Stripper Lit Review

THE DESIGN OF SOUR WATER STRIPPERS

Sour water effluents from petroleum refining pro- cesses originate primarily from the use and subsequent condensation of process steam. Condensation of steam usually occurs simultaneously with the condensation of hydrocarbon liquids and in the presence of a hydrocarbon vapor phase containing H2S. Thus the condensed steam contains H2S which imparts the unpleasant odor characteristic of sour water.

Most sour waters also contain NH, derived either from nitrogen in the process feedstocks, or from ammonia injected into fractionation overhead systems to combat corrosion. In addition to H2S and NH,, sour waters may contain significant amounts of other undesirable pollutants such as Phenols, cyanides and carbon dioxide. Fortunately the principal pollutants, H2S and NH,, can be effectively removed by stripping the sour water with steam or flue gas. Stripping the sour water often removes a good part of the Phenols as well.

The proper design of sour water strippers requires two types of data. Firstly, data are required for predicting the H2S and NH3 content of the origin sour water. Secondly, designing the stripper requires data for obtaining H2S and NH, partial pressures above aqueous solutions of HzS and NH3.

The first type of data involves three-phase systems : (1) a vapor phase containing gaseous hydrocarbons, H2S, and NH,, (2) a hydrocarbon liquid phase, and (3) the aqueous sour water phase. These data should relate the vapor phase H2S content to the sour water H2S content in the presence of a hydrocarbon liquid phase.

The second type of data involves only two phases: (1) a vapor phase containing H20, HzS, and NH,, and (2) an aqueous solution of HzS and NH,. There is no hydrocarbon liquid phase in the stripper.

The main purpose of this paper is to develop the data required and to present a tray-by-tray design for a typical sour water stripper which compares quite well with published operating performances of such strippers.

H2S IN SOUR WATER

Data from the API Manual on Disposal of Refinery ~~~~~~~~~~~

by M. R. BEYCHOK, Process Engineering, Fluor (England) Ltd.

(United States)

Wastes are presented in Table I. These and other data are plotted in Figure 1. These data indicate that:

(a) The equilibrium attained is such that the sour water contains excess NH3 as evidenced by the alkaline pH values

(b) The average equilibrium attained is such that the sour water contains NH, and H2S in the molar ratio of 1-5

(c ) The H2S contents of the sour waters are too high to be explained by assuming equilibrium in a pure HzS-pure HzO system (this point is discussed in more detail below).

The source datal for Table I do not include the type of process origin or the temperature or the HzS partial pressure of the sour water process origin. However, it can be assumed that most refinery sources of sour water operate somewhere within the following range :

Temperature About 100F Total pressure 1.0 to 30 psig H2S in vapor phase 0.5 to 10 volume %

Thus the typical H2S partial pressure at the process source would be about 0.1 to 4-4 psia (allowing for 1-0 psia water vapor pressure at 100F).

The API data in Table I indicate that the H2S content of sour waters ranges from 275 to 11,000 ppm. This range cannot be explained by assuming pure HzS- pure H20 equilibrium. Using Henrys Law as being valid at low solute concentrations, the H2S-H20 equilibrium at 100F is approximately :

ppm HzS in solution = (i80)(H2S P.P., psia)

TABLE I

ANALYSES OF REFINERY SOUR WATERS

275-500 1 OO-700 7.5-8.0 1500 1 o00 8.0 5000 5000 - 7200 4800

3000-1 1,000 7500-9000 4000

1600-2500 8000 1876 I020 3000

5000 3600

3500-6000 6100-7000

2000-2300 3000

5000 1480 748

2400

- 8.3 -

8.4-8.8 8.7

8.5-9.0 8-5

9.0-9.5 8-5 8.4

314 Some Special Aspects of Water Conservation and Air Pollution

If the H,S-H,O equilibrium were valid, an H,S partial pressure of 4.4 psia would dissolve only 810 ppm of H,S in the sour water. To dissolve 11,000 ppm of H,S would require a 60 psia H2S partial pressure which is certainly far above the origin H2S partial pressures for the data in Table I.

Assuming that NH, is present and assuming the validity of Henry's Law for the pure NH3-pure H 2 0 system, the equilibrium at 100F is approximately :

ppm NH3 in solution = (38,000)(NH3 P.P., psia)

If the NH3-H20 equilibrium were valid, the dissolved NH, range of 100 to 7000 ppm in Table I would require an NH, partial pressure range of only 0.0026 to 0.18 psia. Hence, the NH3 contents of sour waters can be explained even by very low concentrations of NH3 in the origin vapor phase.

It therefore appears that the sour water origin systems could be better explained by assuming that H2S

has reacted with NH3 and the sour water contains dissolved NH4SH and excess NH3. This would explain the average sour water NH3/H2S molar ratio of 1.5 as obtained from Fig. 1. Such systems have been studied by Van Krevelin and his co-workers2 who presented the equilibrium for the aqueous system :

NH3 + H2S % NHZ + SH- Their work covered the following ranges:

Mols NH3/mol H2S = 1.5 to 6.0 Temperature = 68-140F Total NH3 in solution = 4700 to 59,000 ppm

Therefore, their data cover the range of typical refinery sour waters in Table I. From their work, Van Krevelin and co-workers derived the following :

r.

3 10" -9.089 s H,S P.P., mm Hg = (4s) - 1

Fig. I .

315 Some Special Aspects of Water Conser vation and Air Pollution

where AIS = NH3/H2S molar ratio S = H2S in gm-mols/liter A = NH3 in gm-mols/liter a = 1.63 at 100F

By using the average molar ratio of 1.5 for AIS derived from Fig. 1 and by conversion to more convenient units, the following is obtained :

H2S P.P., psia = ___ 101.63-S/382,000 8 75,000

where : S = H2S in solution, ppm

The equilibrium solubilities of H2S in H20 at 100F (assuming that Henrys Law is valid), and in aqueous NH4SH at 100F for an NH3/H2S molar ratio of 1-5 (assuming the work of Van Krevelin and co-workers is valid) are plotted in Fig. 2. It is apparent from Fig. 2 that the API data in Table I cannot be explained by assuming that the aqueous NH4SH equilibrium is valid. The aqueous NH4SH equilibrium predicts much higher concentrations of H2S in sour waters than are indicated in Table I.

It is doubtful that any single type of equilibrium between the vapor phase H,S and the aqueous phase H2S can be assumed as valid for the system existing in a process origin vessel. The overall equilibrium is probably somewhere between that of an H2S-H20 system and that of an H2S-NH4SH system . . . and is the cumulative result of:

(a) The H2S vapor-aqueous NH4SH equilibrium (b) The H2S vapor-hydrocarbon liquid equilibrium

which would be just below that of H2S-H20 if plotted on Fig. 2)

Fig. 2.

TABLE II

REFINERY DATA ON SOUR WATER HiS CONTENT VERSUS H2S PARTIAL PRESSURE

H2S in H2S H2S fn vapor Total partia[ sour phase pressure pressure water

11.0 16.2 1.67 2960 8.2 26.7 2.1 1 15,600

14.0 50.0 6.85 14,000 0.6 39.5 0 2 3 750 1.75 28.0 0.47 3000 4.41 19.7 0.82 300 3.65 74.7 2.69 3700 3.86 48.7 1.84 6500

(Water partial pressure assumed as 1.0 psia in all cases)

(volume %) ( P . W ( P W ppm wt.)

(c) The distribution equilibrium of dissolved H,S between hydrocarbon liquid and the sour water

( d ) The degree to which the above equilibria are actually attained

(e) The effect of other components such as mercaptans, Phenols, cyanides, chlorides etc.

Table II presents actual refinery data for eight different process origin sour water vessels where H2S vapor phase partial pressures and sour water H2S contents were determined and where hydrocarbon liquid phases were present. These data were obtained from vessels operating at 90-120F. Unfortunately, a literature search failed to disclose other such data. The writer is indebted to Mr L. V. Sorg of the American Oil Co. and to Mr C . Murdoch of the Mobil Refining Co. of South Africa for their help in obtaining most of these data.

The eight data points from Table II are plotted on Fig. 2. A line has been drawn through these points and labeled Sour Water Equilibrium. Admittedly, much additional data of this type would be required to fully prove that this line accurately represents the equilibrium existing in process origin sour water vessels. However, until further data becomes available, this line will be taken as valid for predicting the H2S content of sour water effluents from process vessels operating at about 100F and wherein a hydrocarbon liquid phase exists.

PARTIAL PRESSURES OF H2S AND NH, ABOVE AQUEOUS NH4SH SOLUTIONS

Assuming the effect of components other than H2S and NH, is negligible, the aqueous NH4SH equilibrium data of Van Krevelin and co-workers should be valid for use in designing sour water strippers. Once the sour water has been separated from the hydrocarbon liquid in the origin vessel, then aqueous NH4SH solution data are applicable.


The NH3 and H2S are present in solution as the salt of a weak base and a weak acid. This salt undergoes considerable hydrolysis to reform free NH3 and free H2S which, being gases, tend to escape from solution. The over-all equilibrium can be depicted as :

NH, H,S (Vapor phase) P.P. P.P.

NH,++SH- % NH,+H,S

(Aqueous phase) I I

Van Krevelin et al. started with an inverse form of the usual hydrolysis equation :

( N H m H - ) (free NH,)(free H2S) K =

And they then defined: A = total NH3 present in solution

S = total H2S present in solution) = free NH3 + NH: (both in solution) = free H2S + SH- (both in solution) =approx. SH- (this approximation will be dis-

cussed in more detail)

Fig. 4.

And since the mols of NH: must equal the mols of SH-, and if the approximation that S = SH- is valid :

Fig. 3. (NHi)(S) K =

( A - S)(free H2S)

__ ... . . - .----I

Some Special Aspects of Water Conservation and Air Polluiion 317

Some Special Aspects of Water Conservation and Air Pollution 319

322 Some Special Aspects of Water Conservation ana Air Pollution

And since the free H2S in solution depends on the vapor phase H2S partial pressure tending to keep it in solution the term free H2S was arbitrarily replaced by the term H2S P.P.. Finally, by rearranging:

(3)

The correlation constant K was then evaluated experimentally and it was found that:

1 / K = 10s (4)

Fig. 3 presents values of the temperature dependent correlation factor a as determined by Van Krevelin et al.

Van Krevelin and co-workers also developed and experimentally confirmed the following expression for the NH3 partial pressure :

( 5 )

Fig. 4 presents values of Ho as used by Van Krevelin et al. in evaluating the numerical constants in equation 5 above. The broken-line extrapolations to higher temperatures in both Figs. 3 and 4 were assumed valid by the writer of this paper.

The above equations can be further rearranged and converted to more convenient units :

1 o/a-s382,000 S H2S P.P. (psia) = (1-75 x lo6) (s- 1)

(6)

(1 - $ ) A 10(151 X 10-6)(l-z)4S)A

(8.8 x 105)H0 (7) NH3 P.P. (psia) =

where: A / S = total NH3/total H2S molar ratio SIA = total H,S/total NH3 molar ratio

A =total NH3 in solution, ppm S = total H2S in solution, ppm

HO=Henrys coefficient for NH3 in pure water(see Fig. 4), (gmols/liter)/mm Hg

a = Correlation factor (see Fig. 3), with Van Krevelins original negative signs changed to positive signs for use herein

Equations 6 and 7 were used to obtain Figs. 5a-5e and Figs. 6a-6e from which the partial pressures of H2S and NH3 can be obtained for various values of temperature and various NH3/H2S molar ratios. These ten charts are the basis for the tray-by-tray design of sour water strippers.

Some cautionary discussion regarding equations 6 and 7 is necessary:

(a) The values of Ho used by Van Krevelin et al. may not be accurate absolute values. However, they were used in the experimental evaluation of the numerical constants in equation 7. Hence the Ho values from Fig. 4 must be used or equation 7 will not be valid.

(b) The ratios AIS and SIA are molar ratios whereas the individual terms A and S are in units of ppm wt. Any further rearrangement of equations 6 and 7 to eliminate the ratios must include the fact that (AIS molar) = 2 ( A / S weight).

(c) Equation 6 will indicate an infinite value of H2S partial pressure for an NH3/H2S molar ratio of 1-0, and a negative value of H2S partial pressure if the NH3/H2S molar ratio is less than 1.0. This is obviously incorrect and a consequence of assuming that the total H2S present in solution is in the form of SH- and that therefore S = approx. SH-. At the conditions of Van Krevelins experimental work, all of the NH3/H2S molar ratios were 1.5 or more and the H2S partial pressures were 7.0 psia or less. Therefore the amount of free H2S remaining in solution was indeed negligible and equation 6 was valid. In summary, equation 6 is only valid for solutions containing equal amounts of NHZ and SH- plus free NH3 and a negligible amount of free H2S. Equations 6 and 7 should probably be used only for NH3/H2S molar values of 1.5 or more. This limitation on the work of Van Krevelin and co-workers has not been recognized in other references and texts which discuss their work.

Figs. 5a-5e and 6a-6e can be used to obtain H2S and NH3 partial pressure directly if given their solution concentrations. However, trial-and-error is required if the partial pressures are given and the solution concentrations are sought, since the correct NH3/H2S ratio in solution must be found.

TYPES OF SOUR WATER STRIPPERS

Most refineries have facilities for stripping H2S, and, in some cases, NH3 from sour waters. Various stripping methods are used but most of them involve downward flow of sour water in a trayed or packed tower while an ascending flow of stripping steam or gas removes the H2S and, in some cases, the NH,. The stripping medium may be steam, flue gas or fuel gas. The sour water may or may not be acidified with H2S04 or HC1 prior to stripping. The operating conditions may vary from 1 to 50 psig and from 100 to 270F.

The H2S is less soluble in water than is NH3 and therefore H2S is more readily stripped. For example, in pure water at 100F, the Henrys Law coefficient for


Fig. da.


H2S is 180 ppmlpsia and for NH3 is 38,000 ppm/psia. To efficiently remove about 90% of the NH,, a temperature of 230F or higher is required. Obviously, at the same stripping gas rate, 90% of the H2S could be removed at much lower temperatures.

The use of mineral acids to acidify the sour water fixes the NH3 as NH4C1 or (NH4)$04 so that free NH, formed by hydrolysis is nil and NH, stripping is prevented. This releases the H2S and 90% or more of the H2S can then be stripped out even at 100F.

The use of flue gas for stripping introduces acidic CO2 which tends to fix the NH, as NH4HC03. Al- though flue gas does not fix the NH3 as effectively as does mineral acid, it gives good H2S removal at lower temperatures than would be needed in a conventional steam stripper.

Fuel gas, or any inert gas, can also be used to strip H2S from sour water. However, if NH3 removal is desired, then a temperature of 230F or more is required

to operate at reasonable stripping gas rates. Thus steam must be used to heat the water if NH, removal is wanted. This is true whether the stripping medium is fuel gas, flue gas or additional steam.

The majority of installed sour water strippers use steam as a heat medium and as stripping gas. This paper is limited to the design of such steam strippers. For those who are interested, the comparative per- formance of different stripper types (i.e. steam, mineral acid, flue gas and fuel gas) are discussed in the writers recently published book7.

DESIGN OF STEAM STRIPPERS (WITHOUT USING ACIDIFICATION)

Fortunately the more commonly used steam strippers lend themselves to rigorous tray-by-tray design. Due to the work of Van Krevelin and

TABLE In SOUR WATER STRIPPERS

(no acidification) -+Acid Gar

@JJ-

0 Raw Feed

0 - -0 Wit hou t al Overhead L Condenser

: IL

al

c

I Bottoms

o

With Overhead Condenser

Steam

Raw Feed

la l b 2a 2b 3 4 5 6

Reference [l],p. 76 [il. Fig. 7

[51 i61 [i I. p . 76 [il, p.39 [il, p . 76 [il, p . 39 and [3] and [4]

(1) Raw feed flow, gpm 150 200 60 60 204 35

(2) Reflux flowrate, gpm 32 33 None None 10 3

113 HZS, PPm 10,000 8000 1500 1876 5000 4600 NH3, ppm 5000 5000 1 O00 1480 5000 3800

- 200 226 Temperature, F 200 - (3) Tower feed flow, gpm 182 233 60 60 214 38

Temperature, F 200 200 195 195 240 160 Tower bottom, F 268 268 230 230 250 240

15 (4) Bottoms flow, gpm 195 250 62 62 21 6 41

NH3, ppm 280 200 300 194 600 227 Heating steam, lbs/hr 6500 8300 1100 1100 1130 1600 Stripping steam, Ibs/hr 15,500 16,700 1400 2400 4870 1400

lbs/gal* 1.41 1.2 0.39 0.67 0.38 0.62 (5) Total steam, lbs/hr 22,000 25,000 2500 3500 6000 3000 H2S removal, % 99.4 99.85 99.86 100.0 98.94 99.0 NH3 removal, % 94.0 95.00 69.00 86.5 87.30 93.0 Trays (or packing) 13 13 (12 ft.) (12 ft.) 6 6

- - - Temperature, F - -

-

Tower bottom, psig 30 22 6 6-7 -

His, ppm 50 10 2 O 50 39

* Pounds of steam per gallon of tower feed (i.e. raw feed plus reflux).

89 41 1 O0

3000 7350 1500 9800

11 7.7 140

1 O 0 48.7 150 207 220 240

10 105 52.7

5 Trace 790

3700 1500 4700 4500 0.78 1.54 8400 6000

99.80 100.0 - 89.6

12 12

-

-

-

-


co-workers, the partial pressure data from Figs. 5a-5e and 6a-6e are applicable to such designs.

Table III presents actual operating data for six steam strippers as gathered from the literature. Refinery designers have long used a rule-of-thumb for sour water strippers: Heat the water with steam to its boiling point at the desired operating pressure and provide additional stripping steam of about 0-5 pounds per gallon of tower feed. Table III confirms that this is a fairly good approximation. But if 95f % NH3 removal is desired, it will not provide an adequate design.

The stripping steam rate, rigorously, is a function of: (u) The operating temperature and pressure. (b) The NH, and H2S content of the raw feed. (c) Whether or not refluxing of overhead condensate

is provided. The overhead reflux is highly enriched in NH3 and H2S which increases the required stripping. When refluxing is provided, the choice of reflux accumulator temperature and pressure

~

Fig. 7.

becomes important since it affects the NH3 and H2S content of the reflux.

(d) The number of trays provided and their efficiency. A detailed tray-by-tray design of a steam stripper is

presented in Example 1 in the appendix, using the partial pressure data developed in Figs. 5a-5e and 6a-6e. The calculated results are in excellent agreement with the actual plant performances in Table III. 99% H2S removal and 97 % NH3 removal, from a feed containing 10,000 ppm H2S and 7500 ppm NH,, is obtained at a stripping steam rate of 0-9 lbs/gal of tower feed and with 4 theoretical trays using a non-refluxed design. Assum- ing 50 % tray efficiency, 4 theoretical trays are equivalent to 8 actual trays. The plants in Table III have from 6-13 actual trays. One plant uses 12ft. of packing which is about 6 theoretical trays if an HETP of 2 ft. is assumed.

Fig. 7 summarizes the composition and temperature profiles calculated in Example 1. It is interesting to note that most of the heating steam is condensed on the top tray and hence serves as additional stripping steam for the lower trays. It also appears that most of the H2S is removed on the top tray whereas the NH3 removal is fairly evenly distributed among all the trays. Fig. 7 indicates that additional theoretical trays will accom- plish very little.

If a design similar to Example 1 were made for a refluxed stripper, using a stripping steam rate of 0.9 lbslgal of tower feed (including reflux returned) and using 4 theoretical trays, the calculated gas removals would be lower than in the non-refluxed stripper. Such a comparison is made in the writers book7 by present- ing two such design calculations. The comparison strongly suggests that stripping steam rates should be based on the amount of NH3 in the total tower feed rather than based on the volumetric flow of tower feed.

REMOVAL OF OTHER CONTAMINANTS IN SOUR WATER

There are a few data available7 on the removal of Phenols in sour water strippers :

Type of stripper % phenol removal Steam (non-refluxed) 35 Steam (refluxed) O

Steam (acidified and refluxed) O Flue gas 24

Steam (refluxed) 24 Steam (refluxed) 30 Steam (non-refluxed) 25

Without more data, it would be unwise to assume more than about 20 % phenol removal in asour water stripper. Phenol removal will vary with the temperature and partial pressure conditions within the stripper as well as with the volatility of the particular Phenols present.


There are no data from which to draw any conclu- sions regarding removal of other contaminants such as mercaptans, ammonium chloride, cyanides, carbon dioxide, etc.

SUMMARY

Based on available data, correlations and methods have been presented herein for predicting the H,S and NH3 contents of sour waters, for obtaining H2S and NH3 partial pressures above their aqueous solutions and for the tray-by-tray design of sour water strippers.

The ever-growing emphasis on reduction of aqueous pollution from refineries makes it important to utilize more accurate design methods. It is hoped that this paper contributes toward that end.

The writer is indebted to the publishers of his book7, John Wiley and Sons Ltd., for their permission to extract sections of the book for presentation herein.

Feed H2S Bottoms water = feed + condensed heating steam

= (50,000)(10,000) lop6 = 500 lbs/hr

= 50,000 + 50,000 (30)/950 = 51,580 lbs/hr

Bottoms NH3 = (0.05)(375) = 18.8 lbs/hr = (18~8/51,800) lo6 = 363 ppm

Bottoms H2S = (0.01)(500) = 5.0 lbs/hr = (5.0/5 1,800) lo6 = 97 ppm

Overhead NH3 = (375 - 18.8)/17 = 20.9 mols/hr Overhead H2S = (500 - 5.0)/34 = 14-55 mols/hr

Thus, as a preliminary flow sheet:

14.55

REFERENCES 1. Manual on Disposal of Refinery Wastes, Vol. III, Chemical

Wastes. New York, American Petroleum Institute, 4th edn, 19hn

2. VNKREVELIN, D. W., HOFTIJZER, P. J., and HUNT-

3. GOTHARD. N. J.. and FOWLER. J. A., Indust. Engng. JENS, F. J., Rec. Trav. Chim., 1949, 68, 191-216.

Chem., 1952, 44 (3), 503-507.

Waste Conf., Purdue Univ., 1961, 292-302. 4. PURSELL, W. L., and MILLER, R. B., Proc. 16th Indust.

5. HARRIS, A. J., J. Water Pollution Control Fed., 1963,35 (9),

6. ALBRIGHT, J. C., Petrol. Process., Nov. 1948, 1116. 7. Aqueous Wastes from Petroleum and Petrochemical Plants,

Milton R. Beychok, London, John Wiley and Sons Ltd., 1967.

1154-1 165.

APPENDIX

Example I : Design of sour water stripper (without reflux) A sour water stream of 100 gpm at 100F contains

10,000 ppm H2S and 7500 ppm NH3. Design a steam stripper to remove about 99% of the H2S and about 95 % of the NH3. The tower overhead will be routed to a furnace firebox and must operate at a pressure of about 19 psia.

Assume the use of 8 actual trays at 0.25 psi dP each and design for 21 psia as tower bottoms pressure.

Assume tower bottoms operates at 0.5-1.0"F below boiling point of water at 21 psia (say 230F) due to residual NH3 and H2S partial pressures.

With tower bottoms at 230"F, assume that heat exchange will heat the feed to 200F and cool the bottoms to about 130F. Heating steam, injected with the stripping steam, will supply heat required to operate at 230F tower bottoms temperature.

Preliminary material balance : Feed water = 100 gpm = 50,000 lbs/hr Feed NH3 = (50,000)(7500) = 375 lbs/hr

Check bottoms temperature, using Figs. 5e and 6e:

Bottoms NH3/H2S molar ratio = 7.3 H2S P.P. at 97 ppm H2S = 0.014 psia NH3 P.P. at 363 ppm NH3 = 0.107 psia H20 P.P. at 230F =20.78 psia Total pressure = 20.901 psia (This checks assumed 21 psia and 230F)

~

Calculate top tray temperature : Assume stripping steam rate of 0.9 lbs/gal of feed. Thus, S = 0.9(100)(60) = 5400 lbs/hr = 300 mols/hr

And, overhead vapor : P.P.

NH3 = 20.9 = (20*9/335*5)(19) = 1.18 H2S = 14.55 = (14.55/335*5)(19) = 0.82 HZO = 300.0 = (300/335*5)(19) = 17.00

335.45 19.00

mols/hr (Psi4

__

Some Special Aspects of Water Conservation and Air Pollution

224OF LI

331

v 2

Thus, top tray is at condensation temperature of steam at 17.00 psia P.P., or approx. 219F.

Theoretical tray I : F

I 200'F .=.; 219OF Heating steam condensed on tray 1 = 50,000(19/950)

= 1000 lbs/hr. Hence, heating steam in V2 = 1000 lbs/hr = 55

mols/hr. The overhead vapor, O, is in equilibrium with the

exit liquid, LI. Use Figs. 5c and 6c (the 220F P.P. charts) to find equilibrium liquid composition having NH3 P.P. of 1-18 psia and H2S P.P. of 0.82 psia. Trial- and-error is required until the correct NH3/H2S molar ratio is found for the liquid. As a first trial, choose a ratio higher than in the feed since the ratio increases as the liquid flows downward.

Assume NH3/H2S = 4.0 molar = 2.0 wt. NH3 = 5300 ppm at 1-18 psia (Fig. 6c)

H2S = 5300/2-0 = 2650 ppm = 0.65 psia (Fig. 5c) (This does not check 0.82 psia H2S P.P. in overhead) Assume NH3/H2S = 3-66 molar = 1-83 wt.

NH3 = 5500 ppm at 1-18 psia (Fig. 6c) H2S = 5500/1.83 = 3000 ppm = 0.83 psia (Fig. 5c) (This does check 0.82 psia H2S P.P. in overhead)

Thus, LI composition: 5500 ppm NH3 = 5500(51,000)

3000 ppm H2S = 3000(51,000) = 280 lbs/hr = 16.5 mols/hr

= 153 lbs/hr = 4.5 mols/hr Since V2 = O + LI - F + condensed heating steam, V2 Composition:

P.P. mols/hr (psia)

NH3 = 20.9 -I- 16.5 -22 = 15.4= 0.80 H2S = 14.55 + 4.5 - 14.7= 4*35= 0.226

= 355.0 = 18.474 H20 == 300 + 55 374.75 19.50

~~

Assuming 1 theoretical tray is about 2 actual trays, the total pressure of V2 is 19.50 psia. Partial pressures for V2 components were calculated from mol fractions and the total pressure of 19.50 psia.

Theoretical tray 2:

Tray 2 is at condensation temperature of steam at

Heating steam condensed on trays 1 and 2

Hence, heating steam in V3

18.47 psia, or 224F. Use Figs. 5d and 6d.

= 50,000(24/950) = 1260 lbs/hr

= 1260 lbs/hr = 70 mols/hr

Liquid in equilibrium with vapor, V2: Assume NH3/H2S = 5.40 molar = 2-70 wt. NH3 = 3050 ppm at 0.80 psia (Fig. 6d) H2S = 3050/2.70 = 1130 ppm = 0.225 psia (Fig. 5d) (This does check 0.226 psia H2S P.P. in V . )

Thus, L2 composition: 3050 ppm NH3 = 3050(51,260) lop6 = 156.5 lbs/hr

= 9.2 mols/hr 1130 ppm H2S = 1130(51,260) = 58 lbs/hr

= 1.71 mols/hr Since V3 = V2 + L2 - LI + condensed heating steam,

V3 Composition: P.P.

NH3 15.4 + 9.20 - 16.5 = 8.10 == 0.426 H2S = 4.35 + 1.71 - 4.5 = 1-56 = 0.082 H20 = 300 + 70

mols/hr (psia)

= 370.0 = 19.492 379.66 20.00 ~ _ _ _

Partial pressures for V3 components were calculated from mol fractions and 20.0 psia total pressure.

Theoretical tray 3: 4 I

Steam condensation temperature at 19.49 psia is

Heating steam condensed on trays 1,2 and 3

Assume NH3/H2S = 6.3 molar = 3-15 wt.

H2S = 1600/3.15 = 508 ppm = 0.084 psia (Fig. 5d) (This checks 0.082 psia H2S P.P. in V,)

226F. Use Figs. 5d and 6d.

= 50,000(26/950) = 1370 lbs/hr = 76 mols/hr

NH3 = 1600 ppm at 0.426 psia (Fig. 6d)


_ _ _ ~ 380.33 20.50

Partial pressures for V4 components calculated from mol fractions and 20.5 psia total pressure.

Theoretical tray 4:

Steam k5 - 230F

Tray temperature is 229F. Use Figs 5e and 6e.

Heating steam in V, = 1530 lbs/hr = 85 mols/hr.

Assume NH3/H2S = 7.0 molar = 3.5 wt. NH3 = 185 ppm at 0.054 psia (Fig. 6e)

H2S = 185/3.5 = 53 ppm = 0.009 psia (Fig. se) Assume NH3/H2S = 6.8 molar = 3.4 wt.

NH3 = 690 ppm at 0.200 psia (Fig. 6e) H2S = 690/3.4 = 203 ppm = 0.034 psia (Fig.

Thus, L4 composition: 690 ppm NH3 = 35.6 lbs/hr = 2.09 mols/hr 203 ppm H2S = 10.5 lbs/hr = 0.309 mols/hr

Thus, Bottoms Composition: 5e) 185 ppm NH3 = 9.6 lbs/hr = 2-6 % of feed NH3

53 ppm HzS = 2.8 lbs/hr = 0.6 % of feed H2S Therefore, at a stripping steam rate of 0.9 lbs/gal

of feed, NH3 removal calculates as 97.4% and H2S removal calculates as 99.4 %.

Date post:	16-Nov-2015
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