolithic catalyst with conventional pellet shapes for natural gas reforming, Journal of Industrial & Engineering Chemistry 17 (2011) 767–776. [14] J.J. Heiszwolf, L.B. Engelvaart, M.G.V.D. Eijnden, M.T. Kreutzer, F. Kapteijn, J.A. Moulijn, Hydrodynamic aspects of the monolith loop reactor, Chemical Engineering Science 56 (2001) 805–812. [15] C.O. Vandu, J. Ellenberger, R. Krishna, Hydrodynamics and mass transfer in an upflow monolith loop
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reactor: influence of vibration excitement, Chemical Engineering Science 59 (2004) 4999–5008.
[16] C.N. Satterfield, F. Özel, Some Characteristics of Two-Phase Flow in Monolithic Catalyst Structures,
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Industrial and Engineering Chemistry Fundamentals 16 (1977) 61–67.
[17] M. Xu, H. Huang, X. Zhan, H. Liu, S. Ji, C. Li, Pressure drop and liquid hold-up in multiphase monolithic reactor with different distributors, Catalysis Today 147 (2009) S132–S137.
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[18] G. Li, X. Yang, G. Dai, CFD simulation of effects of the configuration of gas distributors on gas–liquid
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flow and mixing in a bubble column, Chemical Engineering Science 64 (2009) 5104–5116.
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[19] S. Roy, M. Al-Dahhan, Flow distribution characteristics of a gas–liquid monolith reactor, Catalysis Today 105 (2005) 396–400.
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[20] K. Mishima, T. Hibiki, Some characteristics of air-water two-phase flow in small diameter vertical
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tubes, International Journal of Multiphase Flow, 22 (1996) 703–712. [21] M. Xu, H. Liu, C.Y. Li, Y. Zhou, S.F. Ji, Connection Between Liquid Distribution and Gas-Liquid Mass Transfer in Monolithic Bed, Chinese Journal Chemical Engineering, 19 (2011) 738–746.
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[22] M.L. Jackson, C.C. Shen, Aeration and mixing in deep tank fermentation systems, AIChE Journal 24 (1978) 63–71.
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[23] T.B. Anderson, R. Jackson, A fluid mechanical description of fluidized beds, Industrial and Engineering Chemistry Fundamentals 6 (1967) 527–539.
[24] G. Jourdan, L. Houas, O. Igra, J.-L. Estivalezes, C. Devals, E.E. Meshkov, Drag Coefficient of a
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Sphere in a Non-Stationary Flow: New Results, Proceedings Mathematical Physical & Engineering Sciences 463 (2007) 3323–3345.
[25] G. Bercic, A. Pintar, The role of gas bubbles and liquid slug lengths on mass transport in the Taylor flow through capillaries, Chemical Engineering Science 52 (1997) 3709–3719.
21
2 1
3
3
water 6
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air
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5
4
Liquid
8
9
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7
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100 mm
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200 mm
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U
Gas
Fig.1. Schematic overview of the cold model experimental setup.
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1-monolithic bed; 2-distributor; 3-rotameter; 4-gas outlet; 5-pump; 6-liquid reservoir; 7-packed bed
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distributor; 8-porcelain ring distributor; 9-foam distributor.
22
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(b)
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(a)
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Fig. 2. Schematic overview of computational domain (a), and grid details (b) and (c) for
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different distributors.
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dh
1/2LL
LB
δf
LUC
y
(a)
(b)
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Lf
(c)
(d)
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x
details (b)-(d) for mass transfer investigation.
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Fig. 3. Schematic overview of computational domain of Taylor unit cell (a) and grid
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(b) Grid details for LUC = 7.55 mm, δf = 19.23 μm, dh = 1.1 mm, and UB = 0.40 m·s-1; (c) Close-up view of
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the grid near gas bubble; and (d) Close-up view of the grid near the surface and film.
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0.20 0.18
0.14 0.12 0.10 0.08 0.06 0.04 0.02 0.04
0.06
0.08
0.10
0.12
0.14
0.16
0.20
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Experimental kLa (s-1)
0.18
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0.02
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Simulated kLa (s-1)
0.16
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Fig. 4. Comparison of mass transfer coefficients (kLa) for methane dissolved in water
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M
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between simulated values obtained in this work and experimental data in reference [25].
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(A)
(b)
(c)
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(a)
(e)
U
(d)
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M
A
N
(B)
(b)
(c)
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(a)
(d)
(e)
A
Fig. 5. The influence of superficial gas velocity (A) and liquid velocity (B) on the gas-liquid phase distribution with packed bed distributor obtained by CFD simulation. (A) UL,s = 0.0592 m·s-1: (a) UG = 0.0508 m·s-1; (b) UG = 0.0801 m·s-1; (c) UG = 0.1094 m·s-1; (d) UG = 0.1386 m·s-1; (e) UG = 0.1679 m·s-1. (B) UG,s = 0.1386 m·s-1: (a) UL = 0.0392 m·s-1; (b) UL = 0.0592 m·s-1; (c) UL = 0.0705 m·s-1; (d) UL = 0.0862m·s-1; (e) UL = 0.1019 m·s-1. 26
(A)
(c)
(b)
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(a)
U
(d)
M
A
N
(B)
(e)
(c)
(b)
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(a)
(d)
(e)
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Fig. 6. The influence of superficial gas velocity (A) and liquid velocity (B) on the gas-liquid phase distribution with foam distributor obtained by CFD simulation. (A) UL,s = 0.0592 m·s-1: (a) UG = 0.0508 m·s-1; (b) UG = 0.0801 m·s-1; (c) UG = 0.1094 m·s-1; (d) UG = 0.1386 m·s-1; (e) UG = 0.1679 m·s-1. (B) UG,s = 0.1386 m·s-1: (a) UL = 0.0392 m·s-1; (b) UL = 0.0592 m·s-1; (c) UL = 0.0705 m·s-1; (d) UL = 0.0862m·s-1; (e) UL = 0.1019 m·s-1. 27
(A)
(a)
(b)
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(c)
(e)
U
(d)
M
A
N
(B)
(a)
(c)
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(b)
(d)
(e)
A
Fig. 7. The influence of superficial gas velocity (A) and liquid velocity (B) on the gas-liquid phase distribution with porcelain ring distributor obtained by CFD model. (A) UL,s = 0.0592 m·s-1: (a) UG = 0.0508 m·s-1; (b) UG = 0.0801 m·s-1; (c) UG = 0.1094 m·s-1; (d) UG = 0.1386 m·s-1; (e) UG = 0.1679 m·s-1. (B) UG,s = 0.1386 m·s-1: (a) UL = 0.0392 m·s-1; (b) UL = 0.0592 m·s-1; (c) UL = 0.0705 m·s-1; (d) UL = 0.0862m·s-1; (e) UL = 0.1019 m·s-1.
28
0.45
0.80
(a)
(b)
0.75
0.40 0.70 0.65
0.35 0.30
L
L
0.60 0.55 0.50 0.25 0.45
0.35 0.15 0.05
0.10
0.15
0.20
0.30
0.25
0.05
0.10
UG,S (m·s-1)
0.20
0.25
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UG,S (m·s-1)
0.60
(c)
0.55
U
0.50 0.45 0.40
N
L
0.15
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0.40
0.20
A
0.35
0.25 0.20
0.10
UG,S
0.15
0.20
0.25
(m·s-1)
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ED
0.05
M
0.30
Fig. 8. The influence of superficial gas and liquid velocities on the experimental
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measured liquid holdup in the 400 cpsi monolithic bed with three types of distributors. (a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor. ▲, UL,s = 0.1019 m·s-1; ▼, UL,s = 0.0862 m·s-1; ★, UL,s = 0.0705 m·s-1; ●, UL,s = 0.0594 m·s-1; ■, UL,s =
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0.0392 m·s-1.
29
0.70
0.60
(a)
(b)
0.65
0.55
0.60
0.50 0.55
L
L
0.45 0.40
0.50
0.35
0.40
0.30
0.35
0.10
0.15
UG,S
0.20
0.25
0.05
0.10
(m·s-1)
-1
UG,S (m·s )
0.70
0.20
0.25
(c)
0.65
U
0.60
N
0.55
L
0.15
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0.05
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0.45
0.50
A
0.45
M
0.40 0.35 0.30
0.10
0.15
0.20
0.25
UG,S (m·s-1)
PT
ED
0.05
Fig. 9. The influence of superficial gas and liquid velocities on the experimental
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measured liquid holdup in the 600 cpsi monolithic bed with three types of distributors. (a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor. ▲, UL,s = 0.1019 m·s-1; ▼, UL,s = 0.0862 m·s-1; ★, UL,s = 0.0705 m·s-1; ●, UL,s = 0.0594 m·s-1; ■, UL,s =
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0.0392 m·s-1.
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16000
8000
(b)
14000
7000
12000
6000
∆P/L (Pa·m-1)
10000 8000 6000
5000 4000 3000 2000
4000 0.05
0.10
0.15
0.20
1000
0.25
0.05
0.10
UG,S (m·s-1)
0.25
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(c)
U
12000 10000
N
8000 6000
A
∆P/L (Pa·m-1)
0.20
UG,S (m·s-1)
16000 14000
0.15
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∆P/L (Pa·m-1)
(a)
2000 0.05
M
4000
0.10
0.15
0.20
0.25
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UG,S (m·s-1)
PT
Fig. 10. The influence of superficial gas and liquid velocities on the experimental measured pressure drop in the 400 cpsi monolithic bed with three types of distributors.
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(a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor.
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▲, 0.1019 m·s-1; ▼, 0.0862 m·s-1; ★, 0.0705 m·s-1; ●, 0.0594 m·s-1; ■, 0.0392 m·s-1.
31
20000
9000
(a)
(b)
18000
8000 7000
∆P/L(Pa·m-1)
∆P/L (Pa·m-1)
16000 14000 12000 10000
6000 5000 4000
8000
2000
4000 0.05
0.10
0.15
0.20
0.25
0.05
0.10
16000
(c) 14000
0.20
0.25
U
12000
N
10000 8000 6000
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∆P/L(Pa·m-1)
0.15
UG,S (m·s-1)
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UG,S (m·s-1)
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3000 6000
2000 0.05
M
4000
0.10
0.15
0.20
0.25
ED
UG,S (m·s-1)
PT
Fig. 11. The influence of superficial gas and liquid velocities on the experimental measured pressure drop in the 600 cpsi monolithic bed with three types of distributors.
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(a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor.
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▲, 0.1019 m·s-1; ▼, 0.0862 m·s-1; ★, 0.0705 m·s-1; ●, 0.0594 m·s-1; ■, 0.0392 m·s-1.
32
0.55
0.55
0.45
0.50
0.40
0.45
kLa (s-1)
kLa (s-1)
0.50
0.60
(a)
0.35 0.30
(b)
0.40 0.35 0.30
0.25
0.25 0.20 0.15
0.15 0.05
0.10
0.15
0.20
0.25
0.05
UG,S (m·s-1)
0.10
0.15
0.20
0.25
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UG,S (m·s-1)
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0.20
Fig. 12. The influence of superficial gas and liquid velocities on the mass transfer
U
coefficient in the 400 cpsi (a) and 600 cpsi (b) monolithic beds with packed bed
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distributor and foam distributor.
, 0.0705 m·s-1; ● and , 0.0594 m·s-1; ■ and □, 0.0392 m·s-1.
A
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PT
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M
0.0862 m·s-1; ★ and
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Solid scatters: packed bed distributor; Empty scatters: foam distributor. ▲ and ∆, 0.1019 m·s-1; ▼ and ▽,
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0.6
kLa (s-1)
0.5
0.4
0.3
0.08
0.10
0.12
0.14
0.16
0.18
0.20
-1
SC R
UG,S (m·s )
IP T
0.2
Fig. 13. The experimental and simulated mass transfer coefficients in the 600 cpsi
U
monolithic bed with packed bed distributor.
A
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PT
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M
A
m·s-1; ▲, UL,S = 0.0862 m·s-1; ▼, UL,S = 0.1019 m·s-1.
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Solid lines, CFD simulation; Scattered points, experimental data. ■, UL,S = 0.0594 m·s-1; ●, UL,S = 0.0705
34
S0255-2701(17)30896-6 https://doi.org/10.1016/j.cep.2018.01.026 CEP 7181
To appear in:
Chemical Engineering and Processing
Received date: Revised date: Accepted date:
6-9-2017 27-1-2018 29-1-2018
Please cite this article as: Yanyan Guo, Chengna Dai, Zhigang Lei, Hydrodynamics and mass transfer in multiphase monolithic reactors with different distributors: An experimental and modeling study, Chemical Engineering and Processing https://doi.org/10.1016/j.cep.2018.01.026 This is a PDF file of an unedited manuscript that has been accepted for publication. As a service to our customers we are providing this early version of the manuscript. The manuscript will undergo copyediting, typesetting, and review of the resulting proof before it is published in its final form. Please note that during the production process errors may be discovered which could affect the content, and all legal disclaimers that apply to the journal pertain.
Hydrodynamics and mass transfer in multiphase monolithic reactors with different distributors: An experimental and modeling study
Yanyan Guo, Chengna Dai, Zhigang Lei* State Key Laboratory of Chemical Resource Engineering, Beijing Key Laboratory of Energy
Corresponding author. E-mail address: [email protected] (Z. Lei).
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*
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Environmental Catalysis, Beijing University of Chemical Technology, Box 266, Beijing, 100029, China
Graphical abstract
The hydrodynamic and mass transfer performances in monolith bed with different
A
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A
N
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distributors were studied by the combination of experiments and CFD simulation.
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Highlights o The gas-liquid distribution of monolith bed with different distributors was studied; The hydrodynamic and mass transfer experiments and CFD simulation were carried
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out;
Packed bed distributor was superior to the foam and porcelain ring distributors.
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Abstract This work focuses on the hydrodynamic and mass transfer performances in
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multiphase monolithic reactor with different distributors by the combination of experiments
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and CFD simulation. The experimental results showed that the pressure drop in the monolithic bed with packed bed distributor is lower than that with other distributors, and that
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the liquid holdup follows the trend of packed bed distributor > foam distributor > porcelain
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ring distributor. Moreover, the gas-liquid mass transfer coefficients increase with the increase of superficial velocities of both gas and liquid phases, and mainly depend on the superficial
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liquid velocity. The CFD simulation showed that the gas-liquid phase distribution is mainly dependent on the superficial velocity of liquid phase, and the changes of superficial gas
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velocity don’t have an obvious effect on the gas-liquid phase distribution. Furthermore, the experimental and simulated results showed that the mass transfer coefficients in the monolithic bed with packed bed distributor is much higher than those with foam distributor and porcelain ring distributor.
2
Keywords: liquid holdup; pressure drop; mass transfer; gas-liquid phase distribution; CFD
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simulation; distributor
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1. Introduction In recent years, the multiphase monolithic reactors have received wide attention and are regarded as an intensification replacement of conventional multiphase reactors (e.g., trickle beds and slurry bubble columns) in the catalysis community, due to the low pressure drop and
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high mass transfer coefficients [1-10]. Gao et al. [11] prepared a novel cordierite monolithic catalyst for CO coupling to oxalate, and the result showed that the Pd efficiency of the
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monolithic catalyst (274 h-1) is much higher than that of a traditional pellet catalyst in packed bed reactor (46 h-1), which is attributed to the high Pd dispersion on the surface of the
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monolithic catalysts. Guo et al. [12] demonstrated that the monolithic reactor has higher
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propylene selectivity and methanol conversion than the random packed bed for the
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methanol-to-propylene process. Zamaniyan et al. [13] reported that tube fitted bulk
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monolithic reactor has superior advantages over the conventional packed reactor, especially
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with the low pressure drop and high effective usage of the catalyst in the natural gas reforming process. However, the channels of monolithic reactors are independent, and there
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is no redistribution for gas and liquid in the monolithic bed. Thus, the gas-liquid distributor at
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the entrance of the reactor plays an important and crucial role in the multiphase monolithic reactor performance, and a uniform gas-liquid phase distribution among monolithic channels is essential for reactor scale-up.
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The main function of the distributor is to disperse the gas and liquid phases. It has a
crucial effect on the shape of liquid drop and the lengths of bubble and liquid slug in the channels [14,15]. Many efforts have been done on the effect of distributor design on the gas-liquid phase distribution, gas and liquid holdups, and gas-liquid mass transfer by
4
experiments and CFD simulation. Satterfield and Ozel investigated the relation between different distributors and pressure drop in monolithic beds, and found that the design of suitable distributors plays a vital role on the bed pressure drop [16]. Xu et al. [17] investigated the pressure drop and liquid holdup in multiphase monolithic reactor with
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different distributors, with the result that the packed bed distributor has better gas-liquid distribution capability and the design of distributor has a crucial influence on the liquid
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holdup, pressure drop, and mass transfer. Li et al. [18] investigated the effect of gas
distributors on the gas holdup and mixing characteristics by CFD simulation, and found that
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the gas distributor is conductive to improving the gas holdup. Roy and Al-Dahhan [19]
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investigated the liquid phase distribution in a monolithic bed with three different distributors
A
(i.e., nozzle, foam, and showerhead) under Taylor flow regime, nozzle being the most
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suitable distributor for the monolithic bed. However, the studies on the effect of distributor
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types (i.e., porcelain ring distributor, foam distributor, and packed bed distributor) on the hydrodynamic (i.e., gas-liquid phase distribution, pressure drop, and liquid holdup) and mass
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transfer performances are still very limited.
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Thus, this work is on addressing on the dynamic intensification of monolithic bed with different distributors from the following three aspects: (i) determining the optimal type of distributor by measuring the liquid holdup, pressure drop, and mass transfer in the monolithic
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bed with different distributors and different channel sizes; (ii) establishing the CFD model to simulate the gas-liquid phase distribution at the entrance of monolithic bed with different distributors and to compared with the experimental results; and (iii) using the CFD model based on a Taylor unit to obtain the mass transfer coefficients in the monolithic bed with
5
different distributors, and to compare between the experimental and simulated mass transfer coefficients. 2. Experimental section 2.1. Experimental setup
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The cold model experimental setup mainly consisted of a monolithic reactor (D = 100 mm, and H = 600 mm) and the measurement system for liquid holdup, pressure drop and
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gas-liquid mass transfer. The cold model setup was carried out in the continuous co-current
down-flow mode, with air and water introduced at the top through the distributor. In this
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work, three distributors (i.e., porcelain ring distributor, foam distributor, and packed bed
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distributor) were tested comparatively on the liquid holdup, pressure drop and gas-liquid
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mass transfer. The packed bed distributor (200 mm in height) consisted of glass beads with
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the diameter of 1 mm, and was positioned straightly on the top of monolithic bed, with a
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close wire netting to prevent glass beads entering into the channels of monolithic bed. For porcelain ring distributor the diameter, average length, and height of the porcelain ring are 2
PT
mm, 3 mm, and 200 mm, respectively. The foam distributor consisted of two pieces of foam
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packing (50 mm in height and 100 mm in diameter). Both porcelain ring distributor and foam distributor were positioned straightly on the top of monolithic bed. To make sure that liquid phase spreads uniformly over the packed bed distributor, a
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plate (100 mm in diameter) containing 210 holes (2 mm in diameter) was put on the top of packed bed distributor. Below the distributors, four blocks of cordierite monoliths ( = 95 mm; H = 150 mm; and diameters of channel = 1.1 mm or 1.3 mm) were stacked as the column bed, and sealed inside the column by Teflon tapes to prevent by-pass flows of gas or
6
liquid near the column wall. 2.2. Experimental procedure Hydrodynamics and mass transfer coefficients were measured with the superficial gas and liquid velocities in the range of 0.0508–0.2344 m·s-1 and 0.0392–0.1019 m·s-1,
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respectively. According to the flow regime map for a capillary two-phase flow developed by Mishima and Hibiki [20], the flow investigated in this work is within the Taylor flow regime.
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Pressure drop between the top and the bottom of monolithic bed was measured by the differential pressure transducer (MDM4951 DP3E 0.5%, MICRO SENSOR Co. Ltd.,
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China), which is capable for pressure measurements in the range from 0 to 6 kPa.
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The liquid holdup measurement was performed following two steps. First, the liquid
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holdup in the reactor as a whole, including the contributions of distributor, separator and
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monolithic bed, was measured by collecting and weighting the liquid phase after the liquid
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and gas feedings were shut down simultaneously. Then, a blank experiment, where only one block of monolith was stacked within the reactor in place of four blocks of monoliths in the
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previous case, was carried out in the similar manner. Thus, a net weight value of liquid phase
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was obtained by the subtraction of the weight obtained in the blank experiment from that obtained in previous case. That is, the net weight is the part of liquid retaining in the blocks of monolith, excluding the contributions of distributor and separator. The liquid holdup was
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measured three times in each experiment, and the repetition was found to be within 5%. The mass transfer coefficients (kLa) were obtained following two steps similar as the liquid holdup measurement. First, the kLa coefficients in the reactor as a whole, including the contributions of distributor, separator and monolithic bed, were measured by stationary
7
oxygen absorption technique, in which the amount of oxygen dissolved in water was measured with Luminescent dissolve oxygen (LDO) probe (Hach, model sc100TM). Then, a blank experiment, where two blocks of monolith is in place of four blocks of monoliths, was carried out in the similar manner. The net kLa was obtained by substracting the kLa achieved
e
e '
(1)
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QL / A CL,O2 kL a = ln C Z L,O2
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in the blank experiment from that obtained in the previous case.
where QL (m3·s-1) is the volumetric liquid flow rate, A (m2) and Z (m) are the
L,O2 e
(mol·L-1) and
C L,O2
' e
(mol·L-1) are the oxygen concentrations at the exit of the
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C
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cross-section area of monolithic bed and the height of two blocks of monolith, respectively.
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column packed with four blocks and two blocks of monoliths, respectively. The details on the
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derivation of Eq. (1) can be obtained in previous publication [21].
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The obtained kLa(T) values are converted to 20 C by the following equation [22]: kL a (20) = kL a θ 20T
(2)
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where θ ( θ = 1.02) and T (C) are the temperature coefficient and experimental
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temperature, respectively. All the kLa as mentioned below are the corrected ones kLa(20). 3. CFD simulation
3.1. Modeling of hydrodynamic performance
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3.1.1. Model description The computational domain is shown in Fig. 2(a). The distributor was positioned straightly on the top of monolithic bed. The gas phase was introduced from the side of the column, and the liquid inlet was positioned at the top. The hydraulic diameters of gas and
8
liquid inlets are 25 mm and 27 mm, respectively. 3.1.2. Governing equations The Eulerian-Eulerian approach was selected to describe the gas-liquid two-phase flow using the Fluent software (version 14.5), and a steady state model with standard k–ε model to
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describe the turbulence viscosity was used. The governing equations are given as follows:
(hi i ui ) 0 Mass balance for species:
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iui hi w i Di w
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Continuity equation:
(3)
(4)
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N
where ρi (kg·m-3), hi and ui (m·s-1) are density, volume fraction and mean velocity of phase i
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(liquid or gas), respectively; w is the mass fraction; and Di (m·s-2) is the diffusion coefficient in phase i.
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hL hG 1
(5)
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where hL and hG are the gas and liquid volume fractions as calculated by the Anderson and Jackson method [23].
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Momentum balance equation:
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2 T hi uu hi eff uI hiP FLG hi ( FLS g ) 3
eff t
(6) (7)
where μ, μt and μeff (kg·m-1·s-1) are the corresponding molecular, turbulent and effective viscosities, respectively; and μt was obtained from the k–ε model equations; I is the unit tensor; FLG (N) is the interface drag force between gas and liquid phases, which was
9
calculated by the Schiller–Naumann method [24]; FLS (N) is the flow resistance created by solid phase; and P (Pa) is system pressure. The homogeneous and isotropic porous media model was used for distributors (i.e., porcelain ring distributor, foam distributor, and packed bed distributor), and a source term (Si)
Si
a
ui C2
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is added at the right side in the momentum balance equation:
1 u ui 2
(8)
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where a (m2) and C2 (m-1) are the permeability and inertial resistance coefficient of porous
d P2 3 a 150 (1 )2
N
3.5 (1 ) dP 2
(9)
(10)
A
C2
U
media, respectively.
M
where dp = 1.0 mm, 4.0 mm and 2.0 mm are the mean diameters of glass beads particles,
ED
porcelain ring particles and foam particles in the distributors, respectively; and δ = 0.375, 0.78 and 0.85 are the void fractions of glass beads particles, porcelain ring particles and foam
PT
particles, respectively.
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3.1.3. k–ε model
In the k–ε model, the turbulence viscosity can be calculated by
t Cμ
k
(11)
2
A
The turbulent kinetic energy k equation is uk t k
2 T k eff uI 3
The turbulent dissipation rate ε equation is
10
(12)
u t ε
2 2 T ε C u u uI C 1 eff 2 3 k k
(13)
where the default values were used ( Cμ 0.09 , k 1.0 , ε 1.3 , C1 1.44 , and
C2 1.92 ). I is the turbulent intensity and can be calculated by 1 u' 0.16Re 8 uavg
(14)
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I
where u ' (m·s-1) and uavg (m·s-1) are the root-mean-square of the velocity and the mean
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flow velocity, respectively; Re is Reynolds number. 3.1.4. Boundary conditions
U
The liquid and gas in the inlet were set to be uG = uG,inlet and uL = uL,inlet in the direction
N
normal to boundary. The liquid and gas at the outlet were set to be pressure-out boundary
A
condition. The walls of monolithic channel were set to be no-slip boundary condition. The
M
interior of distributor in the simulation was set as porous media.
ED
The geometrical structure was dealt with the structured grid generation technique, as shown in Figs. 2(b) and (c). The grid independency was validated to ensure the numerical
PT
accuracy with different grid numbers, and the final refined grids number is up to 430,000. In
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the simulation the first-order upwind differencing scheme was chosen for the solution of all differential equations, and the SIMPLE (Semi-Implicit Method for Pressure Linked Equations) algorithm was selected for pressure-velocity coupling. The residual of 10-3 was set
A
for the convergence criterion. 3.2. Modeling of mass transfer in Taylor flow The Taylor unit model consists of a gas bubble and two halves of liquid slugs on both ends of the bubble. The Taylor gas bubble is considered to be a “void” space, which acts as a
11
free surface with the surrounding liquid, as shown in Fig. 3(a). Mass transfer was simulated at 25 C and 101 kPa, and the diameter of monolithic channel was 1.1 mm. The calculation procedure of liquid phase mass transfer coefficients is divided into two steps. The velocity field of liquid flow was first calculated until it was stable. Then, the mass transfer was
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calculated, and the concentration of air at the surface of gas bubble is set to be 1. The volume-averaged mass and momentum conservation equations are given as follows.
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Continuity equation:
(u L ) 0
N
u L ( L u L u L L (u L (u )T ) P L g t
(16)
A
L
U
Momentum balance equation:
(15)
M
Mass transfer convection diffusion equation:
(17)
ED
( LC ) ( L vLC DC ) 0 t
where D (= 0.260 cm2·s-1) is the diffusion coefficient of air in water at 25 C and 101 kPa; C
PT
is the concentration of the air; and ρL (= 1000 kg·m-3) and vL are the density and volume of
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water, respectively.
The periodic boundary conditions for liquid and gas at the inlet and outlet were set (ptop
= pbottom, uL,top uL,bottom , and Ctop = Cbottom). In order to correct for the periodic boundary
A
condition, the source term of hydrostatic pressure drop ρLg is added to the momentum balance equation (16) in the y direction. Simulations were carried out in a reference frame, where the system moves upward with the bubble velocity uB and the gas bubble remains stationary. No-slip boundary condition was applied for the wall of monolithic channels, the
12
velocity of which was set to be uwall = –uB. Free-slip was set for the gas bubble surface: du⊥ /dn = 0 (u⊥ is the velocity component in the direction of bubble surface, and n is the normal direction to bubble surface). The geometrical structure was meshed with structured grids. The grid size less than 1
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μm was used close to the surface, with an exponential increase away from the surface as shown in Figs. 3(b) and (c). The grid independency was validated to ensure the numerical
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accuracy with different grid numbers and sizes, and the final refined grids number was 20,000. The third-order MUSCL (Monotonic Upstream-Centered Scheme for Conservation
U
Laws) differencing scheme was used to deal with the momentum and mass transfer
N
convection diffusion equations. The SIMPLE (Semi-Implicit Method for Pressure Linked
A
Equations) algorithm was used to deal with the pressure-velocity coupling. The residuals less
M
than 10-5 were chosen for the convergence criterion, and 1×10-5 s was taken as the time step.
ED
The average liquid velocity (uL,domain) in computational domain can be calculated by
vu
top i i
v
(18)
top i
PT
uL ,domain
where vi (m3) and ui (m·s-1) are the volume and the vertical velocity of cell i, respectively.
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The superficial liquid velocity (UL) can be calculated by U L (uL,domain uwall )(1 G )
(19)
A
where εG is gas holdup and calculated by
G
bubble volume LUC d h2 / 4
where LUC (m) and dh (m) are unit cell length and hydraulic diameter of the unit. The superficial gas velocity (UG) is written as
13
(20)
U G GU B
(21)
In this work, the mass transfer coefficient in liquid phase (kLa) is calculated by kLa
(CL,system,t t CL,system,t ) / t
(22)
CL,s CL,system,t
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where CL,s (= 1) is the concentration of air in water phase at bubble surface. CL,system,t and CL,system,t+Δt are the average liquid concentrations in computational domain at two consecutive
CL,system,t
top viCL ,i top vi
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time steps (time t and t+Δt), and can be derived from
(23)
U
To validate the Taylor unit model established in this work, the mass transfer coefficients
N
(kLa) for methane dissolved in water were first simulated. As shown in Fig. 4, the simulated
ED
4.1. Gas-liquid phase distribution
M
4. Results and discussion
A
values agree well with literature data [25], confirming the reliability of our Taylor unit model.
Figs. 5-7 show the effect of superficial velocities of gas and liquid phases on the
PT
gas-liquid phase distribution in the monolithic bed with packed bed distributor, foam
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distributor, and porcelain ring distributor, respectively. It can be seen that the gas-liquid phase distribution is mainly dependent on the superficial liquid velocity, and the increase of superficial gas velocity doesn’t have an obvious effect on the distribution. That is, the
A
gas-liquid phase distribution isn’t affected appreciably by the superficial gas velocity. Compared to foam distributor and porcelain ring distributor, the gas-liquid phase distribution in the monolithic bed with packed bed distributor is more uniform, which is attributed to the uniformity of glass beads in packed bed distributor.
14
4.2. Liquid holdup Figs. 8 and 9 show the effect of superficial gas velocity and superficial liquid velocity on the liquid holdup with different distributors and channel densities as obtained from experiments. With the increase of superficial gas velocity, the liquid holdup in monolithic bed
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decreases; while with the increase of superficial liquid velocity, the liquid holdup increases. Evidently, the liquid holdup is affected by the superficial liquid velocity, indicating that the
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gas-liquid phase distribution is mainly influenced by the superficial liquid velocity. On the other hand, under the same superficial gas and liquid velocities, the liquid holdup in the 600
U
cpsi monolithic bed is higher than that in the 400 cpsi monolithic bed for the three
N
distributors, which is attributed to the larger surface area leading to better uniform dispersion
A
of gas and liquid phases in the 600 cpsi monolithic bed. Among the three distributors
M
investigated in this work, the porcelain ring distributor has the lowest liquid holdup, while the
ED
packed bed distributor has the highest liquid holdup for both 400 cpsi and 600 cpsi monolithic beds under the same operating conditions, due to the worst distribution
PT
characteristics for porcelain ring distributor and the uniformity of glass beads in packed bed
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distributor. The liquid holdup follows the trend of packed bed distributor > foam distributor > porcelain ring distributor. 4.3. Pressure drop
A
The pressure drop reflects the energy loss, which is caused by the fluid flowing through
the monolithic bed. It directly relates to the momentum transfer, and is crucial to determine the energy dissipation. Figs. 10 and 11 show the influence of superficial gas and liquid velocities on the experimental measured total pressure drop with different distributors. The
15
results show that with the increase of the superficial gas and liquid velocities, pressure drop increases. In addition, pressure drop follows the trend of porcelain ring distributor > foam distributor > packed bed distributor. The uniformity of packed bed distributor leads to a decrease in pressure drop. Thus, it is advisable to select the packed bed distributor from the
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aspect of pressure drop. 4.4. Liquid phase mass transfer
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The porcelain ring distributor exhibits the largest pressure drop with the highest energy loss among the three types of distributors investigated. Therefore, the mass transfer
U
performance of other two types of distributors (i.e., packed bed distributor and foam
N
distributor) was further studied here. The experimental mass transfer coefficients in the 400
A
cpsi and 600 cpsi monolithic beds with packed bed distributor and foam distributor under
M
different superficial gas and liquid velocities are illustrated in Fig. 12. It can be seen that the
ED
mass transfer coefficients in the 600 cpsi monolithic bed are higher than those in the 400 cpsi monolithic bed for both foam distributor and packed bed distributor, which is attributed to
PT
larger surface area in the 600 cpsi monolithic bed leading to better gas-liquid distribution.
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When the liquid velocity is low, mass transfer coefficients increase slowly with the increase of superficial gas velocity. However, when the superficial liquid velocity is high, mass transfer coefficients increase quickly with the increase of superficial gas velocity. It seems
A
that mass transfer is mainly affected by the superficial liquid velocity, which should play a vital role on the gas-liquid phase distribution. Significant difference of mass transfer coefficients between packed bed distributor and foam distributor at high liquid velocity region arises. This indicates that distribution performance is particularly important at high
16
liquid velocity region. Compared to foam distributor, the mass transfer with packed bed distributor is enhanced. Therefore, it is advisable to select the packed bed distributor when considering both pressure drop and mass transfer together. Fig. 13 shows the experimental mass transfer coefficients in the 600 cpsi monolithic bed
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with packed bed distributor under different superficial gas and liquid velocities, along with the simulated results by CFD model. It can be seen that both simulated and experimental
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results exhibit the same trend. That is, the mass transfer coefficients increases with the increase of superficial gas and liquid velocities. The simulated mass transfer coefficients are
U
higher than experimental data, but at high liquid velocity region the discrepancy becomes
N
much smaller, which can be attributed to the poor gas and liquid distribution at low liquid
A
velocity region.
M
5. Conclusion
ED
Three types of distributors, i.e., packed bed distributor, foam distributor, and porcelain ring distributor, were adopted in this work to investigate the hydrodynamic (i.e., gas-liquid
PT
phase distribution, pressure drop, and liquid holdup) and mass transfer performances in
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monolithic bed by the combination of experiments and CFD simulation. It was found that the gas-liquid phase distribution in the monolithic bed with packed bed distributor is the most uniform, and the superficial liquid velocity plays an important role on the gas-liquid phase
A
distribution. The liquid holdup follows the trend of packed bed distributor > foam distributor > porcelain ring distributor. The pressure drop follows the trend of porcelain ring distributor > foam distributor > packed bed distributor. Therefore, one advantage of using packed bed distributor is brought out in decreasing pressure drop, which means that energy
17
consumption can be saved. Another advantage is the higher mass transfer coefficients than those of other distributors. It is known that the distributor design is crucial to the reactor hydrodynamic and mass transfer performances. The packed bed distributor is considered as the most suitable when considering the holdup, pressure drop and mass transfer performances
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together. The good hydrodynamic and mass transfer performances of packed bed distributor can be attributed to the uniform gas-liquid phase distribution. Moreover, when the superficial
SC R
gas and liquid velocities are high, the experimental data agree well with the simulated results.
This may be attributed to the better gas-liquid distribution at high liquid velocity region,
U
which is close to the model assumption made in CFD simulation.
N
Acknowledgments
A
This work is finally supported by the National Natural Science Foundation of China under Grants
A
CC E
PT
ED
M
(Nos. 21476009 and U1462104), and State Key Laboratory of Heavy Oil Processing (SKLHOP201502).
18
Nomenclature kL = liquid-phase mass transfer coefficient [m·s-1] u = velocity [m·s-1] k = turbulent kinetic energy [J·kg-1] a = specific geometrical area [m-1]
UL = superficial liquid velocity referred to the column cross section [m·s-1] UG = superficial gas velocity referred to the column cross section [m·s-1]
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UB = bubble velocity [m·s-1] S = source term [Pa·m-1] uL,domain = the average liquid velocity [m·s-1]
U
= turbulent dissipation rate [J·kg-1·s-1]
N
ρ = density [kg·m-3]
A
σ = surface tension [N·m-1]
A
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PT
ED
δf = the thickness of liquid film [mm]
M
LB = the length of bubble [mm]
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g = acceleration due to gravity [m·s-2]
19
References [1] F. Kapteijn, J.J. Heiszwolf, T.A. Nijhuis, J.A. Moulijn, Monoliths in multiphase catalytic processes—aspects and prospects, Cattech 3 (1999) 24–41. [2] T.A. Nijhuis, M.T. Kreutzer, A.C.J. Romijn, F. Kapteijn, J.A. Moulijn, Monolithic catalysts as efficient three-phase reactors, Chemical Engineering Science 56 (2001) 823–829.
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[3] T. Bauer, R. Guettel, S. Roy, M. Schubert, M. Al-Dahhan, R. Lange, Modelling and simulation of the monolithic reactor for gas–liquid–solid reactions, Chemical Engineering Science 83 (2005) 811–819.
[4] R.P. Fishwick, R. Natividad, R. Kulkarni, P.A. McGuire, J. Wood, J.M. Winterbottom, E.H. Stitt,
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Selective hydrogenation reactions: a comparative study of monolith CDC, stirred tank and trickle bed reactors, Catalysis Today 128 (2007) 108–114.
[5] A. Cybulski, A. Stankiewicz, R.K. Edvinsson Albers, J.A. Moulijn, Monolithic reactors for fine
U
chemicals industries: a comparative analysis of a monolithic reactor and a mechanically agitated
N
slurry reactor, Chemical Engineering Science 54 (1999) 2351–2358.
A
[6] R.M. Heck, S. Gulati, R.J. Farrauto, The application of monoliths for gas phase catalytic reactions,
M
Chemical Engineering Journal 82 (2001) 149–156.
[7] C.G. Visconti, E. Tronconi, G. Groppi, L. Lietti, M. Iovane, S. Rossini, R. Zennaro, Monolithic
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catalysts with high thermal conductivity for the Fischer–Tropsch synthesis in tubular reactors, Chemical Engineering Science 171 (2011) 1294–1307.
PT
[8] P. Heidebrecht, M. Pfafferodt, K. Sundmacher, Multiscale modelling strategy for structured catalytic reactors, Chemical Engineering Science 66 (2011) 4389–4402.
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[9] V. Tomasic, F. Jovic, State-of-the-art in the monolithic catalysts/reactors, Applied Catalysis A: General 311 (2006) 112–121.
[10] J.E. Antia, R. Govind, Conversion of methanol to gasoline-range hydrocarbons in a ZSM-5 coated
A
monolithic reactor, Industrial & Engineering Chemistry Research 34 (1995) 140–147.
[11] X. Gao, Y. Zhao, S. Wang, Y. Yin, B. Wang, X. Ma, A Pd–Fe/α-Al2O3/cordierite monolithic catalyst for CO coupling to oxalate, Chemical Engineering Science 66 (2011) 3513–3522.
[12] W. Guo, W. Xiao, M. Luo. Comparison among monolithic and randomly packed reactors for the methanol-to-propylene process, Chemical Engineering Journal S207–208 (2012) 734–745. [13] A. Zamaniyan, A.A. Khodadadi, Y. Mortazavi, H. Manafi, Comparative model analysis of the
20
performance of tube fitted bulk monolithic catalyst with conventional pellet shapes for natural gas reforming, Journal of Industrial & Engineering Chemistry 17 (2011) 767–776. [14] J.J. Heiszwolf, L.B. Engelvaart, M.G.V.D. Eijnden, M.T. Kreutzer, F. Kapteijn, J.A. Moulijn, Hydrodynamic aspects of the monolith loop reactor, Chemical Engineering Science 56 (2001) 805–812. [15] C.O. Vandu, J. Ellenberger, R. Krishna, Hydrodynamics and mass transfer in an upflow monolith loop
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reactor: influence of vibration excitement, Chemical Engineering Science 59 (2004) 4999–5008.
[16] C.N. Satterfield, F. Özel, Some Characteristics of Two-Phase Flow in Monolithic Catalyst Structures,
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Industrial and Engineering Chemistry Fundamentals 16 (1977) 61–67.
[17] M. Xu, H. Huang, X. Zhan, H. Liu, S. Ji, C. Li, Pressure drop and liquid hold-up in multiphase monolithic reactor with different distributors, Catalysis Today 147 (2009) S132–S137.
U
[18] G. Li, X. Yang, G. Dai, CFD simulation of effects of the configuration of gas distributors on gas–liquid
N
flow and mixing in a bubble column, Chemical Engineering Science 64 (2009) 5104–5116.
A
[19] S. Roy, M. Al-Dahhan, Flow distribution characteristics of a gas–liquid monolith reactor, Catalysis Today 105 (2005) 396–400.
M
[20] K. Mishima, T. Hibiki, Some characteristics of air-water two-phase flow in small diameter vertical
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tubes, International Journal of Multiphase Flow, 22 (1996) 703–712. [21] M. Xu, H. Liu, C.Y. Li, Y. Zhou, S.F. Ji, Connection Between Liquid Distribution and Gas-Liquid Mass Transfer in Monolithic Bed, Chinese Journal Chemical Engineering, 19 (2011) 738–746.
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[22] M.L. Jackson, C.C. Shen, Aeration and mixing in deep tank fermentation systems, AIChE Journal 24 (1978) 63–71.
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[23] T.B. Anderson, R. Jackson, A fluid mechanical description of fluidized beds, Industrial and Engineering Chemistry Fundamentals 6 (1967) 527–539.
[24] G. Jourdan, L. Houas, O. Igra, J.-L. Estivalezes, C. Devals, E.E. Meshkov, Drag Coefficient of a
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Sphere in a Non-Stationary Flow: New Results, Proceedings Mathematical Physical & Engineering Sciences 463 (2007) 3323–3345.
[25] G. Bercic, A. Pintar, The role of gas bubbles and liquid slug lengths on mass transport in the Taylor flow through capillaries, Chemical Engineering Science 52 (1997) 3709–3719.
21
2 1
3
3
water 6
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air
IP T
5
4
Liquid
8
9
PT
7
ED
100 mm
M
200 mm
A
N
U
Gas
Fig.1. Schematic overview of the cold model experimental setup.
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1-monolithic bed; 2-distributor; 3-rotameter; 4-gas outlet; 5-pump; 6-liquid reservoir; 7-packed bed
A
distributor; 8-porcelain ring distributor; 9-foam distributor.
22
IP T SC R (c)
(b)
U
(a)
N
Fig. 2. Schematic overview of computational domain (a), and grid details (b) and (c) for
A
CC E
PT
ED
M
A
different distributors.
23
dh
1/2LL
LB
δf
LUC
y
(a)
(b)
IP T
Lf
(c)
(d)
U
SC R
x
details (b)-(d) for mass transfer investigation.
N
Fig. 3. Schematic overview of computational domain of Taylor unit cell (a) and grid
A
(b) Grid details for LUC = 7.55 mm, δf = 19.23 μm, dh = 1.1 mm, and UB = 0.40 m·s-1; (c) Close-up view of
A
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PT
ED
M
the grid near gas bubble; and (d) Close-up view of the grid near the surface and film.
24
0.20 0.18
0.14 0.12 0.10 0.08 0.06 0.04 0.02 0.04
0.06
0.08
0.10
0.12
0.14
0.16
0.20
U
Experimental kLa (s-1)
0.18
SC R
0.02
IP T
Simulated kLa (s-1)
0.16
N
Fig. 4. Comparison of mass transfer coefficients (kLa) for methane dissolved in water
A
CC E
PT
ED
M
A
between simulated values obtained in this work and experimental data in reference [25].
25
(A)
(b)
(c)
SC R
IP T
(a)
(e)
U
(d)
ED
M
A
N
(B)
(b)
(c)
CC E
PT
(a)
(d)
(e)
A
Fig. 5. The influence of superficial gas velocity (A) and liquid velocity (B) on the gas-liquid phase distribution with packed bed distributor obtained by CFD simulation. (A) UL,s = 0.0592 m·s-1: (a) UG = 0.0508 m·s-1; (b) UG = 0.0801 m·s-1; (c) UG = 0.1094 m·s-1; (d) UG = 0.1386 m·s-1; (e) UG = 0.1679 m·s-1. (B) UG,s = 0.1386 m·s-1: (a) UL = 0.0392 m·s-1; (b) UL = 0.0592 m·s-1; (c) UL = 0.0705 m·s-1; (d) UL = 0.0862m·s-1; (e) UL = 0.1019 m·s-1. 26
(A)
(c)
(b)
SC R
IP T
(a)
U
(d)
M
A
N
(B)
(e)
(c)
(b)
CC E
PT
ED
(a)
(d)
(e)
A
Fig. 6. The influence of superficial gas velocity (A) and liquid velocity (B) on the gas-liquid phase distribution with foam distributor obtained by CFD simulation. (A) UL,s = 0.0592 m·s-1: (a) UG = 0.0508 m·s-1; (b) UG = 0.0801 m·s-1; (c) UG = 0.1094 m·s-1; (d) UG = 0.1386 m·s-1; (e) UG = 0.1679 m·s-1. (B) UG,s = 0.1386 m·s-1: (a) UL = 0.0392 m·s-1; (b) UL = 0.0592 m·s-1; (c) UL = 0.0705 m·s-1; (d) UL = 0.0862m·s-1; (e) UL = 0.1019 m·s-1. 27
(A)
(a)
(b)
SC R
IP T
(c)
(e)
U
(d)
M
A
N
(B)
(a)
(c)
CC E
PT
ED
(b)
(d)
(e)
A
Fig. 7. The influence of superficial gas velocity (A) and liquid velocity (B) on the gas-liquid phase distribution with porcelain ring distributor obtained by CFD model. (A) UL,s = 0.0592 m·s-1: (a) UG = 0.0508 m·s-1; (b) UG = 0.0801 m·s-1; (c) UG = 0.1094 m·s-1; (d) UG = 0.1386 m·s-1; (e) UG = 0.1679 m·s-1. (B) UG,s = 0.1386 m·s-1: (a) UL = 0.0392 m·s-1; (b) UL = 0.0592 m·s-1; (c) UL = 0.0705 m·s-1; (d) UL = 0.0862m·s-1; (e) UL = 0.1019 m·s-1.
28
0.45
0.80
(a)
(b)
0.75
0.40 0.70 0.65
0.35 0.30
L
L
0.60 0.55 0.50 0.25 0.45
0.35 0.15 0.05
0.10
0.15
0.20
0.30
0.25
0.05
0.10
UG,S (m·s-1)
0.20
0.25
SC R
UG,S (m·s-1)
0.60
(c)
0.55
U
0.50 0.45 0.40
N
L
0.15
IP T
0.40
0.20
A
0.35
0.25 0.20
0.10
UG,S
0.15
0.20
0.25
(m·s-1)
PT
ED
0.05
M
0.30
Fig. 8. The influence of superficial gas and liquid velocities on the experimental
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measured liquid holdup in the 400 cpsi monolithic bed with three types of distributors. (a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor. ▲, UL,s = 0.1019 m·s-1; ▼, UL,s = 0.0862 m·s-1; ★, UL,s = 0.0705 m·s-1; ●, UL,s = 0.0594 m·s-1; ■, UL,s =
A
0.0392 m·s-1.
29
0.70
0.60
(a)
(b)
0.65
0.55
0.60
0.50 0.55
L
L
0.45 0.40
0.50
0.35
0.40
0.30
0.35
0.10
0.15
UG,S
0.20
0.25
0.05
0.10
(m·s-1)
-1
UG,S (m·s )
0.70
0.20
0.25
(c)
0.65
U
0.60
N
0.55
L
0.15
SC R
0.05
IP T
0.45
0.50
A
0.45
M
0.40 0.35 0.30
0.10
0.15
0.20
0.25
UG,S (m·s-1)
PT
ED
0.05
Fig. 9. The influence of superficial gas and liquid velocities on the experimental
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measured liquid holdup in the 600 cpsi monolithic bed with three types of distributors. (a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor. ▲, UL,s = 0.1019 m·s-1; ▼, UL,s = 0.0862 m·s-1; ★, UL,s = 0.0705 m·s-1; ●, UL,s = 0.0594 m·s-1; ■, UL,s =
A
0.0392 m·s-1.
30
16000
8000
(b)
14000
7000
12000
6000
∆P/L (Pa·m-1)
10000 8000 6000
5000 4000 3000 2000
4000 0.05
0.10
0.15
0.20
1000
0.25
0.05
0.10
UG,S (m·s-1)
0.25
SC R
(c)
U
12000 10000
N
8000 6000
A
∆P/L (Pa·m-1)
0.20
UG,S (m·s-1)
16000 14000
0.15
IP T
∆P/L (Pa·m-1)
(a)
2000 0.05
M
4000
0.10
0.15
0.20
0.25
ED
UG,S (m·s-1)
PT
Fig. 10. The influence of superficial gas and liquid velocities on the experimental measured pressure drop in the 400 cpsi monolithic bed with three types of distributors.
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(a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor.
A
▲, 0.1019 m·s-1; ▼, 0.0862 m·s-1; ★, 0.0705 m·s-1; ●, 0.0594 m·s-1; ■, 0.0392 m·s-1.
31
20000
9000
(a)
(b)
18000
8000 7000
∆P/L(Pa·m-1)
∆P/L (Pa·m-1)
16000 14000 12000 10000
6000 5000 4000
8000
2000
4000 0.05
0.10
0.15
0.20
0.25
0.05
0.10
16000
(c) 14000
0.20
0.25
U
12000
N
10000 8000 6000
A
∆P/L(Pa·m-1)
0.15
UG,S (m·s-1)
SC R
UG,S (m·s-1)
IP T
3000 6000
2000 0.05
M
4000
0.10
0.15
0.20
0.25
ED
UG,S (m·s-1)
PT
Fig. 11. The influence of superficial gas and liquid velocities on the experimental measured pressure drop in the 600 cpsi monolithic bed with three types of distributors.
CC E
(a) porcelain ring distributor; (b) packed bed distributor; and (c) foam distributor.
A
▲, 0.1019 m·s-1; ▼, 0.0862 m·s-1; ★, 0.0705 m·s-1; ●, 0.0594 m·s-1; ■, 0.0392 m·s-1.
32
0.55
0.55
0.45
0.50
0.40
0.45
kLa (s-1)
kLa (s-1)
0.50
0.60
(a)
0.35 0.30
(b)
0.40 0.35 0.30
0.25
0.25 0.20 0.15
0.15 0.05
0.10
0.15
0.20
0.25
0.05
UG,S (m·s-1)
0.10
0.15
0.20
0.25
SC R
UG,S (m·s-1)
IP T
0.20
Fig. 12. The influence of superficial gas and liquid velocities on the mass transfer
U
coefficient in the 400 cpsi (a) and 600 cpsi (b) monolithic beds with packed bed
N
distributor and foam distributor.
, 0.0705 m·s-1; ● and , 0.0594 m·s-1; ■ and □, 0.0392 m·s-1.
A
CC E
PT
ED
M
0.0862 m·s-1; ★ and
A
Solid scatters: packed bed distributor; Empty scatters: foam distributor. ▲ and ∆, 0.1019 m·s-1; ▼ and ▽,
33
0.6
kLa (s-1)
0.5
0.4
0.3
0.08
0.10
0.12
0.14
0.16
0.18
0.20
-1
SC R
UG,S (m·s )
IP T
0.2
Fig. 13. The experimental and simulated mass transfer coefficients in the 600 cpsi
U
monolithic bed with packed bed distributor.
A
CC E
PT
ED
M
A
m·s-1; ▲, UL,S = 0.0862 m·s-1; ▼, UL,S = 0.1019 m·s-1.
N
Solid lines, CFD simulation; Scattered points, experimental data. ■, UL,S = 0.0594 m·s-1; ●, UL,S = 0.0705
34