p.1 

OAR Box 1870 

Prepped by Ollie Stewart 

Document Number: 

2DII-A-33 

Docket Number: 

A-90-19 

Printed 11/28/2008 7:05:04 AM Sheet 21 of 67 
P.2 

AIR POLLUTION 

ENGINEERING MANUAL 

SECOND EDITION 

Compiled and Editsd 

by 

John A. Doniolion 

AIR POLLUTION CONTROL DISTRICT 

COUNTY OF LOS ANGELES 

A-90-19 

ll-A-33 

ENVIRONMENTAL PROTECTION AGENCY 

Office of Air ond Water Programs 

Office of Air Quality Planning and Standards 

Research Triangle Park, N.C 27711 

May 1973 
Gas Absorption Equipment 207 

P.3 

16. Horizontal position of a condenser distributes the vapor better and permits 

easier removal of the condensate (Kern, 

1950). 

17. In the horizontal condenser, it is necessary to prevent cooled-condensate from 

forming liquid pools and impeding the 

flow of vapors (Kern, 1950). 

IS. Selection of which material should pass 

through tubes cannot be decided by 

fixed rules, because of factors at a variance with one another. When corrosive 

condensate is encountered, condensation 

within the tubes rather than the shell is 

usually desirable (Nelson, 1959). 

APPLICATIONS 

Condensers have been used successfully 

(either separately or with additional control 

equipment) on the following processes or 

equipment: 

REFINERY AND PETROCHEMICAL 

Alkylation unit accumulator vents 

Amine stripper units 

Butadiene accumulator vents 

Coker blowdown 

Ketone accumulator vents 

Lube oil rerefining 

Polyethylene gas preparation accumulator vents 

Residium stripper unit accumulator vents 

Storage equipment 

Styrene-processing units 

Toluene recovery accumulator vents 

Udex extraction unit 

CHEMICAL 

Manufacture and storage of ammonia 

Manufacture of Cooper naphthenates 

Chlorine solution preparation 

Manufacture of ethylene dibromide 

Manufacture of detergent 

Manufacture of insecticide 

Manufacture of latex 

Manufacture of nitric acid 

Manufacture of phthalic anhydride 

Resin reactors 

Soil conditioner formulators 

Solvent recovery 

Thinning tanks 

MISCELLANEOUS 

Aluminum fluxing 

Asphalt manufacturing 

Blood meal driers 

Coal tar-dipping operations 

Degreasers 

Dry cleaning units 

Esterfication processes 

Pectin preparation 

Rendering cookers (animal waste) 

Vitamin formulation 

GAS ABSORPTION EQUIPMENT 

Gas absorption is the mechanism whereby one 

or more constituents are removed from a gas 

stream by dissolving them in a selective liquid 

solvent. This is one of the major chemical 

engineering unit operations and is treated extensively in all basic chemical engineering textbooks. These texts deal with gas absorption as 

a method of recovering valuable products from 

gas streams, for example, in petroleum production, natural gasoline is removed from wellhead 

gas steams by absorption in a special hydrocarbon oil. Absorption is also practiced in industrial chemical manufacturing as an important 

operation in the production of a chemical compound. For example, in the manufacture of 

hydrochloric acid, one step in the process involves the absorption of hydrogen chloride gas 

in water. 

From an air pollution standpoint, absorption is 

useful as a method of reducing or eliminating 

the discharge of air contaminants to the atmosphere. Even in this application, absorption 

can yield profits to the user. For example, 

it can be employed to remove hydrogen sulfide 

from process gas streams in a petroleum re -

finery to meet air pollution regulations. With 

further processing, this hydrogen sulfide can 

be converted to elemental sulfur, a valuable 

product. 

The gaseous air contaminants most commonly 

controlled by absorption include sulfur dioxide, 

hydrogen sulfide, hydrogen chloride, chlorine, 

ammonia, oxides of nitrogen, and light hydrocarbons. 

K*"<Wr*bSJJ_l 
P.4 

208 CONTROL EQUIPMENT FOR GASES AND VAPORS 

In other examples, such as solvent recovery, 

desorption or stripping may be practiced after 

absorption not only to recover a valuable absorbed constituent but also to recover valuable 

solvent for reuse. Sometimes, after absorption, solute and solvent are not separated but 

are used as a product or intermediate compound in chemical manufacture. 

Treybal (1955) lists some important aspects 

that should be considered in selecting absorption solvents. 

1. The gas solubility should be relatively high 

so as to enhance the rate of absorption and 

decrease the quantity of solvent required. • 

Solvents similar chemically to the solute 

generally provide good solubility. 

2. The solvents should have relatively low 

volatilities so as to reduce solvent losses. 

3. If possible, the solvents should be noncorrosive so as to reduce construction 

costs of the equipment. 

4. The solvents should be inexpensive and 

readily available. 

5. The solvents should have relatively low 

viscosities so as to increase absorption 

and reduce flooding. 

6. If possible, the solvents should be nontoxic, 

nonflammable, chemically stable, and have 

low freezing points. 

GENERAL TYPES OF ABSORBERS 

PACKED TOWER DESIGN 

A packed tower is a tower that is filled with 

one of the many available packing materials, 

as shown in Figure 142. The packing is designed so as to expose a large surface area. 

When this packing surface is wetted by the 

solvent, it presents a large area of liquid 

film for contacting the solute gas. 

GAS OUT 

LIQUID OlSTRISurOR 

LIQUIO 

RE-0ISTR18UTOR 

M C I INO SUPPORT 

au IN 

nauio OUT 

Figure 142. Schematic diagram 

of a packed tower (Treybal, 1955 

P. 134). 

• Wl L. 

Gas absorption equipment is designed to provide 

thorough contact between the gaa and liquid solvent in order to permit interphase diffusion of 

the materials. The rate of mass transfer between the two phases is largely dependent upon 

the surface exposed. Other factors governing 

the absorption rate, such aa solubility of the 

gas in the particular solvent and degree of 

chemical reaction, are characteristic of the 

constituents involved and are more or less independent of the equipment used. This contact 

between gas and liquid can be accomplished by 

dispersing gas in liquid or vice versa. 

Absorbers that disperse liquid include packed 

towers, spray towers or spray chambers, and 

venturi absorbers. Equipment that uses gas 

dispersion includes tray towers and vessels 

• with sparging equipment. 

Usually the flow through a packed column is 

counter current, with the liquid introduced at 

the top to trickle down through the packing 

while gas is introduced at the bottom to pass 

upward through the packing. This results 

in highest possible efficiency, since, as the 

solute concentration in the gaa stream decreases as it riaea through the tower, there 

is constantly fresher solvent available for contact. This gives maximum average driving 

force for the diffusion procesa throughout the 

entire column. 

In concur-rent flow, where the gaa stream and 

solvent enter at the top of the column, there 

is initially a very high rate of absorption that 

constantly decreases until, with an infinitely 

tall tower, the gaa and liquid would leave in 
P.5 

i_____M_H-l 

Gas Absorption Equipment 

equilibrium. Concurrent flow is not often 

used except in the case of a very tall column 

built in two sections, both located on the 

ground, the second section using concurrent 

flow merely as an economy measure to obviate the need for constructing the large gas 

pipe from the top of the first section to the 

bottom of the second. Moreover, for an 

operation requiring an exceptionally high 

solvent flow rate, concurrent ilow might be 

used to prevent flooding that could occur in 

countercurrent flow. 

Packing Materials 

The packing should provide a large surface 

area and, for good fluid flow characteristics, 

should be shaped to give large void space 

when packed. It should likewise be strong 

enough to handle and install without excessive breakage, be chemically inert, and be 

in- :pensive. 

Rock and gravel have been used but have 

disadvantages of being too heavy, having 

small surface areas, giving poor fluid flow 

and, at times, not being chemically inert. 

Coke lumps are also used sometimes and 

here the weight disadvantage is not present. 

Owing to its porosity, coke has a large 

surface area per unit volume. The exposed 

surface is not, however, as large as might 

be expected since the pores are so small 

that they become filled or filmed over by 

the solvent, which considerably reduces 

the effective surface. 

Generally, packing in practice consists of 

various manufactured shapes. Raschig 

rings are the most common type, consisting 

of hollow cylinders having an external diameter equal to the length. Other shapes 

include Berl saddles, Intalox saddles, 

Lessing rings, cross-partition rings, 

spiral-type rings, and drip-point grid 

tiles. Figure 143 shows several common 

shapes. Physical characteristics of these 

various types of packings have been determined experimentally and compiled in 

tables by Leva (1953). 

Packing may be dumped into the column randomly, or regularly shaped packing may be 

manually stacked in an orderly fashion. Randomly dumped packing has a higher specific 

surface contact area and a higher, gas pressure 

drop across the bed. The stacked packings have 

an advantage of lower pressure drop and higher 

possible liquid throughout, but the installation 

cost is obviously higher. Table 57 and Figure 

144 list typical packing costs and packed-tower 

installed prices for 1959. 

fkll RING 

Figure 143. Common tower packing materials 

(Teller, 1960. p. 122). 

Table 57. COSTS OF REPRESENTATIVE 

TOWER PACKINGS (Teller, 1960) 

Packing 

Raschig rings, ceramic 

Raschig ringa. carbon 

Berl saddles, ceramic 

Intalox saddles, ceramic 

Intalox saddles, carbon 

Tellerettes, polyethylene 

Low density 

High density 

Pall rings, ccramu (BASF) 

Pall rings, polypropylene 

Pall rings, sfainless steel 

Cost of packing 

l/2-in. 

11. 70 

16.90 

24.30 

- J . 55 

-

-

• 

-

41.00 

IA6. 50 

1-in. 

6.50 

9.60 

9.90 

9.40 

IB.60 

16.00 

.3.00 

5.00 

.6.0 0 

96.00 

$/(t 3(l959basis) 

l-l/.-in . 

5.05 

3.00 

7. 50 

7. 15 

19.40 

-

-

-

20.75 

3).00 

2-in. 

4.35 

6.60 

7. 70 

7. JO 

-

-

-

• -

13. 50 

69.00 

Liquid Distribution 

Since the effectiveness of a packed tower depends on the availability of a large, exposed, 

liquid film, then obviously, if poor liquid distribution prevents a portion of the packing from 

being irrigated, that portion of the tower is 

ineffective. Poor distribution can be due to 

improper introduction of the liquid at the top of 

the tower and to channeling within the tower. 
P.6 

210 ' CONTROL EQUIPMENT FOR GASES AND VAPORS 

10 000 

S : JOS 

_ 100 

] 

_ _ . H= i 

i 

<+y 

- __ "Y 

^ 

j j 

= 

_-

fcffjjj 

i Mi 

r 

i 

X — -+-

r 

3 4 5 10 20 30 40 SO 100 

OIABETER inches 

Figure 144. Packed-tower costs, 1959, with Raschig rings as 

packing (Tel ler, 1960. p. 123). 

At least five points of introduction of liquid per 

square foot of tower cross-section must generally be provided to ensure complete wetting. The 

liquid rate must be sufficient to wet the packing 

but not to flood the tower. Treybal (1955) states 

that a superficial liquid velocity of at least 800 

pounds of liquid per hour per square foot of 

tower cross-section is desirable. 

Solid-cone spray nozzles make excellent distributors but may plug if solid particles are 

present in the solvent. In randomly packed 

towers, the liquid tends to channel toward the 

walls, because of the usually lower packing 

density near the walls. In tall towers this 

channeling is controlled by liquid redistributors at intervals of 10 to 15 feet. Moreover, 

this effect is minimized if the packing pieces 

are less than one-eighth the diameter of the 

tower. 

Tower Capacity 

The terms used to indicate capacity of a 

packed column or tower are load point and 

flood point. For a given packing and liquid 

rate, if gas pressur e drop is plotted against 

gas velocity on a logarithmic scale, there 

are two distinct breakpoints where the slope 

of the curve increases. At low gas velocities the curve is almost parallel to that obtained with dry packing, but above the breakpoints, the pressure drop increases more 

rapidly with increased gas velocity. The lower of these two breaks is the load point and the 

higher one the flood point. 

As gas velocity increases above the load 

point, the liquid holdup in the bed increases 

until, at the second breakpoint, the flood 

point, most of the void space in the tower 

is filled with liquid and there is liquid 

entrainment in the gas stream. Of course, 

at this point there is an excessive pressure 

drop. Columns should seldom be operated 

above the load point, but since the load point 

is sometimes more difficult to establish than 

the flood point, it is common practice to de -

sign for 40 to 70 percent of the flood point. 

In general, flooding velocities ar e considerably 

higher for stacked packing than for dumped 

packing. The plot of Lobo (Figure 145) can 

• v «•.. .•-•/•__• 
P.7 

Gas Absorption Equipment 11 

1.0 

0 1 

I 

% 

, Ol 

la 

~ 0 01 

.OoM 

.oO . 

.00 2 

0 001 0. 01 o m .t >».t 1 .0 I 

• *v . 

0 1 V T 1 0 I I . 1 i 7 .10. 0 

Figure 145. Correlation for flooding rate in randomly packed 

towers (Lobo, 1945, p. 693). 

be used to determine flow rates that will cause 

flooding. This curve is based on measurements 

with several liquids and gases on a variety of 

packings. 

For many years packed towers were designed 

on the same basis as plate or tray towers. 

The number of theoretical plates or trays required for a given degree of separation was 

calculated and this quantity multiplied by a 

figure called height equivalent to a theoretical 

plate (HETP). This HETP was an experimentally determined figure varying widely with 

packing, flow rates of each fluid used, and 

concentration of solute for any specific system. 

Experimental evaluation of these variables 

made use of this system too cumbersome and 

it is now rarely used. Design procedures now 

employ the concept of the transfer unit. The 

major design items to be calculated are the 

column diameter, number of transfer units, 

the height oT a transfer unit, and the system 

pressure drop. These will be discussed individually. 

Tower Diomstor 

As mentioned previously, gas velocity is limited 

by flooding conditions in the tower; By use of 

the design gas volume, design solvent flow rate, 

and type of packing, the tower diameter can be 

computed by using Lobo's correlation in Figure 

145. Packing factors are obtained from Figure 

146. The procedure is as follows: 

1. Calculate the factor (

P \ 0'

5 

where 

L' = liquid flow rate, lb/hr 

V = gas flow rate, lb/hr 

p_ = gas density, lb/ft 3 

p. = liquid density, lb/ft 
P.8 

212. CONTROL EQUIPMENT FOR GASES AND VAPORS 

10 000 

I 000 

100 

PAC) 

S»0l 

0% 

•-.< 

ING FACTOR f 

LES. a i ] = 

_V^S 

••. > 

• * 

OR 1 in. INT 

90 

7 ^ 

A10X 

i » -

0 5 1 0 I S 2 0 2 5 

NOMINAL PACKING SIZE "lcties 

3 0 3 5 

Figure 148. Packing factors for Raschig rings and saddles 

(lobo, 1945. p. 693). 

2. Using the calculated value in (1), obtain 

from Figure 145 the value of 

n?)(>n 

g

c

 PG P L 

where 

G' = gas flow rate, lb/sec-ft of tower 

cross-section 

— = packing factor from Figure 146. 

p = liquid density, lb/ft 

L 

ji' = liquid viscosity, centipoises 

g = gravitational constant, 32.2 ft/sec . 

Solve for G', the superficial mass gas velocity at flood point from the factor determined in (2). 

Calculate S, the tower cross-section area 

in ft^ for fraction of flooding velocity selected, f, by the equation 

p_ = gas density, lb/ft' G 

S = (G«)(f)(3.600) 

(92) 

ramm m 
P.9 

Gas Absorption Equipment 213 

Calculate the tower's inside diameter, Dc, 

by the equation 

Dc •w 

0.5 

(93) 

Tower diameter should be calculated for 

conditions at both top and bottom of the 

tower. The tower is designed to the larger 

diameter. 

Numb.r of Tronsf.r Units (NTU) 

A transfer unit is a measure of the difficulty of 

the mass transfer operation and is a function of 

the solubility and concentrations of the solute 

gas in the gas and liquid streams. It is expressed as NQQ or NQL, depending upon whether 

the gas film or liquid film resistance controls 

the absorption rate. The gas film resistance 

usually controls when solubility of solute in solvent is high and conversely, the liquid film controls when the solubility is low. 

In air pollution control work where, in general, 

a relatively small concentration of solute is to 

be removed from an airstream, a solvent in 

which the solute gas is highly soluble is usually 

selected in order to obtain the highest possible 

economic separation. Thus, for the majority 

of cases encountered, the gas film resistance 

will be controlling. 

One of the most widely used methods of determining the number of transfer units is that proposed 

by Baker (1935), which is based upon an operating 

diagram consisting of an equilibrium curve and 

an operating line. For a given gas-liquid system, 

if the temperature is constant and the gas partial 

pressure is varied, the gas concentration in the 

liquid changes to an equilibrium concentration 

at each partial pressure. If the system consists 

of a soluble gas to be removed, an insoluble 

carrier gas, and a solvent, then, as the amount 

of soluble gas in the system increases, the 

equilibrium concentration of the soluble gas in 

the liquid .increases but not proportionally. 

These equilibrium conditions can exist for an 

infinite number of concentration states and, when 

plotted on X-Y coordinates, become the equilibrium curve. The operating line represents the 

concentrations of solute in the gas stream and 

in the liquid phase at various points in the tower. 

When plotted as moles solute per mole solvent 

versus moles solute per mole gas on X-Y coordinates, the jesult is a straight line. Thus, 

when the composition of the inlet gas and the 

desired or required degree of absorption are 

known, the points on the operating line for each 

end of the column can be calculated. The operating line is the straight line connecting the 

two points. For absorption to occur, the operating line must lie above the equilibrium curve 

on the diagram. The relative position of the 

operating line and equilibrium curve indicates 

how far the tower conditions are from equilibrium. The more widely separated the lines, 

the further the tower conditions are from equilibrium and the greater is the driving force 

for the absorption operation. 

Figure 147 illustrates the graphical method 

of determining the number of transfer units 

for a counter current packed tower with the 

gas film controlling the absorption rate. The 

equilibrium curve (line AB) for the particular 

gas-liquid system is plotted from experimental 

data, which, for most common systems, has 

been determined. Much of these data can be 

located in the International Critical Tables 

and in Perry (1950). The operating line is 

a straight line drawn between points D and C. 

D is the point representing the concentrations of solute in the gas stream and in the 

liquid stream at the gas inlet and liquid outlet (bottom of the tower). Point C corresponds 

to these concentrations at the top of the column. 

Line EF is drawn so that all points on the line 

are located midway on a vertical line between 

the operating line and equilibrium curve. 

Starting at point C on the operating line (conditions at the top of the column), draw a horizontal line CH so that CG = GH. Then draw a 

vertical line HJ back to the operating line. 

J, 

X' 

1 

y / 

X*0> 

t 

r*B 

X = SOLUTE moles/SOIVENT mole 

Figure 147. Graphical determination of 

the number of transfer units. 

r

-**S*mm 
P.10 

214 CONTROL EQUIPMENT FOR GASES AND VAPORS 

The step CHJ represents one gas transfer unit. 

This stepwise procedure is continued to the 

end of the operating line (conditions at the 

bottom of the column). Two gas transfer units 

(NOG)

 ar e

 shown in Figure 147. 

If the liquid film resistance is the controlling 

factor in the transfer of solute to solvent, draw 

the line EF so that all points on the line are 

located midway on the horizontal axis between 

the operating line and equilibrium curve. Then, 

starting at point D on the operating line, draw 

a vertical line DK so that DL = LK. The step 

is completed by drawing a line KJ back to the 

operating line. This procedure is then continued to point C on the operating line. Figure 

147 does not accurately indicate the number of 

liquid transfer units since the line EF was 

drawn for the case where the gas film resistance controls. 

Height of a Transfer Unit 

Generalized correlations ar e available for 

computing the height of a transfer unit and 

are expressed as H Q and Hj_ for heights of gas 

and liquid transfer units respectively. These 

use experimentally derived factors based on the 

type of packing and the gas and liquid flow rates 

as shown in equations 94 and 95. 

where 

H, 

© G " 

.0.5 

P0 D C G 

H = height of a gas transfer unit, tt 

G = superficial gas rate, lb/hr-ft 

L = superficial liquid rate, lb/hr-ft 

o = a packing constant from Table 59 

|3 = a packing constant from Table 53 

7 = a packing constant from Table 53 

H_ = gas viscosity, lb/hr-ft 

p_ = gas density, .lb/ft G 

D_ = gas diffusivity, ft /hr. G 

"""(*%• ) 

(94) 

The group | -—^— | is known as the Schmidt 

number as shown in Table 59. 

Table 58. CONSTANTS FOR USE IN DETERMINING GAS FILM'S 

HEIGHT OF TRANSFER UNITS (Treybal, 1955, p. 239) 

Packing 

Raschig rings 

3/8 in. 

1 in. 

1-1/2 in. 

2 in. 

Berl saddles 

1/2 in. 

1 in. 

1-1/2 in. 

3-in. partition rings 

Spiral rings (stacked 

staggered) 

3-in. single spiral 

3-in. triple spiral 

Drip-point grids 

No. 6146 

No. 6295 

a 

2. 32 

7.00 

6.41 

17.30 

2.58 

3.82 

32.40 

0.81 

1.97 

5.05 

650 

2.38 

15.60 

3.91 

4.56 

(3 

0.45 

0.39 

0.32 

0.38 

0.38 

0.41 

0.30 

0.30 

0.36 

0.32 

0.58 

0.35 

0.38 

0.37 

0. 17 

. 

0.47 

0.58 

0.51 

0.66 

0.40 

0.45 

0.74 

0.24 

0.40 

0.45 

1.06 

0.29 

0.60 

0.39 

0.27 

Range of 

G' 

200 to 500 

200 to 800 

200 to 600 

200 to 700 

200 to 700 

200 to 800 

200 to 700 

200 to 700 

200 to 800 

200 to 1,000 

150 to 900 

130 to 700 

200 to 1,000 

130 to 1,000 

100 to 1,000 

L' 

500 to 1,500 

400 to 500 

500 to 4, 500 

500 to 1,500 

1,500 to 4,500 

500 to 4, 500 

500 to 1,500 

1,500 to 4,500 

400 to 4, 500 

400 to 4, 500 

3,000 to 10,000 

3, 000 to 10, 000 

500 to 3, 000 

3,000 to 6,500 

2,000 to 11,500 

mr>-
P.11 

Ga s Absorptio n Equipmen t 2 1 : 

.1 

I.• 

Table 59 DIFFUSION COEFFICIENTS OF 

GASES AND VAPORS IN AIR AT 25°C AND 

1 ATM (Perry , 1950) 

Substanc e 

•/Ammoni a ] 

Carbon dioxide I 

Hydrogen

 ; 

Oxygen 

Water 

• Carbon disulfide! 

^ Ethyl ethe r ! 

Methanol 

• Ethyl alcohol 

Propy l alcohol 

v Butyl alcohol 

Amyl alcohol 

v

 Hexyl alcohol 

Formi c acid 

J Aceti c acid 

Propioni c acid 

i-Butyri c acid 

Valeri c acid 

i-Caproi c acid 

Diethyl amin e 

Butyl amin e 

Aniline 

Chlor o benzen e 

Chlor o toluene 

Propy l bromid e 

Propy l iodide 

v Benzen e 

Toluene 

Xylene 

/ Ethyl benzen e 

Propy l benzen e 

Diphenyl 

n-Octan e 

Mesitylen e 

D, cm /sec 

0. 236 

0. 164 

0. 410 

0. 206 

0. 256 

0. 107 

0. 093 

0. 159 

0. 119 

0. 100 

0. 090 

0. 070 

0. 059 

0. 159 

0. 133 

0. 099 

0. 031 

0. 067 

0.060 

0. 105 

0. 101 

0. 072 

0. 073 

0.065 

0. 105 

0.096 

0.088 

0.084 

0.071 

0. 077 

0.059 

0.068 

0.060 

0.067 

HCL 

HP 

wher e 

H. 

0X77 

0 . \X°\ 

H. fPD 

0 

0 

0 

0 

0 

1 

1 

0 

1 

1 

1 

2 

2 

0 

1 

1 

1 

2 

2 

1. 

1 

2 . 

2 . 

2 . 

1. 

1. 

1. 

1. 

2 . 

2 . 

2 . 

2 . 

2 . 

2 . 

. 6 6 

. 94 

.ZZ 

75 

60 

45 

66 

97 

30 

55 

72 

21 

60 

97 

16 

56 

91 

31 

58 

47 

53 

14 

12 

33 

47 

61 

76 

8 4 

18 

01 

62 

2 8 

58 

31 

(*)" (X) 

p. SI 

i. is 

5 

P, 7? 

(95) 

0 

n 

= height of a liquid transfer unit, ft 

2 

= superficia l liquid rate, lb/hr-ft 

= liquid viscosity , lb/hr-ft 

= a packing constant, Table 60 

= a packing constant, Table 60 

= liquid density, lb/ft 

Tabl e 60. CONSTANTS FOR USE IN 

DETERMINING LIQUID FILM'S HEIGHT OF 

TRANSFER UNITS (Treybal, 1955, p. 237) 

Packing 

Raschig rings 

3/3 in. 

1/2 in. 

1 in. 

1-1/2 in. 

2 in. 

Berl saddles 

1/2 in. 

1 in. 

1-1/2 in. 

3-in. partition rings 

Spiral rings (stacked! 

staggered) j 

0. 0013210. 46 

0. 00357 I 0. 35 

0.0100 ^0.22 

0.0111 ;0.22 

0.0125 0.22 

0.00666^0.23 

0.0053SJ0.23 

0.00625 10.23 

0.0625 J0.09 

Range of L 

400 to 15 

400 to 15 

400 to If 

400 to 15 

400 to 1 : 

000 

000 

000 

000 

000 

400 to 1 

400 to I 

400 to 1 

5. 000 

i. 000 

-. 000 

3, 000 to 14. 000 

3-in. single spiral 

3-in. triple spiral 

Drip-point grids 

No. 6146 

No. 6295 

0 

0 

0 

0 

00909|0 

01 16 '0 

1 

0154 lo 

00725 lo 

23| 

23! 3 

23 3 

31 1 2 

1 

400 to 1* 

000 to 14 

500 to 30 

500 to 12. 

000 

000 

000 

000 

The grou p (ft)-

s the Schmidt numbe r 

a s shown in Tabl e 61 . Eac h of thes e empirica l 

equations neglects the effect of the othe r film's 

resistance . Actually, however, even in the 

cas e of absorbin g highly soluble ammoni a in 

water, experimenta l results have shown that 

the liquid film resistanc e is significant. The 

height of an overall gas transfe r unit, H_._, 

is determine d by the following equation, which 

takes into account the liquid film resistance . 

H

oc

 = Ho * (ml 

* m / 

(96) 

wher e 

0° slope of the equilibrium curv e 

gas rate , lb-moles/h r 

liquid rate , lb-moles/hr . 

D = liquid diffusivity, ft /hr . 

Thes e equations ar e widely accepted for desig n 

purposes. Another group of equations listed 

by Leva (1953) include overall gas and liquid 

transfe r coefficients, Kg a and K^a. Still mor e 

recently, Cornell et al. , (1960) published corre -

lations based on a mass of experimental dat a reported up to 1957. 

* . »7.!»_A 
P.12 

216 CONTROL EQUIPMENT FOR GASES AND VAPORS 

Tabl e 61 . DIFFUSION COEFFICIENTS IN 

LIQUIDS AT 20°C (Perry , 1950) 

Solute 

o2 

co2 N 2 0 

NH3 

CI2 

Br 2 

H2 

N2 

HCl 

H2S 

H 2 S 0 4 

HNO3 

Acetylene 

Aceti c acid 

Methanol 

Ethanol 

Propano l 

Butanol 

Allyl alcohol 

Phenol 

Glycero l 

Pyrogallo l 

Hydroquinone 

Ure a 

Resorcino l 

Urethan e 

Lactos e 

D x 105 

(cm 2/sec ) x 103 

1. 30 

1. 77 

1. 51 

1.76 

1.22 

1.20 

5. 13 

1.64 

2.6 4 

1. 41 

1. 73 

2.6 0 

1. 56 

0.88 

1.28 

1.00 

0.87 

0.77 

0. 93 

0.8 4 

0.72 

0. 70 

0. 77 

1.06 

0. 80 

0.92 

0.4 3 

___. 

pD 

553 

559 

665 

570 

324 

340 

196 

613 

381 

712 

530 

390 

645 

1, 1 40 

735 

1,005 

1, 150 

1, 310 

1,080 

1,200 

1,400 

1, 440 

1, 300 

946 

1,260 

1,090 

2,34 0 

Maltos e 

Glucos e 

Mannitol 

Raffinose 

Sucros e 

Sodium chlorid e 

Sodium hydroxide 

co2

b 

Phenol " 

Chloroform0 

Phenol c 

Chloroform0 

Aceti c acid c 

Ethylene dichloride 0 

0.4 3 

0.60 

0. 58 

0.37 

0. 45 

1.35 

1.51 

3.40 

0.80 

1.23 

1. 54 

2. 11 

1.92 

2.4 5 

2, 340 

1,730 

2,72 0 

2,23 0 

745 

665 

445 

1,900 

1,230 

479 

350 

384 

301 

a

Solven t is wate r exclept where indicated. 

^Solvent is ethanol. ', 

c

Solvent is benzene. 1 

M 0 r 

H P i 

Pressure Drop Through Pock 

-.11 

ing 

JL. 

5"/77' 

Treybal (1955) state s that pressur e drop data of 

various investigator s varie s widely even for the 

same packing and flow rates. Thes e discrep -

ancie s wer e probably due to difference s in packing density. Moreover , not enough work has 

been done on liquids of high viscosit y for proper 

evaluation, though it is recognized that, at equal 

mass flow rates, high-viscosit y liquids caus e 

greate r ga s pressur e dro p than thos e of low viscosity do. 

Leva's empirica l relatio n applie s below the Load 

point. This is as tollows: 

• 1 

_____ 

Z m. (10- 3

)(10 n L '

/ P L , 9L 

1

 P. 

(97) 

wher e 

P 

Z 

m 

pressur e drop , lb/ft 

packed height of tower, ft 

pressur e dro p constant from Table 62 

n = pressur e dro p constant from Table 62 

L' = superficia l mass liquid velocity, lb / 

hr-ft 2 

G

1

 = superficia l mass ga s velocity, lb/hr -

ft2 

p = liquid density , lb/ft 

p_ = ga s density , lb/ft . 

Illustrative Problem 

The following exampl e illustrate s the precedin g 

principle s of packed towe r design . Knowing the 

amoun t of solut e in the gas stream , the total flow 

rat e of the ga s stream , the most suitabl e solvent, 

an acceptabl e packing, and the desire d degre e of 

absorption , calculat e the towe r dimensions. 

Given: 

Design a packed towe r to remov e 95% of the am -

moni a from a gaseou s mixtur e of 13% by volume 

of ammoni a and 90% b y volume of air. The gas 

mixtur e consists of 80 lb-moles/h r at 68 "F and 

1 atm . Wate r containing no ammoni a is to be 

used a s solvent and the packing will be 1-inch 

Raschi g rings. The towe r will be designe d to 

operat e at 60% of the flood point, and isotherma l 

conditions at 68°F will be assumed . The wate r 

will not be recirculated . 

Problem : 

Determin e wate r flow rate , towe r diameter, 

packed height, and towe r pressur e drop . 
P.13 

Gas Absorption Equipment 217 

Table 62. PRESSURE DROP CONSTANTS FOR TOWER 

PACKING (Treybal, 1955) 

Packing 

Raschig rings 

Berl saddle s 

Intalox saddle s 

Drip-poin t gri d 

tile s 

Nomina l 

. size , 

in. 

1/2 

3/4 

1 

1-1/2 

2 

1/2 

3/4 

1 

1-1/2 

1 

1-1/2 

No. 6146 

Continuous 

flue 

Cross flue 

No. 6295 

Continuous 

flue 

Cross flue 

m 

139 

32. 90 

32. 10 

12.03 

11. 13 

60. 40 

24. 10 

16. 01 

3.01 

12. 44 

5.66 

1. 045 

1. 218 

1.088 

1. 435 

n 

0.00720 

0.00450 

0.00434 

0.00398 

0.00295 

0. 00340 

0. 00295 

0.00295 

0.00225 

0.00277 

0.00225 

0.00214 

0.00227 

0. 00224 

0.00167 

Range of L', 

lb/hr-ft 2 

300 to 8, 600 

1,300 to 10,800 

360 to 27,00 0 

720 to 13,000 

720 to 21,000 

300 to 14,100 

360 to 14,400 

720 to 78,80 0 

720 to 21,60 0 

2, 520 to 14, 400 

2,520 to 14,400 

3,000 to 17,000 

300 to 17,500 

850 to 12,500 

900 to 12, 500 

Range 

of P/Z, 

lb/ft 2-ft 

0 to 2. 6 

0 to 2. 6 

0 to 2. 6 

0 to 2.6 

0 to 2. 6 

0 to 2. 6 

0 to 2. 6 

0 to 2. 6 

0 to 2. 6 

0 to 2. 6 

0 to 2. 6 

0 to 0. 5 

0 to 0. 5 

0 to 0. 5 

0 to 0. 5 

Solution: 

1. Calculate the water rate: : 

a. Equilibrium data for the syste m ammonia -

water ar e as follows: 

X 0.0206 0.0310 0.0407 0.0502 0.0735 0.0962 

Y 0.0153 0.0240 0.032 9 0.0418 0.0660 0.0920 

Plot the equilibrium curve as shown in Fig -

ur e 148: 

The curve is straight approximately to the 

point P, with a slope of about 0. 75. Above 

point P, the slope is variabl e and higher 

than 0.75 . Us e 0.7 5 as the slope , m, of 

the equilibrium curve . 

b. When the temperatur e ris e of the solvent 

is negligible , apply the relation 

G (m) 

m 

L 

m 

= 0.70 

wher e 

m = slope of equilibrium curve = 0. 75 

gas rate = 80 lb-moles/h r 

liquid rate, lb-moles/h r 

^il i = 85 . 81b . moles/hr . 

2. Fro m the given gas flow rate, the calculated 

liquid rate, and the degre e of absorption de -

sired (95% of ammonia), tabulate gas and 

liquid flow rate s at both ends of the tower: 

Density, 

lb-moles/h r lb/hr lb/ft 3 

Inlet gas (bottom) 80 

Outlet gas (top) 72. 4 

Inlet wate r (top) 85. 8 

Outlet liquor (bottom) 93. 4 

2,221 0.0720 

2,092 0.075 0 

1,542 62. 4 

1,671 62. 4 

3. Calculate the towe r diameter : 

a. Us e conditions at top of tower: 

0 . 5 

= 0.0256 
P.14 

213 CONTROL EQUIPMENT FOR GASES AND VAPORS 

0. 09 

0 05 

0.04 

0 

<^ 

J / 

Vs f f4A 

^ A 

r—/\ 

/ 

& 

S<y 

w-

» 

) 

/ 

/ 

0 0 01 0 02 0 113 0 04 0 05 0 06 00 7 0 09 0 09 0 10 

X -- no les NH, mole H20 • 68 °F 

Figure 148. Equilibrium curve for ammoniawater system. 

(2) From Figure 145, the relationship 

•'\7) (0(^(H 2 ) 0 . 2 

(gc )(PG«PL) 

= 0. 19 

(3) From Figure 146, the packing factor!—r) = 160 (tf 

(4) From the relationship in (2), calculate the 

superficial mass gas velocity (G ) at flooding: 

(2) From Figure 145, the relationship 

M (G^W/-2 

(gc)(pQ)(pL) 

(3) From Figure 146, the 

0. 19 

packing factor/-^ \ = lfco 

(4) O' = (0.19)(32.2M0.0720M62.4) = . ^ ^ ^ 

(160)(1) * 

(5) GJ = 0.415(0.60) = 0.249 Ib/sec-ft' 

£

 I 

(6) S = 

2,221 lb/hr 

(0.249 lb/sec-ft )(3, 600 sec/hr) 

2.43 ft' 

c. Select tower with uniform cross-section of 

2.48 ft2 

d. Tower diameter: 

oc . (<^H_.f5 

1.78 ft or 21. 4 in. 

4. Determine the number of transfer units: 

a. Calculate mole fractions of solute in gas 

liquid streams at both ends of the tower. 

' • ( 

0.5 

Q l a M. 19)(32. 2)(0. 0750)(62. 4ft • ^ ... ^ ^ 

(160)(1) 0.2 

(5) At 60% of flooding: 

G' = (G)(0.60) = (0.424)(0.60) = 0. 254 lb/sec-ftf 

(1) Bottom of tower: 

y = — = 0. Ill mole NH./mole air 

X = "£t\ ~ 0. Q88. mole NH3/mole water 

(6) Tower cross section: 

- 2,092 lb/hr 

(0.254 lb/sec-ft )(3, 600 sec/hr) 

= 2. 29 ft 

(2) Top of tower: 

0.4 

*2 • IT -

0.0056 mole NH./mole air 

b. Use conditions at bottom of tower: 

= 0. 0256 

X, = 0 (entering water is NH free) 

b. Plot the operating line from the data in (a) 

on the same graph used for the equilibrium 

curve. 
P.15 

c. By the metho d of Bake r (describe d previously ) 

graphicall y determin e the numbe r of transfe r 

units: 

NTU = 6 

5. Calculat e the height of a transfe r unit: 

a. Ga s transfe r unit: 

0. 5 

• G p / >V. ' 

Hg 

Gas Absorption Equipmen t 219 

3 0 

___2_L/J_c_\ 

L^ ^G D G / 

wher e 

G .

 2 ^211b/

.

h r

 - 396 lb/hr-ft 2 

2. 48 ft 

H = 0.01 

J_J 

/622\ 0-

2 2 

(570)° , 5. = 0.79 ft 

c Overall gas transfe r unit 

G 

H OG H G + m ft) -. 

wher e 

m = slope of equilibrium curv e = 0.75 

G ^ = gas rat e = 30 lb-moles/h r 

L = liquid rat e = 85. 3 lb-moles/hr. 

H _ 1.92 +

 (0

-

7 5 >(30>(0-79 ) . 2 4 7 f t 

OG (35.3) 

L = 

1, 542 

2.4 8 = 622 lb/hr-ft 

a = 7. 00 from Tabl e 58 

|3 = 0. 39 from Tabl e 58 

7 = 0. 58 from Tabl e 58 

(jy...«_ 

H, 

om Tabl e 59. 

(7.00)(896) 0,3 9 (0.66) 0 , 5 

(622) 0'

5 8 

= 1.92 ft 

b . Liquid transfe r unit: 

"^ -(tH*)*" 

wher e 

• = 0. 01 from Table 60 

n = 0. 22 from Table 60 

L = 622 lb/hr-ft 2 

u _ 1 centipoise = 2.4 2 lb/hr-ft 

6. Calculat e the packed towe r height (Z): 

Z - NTU x H o c 

Z = 6 x 2.4 7 = 14.8 ft 

7. Calculat e the towe r pressur e drop 

nL/fl , 

*t =

 mi' 1 0"8

' (-id) 

m-= 570 Trom Table 61 

Ap = 

wher e 

Z 

Ap = pressur e drop , lb/ft 

Z = packe d height = 14.8 ft 

n ij = 32. 10 from Table 62 

n = 0. 00434 from Table 62 

L = 622 lb/hr-ft 2 

fi. = 62. 4 lb/ft 3 

G = 896 lb/hr-ft 2 

P G = Avg gas density = 0. 0736 lb/ft 3. 

(32.ixio- 8mo)( 0-

00434)(622)/62

-

 4 (896) 2

(14.8 ) 

0.0736 

Ap = 57. 2 lb/ft 2 

•fc*. 
P.16 

220* CONTROL EQUIPMENT FOR GASES AND VAPORS 

57.2 lb/ft (1 in. WC) . , . . ,„„ 

Ap =

 s

 .

 L

 = 11.0 in. WC 

5.197 lb/ft 

liquid. The liquid enters at the top of the 

tower, flows across each plate and downward from plate to plate through downspouts. 

PLATE OR TRAY TOWERS 

In. contrast to packed towers, where gas and 

solvent are in continuous contact throughout the 

packed bed, plate towers employ stepwise contact by means of a number of trays or plates 

that are arranged so that the gas is dispersed 

through a layer of liquid on each plate. Each 

plate is more or less a separate stage, and 

the number of plates required is dependent 

upon the difficulty of the mass transfer operation and the degree of separation desired. 

Types of Plates 

The bubble cap plate or tray is most common, 

and most general references deal primarily 

with it when discussing plate towers. Other 

types of plates include perforated trays, 

Turbogrid trays, and Flexitrays. 

A schematic section of a bubble cap tray tower 

is shown in Figure 149. Each plate is equipped 

with openings (vapor risers) surmounted with 

bubble caps. Typical bubble caps are illustrated in Figure 150. The gas rises through 

the tower and passes through the openings in 

the plate and through slots in the periphery 

of the bubble caps, which are submerged in 

3 CAP SUPPORTS 

AT 120°F 

VAPOR RISER 

H010-00WN BAR 

•SLOTS 

Ho 

-CAST TRAT 

CAST CAP 

SHEET METAL CAP 

SHEET METAL TRAT 

Figure ISO. 11 lustratio n of 

some typica l bubble caps. 

s«s our 

SHELL — 

TfUY 

001NSP OUT -

TRAT 

SUPPORT RINQTR»Y 

STIFFIMR 

VAPOR 

RISER -

FR0TH-

/?;^£L#ruauiD IN 

IUIUI CAP 

__mTiRM0im 

"• FEED 

^-LIQUID OUT 

The depth of liquid on the plate, and liquid 

flow patterns across the plate are controlled 

by various weir arrangements, which will be 

discussed in greater detail. 

In perforated plates or sieve trays, the gas 

passes upward through a pattern of holes 

drilled or punched in the trays. Threesixteenth-inch-diameter holes spaced on a 

3/4-inch triangular pitch are commonly used. 

A disadvantage of this type is the tendency of 

liquid to "weep" or leak down through the 

holes instead of through the downspouts at 

low gas velocities. Moreover, the trays 

must be installed perfectly level, or channeling, with resultant loss of efficiency, 

will occur. On the other hand, a perforated 

tray costs only 60 to 70 percent as much as 

a bubble cap plate designed for the same 

throughput. With towers of the same diameter, perforated trays supposedly have 

a capacity 10 to 40 percent greater than 

that of bubble cap plate towers. 

Figure 149. Schematic diagram 

of a bubble-cap tray tover 

(Treybal, 1955, p. 111). 

With Turbogrid trays, licensed by Shell Development Company, the vapor passes up through the 

spaces between parallel rods or bars, and the 

•WR