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Experimental study and process simulation of acetaldehyde separation from glyoxal reaction mixture using a gas stripping technique
chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
Contents lists available at ScienceDirect
Chemical Engineering Research and Design journal homepage: www.elsevier.com/locate/cherd
Experimental study and process simulation of acetaldehyde separation from glyoxal reaction mixture using a gas stripping technique Zhiyong Zhang a,∗ , Baoyun Xu b , Dishun Zhao c a
Tianjin Key Laboratory of Applied Catalysis Science and Technology and State Key Laboratory for Chemical Engineering (Tianjin University), School of Chemical Engineering, Tianjin University, Tianjin, 300072, China b Shanghai Research Institute of Chemical Industry, Shanghai 200062, China c College of Chemistry and Pharmaceutical Engineering, Hebei University of Science and Technology, Shijiazhuang 050018, China
a b s t r a c t Efficient separation of acetaldehyde from glyoxal (CAS number: 107-22-2) reaction mixture using a gas stripping (GS) technique is investigated experimentally. A non-equilibrium model for describing the GS process is proposed and solved with a Newton-homotopy method. The steady-state performances of the GS column are investigated by a non-equilibrium model. The effects of liquid–gas ratio, feeding temperature and theoretical plate number on the residue acetaldehyde concentration are examined in detail. The simulating results are closed to the experimental data. According to the simulation results, liquid–gas ratio 1.6, feeding temperature 313 K and theoretical Plate 12 are the optimum operation condition for getting an acetaldehyde concentration of below 0.5 wt%. © 2011 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. Keywords: Acetaldehyde; Glyoxal; Gas stripping; Non-equilibrium model
1.
Introduction
Acetaldehyde oxidation with nitric acid is still the major route for the production of high-quality glyoxal. However, the acetaldehyde conversion cannot be completed and the unconverted acetaldehyde should be recovered. Conventional distillation technique for separation suffers from several drawbacks. Since distillation is an energy intensive process, it is generally not cost effective to distill and recycle acetaldehyde at low concentrations, and furthermore, glyoxal is heat-sensitive and is darkened if subject to over-heating for a long time. Vacuum distillation is not appropriate for acetaldehyde recovery, because it is very difficult to condense acetaldehyde vapor on the top of the column. In the current commercial processes, the high conversion of acetaldehyde is achieved with the cost of acetaldehyde selectivity of converting to glyoxal. With such obstacles, an efficient separation technology is called for, which should help to improve
∗
the product quality and facilitate the oxidation selectivity (Hotanahalli and Chandalia, 1972; Saha et al., 2000). The GS technique can be simply defined as a unit operation in which liquid and gas are brought into contact with each other with the purpose of transferring volatile substances from liquid to gas. This process has been effectively used in water and wastewater treatment to strip dissolved gases such as hydrogen sulphide, carbon dioxide and ammonia. In other applications, it has successfully been used to strip and reduce the concentration of taste and odour producing compounds and trace volatile organics (Obaid-ur-Rehman and Beg, 1990). Pilot scale air stripping of ammonia from wastewater was conducted at Lake Tahoe, California at 1966. Packed towers, operating in a counter-current mode, have been used for the removal of volatile organic compounds in commercial scale and have been proved to be highly efficient and cost effective (Slechta and Culp, 1967; Ball et al., 1984; Ribeiro et al., 2004; Qureshi and Blaschek, 2001; Namana et al., 2008). The present work focuses on the recovery of acetaldehyde by the GS technique.
Corresponding author. E-mail address: [email protected] (Z. Zhang). Received 14 November 2010; Received in revised form 13 March 2011; Accepted 11 April 2011 0263-8762/$ – see front matter © 2011 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. doi:10.1016/j.cherd.2011.04.010
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chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
Nomenclature DV i,k e E h H ki,k K L M N PV i,j P Q R Sc T V x y
Fick’s binary diffusion coefficient in the gasphase (m2 /s) heat-transfer rate (J/s) residual function for heat balance heat transfer coefficient (J/m2 s K) enthalpy (J/mol) mass transfer coefficient of binary i–k pair (mol/m2 s) phase equilibrium constant liquid flow rate (mol/s) residual function for material balance mass-transfer rate (mol/s) gas pressure of component i on stage j (kPa) total pressure (kPa) residual function for phase equilibrium relation at the interface residual function for mass transfer rate Schmidt number temperature (K) gas flow rate (mol/s) liquid mole fraction of component i gas mole fraction of component i
Greek Letters ˛ gas–liquid effective phase interface (m2 /m3 ) i,j activity coefficient of components i on stage j Dirac delta function ıi,k Subscripts i component j theoretical stage number c number of components Superscripts I phase interface liquid phase L V gas phase
Components
Content (wt%)
Acetaldehyde Glyoxal Acetic Acid Water
8 10 8 74
Experimental
2.1.
Materials
The liquid materials used in the experiments were an aqueous glyoxal mixture oxidized by nitric acid. The composition of the liquid is given in Table 1. High purity nitrogen was used as the stripping gas.
2.2.
counter-flow packed towers have been extensively used for ammonia stripping, although counter-flow gives better efficiencies. In spray towers, water is sprayed into the tower from a series of nozzles normally located at the top of the chamber, while air is blown upwards through the tower. In this work, the experiments were carried out in a atmospheric packed tower with counter current nitrogen flowing from the bottom to the top. The GS column used in this study is shown in Fig. 1, which consists mainly of a stainless steel column of ID 200 mm. The ceramic pall ring packing of ID 25 mm was filled in the column, and the packing height was about 5.2 m. The sample connections and the temperature testers were evenly distributed along the packing bed.
3.
Table 1 – Composition of raw material.
2.
Fig. 1 – Equipment of gas stripping column.
Mathematical models
The non-equilibrium model eliminating the necessity of using plate efficiencies was used to predict the actual performance of the process. The configuration of a non-equilibrium stage is shown as in Fig. 2. This non-equilibrium stage may represent a section of packing in the column. In this model, the mass and energy conservation equations are split into two parts, one for each phase. The equations for each phase are connected by mass and energy balances around the interface and by the assumption that the interface is at the thermodynamic equilibrium. The process of simultaneous mass and energy transfer through the interface is modeled by means of rate equations and transfer coefficients. The mass transfer coefficients for the liquid phase and vapor phase are combined to calculate the overall mass transfer coefficient. Resistances to mass and energy transfer offered by fluid phases can be accounted for by using separate equations for each phase (Kooijman and Taylor, 1995). The following assumptions were adopted to develop a model for the steady-state operation of the GS column.
Equipment
Stripping towers may be classified as packed towers or spray towers. In packed towers, packing is used to provide a large surface area per volume of packing. Both cross-flow and
1 2 3 4
A steady state operation. Mechanical equilibrium, i.e. PV = PL = P, at the stage. Gas–liquid equilibrium is only assumed at the interface. Each phase was perfectly mixed in each stage.
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chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
In the model equations, the mass-transfer rate and energy-transfer rate between the gas and liquid phases are calculated using the explicit method of the generalized Maxwell-Stefan equations. The mass-transfer coefficients for multi-component systems can be calculated directly from the binary mass coefficient as follows: [Kv ] = [Bv ]
−1
(10)
and [KL ] = [BL ]
−1
[ L ]
(11)
where Bvi,i = Fig. 2 – Physical model of non-equilibrium stage. 5 The heat transfer coefficient was assumed to be a constant over the volume of the column. 6 A theoretical stage was determined by the HETP of the packing.
(1)
L Mi,j
(2)
=
L Lj−1 xi,j−1 − Lj xi,j − Ni,j
=0
Bvi,k = −yi
(3)
L ELj = Lj−1 Hj−1 − Lj HjL + eLj = 0
(4)
The mass fluxes Ni,j from the gas phase to the interface are equal to the mass transfer rates from the interface to the liquid. The fluxes through the interface are assumed to be continuous:
V I V kV i,k,j ˛j (yk,j − yk,j ) − yi,j
k=1 c−1
c
k=1
c
−
εV j
(TjV − TjI ) + expεV j
c k=1
L Nk,j Hk,j =0
(13)
xi
c xk
+
kLi,c
k=1
(14)
kLi,k
/ i k=
= −xi
1 kLi,k
L = ıi,k + xi
−
1
,i= / k = 1, 2, . . . , c − 1
kLi,c
∂Lni , ∂xk
(15)
i, k = 1, 2, . . . , c − 1
(16)
The mass transfer coefficient was taken from the correlations provided by Onda et al. (Kakusaburo et al., 1968).
W 0.7 G
kvi,k = a
˛t vm
v (Sci,k )
WL
1/3
2/3
(˛t dp )
−2
˛t Dvi,k P
(17)
RT V
0.4 L −0.5 (Sci,k ) (˛t dp )
gm
1/3
Nk,j = 0
(7)
Nk,j = 0
(8)
k=1
c
i= / k = 1, 2, . . . , c − 1
,
The heat transfer coefficient was calculated using Chilton–Colburn rule, i.e. jH = jD. At the gas–liquid interface, phase equilibrium is assumed. The following relation relates the composition of component in gas phase with the composition of component in liquid phase on stage j:
Energy fluxes between the gas–liquid interfaces:
EIj = hV j ˛j
(6)
k=1
L I L kLi,k,j ˛j (xk,j − xk,j ) − xi,j
1 1 − v kvi,k ki,c
kLi,k
Mass fluxes between the gas and liquid phases:
L Ri,j = Ni,j −
BLi,i =
(5)
EIi,j = eLj − eV j =0
c−1
(12)
kvi,k
and
BLi,k
V V V EV j = Vj+1 Hj+1 − Vj Hj − ej = 0
V = Ni,j − Ri,j
k=1
The enthalpy balances for both gas and liquid phases:
I L V Mi,j = Ni,j − Ni,j =0
c yk
/ i k=
The component material balances for the gas and liquid phase: V V Mi,j = Vi yi,j+1 − Vj yi,j + Ni,j =0
yi + kvi,c
= 0.0051
˛w Lm
L m
L m
I I I Qi,j = Ki,j xi,j − yIi,j = 0
(18)
(19)
K-values were computed from the following equation:
V Nk,j Hk,j − hLj ˛j (TjI − TjL )
Ki,j =
k=1
(9)
i,j PV i,j P
(20)
is the partial preswhere P is the system total pressure, PV i,j sure of component i on stage j and its value can be calculated according to Henry’s law. UNIFAC model have been used to
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chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
Fig. 4 – Liquid acetaldehyde concentration profile at different liquid–gas ratio.
Fig. 3 – Flow chart of the algorithm.
evaluate the activity coefficient for liquid phase i,j . The UNIFAC group contribution method was used for the description of the liquid phase non-ideality and the volume and area parameters of the groups of are quoted from Ma (2005). Other physical properties were evaluated using the methods suggested by Reid et al. (Poling et al., 2001). With setting the minimum of the total square sum of residuals of the above equations as the objective function, the model equations were solved using a Newton-homotopy method in MATLAB environment. The calculation process was shown in Fig. 3.
4.
Results and discussion
Experiments and simulation were conducted in order to obtain an optimum operation condition for acetaldehyde removal. Several parameters, e.g. liquid–gas ratio, feeding temperature, theoretical stage of the column, were studied in order to obtain the optimal operating condition for producing the liquid product with acetaldehyde concentration below 0.5 wt%.
4.1.
Effect of liquid–gas ratio
Liquid–gas ratio, which is the volume flow rate ratio of liquid to gas, played an important role in determining the product purity. To study the effect of the liquid–gas ratio in the process, the liquid–gas ratio was investigated at 1.6, 1.8, 2.25, 3.85. The results of the simulation and the experiment are represented by the curves and the data points, respectively in Figs. 4 and 5. It can be seen from Figs. 4 and 5 that the simulating results and the experimental data agreed very well and the liquid acetaldehyde concentration gradually increased with the increase of liquid–gas ratio. That is to say, decreasing liquid–gas ratio was favorable for acetaldehyde removal. The reason was that a high liquid–gas ratio results in higher capacity of the nitrogen to receive gas. The driving force for stripping the compound off was therefore better represented by the low L/G ratios. In practical production, high liquid–gas ratio led to a lower acetaldehyde removal, while low liquid–gas ratio led to more nitrogen consumption. So it was very important to select an appropriate liquid–gas ratio. In this paper, the liquid–gas ratio of 1.6 can ensure the acetaldehyde concentration meets the requirement.
Fig. 5 – Acetaldehyde concentration in product at different molar ratio of liquid to gas.
4.2.
Effect of liquid feeding temperature
To study the effect of liquid feeding temperature, experiments were conducted at different temperatures, 303 K, 313 K, 323 K, 333 K and 343 K. The liquid acetaldehyde concentration profiles in the column at different temperatures are shown in Fig. 6, and the liquid acetaldehyde concentrations in the product are shown in Fig. 7. In Figs. 6 and 7, the curves were composed of the simulation results, while the data points were the experimental values. It can be seen from Figs. 6 and 7 that the simulating results and the experimental data are in excellent agreement and the liquid acetaldehyde concentration in the column decreases gradually with the increase of feeding temperature. That is to say, increasing the feeding temperature is favorable for acetaldehyde removal. This is because acetaldehyde molecules move faster at higher temperature and are easier to escape from liquid. It can be seen from Fig. 7 that when increasing feeding temperature from 303 K to 343 K, the acetaldehyde concentration in the product decreases from 1.1% to 0%. However, high feeding temperature means more energy consumption, the feeding temperature is designated at 313 K.
4.3.
Effect of theoretical stage of the stripping column
Then, the effect of the theoretical stage of the stripping column on acetaldehyde concentration in the product was
chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
Fig. 6 – Liquid acetaldehyde concentration profile at different feeding temperature.
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Fig. 9 – Acetaldehyde concentration in product at different theoretical stage.
0.55 wt% to 0.25 wt%, while the theoretical stage increases from 8 to 14. So the theoretical stages of the column were selected as 12.
5.
Fig. 7 – Acetaldehyde concentration in product at different feeding temperature.
investigated with the model. The liquid acetaldehyde concentration profiles in the column at different theoretical stages and the liquid acetaldehyde concentrations in the product are shown in Figs. 8 and 9, respectively. In principle, more theoretical stages provide more gas–liquid contacting areas, which are very useful for the separation. It can be seen from Figs. 8 and 9 that the acetaldehyde concentration in the product decreases from
Fig. 8 – Liquid acetaldehyde concentration profile at different theoretical stage.
Conclusion
In this paper, the removal of acetaldehyde from glyoxal mixture using GS technique is investigated. The research results show that the GS technique can achieve high acetaldehyde removal successfully, which can avoid effectively the problems of the conventional separation methods. A non-equilibrium steady-state model is established and solved with Newton-homotopy continuation method. The effects of liquid–gas ratio, feeding temperature and theoretical stage are investigated systematically using the model. A good agreement between calculated curve and the experimental points is observed from the comparison chart of experimental and simulation results. In order to obtain the minimum acetaldehyde concentration of 0.5 wt%, the liquid–gas ratio of 1.6, feeding temperature of 313 K and theoretical stage of 12 are selected as the optimum operation condition.
References Ball, W.P., Jones, M.D., Kavanaugh, M.C., 1984. Mass transfer of volatile organic compounds in packed tower aeration. J. Water Pollut. Control Fed. 56, 127–135. Hotanahalli, S.S., Chandalia, S.B., 1972. Oxidation of acetaldehyde to glyoxal by nitric acid. J. Appl. Chem. Biotechnol. 22, 1243–1252. Kakusaburo, O., Hiroshi, T., Yoshio, O., 1968. Mass transfer coefficients between gas and liquid phases in packed columns. J. Chem. Eng. Jpn. 1, 56–66. Kooijman, H.A., Taylor, R., 1995. A nonequilibrium model for dynamic simulation of tray distillation columns. AIChE J. 41 (8), 1852–1863. Ma, P.S., 2005. Chemical Engineering Thermodynamics. Chemical industry press, Beijing. Namana, S.A., Tureb, I.E., Veziroglub, T.N., 2008. Industrial extraction pilot plant for stripping H2 S gas from Black Sea water. Int. J. Hydrogen Energy 33, 6577–6585. Obaid-ur-Rehman, S., Beg, S.A., 1990. Ammonia removal by air stripping—from origin to present state of technology. J. Environ. Sci. Health Part A 25, 343–365. Poling, B.E., Prausnitz, J.M., O’Connell, J.P., 2001. The Properties of Gases and Liquids. McGraw-Hill, New York.
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Qureshi, N., Blaschek, H.P., 2001. Recovery of butanol from fermentation broth by gas stripping. Renew. Energy 4, 557– 564. Ribeiro, C.P., Lage, P.L.C., Borges, C.P., 2004. A combined gas-stripping vapour permeation process for aroma recovery. J. Membr. Sci. 238, 9–19.
Saha, B., Chopade, S.P., Mahajani, S.M., 2000. Recovery of dilute acetic acid through esterification in a reactive distillation column. Catal. Today 60, 147–157. Slechta, A.F., Culp, G.L., 1967. Water reclamation studies at the South Tahoe public utility district. J. Water Pollut. Control Fed. 39, 787–814.
Contents lists available at ScienceDirect
Chemical Engineering Research and Design journal homepage: www.elsevier.com/locate/cherd
Experimental study and process simulation of acetaldehyde separation from glyoxal reaction mixture using a gas stripping technique Zhiyong Zhang a,∗ , Baoyun Xu b , Dishun Zhao c a
Tianjin Key Laboratory of Applied Catalysis Science and Technology and State Key Laboratory for Chemical Engineering (Tianjin University), School of Chemical Engineering, Tianjin University, Tianjin, 300072, China b Shanghai Research Institute of Chemical Industry, Shanghai 200062, China c College of Chemistry and Pharmaceutical Engineering, Hebei University of Science and Technology, Shijiazhuang 050018, China
a b s t r a c t Efficient separation of acetaldehyde from glyoxal (CAS number: 107-22-2) reaction mixture using a gas stripping (GS) technique is investigated experimentally. A non-equilibrium model for describing the GS process is proposed and solved with a Newton-homotopy method. The steady-state performances of the GS column are investigated by a non-equilibrium model. The effects of liquid–gas ratio, feeding temperature and theoretical plate number on the residue acetaldehyde concentration are examined in detail. The simulating results are closed to the experimental data. According to the simulation results, liquid–gas ratio 1.6, feeding temperature 313 K and theoretical Plate 12 are the optimum operation condition for getting an acetaldehyde concentration of below 0.5 wt%. © 2011 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. Keywords: Acetaldehyde; Glyoxal; Gas stripping; Non-equilibrium model
1.
Introduction
Acetaldehyde oxidation with nitric acid is still the major route for the production of high-quality glyoxal. However, the acetaldehyde conversion cannot be completed and the unconverted acetaldehyde should be recovered. Conventional distillation technique for separation suffers from several drawbacks. Since distillation is an energy intensive process, it is generally not cost effective to distill and recycle acetaldehyde at low concentrations, and furthermore, glyoxal is heat-sensitive and is darkened if subject to over-heating for a long time. Vacuum distillation is not appropriate for acetaldehyde recovery, because it is very difficult to condense acetaldehyde vapor on the top of the column. In the current commercial processes, the high conversion of acetaldehyde is achieved with the cost of acetaldehyde selectivity of converting to glyoxal. With such obstacles, an efficient separation technology is called for, which should help to improve
∗
the product quality and facilitate the oxidation selectivity (Hotanahalli and Chandalia, 1972; Saha et al., 2000). The GS technique can be simply defined as a unit operation in which liquid and gas are brought into contact with each other with the purpose of transferring volatile substances from liquid to gas. This process has been effectively used in water and wastewater treatment to strip dissolved gases such as hydrogen sulphide, carbon dioxide and ammonia. In other applications, it has successfully been used to strip and reduce the concentration of taste and odour producing compounds and trace volatile organics (Obaid-ur-Rehman and Beg, 1990). Pilot scale air stripping of ammonia from wastewater was conducted at Lake Tahoe, California at 1966. Packed towers, operating in a counter-current mode, have been used for the removal of volatile organic compounds in commercial scale and have been proved to be highly efficient and cost effective (Slechta and Culp, 1967; Ball et al., 1984; Ribeiro et al., 2004; Qureshi and Blaschek, 2001; Namana et al., 2008). The present work focuses on the recovery of acetaldehyde by the GS technique.
Corresponding author. E-mail address: [email protected] (Z. Zhang). Received 14 November 2010; Received in revised form 13 March 2011; Accepted 11 April 2011 0263-8762/$ – see front matter © 2011 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. doi:10.1016/j.cherd.2011.04.010
2786
chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
Nomenclature DV i,k e E h H ki,k K L M N PV i,j P Q R Sc T V x y
Fick’s binary diffusion coefficient in the gasphase (m2 /s) heat-transfer rate (J/s) residual function for heat balance heat transfer coefficient (J/m2 s K) enthalpy (J/mol) mass transfer coefficient of binary i–k pair (mol/m2 s) phase equilibrium constant liquid flow rate (mol/s) residual function for material balance mass-transfer rate (mol/s) gas pressure of component i on stage j (kPa) total pressure (kPa) residual function for phase equilibrium relation at the interface residual function for mass transfer rate Schmidt number temperature (K) gas flow rate (mol/s) liquid mole fraction of component i gas mole fraction of component i
Greek Letters ˛ gas–liquid effective phase interface (m2 /m3 ) i,j activity coefficient of components i on stage j Dirac delta function ıi,k Subscripts i component j theoretical stage number c number of components Superscripts I phase interface liquid phase L V gas phase
Components
Content (wt%)
Acetaldehyde Glyoxal Acetic Acid Water
8 10 8 74
Experimental
2.1.
Materials
The liquid materials used in the experiments were an aqueous glyoxal mixture oxidized by nitric acid. The composition of the liquid is given in Table 1. High purity nitrogen was used as the stripping gas.
2.2.
counter-flow packed towers have been extensively used for ammonia stripping, although counter-flow gives better efficiencies. In spray towers, water is sprayed into the tower from a series of nozzles normally located at the top of the chamber, while air is blown upwards through the tower. In this work, the experiments were carried out in a atmospheric packed tower with counter current nitrogen flowing from the bottom to the top. The GS column used in this study is shown in Fig. 1, which consists mainly of a stainless steel column of ID 200 mm. The ceramic pall ring packing of ID 25 mm was filled in the column, and the packing height was about 5.2 m. The sample connections and the temperature testers were evenly distributed along the packing bed.
3.
Table 1 – Composition of raw material.
2.
Fig. 1 – Equipment of gas stripping column.
Mathematical models
The non-equilibrium model eliminating the necessity of using plate efficiencies was used to predict the actual performance of the process. The configuration of a non-equilibrium stage is shown as in Fig. 2. This non-equilibrium stage may represent a section of packing in the column. In this model, the mass and energy conservation equations are split into two parts, one for each phase. The equations for each phase are connected by mass and energy balances around the interface and by the assumption that the interface is at the thermodynamic equilibrium. The process of simultaneous mass and energy transfer through the interface is modeled by means of rate equations and transfer coefficients. The mass transfer coefficients for the liquid phase and vapor phase are combined to calculate the overall mass transfer coefficient. Resistances to mass and energy transfer offered by fluid phases can be accounted for by using separate equations for each phase (Kooijman and Taylor, 1995). The following assumptions were adopted to develop a model for the steady-state operation of the GS column.
Equipment
Stripping towers may be classified as packed towers or spray towers. In packed towers, packing is used to provide a large surface area per volume of packing. Both cross-flow and
1 2 3 4
A steady state operation. Mechanical equilibrium, i.e. PV = PL = P, at the stage. Gas–liquid equilibrium is only assumed at the interface. Each phase was perfectly mixed in each stage.
2787
chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
In the model equations, the mass-transfer rate and energy-transfer rate between the gas and liquid phases are calculated using the explicit method of the generalized Maxwell-Stefan equations. The mass-transfer coefficients for multi-component systems can be calculated directly from the binary mass coefficient as follows: [Kv ] = [Bv ]
−1
(10)
and [KL ] = [BL ]
−1
[ L ]
(11)
where Bvi,i = Fig. 2 – Physical model of non-equilibrium stage. 5 The heat transfer coefficient was assumed to be a constant over the volume of the column. 6 A theoretical stage was determined by the HETP of the packing.
(1)
L Mi,j
(2)
=
L Lj−1 xi,j−1 − Lj xi,j − Ni,j
=0
Bvi,k = −yi
(3)
L ELj = Lj−1 Hj−1 − Lj HjL + eLj = 0
(4)
The mass fluxes Ni,j from the gas phase to the interface are equal to the mass transfer rates from the interface to the liquid. The fluxes through the interface are assumed to be continuous:
V I V kV i,k,j ˛j (yk,j − yk,j ) − yi,j
k=1 c−1
c
k=1
c
−
εV j
(TjV − TjI ) + expεV j
c k=1
L Nk,j Hk,j =0
(13)
xi
c xk
+
kLi,c
k=1
(14)
kLi,k
/ i k=
= −xi
1 kLi,k
L = ıi,k + xi
−
1
,i= / k = 1, 2, . . . , c − 1
kLi,c
∂Lni , ∂xk
(15)
i, k = 1, 2, . . . , c − 1
(16)
The mass transfer coefficient was taken from the correlations provided by Onda et al. (Kakusaburo et al., 1968).
W 0.7 G
kvi,k = a
˛t vm
v (Sci,k )
WL
1/3
2/3
(˛t dp )
−2
˛t Dvi,k P
(17)
RT V
0.4 L −0.5 (Sci,k ) (˛t dp )
gm
1/3
Nk,j = 0
(7)
Nk,j = 0
(8)
k=1
c
i= / k = 1, 2, . . . , c − 1
,
The heat transfer coefficient was calculated using Chilton–Colburn rule, i.e. jH = jD. At the gas–liquid interface, phase equilibrium is assumed. The following relation relates the composition of component in gas phase with the composition of component in liquid phase on stage j:
Energy fluxes between the gas–liquid interfaces:
EIj = hV j ˛j
(6)
k=1
L I L kLi,k,j ˛j (xk,j − xk,j ) − xi,j
1 1 − v kvi,k ki,c
kLi,k
Mass fluxes between the gas and liquid phases:
L Ri,j = Ni,j −
BLi,i =
(5)
EIi,j = eLj − eV j =0
c−1
(12)
kvi,k
and
BLi,k
V V V EV j = Vj+1 Hj+1 − Vj Hj − ej = 0
V = Ni,j − Ri,j
k=1
The enthalpy balances for both gas and liquid phases:
I L V Mi,j = Ni,j − Ni,j =0
c yk
/ i k=
The component material balances for the gas and liquid phase: V V Mi,j = Vi yi,j+1 − Vj yi,j + Ni,j =0
yi + kvi,c
= 0.0051
˛w Lm
L m
L m
I I I Qi,j = Ki,j xi,j − yIi,j = 0
(18)
(19)
K-values were computed from the following equation:
V Nk,j Hk,j − hLj ˛j (TjI − TjL )
Ki,j =
k=1
(9)
i,j PV i,j P
(20)
is the partial preswhere P is the system total pressure, PV i,j sure of component i on stage j and its value can be calculated according to Henry’s law. UNIFAC model have been used to
2788
chemical engineering research and design 8 9 ( 2 0 1 1 ) 2785–2790
Fig. 4 – Liquid acetaldehyde concentration profile at different liquid–gas ratio.
Fig. 3 – Flow chart of the algorithm.
evaluate the activity coefficient for liquid phase i,j . The UNIFAC group contribution method was used for the description of the liquid phase non-ideality and the volume and area parameters of the groups of are quoted from Ma (2005). Other physical properties were evaluated using the methods suggested by Reid et al. (Poling et al., 2001). With setting the minimum of the total square sum of residuals of the above equations as the objective function, the model equations were solved using a Newton-homotopy method in MATLAB environment. The calculation process was shown in Fig. 3.
4.
Results and discussion
Experiments and simulation were conducted in order to obtain an optimum operation condition for acetaldehyde removal. Several parameters, e.g. liquid–gas ratio, feeding temperature, theoretical stage of the column, were studied in order to obtain the optimal operating condition for producing the liquid product with acetaldehyde concentration below 0.5 wt%.
4.1.
Effect of liquid–gas ratio
Liquid–gas ratio, which is the volume flow rate ratio of liquid to gas, played an important role in determining the product purity. To study the effect of the liquid–gas ratio in the process, the liquid–gas ratio was investigated at 1.6, 1.8, 2.25, 3.85. The results of the simulation and the experiment are represented by the curves and the data points, respectively in Figs. 4 and 5. It can be seen from Figs. 4 and 5 that the simulating results and the experimental data agreed very well and the liquid acetaldehyde concentration gradually increased with the increase of liquid–gas ratio. That is to say, decreasing liquid–gas ratio was favorable for acetaldehyde removal. The reason was that a high liquid–gas ratio results in higher capacity of the nitrogen to receive gas. The driving force for stripping the compound off was therefore better represented by the low L/G ratios. In practical production, high liquid–gas ratio led to a lower acetaldehyde removal, while low liquid–gas ratio led to more nitrogen consumption. So it was very important to select an appropriate liquid–gas ratio. In this paper, the liquid–gas ratio of 1.6 can ensure the acetaldehyde concentration meets the requirement.
Fig. 5 – Acetaldehyde concentration in product at different molar ratio of liquid to gas.
4.2.
Effect of liquid feeding temperature
To study the effect of liquid feeding temperature, experiments were conducted at different temperatures, 303 K, 313 K, 323 K, 333 K and 343 K. The liquid acetaldehyde concentration profiles in the column at different temperatures are shown in Fig. 6, and the liquid acetaldehyde concentrations in the product are shown in Fig. 7. In Figs. 6 and 7, the curves were composed of the simulation results, while the data points were the experimental values. It can be seen from Figs. 6 and 7 that the simulating results and the experimental data are in excellent agreement and the liquid acetaldehyde concentration in the column decreases gradually with the increase of feeding temperature. That is to say, increasing the feeding temperature is favorable for acetaldehyde removal. This is because acetaldehyde molecules move faster at higher temperature and are easier to escape from liquid. It can be seen from Fig. 7 that when increasing feeding temperature from 303 K to 343 K, the acetaldehyde concentration in the product decreases from 1.1% to 0%. However, high feeding temperature means more energy consumption, the feeding temperature is designated at 313 K.
4.3.
Effect of theoretical stage of the stripping column
Then, the effect of the theoretical stage of the stripping column on acetaldehyde concentration in the product was
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Fig. 6 – Liquid acetaldehyde concentration profile at different feeding temperature.
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Fig. 9 – Acetaldehyde concentration in product at different theoretical stage.
0.55 wt% to 0.25 wt%, while the theoretical stage increases from 8 to 14. So the theoretical stages of the column were selected as 12.
5.
Fig. 7 – Acetaldehyde concentration in product at different feeding temperature.
investigated with the model. The liquid acetaldehyde concentration profiles in the column at different theoretical stages and the liquid acetaldehyde concentrations in the product are shown in Figs. 8 and 9, respectively. In principle, more theoretical stages provide more gas–liquid contacting areas, which are very useful for the separation. It can be seen from Figs. 8 and 9 that the acetaldehyde concentration in the product decreases from
Fig. 8 – Liquid acetaldehyde concentration profile at different theoretical stage.
Conclusion
In this paper, the removal of acetaldehyde from glyoxal mixture using GS technique is investigated. The research results show that the GS technique can achieve high acetaldehyde removal successfully, which can avoid effectively the problems of the conventional separation methods. A non-equilibrium steady-state model is established and solved with Newton-homotopy continuation method. The effects of liquid–gas ratio, feeding temperature and theoretical stage are investigated systematically using the model. A good agreement between calculated curve and the experimental points is observed from the comparison chart of experimental and simulation results. In order to obtain the minimum acetaldehyde concentration of 0.5 wt%, the liquid–gas ratio of 1.6, feeding temperature of 313 K and theoretical stage of 12 are selected as the optimum operation condition.
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