Ind. Eng. Chem. Res. 1999, 38, 3481-3488
3481
Performance of Absorption Columns Equipped with Low Pressure Drops Structured Packings Pier Massimiliano Launaro† and Alessandro Paglianti*,‡ Consorzio Polo Tecnologico Magona, I-57023 Cecina, Italy, and Department of Chemical Engineering, Industrial Chemistry and Materials Science, University of Pisa, I-56126 Pisa, Italy
During the past few years the industry has demanded structured packings with higher and higher efficiency and with lower and lower pressure drops. Using these internals significant improvements were obtained in the absorption process but also in direct-contact heat transfer. In the present work a new type of packing is investigated both experimentally and theoretically to evaluate gas and liquid mass transfer coefficients, surface area available for mass transfer, and pressure drops. The experimental performances of the packing are tested in absorption and desorption systems to identify relations that allowing reliable design. Introduction During the past few years the attention focused on environmental problems has led to the development of new packings for absorption columns allowing to the attainment of high abatement efficiency and low energy consumption. Using structured packing permits the reduction of operating costs and column dimensions (Brunazzi et al.1), but nowadays higher and higher performances are required. In particular, to reduce operating costs, the industry requires packings with a highly specific surface, and therefore high mass transfer efficiency, but lower and lower pressure drops. Brunazzi and Paglianti,2 analyzing the pressure drops in the most common structured packings, identified two contributions: one due to distributed pressure drops and another, which they called “concentrated pressure drops”, due to the changes of gas direction arising because each element is oriented at a 90° angle to the next one. The first is due to the specific surface and therefore cannot be reduced significantly without reducing mass transfer performances, whereas the second can be reduced avoiding abrupt changes of gas direction. Based on this simple analysis, during the past few years, some of the most important suppliers are looking for new internals that, avoiding abrupt changes of gas direction, maintain high mass transfer performances. The growing problem, therefore, is to suggest reliable models and equations to evaluate surface area, liquid and gas mass transfer coefficients, and pressure drops for these new families of structured packing. To evaluate the effective area available to mass transfer in structured packing, some semiempirical models have been published. Most of them agree in identifying the area available to the mass transfer as a fraction of the geometric surface (Brunazzi et al.,3 Nardini et al.,4 Bravo et al.,5 Billet and Shultes,6 Olujic7) and are based essentially on the study of liquid flow on a inclined layer. To evaluate mass transfer coefficients all published articles agree that the Sherwood number of the gas phase, and therefore the transfer coefficient, is a func* Correspondence should be addressed to Ing. A. Paglianti Department of Chemical Engineering, Via Diotisalvi n. 2, I-56126 Pisa, Italy. E-mail:
[email protected]. † Consorzio Polo Tecnologico Magona. ‡ University of Pisa.
tion of the Reynolds and Schmidt numbers, whereas the published articles partially disagree each other regarding the liquid phase. Recently, Nawrocky and Chuang8 showed that in structured packing the theoretical analysis of the stable liquid rivulet flow can be used to predict with good results the liquid-side mass transfer coefficient. Experimental Apparatus The experimental loop used in this work can operate with columns of 1.5- and 4-in. diameters and 0.985- and 2-m lengths, respectively. Figure 1 is a schematic diagram of the experimental test rig. Two types of experiments have been performed to evaluate gas and liquid mass transfer coefficients. In both cases, compressed air is measured by a set of rotameters and routed to the absorption column, which can work both cocurrently and countercurrently. To evaluate the gasside mass transfer coefficient the liquid stream is sent by a centrifugal pump from the storage tank D2 to the top of the columns; then the liquid that is discharged from the column flows into tank D2. To evaluate liquidside mass transfer the liquid is sent from the storage tank D1 to the top of the columns. The liquid discharged from the columns flows into tank D2. Measurements of absorption/desorption efficiency were performed using two sampling points on the gas line and two sampling points on the liquid line. During each test a sampling before the column inlet and one on the column outlet were done. Each experimental datum was obtained as a mean of three different acquisitions. To evaluate the gas-side mass transfer coefficient, the compressed air was polluted by SO2, whereas the washing solution was an aqueous solution of 4 g/L NaOH. Each gas sample (on the average, 15 L/h) was drawn off the sampling points and after being measured, was sent to two bubblers in a series containing an absorption solution. To evaluate the liquid-side mass transfer coefficient, tank D1 was used to saturate the liquid phase with pure carbon dioxide supplied from a cylinder. During these experiments the solution was fed to the top of the column from tank D1 by a centrifugal pump. After coming into contact with the gas phase, flowing countercurrently, the liquid was discharged from the bottom of the column into tank D2. The desorption efficiency
10.1021/ie990024i CCC: $18.00 © 1999 American Chemical Society Published on Web 07/22/1999
3482 Ind. Eng. Chem. Res., Vol. 38, No. 9, 1999
Figure 2. Schematic drawing of the HelieR from the geometrical point of view.
Figure 1. Schematic diagram of the loop. Symbols: C, absorption column; D, storage tank; DP1, DP cell; FI, flow indicator; PI, pressure indicator; M, mixer; SP, sampling point; P1 P2 P3, valves; CC, CO2 cylinder.
was measured by analyzing the CO2 content in the liquid phase (Brunazzi and Paglianti9). The liquid hold-up measurements were performed by the system of fast closing valves indicated in Figure 1. The closing of valves P1 and P3 was simultaneous with the opening of valve P2. The liquid present in the column flowed down and was measured by the level indicator DP1. Experiments were carried out at atmospheric pressure and room temperature. Liquid flow rates were changed up to 5 m3 m-2 h-1 and gas flow rates up to 31 m3 m-2 h-1. The structured packing analyzed in this work is the HelieR made by Polcon Italiana srl. From the geometrical point of view (see Figure 2) the single element is made by a cylindrical channel filled with an helix. The height of the element is equal to the diameter. The single element can be oriented at a 90° angle to the next one or without any angle discontinuity (see Figure 3). Because of its geometrical properties this structured packing could be used if working liquids are dirty and if they contain suspended solids. The main geometrical characteristics of the packings are reported Table 1. Elements can be linked one to another with a slot system to form tubes of the desired length. Industrial structured packings are obtained by arranging more tubes in a bundle, as a common shell-and-tube heat exchanger. The single element in a tube can either be oriented by an angle of 90° to the next one, which leads to a helix discontinuity, or without any angle discontinuity. A variety of arrangements are possible; helix without any angle discontinuity (in line) after each elements (arrangements 1 × 1), after three or six elements (arrangements 3 × 3 or 6 × 6, respectively).
Figure 3. Schematic drawing of the assembled HelieR. Table 1. Geometric Characteristics of the Tested Packings diameter (inch)
weight single element (gr)
specific area (m2/m3)
void fraction of an element
1 1.5 2 4
3.18 7.34 12.73 50.89
320 210 160 80
0.904 0.936 0.952 0.976
Effective Surface Area Analysis of the correlations reported in the literature concerning the evaluation of effective packing area shows that a good agreement cannot be found in the dependence of this parameter on the superficial velocity, viscosity, density, and surface tension of the liquid. In general, however, all the published correlations assume that the interfacial area increases with liquid load and packing surface area, whereas the influence of the gas is neglected by all the authors for absorption columns operating in conditions below the loading point. Most of the recently published relations (Bravo et al.,5 Nardini
Ind. Eng. Chem. Res., Vol. 38, No. 9, 1999 3483
et al.,4 Brunazzi et al.3) are based on the study of falling film. Nardini et al. 4 and Brunazzi et al.3 showed that the effective surface area available for mass transfer process can be easily computed if the liquid film thickness is evaluated theoretically and the liquid hold-up is measured experimentally. On the basis of the definition of liquid hold-up, the channel fraction occupied by the liquid phase, the ratio between the specific area available for mass transfer process, ae, and the specific geometric surface of the packing, ag, can be computed as:
ae hc de ) ‚ ag 4 δ
(1)
where hc represents the liquid hold-up in the channel, δ represents the liquid thickness, and de the equivalent diameter of the channel. In this work the evaluation of de was done according to the definition suggested by Billet and Schultes:6
4 ag
de )
(3)
xx Fl‚
3µl
( )
‚vl
(4)
2 2
dm‚ωl g + 2 2
2
where the product (dm/2)ωl represents the centrifugal acceleration, dm is the diameter of the packing element, and vl represents the velocity of the liquid film. The liquid velocity is composed by two terms: the axial, val, and the tangential velocity, vtl. The former can be easily computed as:
val )
Usl ‚hc
2π tr
(7)
and P is the height of a single element. The effective velocity of the liquid in the film can be evaluated as:
vl )
Usl
(8)
hc‚‚cos(R)
where R is the slope of the descent line with respect to the vertical axis and can be computed as:
(
)
(dm‚ωl2/2) R ) arctg g
(9)
Finally the ratio between effective and geometrical interfacial area can be evaluated as:
( )( )
de hl ae ) ‚ ag 4
1.5
‚
[x
g2 +
Fl.
(
)
dm 2 ‚ω 2 l
3µl
]
0.5
2
‚cos(R) Usl
‚
(10)
In this work the gas-side mass transfer coefficient was evaluated experimentally with SO2 absorption into aqueous solutions of caustic soda. Working conditions were used to ensure that a surface reaction regime occurred; therefore, it was possible to neglect the resistance to the mass transfer in the liquid phase. In these conditions, the gas-side mass transfer was calculated by referring to the theory of Johnstone and Pigford 11 in the general form:
Shg ) C1‚Reg0.8‚Scg0.33
Reg )
Fg‚deg‚vg µg
(12)
In this work, deg was computed as (Brunazzi and Paglianti2):
deg )
4Ag π‚de
(13)
where Ag, the cross-sectional area available for the gas flow, was evaluated as:
Ag )
(6)
(11)
where the coefficient C1 is equal to 0.054, as suggested by Bravo et al.5 for the usual structured packings, and Reg is defined as:
(5)
where Usl is the superficial liquid velocity and is the void fraction. Revolution velocity, because of the geometrical characteristics of the packing, can be evaluated as:
ωl )
2P val
Mass Transfer Coefficients
hl
Therefore, as suggested by Brunazzi et al.3 and Nardini et al.,4 evaluation of the specific surface the liquid hold-up, hl, and the film thickness is necessary. In the common structured packings, as suggested by Bravo et al.5 the liquid thickness can be computed easily using the equation that describes the flow of a film over an inclined layer (Bird et al.10 ). In this case the equation has to be modified to account for the centrifugal force acting on the liquid film. Therefore under the hypothesis that regime of the liquid is laminar, and neglecting the shear stress at the gas/liquid interface, the film thickness can be computed as:
δ)
tr )
(2)
whereas the liquid hold-up of the channel is a function of the column hold-up hl:
hc )
where the time tr, necessary to complete a revolution, can be evaluated as:
π‚de2 ‚(1 - hc) 4
(14)
To compute the effective gas velocity, vg, it is necessary to evaluate both the tangential and the axial terms. The axial term can be easily computed as:
vag )
Usg ‚(1 - hc)
(15)
3484 Ind. Eng. Chem. Res., Vol. 38, No. 9, 1999
whereas the tangential term has to be computed if the revolution velocity, ωg, is known and integrating the local gas velocity over all the cross-section area. Therefore the mean value of the tangential velocity can be computed as:
vtg )
ωg‚dm 3
(16)
ωg )
vag‚π P
(17)
where
Finally gas velocity can be computed as:
vg )
x( ) ω‚dm 3
2
+ vag2
(18)
In the literature many relations exist that permit the evaluation of the liquid mass transfer coefficient in structured packing (Bravo et al.,5 De Brito et al.,12 Billet and Schultes,6 Brunazzi and Paglianti9): some of these neglect the dependence from the gas velocity and the column length, whereas some others, based on the theoretical analysis on stable rivulet flow (Nawrocki and Chuang8), suggest a dependence of liquid mass transfer coefficient on column height. The approach suggested by Nawrocki and Chuang8 was used in this work. In columns equipped with structured packings, these authors showed that the liquid mass transfer coefficient can be computed as a function of the Graetz and Kapitza numbers, which suggests the following equation:
GzB Shl ) A‚ C Ka
Ka )
σ3‚Fl µ4l ‚g
δ Gz ) Rel‚Scl‚ H Rel )
del‚Fl‚vl µl
µl Scl ) Fl‚Dl
(20)
Z cos(R)
(21)
(22)
(23)
Many relations are available to compute pressure drops in structured packings (Stichlmair et al.,13 Bravo et al.,14 Billet,15 Rocha et al.,16 Robbins,17 Spiegel and Meier18). In this work the model suggested by Brunazzi and Paglianti2 is followed because, as shown by the authors, it allows the prediction of pressure drops based on the simple application of momentum and massconservation equations. These authors divided pressure drops in two terms, distributed and concentrated, because of the changes of direction in gas flow. According to this model and neglecting acceleration and gravitational terms, the pressure drops without liquid flow can be computed as:
( ) (
)
Le Fg‚vg2 dP 1 ) 4fg‚ + Nc‚ ‚ dx de 2 deg‚cos(β)
Le ) 35 de
(28)
The number of changes of flow direction by unit of height can be evaluated as:
1 ne‚P
(29)
where ne is the number of elements filled consecutively in the column without changes of flow direction. The friction factor fg depends on the Reynolds number of the gas and the roughness of the channel walls. In this article the friction factor for both laminar and turbulent conditions was correlated by the following relationship (as suggested by Bravo et al.14):
(24)
The last parameter that has to be defined is the flow distance, H. It represents the path of the liquid phase from the top to the bottom of the column. It is a function of the geometric characteristics of the packings and the column height, Z. According to the analysis suggested to evaluate the surface area, it is possible to compute
(27)
where β is the helix angle with respect to the vertical, Nc is the number of changes in flow direction in a unit height of the packing, deg is the equivalent diameter for the gas phase, and Le/de is the ratio between the equivalent channel length of each flow direction change and the characteristic dimension of the channel. Because the change in the flow direction is approximately 90°, as suggested by Brunazzi and Paglianti2 for common structured packings, it is possible to assume that:
fg ) A1 +
(25)
(26)
Pressure Drops
Nc )
where the liquid film characteristic dimension, del, was computed as:
del ) 4δ
H)
(19)
where
kl‚del Shl ) Dl
the flow distance, H, as a function of the angle R (see eq 9):
B1 Reg
(30)
If liquid flows, concurrently or countercurrently with the gas phase, pressure drops increase because the effective gas velocity and the friction factor increase. Neglecting accelerational and gravitational terms, distributed pressure drops in irrigated cases can be computed as:
1 ) ‚(τ (dP dx ) A
wg‚Sg
g
1 + τi‚Si)‚ cos(β)
(31)
where Sg is the perimeter of the single channel affected by the gas phase, Si is the length of the wetted
Ind. Eng. Chem. Res., Vol. 38, No. 9, 1999 3485
Figure 4. kl‚ae as a function of the number of elements ne; liquid load 31 m3 m-2 h-1.
perimeter, and τi and τwg are the shear stresses at the gas-liquid interface and at the wall, respectively:
ae ‚π‚de ag
Si )
(32)
Sg ) π‚de - Si
(33)
The shear stress at the channel wall is defined as in the dry case,
1 τwg ) ‚fg‚Fg‚vg2 2
(34)
whereas the interfacial shear stress is computed as:
1 τi ) ‚fi‚Fg‚(vg ( vl)2 2
(35)
where (+) is for countercurrent and (-) for concurrent set up. The concentrated term can be evaluated as in the dry case,
dP ) 4‚fm‚Nc‚ ‚ dx de 2
(36)
where the friction factor fm derives from a mean, weighted on the wetted area, between the wall friction factor, fg, and the interfacial friction factor, fi. Friction factor, fi, can be evaluated using the equation suggested by Brunazzi and Paglianti2 as:
[
(
fi ) fg‚ 1 + D1‚
)
del2‚g‚(Fl - Fg) σl
0.3
+
( )( )
El ‚
δ - δ0 µl ‚ de µl0
0.1
]
‚Welγ (37)
where Wel is the Weber number for the liquid phase, µl and µl0 are the viscosity, respectively, of liquid and of water at 20 °C; E1 and γ are two constants. In this article, because shear stress at the gas/liquid interface can be neglected, δ ) δ0 (δ and δ0, are the film liquid thickness with and without gas flow, respectively) and therefore eq 37 becomes:
[
(
fi ) fg‚ 1 + D1‚
First, it is necessary to examine the effect of geometric configuration. This new type of structured packing can be used in different ways, depending on the number of elements ne, assembled without changing the direction of gas flow. It is evident that when ne increases, pressure drops decrease, but unfortunately the wetted surface area decreases too. Therefore it is necessary to look for the setup that permits minimization of pressure drops but maintains high mass transfer efficiency. Figure 4 shows the product kl‚ae as a function of the number of elements ne. The figure shows that best results are obtained for ne ) 3; therefore, this particular setup has been used in the following analysis. As pointed out by several authors to predict properly the wetted surface area, necessary to compute mass transfer efficiency and pressure drops, it is necessary to look for a reliable relation to evaluate liquid holdup. Suess and Spiegel,19 studying columns equipped with structured packing and working with air/water systems, suggested the following empirical correlation: X hl ) Y‚a0.83 g ‚(3600‚Usl)
Le Fg‚vg2
( )
Results and Discussions
)]
del2‚g‚(Fl - Fg) σl
0.3
(38)
(39)
In this work the equation suggested by Suess and Spiegel19 was modified to predict liquid hold-up in the configuration with ne ) 3, where Y and X were evaluated experimentally and assumed the values 7.26‚10-5 and 0.55, respectively. The equation by Suess and Spiegel19 was preferred to others available in the literature because it allows the worker to predict properly the effect of specific surface area on the liquid hold-up. Figure 5 shows a comparison between experimental and computed liquid hold-up values. The experimental data were obtained using two different size of packing, 1.5 and 4 in. The figure shows that if eq 39 is used, an accurate prediction of the experimental values of liquid hold-up can be obtained for both packing sizes. Figure 6 shows a comparison between experimental data of kg‚ae and the proposed model for absorption of SO2 with NaOH solutions. The liquid-phase flow rate changes from 5 to 31 m3 m-2 h-1, whereas the superficial velocity of the gas phase ranges from 0.75 to 2.0 m/s. The root-mean-square error is less than 11.5% for 1.5 in. and it is less then 26% for 4 in. No systematic error arises changing liquid flow rate, gas flow rate, or packing size. The absence of systematic errors with the gas flow rate is particularly important because it
3486 Ind. Eng. Chem. Res., Vol. 38, No. 9, 1999
Figure 5. Liquid hold-up: comparison between experimental and computed data.
Figure 6. kg‚ae: comparison between experimental and computed data obtained by using present model (HelieR; column diameter 1.5 in., 4 in.; air/SO2/NaOH system).
Figure 7. kl‚ae: comparison between experimental and computed data obtained by using present model (HelieR; column diameter 1.5 in.; air/CO2/water system).
justifies the assumption to neglect the influence of gas phase in effective surface area calculation. Figure 7 shows a comparison between the measured and computed values of kl‚ae. No other data have been published, to our knowledge, on this type of structured packings so far; but nowadays, analyzing present experimental data, it is possible to evaluate the constants A, B, and C necessary to compute the liquid-side mass transfer coefficient. The suggested values are A ) 55.15 B ) 0.5, and C ) 0.09. In this case the error mean square is less than 9.5%. Figure 8 shows a comparison between experimental and computed dry pressure drops. Constants A1 and B1 have been evaluated based on the analysis of experimental data performed in this work and assume the values 7.13‚10-3 and 6.16, respectively. Figure 9 shows
a comparison between experimental and computed values of irrigated pressure drops. As for the dry pressure drops, also in this case, the suggested model seems to predict properly experimental data, using D1 ) 3.657. The errors mean square are 7.5% and 16.8% for dry and wetted case, respectively. A very interesting experimental result that has to be pointed out is that flooding phenomena did not occur even at high liquid load, up to 77 m3 m-2 h-1 and high gas flow rate, up to 8000 m3 m-2 h-1. Finally it is necessary to compare performances of HelieR structured packings with the performances of common structured packings. The parameter that has to be taken into account is the specific pressure drops per transfer unit, as suggested by Billet.15 Figure 10 shows the behavior of this parameter as a function of
Ind. Eng. Chem. Res., Vol. 38, No. 9, 1999 3487
Figure 8. Dry pressure drops: comparison between experimental and computed data obtained by using the present model (HelieR; column diameter, 4 in.).
Figure 9. Irrigated pressure drops: Comparison between experimental and computed data obtained by using the present model (HelieR; column diameter 4 in.; liquid load, 5-31 m3 m-2 h-1).
Figure 10. Specific pressure drops per transfer unit: comparison between HelieR 1.5 in., HelieR 4 in., Montz B1-200 (computed according to Billet15) and Mellapak 250 Y; liquid load, 21 m3 m-2 h-1.
the gas capacity factor, Fv, for different packings, at a liquid load of 21 m3 m-2 h-1. It is interesting to notice that 1.5 HelieR structured packings present lower values of pressure drops per transfer unit than common structured packings made of corrugated layers. Nevertheless, with these good performances it is also necessary to underscore that present structured packing shows height of transfer unit 20% higher than height of transfer unit achievable with Mellapack or Montz packing. Conclusions This work reports experimental data on a new type of structured packing. This new type of structured
packing is particularly interesting because it shows lower pressure drops per transfer unit than the common structured packing. This characteristic is very important because it is essential for saving energy and avoiding thermal decomposition of the products in the operation of thermal separation processes. From the experimental point of view two sets of experiments have been performed to measure separately mass transfer efficiency when resistance is concentrated in the gas phase and when it is concentrated in liquid phase. Some experimental data on pressure drops have been performed also. To predict the performances of this new packing a model that allows the evaluation of effective wetted area and pressure drops has been suggested. Further, two
3488 Ind. Eng. Chem. Res., Vol. 38, No. 9, 1999
relations for evaluating gas-side and liquid-side mass transfer coefficients are also proposed. The model is based on a very fundamental approach because it derives from mass and momentum-conservation equations, whereas the equation suggested to compute liquid-side mass transfer coefficient derives from the theoretical analysis on stable liquid-rivulet flow. Acknowledgment The authors thank Ing. G. Petrillo from Polcon Italiana srl, Via F.lli Cervi, 77 Cantalupo (Milan), for having provided the HelieR packing. Notation A ) constant in eq 19 A1 ) constant in eq 30 ae ) effective interfacial area, m2/m3 ag ) packing specific area, m2/m3 Ag ) cross-section area available for the gas, m2 B ) constant in eq 19 B1 ) constant in eq 30 C ) constant in eq 19 C1 ) constant in eq 11 de ) equivalent diameter of the channel, m deg ) gas side characteristic dimension, m del ) liquid film characteristic dimension, m dm ) diameter of the single element, m D1 ) constant in eq 37 D ) Diffusivity, m2/s Fv ) Usg (Fg)0.5, gas capacity factor, m/s‚(kg/m3)0.5 f ) friction factor g ) gravitational constant, m/s2 hc ) liquid hold-up in a channel hl ) liquid hold-up in column H ) flow distance, m kg ) gas-side mass transfer coefficient, m/s kl ) liquid-side mass transfer coefficient, m/s Le ) equivalent channel length, m Nc ) number of flow direction changes in a unit height ne ) number of elements filled consecutively P ) height of a single element, m Sg ) channel perimeter affected by the gas phase, m Si ) channel perimeter wetted by the liquid phase, m tr ) revolution time, s Usg ) superficial gas velocity, m/s Usl ) superficial liquid velocity, m/s va ) axial velocity, m/s vg ) effective gas velocity, m/s vl ) effective liquid velocity, m/s vt ) tangential velocity, m/s X ) constant in eq 39 Y ) constant in eq 39 Z ) column height, m Greek letters R ) angle of the film liquid velocity respect to the vertical axis, ° β ) angle of the gas velocity respect to the vertical axis, ° δ ) liquid film thickness, m ) void fraction µ ) viscosity, kg m-1 s-1 Π ) 3.1415 F ) density, kg/m3 σ ) surface tension of liquid, N/m τ ) shear stress, N/m2 ω ) revolution velocity, 1/ s Subscript and superscript g ) gas phase
i ) interfacial l ) liquid phase m ) mean w ) wall Dimensionless groups Gz ) Rel‚Scl‚δ/H ) Graetz number Ka ) σ3‚Fl/µl4‚g ) Kapitsa number Reg ) deg‚Fg‚vg/µg ) Reynolds number for the gas phase Rel ) del‚Fl‚vl/µl ) Reynolds number for the liquid phase Scg ) µg/Fg‚Dg ) Schmidt number for the gas phase Scl ) µl/Fl‚Dl ) Schmidt number for the liquid phase Shg ) kg‚deg/Dg ) Sherwood number for the gas phase Shl ) kl‚del/Dl ) Sherwood number for the liquid phase Wel ) Fl‚vl2‚del/σl ) Weber number
Literature Cited (1) Brunazzi, E.; Paglianti, A.; Petarca, L. Design of Absorption Columns Equipped with Structured Packings. Chim. Ind. 1996, 78, 459. (2) Brunazzi, E.; Paglianti A. A Mechanistic Pressure Drop Model for Columns Containing Structured Packings. AIChE J. 1997, 43, 317. (3) Brunazzi, E.; Nardini, G.; Paglianti, A.; Petarca, L. Interfacial Area of Mellapak Packing: Absorption of 1,1,1-Trichloroethane by Genosorb 300. Chem. Eng. Technol. 1995, 18, 248. (4) Nardini, G.; Paglianti, A.; Petarca, L.; Viviani, E. Sulzer BX Gauze: Fluodynamics and Absorption of Acid Gases. Chem. Eng. Technol. 1996, 19, 20. (5) Bravo, J. L.; Rocha, J. A.; Fair, J. R. A Comprehensive Model for the Performance of Columns Containing Structured Packings. I. Chem. E. Symp. Ser. 1992, 128, A439. (6) Billet, R.; Schultes, M. Predicting Mass Transfer in Packed Columns. Chem. Eng. Technol. 1993, 16, 1. (7) Olujic′ Z. Simulation of Structured Packing Performance. Chem. Biochem. Eng. Q. 1997, 11, 31. (8) Nawrocki, P. A.; Chuang, K. T. Carbon Dioxide Absorption into a Stable Liquid Rivulets. Can. J. Chem. Eng. 1996, 74, 247. (9) Brunazzi, E.; Paglianti, A. Liquid-film Mass-Transfer Coefficient in a Column Equipped with Structured Packings. Ind. Eng. Chem. Res. 1997, 36, 3792. (10) Bird, R. B.; Stewart, W. E.; Lightfoot, E. N. Transport Phenomena; Wiley: New York, 1960. (11) Johnstone, H. F.; Pigford, R. L. Distillation in a WettedWall Column. Trans. Am. Inst. Chem. Eng. 1942, 38, 25. (12) Henriques de Brito, M.; von Stockar, U.; Menendez Bangerter, A.; Laso, M. Effective Mass-Transfer Area in a Pilot Plant Column Equipped with Structured Packings and with Ceramic Rings. Ind. Eng. Chem. Res. 1994, 33, 647. (13) Stichlmair, J.; Bravo, J. L.; Fair, J. R. General Model for prediction of Pressure Drop and Capacity of Countercurrent Gas/ Liquid Packed Columns. Gas Sep. Purif. 1989, 3, 19. (14) Bravo, J. L.; Rocha, J. A.; Fair, J. R. Pressure Drop in Structured Packings. Hydrocarbon Proc. 1986, 56, 45. (15) Billet, R. Packed Column Analysis and Design; RuhrUniversita¨t Bochum, Department of Thermal Separation Processes: Germany, 1989. (16) Rocha, J. A.; Bravo, J. L.; Fair, J. R. Distillation Columns Containing Structured Packings: A Comprehensive Model for Their Performance. I. Hydraulic Models. Ind. Eng. Chem. Res. 1993, 32, 641. (17) Robbins, L. A. Improve Pressure Drop Prediction with a New Correlation. Chem. Eng. Prog. 1991, 87, 87. (18) Spiegel, L.; Meier, W. A. Generalized Pressure Drop Model for Structured Packings. Ind. Chem. Eng. Symp. Ser. 1992, 128, B85. (19) Suess, P.; Spiegel, L. Holdup of Mellapak structured packings. Chem. Eng. Proc. 1992, 31, 119.
Received for review January 8, 1999 Revised manuscript received May 17, 1999 Accepted May 25, 1999 IE990024I
