Ullmann's Enc. of Industrial Chemistry

Sulfuric Acid and Sulfur Trioxide 1

Sulfuric Acid and Sulfur Trioxide HERMANN MüLLER, Lurgi Metallurgie GmbH, Frankfurt/Main, Federal Republic of Germany

1. 2. 2.1. 2.2. 3.

3.1. 3.2.

33. 3.4. 4. 4.1. 4.1.1.

4.1.2. 4.1.3. 4.1.3.1. 4.1.3.2.

4.1.3.3. 4.1.3.4. 4.1.4.

4.1.4.1.

4.1.4.2.

4.1.4.3.

4.1.4.4. 4.1.5.

Introduction ............................. Properties ................................. Physical Properties .................... Chemical Properties .................. Development of the Sulfuric Acid Industry .................................... Early Development .................. Further Development of the Nitro- gen Oxide Process .................... Ascendency of the Contact Process Raw Materials Usage ................ Production ............................... Production by Contact Processes Reaction Kinetics and Thermody- namics ............................................. Catalysts ........................................... Process Summary ........................... Gas Drying ..................................... Catalytic Oxidation of Sulfur Diox- ide ..................................................... Absorption of Sulfur Trioxide . Acid Cooling Practical Versions of the Contact Process Double-Absorption Process Based on Sulfur Combustion Double-Absorption Processes Based on Metallurgical Gases Ordinary Single-Absorption Processes ........................................ Wet-Catalysis Processes ................ Tail-Gas Treatment ........................

2 3 3 7

9 9

9 9

10 11 11

11 13 14 15

16 20 22

23

24

31

34 35 40

4.1.6. 4.2.

4.3.

4.3.1. 4.3.2.

4.3.3. 4.4. 5. 5.1. 5.2. 6.

6.1. 6.2. 6.3. 7. 7.1. 7.2. 7.3.

8. 8.1. 8.2. 8.3. 9. 9.1. 9.2. 10.

Economic Factors ........................... Production by Nitrogen Oxide Processes ................................. Regeneration of Spent Sulfuric Acid ........................................ Introduction ..................................... Reconcentration to 70 —75 % H2 504 ............................................. Concentration to 93 —98 % H2 SO4

Production of Oleum ................ Construction Materials ............. Metallic Materials .................... Nonmetallic Materials ............. Uses of Sulfuric Acid and Eco- nomic Aspects ........................... Indirect Uses ............................. Direct Uses ............................... Economic Aspects .................... Analytical Techniques ................ Concentration Measurement . Measurement of Impurities .. Analysis of Acid-Plant Gas Streams .................................... Sulfur Trioxide ......................... Properties ................................. Manufacture ............................. Handling and Uses .................... Toxicology ............................... Sulfuric Acid ............................. Sulfur Trioxide ......................... References ...............................

44

45

48 48

50 56 59 60 61 62

63 63 63 64 64 64 64

65 65 65 67 67 68 68 68 68

Based on the corresponding article in Ull-mann, 4th ed., written by ULRICH SANDER, UL-RICH ROTHE, and ROLF KOLA (English Edition: British Sulphur Corporation, 1984)

1. Introduction

Of all the heavy industrial chemicals, sulfuric acid is perhaps the most fundamentally impor-tant, in that it has a number of large-scale uses not only within the chemical industry but in other in-dustries as well. By far the most important user

© 2005 Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim 10.1002/14356007.a25_635

is the phosphate fertilizer industry. Other im-portant applications of sulfuric acid are found in petroleum refining, pigment production, steel pickling, nonferrous metals extraction, and the manufacture of explosives, detergents, plastics, and man-made fibers. Many specialty areas of the chemical industry also use varying amounts of sulfuric acid including the production of dyes, pharmaceuticals, and fluorine chemicals.

The consumption of sulfuric acid has many times been cited as an indicator of the general state of a nation's economy, and although many other indicators (such as energy consumption)

2 Sulfuric Acid and Sulfur Trioxide

might today be regarded as more important, sul-furic acid consumption still follows general eco-nomic trends. For example, the recession that resulted from the "energy crisis" of 1974 was clearly reflected in the pattern of sulfuric acid consumption in the following three years. The recession a few years later was similarly accom-panied by a generally declining trend in sulfuric acid consumption starting in mid-1980.

Sulfuric acid is manufactured from sulfur dioxide. The primary raw material for pro-ducing this intermediate is elemental sulfur (-+ Sulfur ; -+ Sulfur Dioxide). A large involun-tary producer of sulfur dioxide is the nonferrous metals industry, the roasting and smelting pro-cesses of which generate off-gases with a suf-ficiently high concentration of sulfur dioxide to permit direct processing to sulfuric acid as a by-product (-+ Sulfur Dioxide, Chap. 5.5.).

A significant number of sulfuric acid facil-ities have been installed in metallurgical plants for the recovery of SO2, mainly for environmen-tal reasons. To an increasing extent the acid pro-duced in such plants is replacing acid formerly obtained from elemental sulfur (or pyrite). An important task for the years ahead is to estab-lish an infrastructure for distributing this "met-allurgical acid" to consumers, especially in the fertilizer industry.

Pyrite still serves in several countries as a raw material for sulfuric acid production, but produc-tion rates are no longer significant compared to the other sources described.

Sulfur dioxide is also produced in the ther-mal decomposition stage of sulfuric acid re-generation from the heavily contaminated sul-furic acid ("spent acid") created by certain industrial processes in which sulfuric acid is introduced(-+Sulfur Dioxide, Chap. 5.6.1.). A further source of sulfur dioxide for sulfuric acid production is the combustion of waste gases containing hydrogen sulfide (-+ Sulfur Dioxide, Chap. 5.7.).

In theory, there is enough sulfur dioxide to supply all the world's sulfuric acid needs in ex-haust gases released during the combustion of sulfur-containing coal, fuel oil, and other fos-sil fuels in power stations and large industrial boilers. Although it is possible to recover such SO2 in a concentrated form by regenerative flue-gas desulfurization processes, or to convert it directly into H2 SO4 , only a minor amount of

sulfuric acid is prepared by this route. Numer-ous flue-gas desulfurization plants installed re-cently as a way of complying with environmen-tal regulations instead convert the SO2 into gyp-sum, primarily for economic reasons. Unless en-vironmental regulations require desulfurization of combustion gases, flue-gas sulfur dioxide is simply dis charged into the atmosphere — where, ironically, it is then transformed by atmospheric processes into sulfuric acid, which ultimately re-turns to earth dissolved in rainfall.

Essentially all sulfuric acid today is manu-factured by the contact process. The correspond-ing technology is very mature, although impor-tant alterations in the detailed arrangement of the conversion and absorption stages were intro-duced commercially in the 1960s to increase the sulfur dioxide conversion efficiency, primarily in the interest of environmental protection. Some plants are now designed for a sulfur dioxide con-version efficiency exceeding 99.8 %. Neverthe-less, the basic principle of the process remains the same today as when it was first introduced in the 1930s.

Modern plants can be designed to be ex-tremely efficient in terms not only of sulfur diox-ide conversion but also energy recovery. It has been common practice for many years to recover 60 % or more of the total energy released in sul-furic acid production in the form of high-pres-sure steam. The energy efficiency of some plants has recently been very substantially increased, however, by providing for additional recovery of low-level heat from the acid system.

In the attempt to comply with requirements for existing metallurgical complexes, sulfuric acid plants have been developed and constructed for treating gases containing as little as 2 % SO 2 (in a single-catalysis plant; 5 % in a double-catalysis plant — see Section 4.1.4). A current ob-jective is to process sulfur dioxide gas streams with the highest SO2 concentrations possible in the interest of reducing capital and operating costs and increasing further the extent of high-temperature energy recovery. At the same time, quality requirements have become more strict for commercial acid. Given the large quantities of "metallurgical acid" now entering the market it will be necessary to develop new processes for eliminating such impurities as may still be present.

V200

1 O n


Yet another important challenge facing sulfu-ric acid makers is the reprocessing of acid wastes from user industries. It is a peculiarity of sulfu-ric acid that very little of the material actually ends up in the products it is used to make. In-deed, apart from small-volume products such as pharmaceuticals, almost the only end products with sulfur values comparable to sulfuric acid are synthetic detergents - and even they are ulti-mately destined for the drain. Many uses result in so-called spent acid: sulfuric acid in vary-ing states of dilution as well as contamination with organic and inorganic impurities. Environ-mental authorities are becoming increasingly in-tolerant of the sometimes rather casual waste-disposal methods of the past, and the industry is under pressure to accept greater amounts of used acid for reconcentration, purification, or regen-eration. Although the technology in this field is already quite extensive (see Section 4.3), further development can be expected in the use of reco-vered heat from sulfuric acid plants.

2. Properties

2.1. Physical Properties

Pure sulfuric acid [7664-93 -9], H2SO4, Mr 98.08, is a colorless, water-white, slightly viscous liquid, mp 10.4 °C, bp 279.6 °C, and d 145 1.8356. It can be mixed with water in any ratio. Aqueous sulfuric acid solutions are de-fined by their H2SO4 content in weight-percent terms Anhydrous (100 %) sulfuric acid is even today sometimes referred to as "monohydrate," which simply means that it is the monohydrate of sulfur trioxide. Sulfuric acid will dissolve any quantity of S03 , forming oleum ("fuming sulfuric acid"). The concentration of oleum is expressed in weight-percent of dissolved S03 ("free S03 ") in 100 % H2SO4.

The physical properties of sulfuric acid and oleum [1-3] are dependent on the H2SO4 and S03 concentrations, the temperature, and the pressure. Figure 1 shows the densities of sulfuric acid and oleum as a function of temperature and concentration [4]. At constant temperature, the density of sulfuric acid increases steeply with rising H2SO4 concentration, reaching a maxi-mum at about 98 %. From there up to a concen-tration of 100 % the density decreases slightly,

but it rises again in the oleum range up to a con-centration of ca. 60 % free S03 .

2 000 .."•4•..„.1.....:11. e LI rTE (% free SCy

N-----=-____1, -it- 0 --- -..... .. -----11). 20--------- i-..

uifi.iric ,, id 1.800 So(à"--... rz H250r„.....

..-e•

Figure 1. Oleum and sulfuric acid density as a function of temperature and concentration [4]

On account of the clear relationship between density and concentration at defined tempera-tures in the lower concentration range, density measurement provides a quick method for deter-mining concentration up to about 95 % H2SO4.

Hydrometers used for this purpose were for-merly calibrated in "degrees Baumé" (°Bé), and for that reason sulfuric acid concentration was often, and sometimes still is, expressed in °Bé. The density p in g/cm3 is given by the expres-sions

[d12] r= 144.3/ (144.3 —° B e ) (Europe)

[c112] r = 145/ (145 —° B e ) (United States)

The electrical conductivity of sulfuric acid at 20 °C as a function of concentration is shown in Figure 2 (see next page). The pe-culiar shape of the curve is due to the vari-ous states of ionic dissociation present in the system H2 0/H2 SO4/S03 at different concen-trations. Conductivity measurement is thus also

0 25 50 free 5C1 3 , eV,. H SO vitt)/

Oleum

Sulfuric acid

I sI

20 1.0 60 6{1 1 00

0 50 100 wt% free E121,wtie..4.

Ely rl

CDriLenkraPiun-

30:.

20 40 60, 80 Concentration, 1..et°4-

t GO


Concentration --

Figure 2. Electrical conductivity of sulfuric acid and oleum at 20°C

Figure 3. Dynamic viscosity of sulfuric acid [5] a) At 25°C; b) At 45°C; c) At 60°C; d) At 80°C

useful as a method for determining the concen-tration of sulfuric acid.

The dynamic viscosity of sulfuric acid as a function of concentration at various tempera-tures is shown in Figure 3 [5].

The various maxima and minima observed on the freezing point curve of sulfuric acid and oleum, shown in Figure 4 [6], are due to the ex-istence of different sulfuric acid hydrates at dif-ferent temperatures. Whereas the 98.0 — 98.5 % acid used for sulfur trioxide absorption in sulfu-ric acid production (Section 4.1.3) solidifies at about — 1 to + 1 °C, commercial 96 % sulfuric acid (66°Bé) solidifies at ca. — 15 °C. Commer-cial 66°Bé acid in the United States contains only 93.2 wt % H2 SO4 and has a freezing point of ca. — 29 °C. This behavior is obviously of im-portance if sulfuric acid is to be stored or trans-ported under very cold conditions.

Figure 5 is a phase diagram for aqueous so-lutions of sulfuric acid based on boiling point measurements by HAASE and REHSE [7]. The

Figure 4. Freezing point curve for sulfuric acid and oleum [6]

Figure 5. Boiling curves for sulfuric acid at 1013 mbar [7] a) Vapor; b) Liquid

lower curve shows the relationship between the boiling point of sulfuric acid and its concen-tration, while the upper curve shows the sul-furic acid concentration in vapor evolved from acid boiling at a particular temperature. When an aqueous solution of sulfuric acid is boiled, the vapor contains more water than the boil-ing acid, so the concentration of the acid in-

- feeS / \ j /

/ //\

/ \'.. nk in SuEfuric 2[1d Dleurm

I , ‘, 90 95 9B 99

1-1 2S0,„wtp.4 5 IO

15 S0 3 . 0%.

lp -s85


creases and its boiling point rises. This contin-ues until the boiling point reaches a maximum of ca. 339 °C at a sulfuric acid concentration of 98.3 wt % H2SO4. At this point the liquid and vapor phases are identical in composition, corre-sponding to an azeotrope, so the concentration of the boiling acid cannot increase further. As can be seen from Figure 5, if vapor in equilibrium with sulfuric acid of 85 wt %, boiling at about 223 °C, is completely condensed it will contain about 7 wt % H2SO4. At concentrations below ca. 75 wt % H2SO4, essentially nothing but wa-ter evaporates. The boiling behavior of sulfuric acid is especially important with respect to in-dustrial processes for thermal concentration of dilute acid. As noted above, the azeotropic con-centration (98.3 wt % H2SO4) represents the ul-timate limiting concentration that can be reached by this method.

The vapor above concentrated sulfuric acid containing more than 98.3 wt % H2SO4 includes not only a greater proportion of H2SO4 than the liquid but also considerable quantities of S03. The vapor over oleum consists almost entirely of S03 . Figure 6 shows the equilibrium vapor pressures of H2O, H2SO4, and S03 above sul-furic acid at 60 °C in the concentration range

from 85 wt % H2SO4 to 15 % S03 oleum, based on measurements made by LUCINSKIJ [8]. These vapor-pressure curves are all-important in gas drying and S03 absorption, essential steps in the production of concentrated sulfuric acid by the contact process (Section 4.1.3).

Apart from the field of sulfuric acid manufac-ture, the system H20/S03/H2SO4 is also of im-mense importance in connection with the com-bustion of sulfur-containing fuels. It is essential to ensure that the temperature of the combus-tion gas does not drop below the dewpoint [9, 10] prior to discharge, since otherwise there is a danger of corrosion by condensing sulfuric acid. Various formulae have been developed [11-13] for calculating dewpoints theoretically as a func-tion of the total gas pressure as well as the H2 O, S03, and H2SO4 partial pressures [14]. Unfortu-nately, the situation is further complicated by the possible formation of various types of hydrated and associated molecules in the gas phase [15].

The specific heat of sulfuric acid falls as con-centration increases. Figure 7 shows the depen-dence of specific heat on concentration and tem-perature [4]. The standard enthalpy offormation for pure liquid H2SO4 is — 8.305 kJ/kg, and the latent heat of evaporation at the boiling point is ca. 605 kJ/kg [16]. Figure 8 (see next page) shows the enthalpies of liquid sulfuric acid and oleum over a wide concentration range at tem-peratures between 0 °C and the boiling point [17], assuming an arbitrary value of 0 kJ/kg as the enthalpy of pure water at 0 0°C. This diagram

Concpntro.hon 100 ar) 300 Temperature,

Figure 6. Equilibrium vapor pressures over sulfuric acid and oleum [8] Figure 7. Specific heat of sulfuric acid [4]


0 50 100 11 50, w free 50 3 , wi %

I-

Figure 8. Enthalpy diagram for sulfuric acid and oleum [17]

provides a simple method for determining the amount of heat liberated when sulfuric acid or oleum is diluted from one concentration to an-other by addition of water.

The amount of heat produced by diluting con-centrated sulfuric acid with water is consider-able, so rapid mixing is important to ensure that local overheating and boiling are avoided. If one wishes to concentrate dilute acid, a correspond-ing amount of heat, the so-called heat of dehy-dration, must be supplied in addition to such heat as may be required to evaporate the water. Fig-ure 9, which has been calculated from enthalpy values, shows the heat of hydration evolved in the dilution of 98.3 % acid to lower concentra-tions at 20 °C (or, conversely, the amount of heat theoretically required to dehydrate it from lower concentrations up to 98.3 %). The heat of hydra-tion liberated by diluting between intermediate concentrations — from 75 % H2SO4 to 25 %, for example — is simply the difference between val-ues read off the curve opposite the appropriate concentrations.

2.2. Chemical Properties

Sulfuric acid is a strong acid with characteristic hygroscopic and oxidizing properties. Sulfu- ric acid, like the sulfate ion, is chemically and

Concentration, w

Figure 9. Heat of dilution or dehydration of sulfuric acid at 20° C

thermally very stable. The dehydrating effect of concentrated sulfuric acid is due to the formation of hydrates. Several hydrates have been identi-fied in the solid state, and these explain the ir-regular variation of some of the physical prop-erties of sulfuric acid with concentration, such as its freezing temperature (see Fig. 4). Known hydrates are H2SO4 • H2 O (corresponding to 84.5 wt % H2SO4); H2SO4 • 2 H2O (71.3 wt % H2SO4); H2SO4 • 3 H2O (64.5 wt % H2SO4); H2 SO4 • 4 H2 O (57.6 wt % H2SO4); and H2SO4 • 6 H20 (47.6 wt % H2SO4) [1].

Pure sulfuric acid is ionized to only a small extent as expressed by Equations (1) and (2) [1].

2 H2 SO4 H3SO4 + HSO:i (1)

2 H2SO4 ,=• H30+ +HS20,7 (2)

This is the reason why the electrical conductiv-ity of a sulfuric acid solution has its lowest value at 100 % H2 SO4 (see Fig. 2). When pure sulfu-ric acid is diluted with water, dissociation occurs increasingly according to Equation (3).

H2SO4 +H20 H30+ +HSO4—

(3 )

The conductivity rises accordingly. Between 92 wt % and 84.5 wt % H2 SO4, the monohydrate (H2 SO4 • H2 O) exists preferentially in equilib-rium with the ionic species, so the conductivity decreases slightly. At lower H2SO4 concentra-tions the extent of dissociation increases, as does therefore the conductivity.

At high water content the second stage of dissociation becomes increasingly immportant (Eq. 4).

HSO:i + H20 ,=s H30+ + (4)


On account of the diminishing total concentra-tion of sulfuric acid, however, the conductiv-ity reaches a maximum at ca. 30 wt % H 2 SO4 (the exact value depends on the temperature), and decreases steeply down to 0 wt % H2 SO4 [1]. Dilute sulfuric acid is the preferred elec-trolyte for industrial metal electrowinning and electroplating plants on account of its high con-ductivity and the chemical stability of the sul-fate ion. To take advantage of the electrical con-ductivity maximum, sulfuric acid of about 33 % concentration is used in lead storage batteries (-+ Batteries, Chap. 4.).

Dilute sulfuric acid is a strong dibasic acid, so it will dissolve all base metals. Hydrogen is released, and the respective metal sulfates and bisulfates (hydrogensulfates) are formed. Bar-ium and lead are exceptions, not because they do not react in the first place but because they become coated with an insoluble sulfate layer that protects them from further attack by the acid [1]. Hot, concentrated sulfuric acid has an oxi-dizing effect, reacting with precious metals and with carbon, phosphorus, and sulfur, by which it is reduced to sulfur dioxide.

A very important property of sulfuric acid is its ability to decompose the salts of most other acids. Examples of industrial importance include:

1) Production of sodium sulfate and hydrogen chloride from sodium chloride (—> Hydro-chloric Acid, Chap. 3.1.; —> Sodium Sulfates, Chap. 1.3.4.)

2) Decomposition of sulfites to sulfur dioxide 3) Decomposition of phosphate rock (natural

calcium phosphates) to phosphoric acid and calcium sulfate (-+ Phosphoric Acid and Phosphates, Chap. 1.2.2.).

The reactions of concentrated sulfuric acid with organic compounds are frequently domi-nated by its oxidizing and hygroscopic proper-ties [1]. Carbohydrates, for example, are decom-po sed to the point of carbonization. Organic con-densation reactions in which water is eliminated are promoted by sulfuric acid because it effec-tively removes the water as soon as it is formed. Sulfuric acid is therefore frequently used in in-dustry for this purpose. It also exercises a cat-alytic effect on certain reactions involving or-ganic compounds.

Sulfuric acid is thermally extremely stable. Only at very high temperatures is it partially de-composed into its anhydride, sulfur trioxide, and water vapor (Eq. 5).

H2SO4 ,=s SO3 +H20 (5 )

The reverse of this reaction is the overall route by which sulfuric acid is formed in the absorption section of a contact sulfuric acid plant. However, it is not possible in practice to manu-facture sul-furic acid by absorbing sulfur trioxide directly into water, because the sulfur trioxide reacts with water vapor in equilibrium with the liquid near the surface, initially forming sulfuric acid vapor. This quickly condenses as a mist of very fine (submicron) droplets, which are practically im-possible to collect. However, sulfuric acid itself reacts readily with sulfur trioxide to form disul-furic acid (Eq. 6), which can be converted back to sulfuric acid by reaction with water (Eq. 7).

H2SO4 +SO3 •■=s 112 S207 (6)

H2S207 +H20 .= 2 H2SO4 (7)

It is therefore quite feasible to absorb sulfur tri-oxide in sulfuric acid of 98 % or higher concen-tration, over which the partial pressure of water vapor is very low, thus avoiding the problem of mist However, because the vapor pressures of both H2 SO4 and S03 increase steeply at concen-trations above 99 % H2 SO4, the optimum sulfur trioxide absorption efficiency is achieved in 98 —99 % acid.

If sulfur trioxide is produced in a gas stream that already contains moisture, gaseous sulfu-ric acid forms progressively by reaction (5), the thermodynamic equilibrium point of which is shifted towards H2SO4 with decreasing temper-ature [18]. The resulting sulfuric acid can then be condensed in a controlled manner without sig-nificant mist production. This route to sulfuric acid is also exploited in industry.

The vapor pressure of sulfur trioxide over oleum or disulfuric acid is appreciable, so when oleum is exposed to ambient air, which always contains moisture, sulfuric acid mists invariably form. It is this property that gives oleum its fa-miliar name of fuming sulfuric acid.

Pure disulfuric acid, H2 S207, which corre-sponds theoretically to oleum with 44.9 wt % free S03, crystallizes at ca. 35 °C, the maximum of the freezing-point curve in the oleum range


(see Fig. 4). Disulfuric acid is partially dissoci-ated in sulfuric acid solution (Eq. 8).

1-12 S 2 07 + H2 SO4 .,=• 113 S0 14

- + HS 2 07 (8)

It is for this reason that the electrical conductiv-ity of oleum rises slightly as the S03 concen-tration is increased from 0 to about 10 wt % free S03 .

Sulfuric acid is oxidized both by hydro-gen peroxide and anodically to diperoxysulfuric acid, H2 S208 , and the unstable monoperoxysul-furic acid (Caro's acid), H2 S05 (Eq. 9).

11202 + 112 SO4 ,= H2S05 + 112 0 (9)

Since it is a strong oxidant, Caro's acid can ox-idize sulfur dioxide to sulfuric acid, a property that has been exploited in pollution control for sulfuric acid plants (the Peracidox process, Sec-tion 4.1.5).

Nitrogen oxides (NO + NO2) react with sul-furic acid at concentrations above 70 wt % H2 SO4 to give nitrosyl hydrogensulfate, NOHSO4 (see Section 4.2).

In 1827, GAY-LUSSAC introduced a method for absorbing nitrogen oxides from the lead-chamber off-gases. With the further develop-ment by GLOVER in 1859 of a method for re-covering nitrogen oxides from the newly formed acid by stripping with incoming hot gases, it became possible to make the nitrogen oxide-catalyzed process continuous.

As early as 1831, PHILLIPS, in Bristol, Eng-land, had patented the oxidation of sulfur diox-ide to sulfur trimdde over a platinum catalyst at high temperature. Nevertheless, it was only af-ter oleum demand for dye manufacture began to increase — from about 1872 onward — that this invention was adopted by industry, and inten-sive development of the contact process began. Better solid catalysts were then sought, and the chemistry and thermodynamics of the S02/S03 equilibrium were investigated. Systematic stud-ies undertaken by KNIETSCH at BASF on the re-action equilibrium of SO2 oxidation over a plat-inum catalyst, published in 1901 [24], formed an important basis for an understanding of ther-modynamic principles.

3. Development of the Sulfuric Acid Industry

Detailed descriptions of the development of sul-furic acid production procedures can be found in the literature [19-23].

3.1. Early Development

In the late Middle Ages, sulfuric acid was ob-tained in small quantities in glass vessels in which sulfur was burned with saltpeter in a moist atmosphere. Higher rates of production first became possible with the introduction of lead chambers as reaction vessels by ROEBUCK,

in Birmingham, England, in 1746. The next ma-jor step forward came in 1793, when CLEMENT

and DESORMES achieved better results by intro-ducing supplemental air into the lead chamber process. They interpreted this as meaning that the nitrous gases were acting only as facilitators of the process, and that the oxidation itself was being effected by the air (i.e., oxygen).

3.2. Further Development of the Nitrogen Oxide Process

The growth in popularity of the contact pro-cess stimulated new competitive efforts to im-prove the lead chamber process. The mo st signif-icant development was replacement of the lead chambers themselves with acid-irrigated towers, which substantially reduced the specific space requirements. The first so-called tower plant was built by OPL in 1907. However, widespread in-dustrial use of the tower process had to await the development of adequate acid pumps. In 1923, PETERSEN introduced an improved tower pro-cess that remaine competitive with the contact process up to the 1950s. However, a fundamen-tal disadvantage of the nitrogen oxide processes is that product concentration is limited to a max-imum of 70 — 75 %, while the contact process produces concentrated (98 %) acid. With the de-velopment of relatively inexpensive vanadium catalysts for the contact process together with in-creasing demand for concentrated sulfuric acid, the proportionate share of world sulfuric acid output produced in nitrogen oxide process plants declined steadily. Around 1910, these accounted


for ca. 80 % of production in Western Europe and North America. By 1930 the figure had shrunk to about 73 %,by 1950 to 20 – 25 %, and by 1960 to ca. 15 % [23]; in 1980 virtually no acid was be-ing produced in nitrogen oxide process plants in these parts of the world. Nonetheless, the nitro-gen oxide process has continued to be the object of interest and a certain amount of development work, especially for the processing of gases with extremely low SO2 content (see Section 4.2).

3.3. Ascendency of the Contact Process

Platinum remained the predominant catalyst for the contact process until the 1930s. As early as 1913, however, BASF was granted a patent [25] for a catalyst based on vanadium pentox-ide, which eventually succeeded in replacing the platinum catalyst because of its insensitivity to catalyst poisons and its considerably lower cost.

In 1936/37 Lurgi introduced the wet con-tact process for converting moist sulfur dioxide-containing gases over a vanadium catalyst. This made it possible to process hot gases from the combustion of hydrogen sulfide in coking plants directly to sulfuric acid.

In succeeding years a number of factors affected the development of the contact pro-cess. First, the raw material basis of the indus-try changed progressively from mainly roaster gases to sulfur combustion gases containing higher concentrations of sulfur dioxide. Second, plant capacities increased as a result of a great rise in the consumption of sulfuric acid by the fertilizer industry. These and other factors pro-vided a stimulus for the introduction of improve-ments in the individual process steps and in the design of associated equipment (e.g., the shift to tray converters from tube converters).

Double Absorption. In 1960, a patent appli-cation was filed by Bayer [26] for the so-called double-catalysis process, and the first plant us-ing this process, built by Lurgi, started up in 1964 [27]. By incorporating a preliminary SO 3 ab-sorption step ahead of the final catalytic stages, the improved contact process permitted a deci-sive increase in overall SO2 conversion, thus re-ducing SO2 emissions substantially. Because the essential difference between this process and the

ordinary contact process is in the number of ab-sorption stages, itis referred to hereafter as the "double-absorption" process.

Environment and Energy. In the 1970s the principal industrial countries introduced more stringent regulations for environmental protec-tion, which made the use of the double-absorp-tion process more or less mandatory in new plants. Nevertheless, the conventional contact process continues to be used in countries where environmental regulations are less exacting.

On account of the steep rise in energy costs, the main thrust of current development in the contact process is toward increasing the recov-ery and utilization of the very substantial amount of process heat. Indeed, a large, modern sulfu-ric acid plant may be looked upon not just as a chemical plant but also as a thermal power plant [28-32].

3.4. Raw Materials Usage

The principal starting material for sulfuric acid production is sulfur dioxide, which can be ob-tained by different methods from various raw materials (- Sulfur Dioxide, Chap. 3.). Moder-ately concentrated sulfuric acid is also produced by reconcentration and purification of so-called spent or waste sulfuric acid. Reprocessing spent acids, which are generated in large quantities in many processes, and recycling the regenerated acid to the user is becoming increasingly im-portant from an environmental protection stand-point especially in the major industrial countries [33].

Up about 1970, pyrite was the predominant raw material, maintaining a 57 – 62 % share of the continually rising total. Thus, in 1962, out of a total of ca. 3.1 x 10 6 t of sulfuric acid, about 1.9 x 106 t (62 %) was made from pyrite; in 1970, about 2.7 x 106 t was pyrite-based, correspond-ing to ca. 61 % of the total production of ca. 4.4x 106 t [34]. Since 1970, the proportionate role of pyrite has been declining relative to the total, which has continued to rise. In 1979, pyrite accounted for only ca. 24 %.

Since the beginning of the 1970s, production of sulfuric acid from elemental sulfur in the Fed-eral Republic of Germany has grown faster than total acid production, in step with rapid growth


in the production of recovered sulfur from the re-fining of crude oil or purification of natural gas from the gas fields of Northern Germany. The proportion of total sulfuric acid production de-rived from elemental sulfur increased from ca. 30 % in 1970 to ca. 50 % in 1979, when total production stood at about 5 x 10 6 t [34].

Sulfuric acid manufacture based on elemental sulfur and pyrite is, of course, relatively sensi-tive to market conditions, because acid produced from these materials represents a primary prod-uct. The same is not true of sulfuric acid produc-tion based on any of the other sulfur-containing raw materials. In those cases, sulfuric acid is a secondary product, manufactured as a means of disposing of waste from another process. The level of production is therefore dictated not by conditions in the sulfuric acid market, but by conditions in the market for the primary prod-uct. Typical sources of this so-called fatal acid are sulfuric acid plants associated with nonfer-rous metal smelters processing sulfide ores. The quantity of fatal acid produced in the Federal Republic of Germany has been rising constantly, reaching ca. 1.3 x 106 t in 1979 (ca. 26 % of the total).

The raw material basis for sulfuric acid pro-duction in other European countries has fol-lowed a pattern generally similar to that in the Federal Republic of Germany. Pyrite remained the dominant raw material until the 1950s, and it was only with the advent of large quantities of recovered sulfur from the Lacq natural gas fields in France, and later from natural gas oper-ations in Canada, that elemental sulfur assumed its current predominance.

In contrast, the United States industry has been based since the early years of this century on elemental sulfur because of development of the Frasch industry in the Gulf states (—> Sulfur, Chap. 5.5.). In 1978, the various raw materials accounted for the following approximate pro-portions of the total U.S. sulfuric acid production of ca. 36 x 106 t [35] in 1963 and of ca. 280x10 6 t in 1993: Elemental sulfur 79% Nonferrous sulfide ores 9% Waste acid 5% Pyrite and hydrogen sulfide 3% Miscellaneous 4%

4. Production

4.1. Production by Contact Processes

4.1.1. Reaction Kinetics and Thermodynamics

In the contact process, a gas mixture containing sulfur dioxide is passed together with oxygen over a catalyst to oxidize the sulfur dioxide to sulfur trioxide (Eq. 10) [36]:

SO2 +1/202 ,=s S03 AH ° = —99.0 kJ (10)

The sulfur trioxide is then absorbed in sulfuric acid where it reacts with added water to form more sulfuric acid (Eq. 11).

S03 (g)+1-120 (1) .= H2SO4 (1) AH ° = —132.5 kJ (11)

The position of equilibrium in the gas-phase exothermic oxidation of sulfur dioxide to sul-fur trioxide (Eq. 10) depends on the prevailing temperature, total pressure, and concentrations (partial pressures) of the reactants. Thermody-namic equilibrium is determined by the equilib-rium constant Kp according to the law of mass action:4

p (SO3) Kp =

p (SO2) p (02) 115

On account of the negative reaction enthalpy of sulfur dioxide oxidation, both Kp and the SO2 equilibrium conversion decrease with ris-ing temperature. The classical relationship bet-ween Kp (in atm –03 ) and temperature was de-veloped empirically by BODENSTEIN and POHL [37], although more recent data [16,38, 39] show deviations.

log Kp = 5186.5

+0.611 logT-6.75 T

where T= temperatue in K

Influences on the Conversion Equilibrium. Any increase in the overall pressure will increase the extent of conversion at equilibrium, because the reaction leads to a decrease in volume.

The maximum possible equilibrium sulfur dioxide conversion at a given temperature T and total pressure p depends upon the SO2 and 02 concentrations of the supplied gases. If the sulfur


dioxide concentration is 2a vol % and the oxy-gen concentration b vol %, the fraction x of sul-fur dioxide oxidized to sulfur trioxide at equilib-rium can be calculated from the following equa-tion based on the law of mass action:

1 100 — ax K p = / 1—xx V b — ax

The appropriate value of Kp is determined from the equation of BODENSTEIN and POHL. The dependence of SO2 equilibrium conversion on temperature and pressure is shown in Figure 10 for the example of sulfur combustion gas con-taining 10 vol % S02 .

80

f 60

o

c

20

0 400 .50 500 550 600 650 70:::

Temperture., —

Figure 10. Theoretical conversion equilibrium in the oxida-tion of SO 2 to S03 as a function of temperature and pressure (feed gas composition: 10 vol % S02, 10.9 vol % 02) a) 10 bar; b) 8 bar; c) 5 bar; d) 1.3 bar

In accordance with the law of mass action, in-creasing the oxygen partial pressure will also in-crease the degree of conversion. However, when air alone is used as the source of oxygen in S02 production, as is usually the case, the oxygen and sulfur dioxide concentrations are inversely re-lated, so the greater the oxygen concentration in the combustion gases the lower will be the sulfur dioxide content. The essential factor determin-ing the attainable SO2 conversion is thus the vol-umetric 02/S02 ratio in the feed gases. Whereas

sulfur dioxide oxidation requires a stoichiomet-ric 02/S02 ratio of only 0.5: 1, in industry it is normal practice to use a ratio of at least 1: 1. The presence of excess oxygen not only raises the SO2 equilibrium conversion but is also an essential prerequisite for maintaining the activ-ity of the vanadium catalyst. There are, however, practical limits on the amount of extra air that can be added, because nitrogen present in the air dilutes the sulfur dioxide to the point where the economics of the process are impaired. Al-though it would be technically feasible to avoid nitrogen dilution by using oxygen instead of air, as is sometimes done in pyrometallurgical pro-cesses that produce high-strength byproduct sul-fur dioxide gas streams, it is usually difficult to justify in a sulfur-burning installation [40].

Another practical expedient for improving the conversion of sulfur dioxide to sulfur triox-ide is to remove, at an intermediate stage in the process, the sulfur trioxide already formed. In a double-absorption type sulfuric acid plant this is accomplished by routing the reaction gases af-ter two of three stages of catalytic conversion through an intermediate absorption stage and then through one or two subsequent catalytic conversion stages. Because of the large (100 %) stoichiometric oxygen excess in the original feed gas and the diminished sulfur dioxide concentra-tion, the 02/S02 ratio at this point is about six times more favorable than at the start.

Influences on Reaction Rate. In an indus-trial plant, actual sulfur dioxide conversion never reaches the theoretical equilibrium value. Gas-phase oxidation of sulfur dioxide is kinetically inhibited, and virtually impossible at any tem-perature without a catalyst. At ordinary temper-atures the reaction is so slow that, in practical terms, it does not occur at all. Increasing the temperature increases the rate of reaction, but si-multaneously shifts the position of equilibrium in an unfavorable way — away from sulfur triox-ide and toward sulfur dioxide and oxygen. With-out a catalyst, the temperature required to make the system react at a practical rate is so high that conversion is very poor. Even with present-day catalysts, a temperature of ca. 400 °C is neces-sary to initiate a self-sustaining reaction.

The reaction mechanism varies depending on the catalyst used. Reaction with a platinum cata-lyst involves heterogeneous gas — solid catalysis

100

90

80

400 5n0 600 : G G


[22, 23, 26]. In contrast, according to present un-derstanding, oxidation over a vanadium catalyst is a homogeneous reaction that takes place in a liquid melt of active components on both the ex-ternal and internai surfaces of an inert solid cata-lyst carrier [41, 42]. The reaction mechanism and the chemical structures of the active components have not yet been clearly defined. According to the model of MARS and MAESSEN [43], reaction in the melt takes place by way of the intermedi-ate steps shown in Equations (12) and (13).

SO2 + 2 V 5+ + 02— ,= 2 V4+ + SO3 (12)

0.5 02 + 2 V4+ —> 2 V5+ + 02— (13)

The validity of kinetic equations derived from this and other reaction models is limited to cer-tain temperature ranges [44]. However, the rate of the catalytic oxidation of sulfur dioxide de-pends not only on the chemical mechanism but also on mass and heat transfer at the gas — liquid interface of the catalyst [42, 45].

Ternperal'ure.

Figure 11. Comparison of (a) theoretical equilibrium SO2 conversion for sulfur-burner gases (10 % S02, 10.9 % 02) with (b) actual SO2 conversion attained over a specific cat-alyst

Other purely technical parameters, such as the gas velocity, gas distribution, and residence time in the catalyst bed, help to determine how closely sulfur dioxide conversion in practice will

approach the theoretical equilibrium. To esti-mate the sulfur dioxide conversion that can be achieved in reality in comparison with the ther-modynamic equilibrium conversion, a correc-tion function is used that takes into account the influences of the individual variables. Figure 11 is a plot against temperature of the theoretical sulfur dioxide equilibrium conversion and a typ-ical observed sulfur dioxide conversion func-tion. Actual conversion characteristics are sig-nificantly influenced by the specific catalyst ac-tivity, which must be determined empirically for each individual catalyst.

4.1.2. Catalysts [177]

Apart from catalyst activity, other factors includ-ing thermal stability, service life, and mechan-ical strength are of practical importance [20— 23, 42, 44-47]. Of all substances tested for cat-alytic activity toward sulfur dioxide oxidation, only vanadium compounds, platinum, and iron oxide have proven to be technically satisfactory. Today, vanadium pentoxide is used almost ex-clusively.

Commercial catalysts contain 4 — 9 wt % vanadium pentoxide, V2 05 , as the active com-ponent, together with alkali-metal sulfate pro-moters. Under operating conditions these form the liquid melt in which the oxidation of sulfur dioxide is thought actually to take place. Potas-sium sulfate is used most often in a K/V mo-lar proportion of ca. 2.5 — 3.5. Some catalysts also contain sodium sulfate to reduce the melt-ing point. The carrier material is silica in the form of diatomaceous earth, silica gel, or zeo-lites, all of which present especially large spe-cific surface areas. Cesium-doped catalysts have also been developed and installed in various fa-cilities. Cesium sulfate as a promoter reduces the melting point of the active components, result-ing in significantly lower temperature limits for sustainable stable activity.

The catalyst components are mixed together to form a paste, which is then usually extruded into solid cylindrical pellets or rings. These are dried and baked at elevated temperature. Other catalyst forms less common in industry include spheres and tablets. Pellet-type catalysts were used almost exclusively until the mid 1980s, but plants today are usually equipped with ring-


shaped (or "star-ring") catalysts. The advantage of a ring-type catalyst is a lower pressure drop; compared to pellet-type catalysts the pressure drop is reduced by more than half. Furthermore, a ring-shaped catalyst is less sensitive to dust blockages. The relationship linking gas velocity with pressure-drop for different catalyst forms is illustrated in Figure 12.

1000

8411 600 544 -400 340

200 - E 15 0

-471

▪ 30 L O_ 20

Figure 12. Pressure drops for various catalyst types per me-ter of catalyst depth (BASF) a) Pellets (6 mm); b) Rings (10/5 mm); c) Star-rings (11/4 mm)

Operating Temperature Range. An impor-tant property of the vanadium catalyst is the low-temperature limit at which stable opera-tion is possible under fixed gas conditions. This temperature is ca. 410 — 430 °C for a conven-tional catalyst and ca. 380 — 390 °C fora cesium-doped catalyst. Low-temperature activity de-pends mainly on the melting point and the chem-ical properties of the mixture of active con-stituents.

The upper operating-temperature limit is de-termined by the thermal stability of the catalyst. Above ca. 600 — 650 °C catalyst activity may be lost irreversibly because of damage to the struc-ture of the carrier and reduction of its internai surface.

Service Life. The average service life quoted by most catalyst producers [46] is about ten years. Service life is generally determined not so much by progressive loss of activity as by catalyst losses incurred when filling and emp-tying the reactor and during routine screening. Depending on the dust load of the gas enter-ing the converter, the size and shape of the cata-lyst grains, and the properties of the active melt, dust will accumulate in the catalyst bed over the course of time. This dust eventually increases the gas-pressure drop through the catalyst bed and reduces both gas throughput and SO2 conversion efficiency. That is the reason why the catalyst must be screened from time to time to remove dust [48]. When catalyst is withdrawn, screened, and returned to the reactor, a certain amount is bound to be lost as a result of abrasion. The pre-cise amount depends on the handling method used and the stability of the catalyst. This loss must be compensated by the addition of new cat-alyst.

In contrast to platinum, vanadium catalyst is largely insensitive to catalyst poisons [49]. Flu-orine compounds in elevated concentrations will attack the carrier material, leading to increased abrasion loss. Chlorine compounds, especially at elevated temperatures, will volatilize the vana-dium and consequently decrease the activity. Ar-senic, which may be present in the feed gases that result from roasting arsenical pyrites, will accu-mulate in the catalyst, but it will only cause an observable decrease in activity if the As203 con-centration in the catalyst mass exceeds 15 wt % [20].

Water vapor in the feed gas is not deleterious to a vanadium catalyst so long as the temperature is sufficiently high to prevent condensation of sulfuric acid. At low temperature (during plant stoppages, for example, or when the catalyst is exposed to humid air) there is a danger that wa-ter will be absorbed by the hygroscopic active constituents, and this can impair the mechanical strength of the catalyst.

4.1.3. Process Summary

There are four main process steps in the produc-tion of sulfuric acid from sulfur dioxide-con-taining gases by the contact process: 1) Gas drying

120

0 60 50 40


2) Catalytic conversion of sulfur dioxide to sul-fur trioxide

3) Absorption of sulfur trioxide 4) Acid cooling

The gas-drying stage is not applicable to a plant of the wet-catalysis type Almost without exception, contact plants operate under essen-tially atmospheric pressure; compression is re-quired only for driving the gases through the plant.

4.1.3.1. Gas Drying

Gas drying is an important process step in con-ventional contact plants (in contrast to wet-catalysis plants). It protects cooler parts of the plant, such as heat exchangers, against corrosion by acid condensation, and it safeguards against the formation of sulfuric acid mist, which can be very difficult to absorb. It also protects the cat-alyst from ill effects of acid condensation when the plant is shut down for any reason. Therefore, both the operating performance (especially tail-gas purity) and the service life of the plant de-pend in large measure on an efficient and reliable gas-drying stage.

In sulfur-burning plants the molten sulfur used is dry from the outset, because its melting point is well above the boiling point of water, and any moisture originally present will have been driven off in the melter. The combustion air, however, must be dried. In the usual arrange-ment, filtered air from the atmosphere is drawn through a drying tower by the main blower.

When the feed gas is derived from smelter waste gases or pyrite roasting, cold, humid, sulfur dioxide-containing gas from the gas-cleaning system (-+ Sulfur Dioxide, Chap. 6.) is mixed with such additional air as may be re-quired to bring the 02/S02 ratio to the opti-mum process value before the gas mix enters the drying section. Similar gas-drying equipment is used in both situations.

The gases are in most cased dried in coun-tercurrent with fairly concentrated sulfuric acid in irrigated packed towers [20-22]. The sulfuric acid is circulated. The residual water content of the gases after drying corresponds theoretically to the partial pressure of water vapor above the drying-tower acid at the prevailing temperature

and concentration. For purposes of achieving a high drying efficiency, the temperature of the irrigation acid is normally maintained at 50 —60 °C.

A substantial amount of heat — not simply the heat of dilution of the sulfuric acid but also the heat of condensation of the water—is liberated in the gas-drying stage. For this reason the cir-culated acid is generally cooled by indirect heat exchange before returning to the dryer.

Water Balance. The sulfuric acid concentra-tion of dryer acid is usually between 93 and 98 %, depending on the production conditions and the plant concept. This level is maintained constant by bleeding off part of the dilute acid leaving the dryer and exchanging it for a corre-sponding amount of concentrated acid (98.5 %) from the absorber circuit. When the desired product-acid strength is 93 — 95 % H2 SO4, prod-uct acid can be taken directly from the drying circuit. In the case of a sulfur-burning plant, pro-cess water may even need to be added to the dry-ing acid circuit to prevent the acid concentration from rising, and it is then not necessary to trans-fer dryer acid back to the absorber circuit. Occa-sionally, 98 % acid is used in the drying tower; this arrangement permits a common pump tank to be used for both dryer and absorption circuits instead of the usual separate pump tanks [50].

Water absorbed by the dryer acid is thus used as process water for sulfuric acid formation be-cause of connections between the dryer and ab-sorption systems. When processing metallurgi-cal off-gases, the water content of the feed gas entering the dryer must be controlled by cool-ing in the gas-cleaning plant so that it does not exceed the stoichiometric requirement for pro-duction of H2 SO4 based on the amount of SO3 to be absorbed. Otherwise, the water balance in the contact plant will not be maintained, and the concentration of circulating acid will drop below the minimum level required for proper plant op-eration.

The permissible water content for feed gases is determined by the concentration of the prod-uct acid, the sulfur dioxide content of the con-tact gases, and the sulfur dioxide conversion ef-ficiency. Figure 13 illustrates how the allowable feed-gas moisture content (represented by the gas temperature at the inlet of the drying tower, to which is is related) varies with sulfur dioxide

Df f eed 12 16

gas, V 0M, -

;"1

2 20 0 I.

S02 content 2 3

60 -

50

4Q


content for two different product-acid strengths. It can be seen that the lower the sulfur dioxide content of the gas, the cooler (i.e., the less moist) the gas must be before it enters the drying tower.

Figure 13. Maximum permissible temperature of moist gas at drying tower inlet of a roaster gas-based double-absorp-tion sulfuric acid plant as a function of gas SO2 content and desired product acid strength a) 93 % H2 SO4 product; b) 98.5 % H2 SO4 product

In the case of metallurgical gases with ex-tremely low SO2 concentrations it is some-times technically difficult and uneconomical to cool the feed gas to the low temperature re-quired to maintain the correct water balance. For such cases Lurgi has developed a predryer-reconcentrator system that has been proven in service [51]. In the predryer, upstream from the main drying tower, surplus water is removed from the metallurgical gas by washing with sul-furic acid of medium concentration (30 — 60 %). This acid does not circulate to the main acid-plant dryer and absorber circuit. Instead, it circu-lates in a closed circuit between the predryer and a reconcentrator immediately downstream from the final S03 absorber in the main acid plant, where itgives up moisture to tail gas released to the atmosphere (see the description of the venturi process in Section 4.3.2). Alternatively, a water refrigeration plant would be required to adjust the gas temperature to the appropriate level.

Drying-Tower Design. The dryers used to-day are, as a general rule, vertical cyclindrical towers. Their steel shells are lined with acid-proof bricks. In the lower part of most such tow-ers a plastic foil (polyisobutylene) is applied bet-ween the steel shell and the bricks in order to prevent acid from penetrating through the bricks to the steel shell. A grate of acid-proof material

supports the packing. Normally, Raschig rings or Intalox saddles of ceramic material are used as packing. The dryer design is determined by a number of interacting parameters, such as the surface character and geometry of the packing, the packing height, the gas velocity, and the ir-rigation rate. The efficiency of moisture absorp-tion by the sulfuric acid depends mainly on the diffusion resistance at the gas — acid interface. The gas flow velocity is best kept well above the region of laminar flow, because this induces turbulence that not only lowers the diffusion re-sistance but also improves distribution of the liq-uid.

A lower limit on the irrigation density is set by the minimum amount of acid required for uniform wetting of the entire packing surface. The maximum level is defined by the flooding limit, which is itself a function of the gas veloc-ity, and must be determined empirically for each packing type.

An important factor in ensuring uniform acid distribution over the entire tower cross-section is the nature of the acid irrigation system. Prefer-ably, this should be located in the upper part of the packing layer, and it should consist of cast iron or stainless steel nozzle tubes with lateral oblique openings directed upward. As a general rule, a wire-mesh filter of plastic or stainless steel is installed above the packing layer to sep-arate entrained acid droplets.

Normally, a residual moisture content of 50 mg/m3 in the dry gas is considered satisfac-tory.

4.1.3.2. Catalytic Oxidation of Sulfur Dioxide

The reactor in which sulfur dioxide is oxidized catalytically to sulfur trioxide is known as the converter. It is the heart of the sulfuric acid plant.

In the design and construction of any con-verter intended to assure maximum sulfur diox-ide conversion, proper attention to removal of the very considerable reaction heat is of vital im-portance. The reaction is generally carried out under adiabatic conditions, so the temperature of the solid catalyst bed rises, thereby determin-ing, and at the same time limiting, the attainable level of SO2 conversion consistent with thermo-dynamic equilibrium (see Fig. 10). To achieve a

100

80

20

10

0 400 1450 500 5.5.4 600 650

Cal alysIP bed outlet tempereure, 700


high final SO2 conversion, the total catalyst mass is divided up into several catalyst beds (trays), and hot gas leaving each bed is cooled to the min-imum working temperature of the catalyst before it enters the next bed. Tubular converters oper-ating at nearly isothermal conditions have now almost completely fallen out of use in sulfuric acid plants. Isothermal fluidized bed converters are under development, however, and some have reached the point of industrial use.

Figure 14. Comparison of reaction profiles and SO2 conver-sion for 4-bed normal contact (single absorption) and (2 + 2) double-absorption processes (feed gas: 8.5 vol % S02) a) Double-absorption process equilibrium curve after inter-mediate absorption; b) Equilibrium curve for normal contact process; c) Adiabatic reaction in bed 1; d) Adiabatic reaction in bed 2; e) Cooling and intermediate absorption; f) Cooling; g) Bed 3; h) Bed 4

Figure 14 shows the reaction profile fora con-verter comprising four beds and operating under adiabatic conditions, together with the SO2 con-version attainable in each bed within the gen-eral limits of actual conversion characteristics. A profile for the normal contact process with-out intermediate absorption (single absorption) is compared with that for a double-absorption process with intermediate absorption after the second bed. It is evident that the overall conver-sion ultimately obtained in the double-absorp-tion process is substantially greater than in a

single-absorption process with the same number of catalyst beds. In the single-absorption pro-cess, the maximum achievable SO 2 conversion with a typical four-bed converter is ca. 98 %. (The exact figure depends on the feed-gas com-position.) This contrasts with a final SO 2 conver-sion > 99.5 % in the double-absorption process.

The normal configuration for a double-ab-sorption plant using a fixed-bed converter is ei-ther (2 + 2) or (3 + 1). That is to say, the inter-mediate S03 absorber can be placed after either the second or the third bed. If an SO2 conver-sion efficiency > 99.7 % must be guaranteed, a five-bed converter in a (3 + 2) or (4 + 1) config-uration may be preferred [52]. Further improve-ments in conversion efficiency, especially in a single-catalysis plant, can be achieved with ce-sium doping (see Section 4.1.2). The resulting considerable increase in low-temperature activ-ity permits a conversion efficiency of ca. 99 % in a single-catalysis plant and > 99.7 % in a double-catalysis plans. However, the applicabil-ity of such catalysts is limited by high cost.

The double-absorption method has another significant advantage as well, in that it can pro-cess feed gases with a higher sulfur dioxide content and correspondingly lower 0 2/S02 ra-tio than the single-absorption process. On ac-count of a lower specific gas-flow rate fora given nominal H2SO4 production capacity, equipment for the double-absorption process can also be smaller.

Converter Design. The process design for a converter [22, 45] requires careful optimization because of the large number of interacting pa-rameters. It has been considerably simplified by the development of appropriate computer pro-grams [53-55]. The most important variables are the sulfur dioxide concentration and gas-flow rate, the number of beds, the specific catalyst quantity and its distribution between the individ-ual beds, the gas-pressure drop, and the gas-inlet temperatures at the individual beds. In arriving at a final design, due consideration must be given to the relationship between equipment costs and energy costs.

The specific catalyst quantity required for production of 1 t/d of sulfuric acid is ca. 200 —260 L fora "normal" contact (single-absorption) plant and about 150 — 200 L fora double-absorp-tion plant. The distribution of the catalyst bet-


ween the individual beds can vary widely as a function of the gas concentration and the way the gas is routed. The catalyst bed height may vary from ca. 200 mm to 1000 mm Different catalysts may also be used in different beds; the choice depends on the thermal stresses prevail-ing at each stage. The preferred grain size de-pends on pressure-drop considerations and the permissible dust load.

The optimum method of cooling the reac-tion gas between the catalyst beds is a func-tion of the composition and initial temperature of the feed gas, and thus its origin. It is nor-mally preferable to use indirect heat exchange. In a "cold-gas" plant running on off-gases from a metallurgical or spent-acid decomposition plant, virtually all the surplus reaction heat is trans-ferred to the feed gas in a series of gas — gas heat exchangers to raise the gas temperature to the reaction temperature of the first catalyst bed. Modem plants processing gases with higher SO 2 loads (> 8 vol %) are equipped with boiler ele-ments as a way of removing excess reaction heat. This improves the overall energy efficiency and maintains the gas temperature at the inlet of the absorption tower within reasonable limits. In a sulfur-burning plant, where the feed gas is al-ready hot, surplus reaction heat from the cat-alytic section is recovered in steam generators and boiler feed-water preheaters (economizers). In all double-absorption plants, however, gas passing from the converter to the intermediate absorber and back is cooled and reheated in in-direct gas — gas heat exchangers. Direct injection of cold feed gas or quench air may be employed to a limited extent under certain circumstances, either after the first catalyst bed for the purpose of limiting the outlet gas temperature and thus the thermal stress on the first heat exchanger or steam generator, or in later stages if the amount of available heat is insufficient to warrant provi-sion of an indirect heat exchanger or feed-water preheater.

The idea of mounting at least some of the heat exchangers inside or around the converter shell has become attractive once again even though it was abandoned in the 1970s and 1980s. Convert-ers fabricated entirely of stainless steel provide sufficient flexibility from a design and construc-tion standpoint to permit such heat exchangers to be incorporated into the converter vessel. The advantage of this arrangement is of course elim-

ination of some of the gas ducts, including ex-pansion joints, supports, etc. However, no gen-eral rule is available to cover all types of plants; the design engineer must consider carefully the relative advantages of fewer gas ducts versus a more complicated converter vessel.

Normally, the converter is designed as a ver-tical cylindrical vessel, with the catalyst beds mounted above one another in separate, virtu-ally gas-tight compartments. Gases enter and leave through lateral nozzles, passing upward and downward through the beds.

The catalyst mass is supported by a metal-lic bed grate. A base layer of ceramic packing prevents direct contact between the catalyst and the grate, which could give rise to corrosion. The catalyst bed is itself covered with another layer of packing. This serves the dual purpose of helping to ensure uniform gas and temperature distribution over the surface of the catalyst and preventing the catalyst from blowing around and leaving cavities in the bed, which might also ad-versely affect gas and temperature distribution in the main body of the catalyst bed. The con-verters are usually lined with ceramic material in the catalyst areas.

Three principal types of converter are in use: stainless steel (with or without heat exchangers), steel, and brick-lined [56].

The stainless-steel converter (Fig. 15, see next page), consisting of the shell, the tray sep-arators, and the trays, is fabricated entirely of stainless steel or heat-resistant steel. A central tube is often used to support the trays. Heat ex-changers (also made of stainless steel) can either be incorporated into the central tube or arranged around the shell.

The steel converter (Fig. 16, see next page) corresponds in principle to the stainless-steel converter, but boiler plate or equivalent is used as the construction material for the shell, the sepa-rations, and trays for those layers with lower op-erating temperatures. A steel converter is more economical compared with a stainless-steel con-verter and is appropriate for treatment of gases with lower SO2 loads.

The brick- lined converter (Fig. 17, see fol-lowing page) is completely lined internally with acid-proof bricks. The compartment separators are self-supporting domed structures made of


Outlet lbed 1}

Figure 15. Lurgi converter in stainless steel with integrated heat exchangers shaped bricks. These carry brick columns that support the catalyst-bed grates. The brick-lined converter represents a conservative design guar-anteeing long lifetime together with a high "ther-mal inertia." This facilitates operation with fluc-tuating gas loads as well as start-up after idle periods.

Figure 16. Four-bed all-steel converter a) Inlet (bed 1); b) Quench inlet; c) Inlet (bed 2); d) Inlet (bed 3); e) Inlet (bed 4); f) Catalyst; g) Outlet (bed 2); h) Out-let (bed 3); i) Outlet (bed 4); j) Central bed support tube

The cast-iron converter that for many years constituted the standard design – specifically in the United States – has now been practically eliminated from sulfuric acid plants.

Tubular converters (—> Tubular Reactors), which were in common use in sulfuric acid plants until the 1960s, have the inherent advan-tage over bed converters that reaction conditions can be made almost isothermal, but they are quite unsuitable for use in today's large-capacity acid plants on account of several seri-ous disadvan-tages. In particular, in order to ensure that heat is dissipated efficiently from the catalyst it is nec-essary to use a large number of narrow-diameter tubes, even for small plant capacities. It is diffi-cult to change the catalyst charge in such tubes, gas distribution is difficult to control, and the risk of corrosion is increased because catalyst is in direct contact with the tube material.


Figure 17. Four-bed brick-lined converter A) Side elevation; B) Plan a) Inlet (bed 1); b) Catalyst bed grate; c) Compartment sep-arator; d) Outlet (bed 2); e) Inlet (bed 3); f) Outlet (bed 4); g) Outlet (bed 1); h) Inlet (bed 2); i) Brick supporting col-umn; j) Outlet (bed 3); k) Inlet (bed 4); 1) Catalyst grate; m) Brick columns

These shortcomings are not shared by the fluidized -bed converter, which also operates un-der virtually isothermal conditions. The good heat-transfer characteristics of a fluidized bed are well known (—> Fluidized-Bed Reactors). It is possible to conduct sulfur dioxide oxidation

in this way at a constant temperature in the op-timum operating range of the catalyst. The re-leased heat of reaction is removed from the flu-idized bed by immersed tube bundles that gener-ate steam. One great advantage of the fluidized-bed converter is that cold feed gases from met-allurgical processes or other sources need not be preheated to the reaction temperature; instead the conditions in the fluidized bed are simply adjusted to provide any heat necessary to the in-coming gas. A further attraction of this type of converter is that it can process feed gases with much higher sulfur dioxide concentrations rela-tive to a fixed-bed converter [57].

Because of the favorable thermodynamic conditions in a fluidized-bed converter, the con-version attainable in a single stage may equal that achieved in the first three beds of a con-ventional fixed-bed converter. The fluidized-bed converter is therefore ideally suited for use ahead of the intermediate S03-absorption stage in a double-absorption plant. However, itdoes re-quire an extremely abrasion-resistant catalyst [58]. Bayer AG has developed a spherical cat-alyst for this purpose [59], and has used it in two fluidized-bed converter plants since 1971 and 1976, respectively, apparently with success [60].

4.1.3.3. Absorption of Sulfur Trioxide

Sulfur trioxide formed by the catalytic oxidation of sulfur dioxide is absorbed in sulfuric acid of at least 98 % concentration, in which it reacts with existing or added water to form more sul-furic acid [20-23]. The optimum concentration of the absorber acid corresponds to the azeotrope (see Section 2.1), where the partial pressures of S03, H2SO4, and water vapor are all at a min-imum. At lower acid concentrations the water-vapor partial pressure is higher, and there is a correspondingly greater risk that sulfuric acid mist will form as a result of direct reaction of sulfur trioxide in the gas phase with water va-por above the acid. At acid concentrations above the azeotropic point the tail gas will contain in-creased amounts of sulfur trioxide and sulfuric acid on account of their higher partial pressures.

In the original contact (single-absorption) process, process gas passes through an the con-verter beds before the sulfur trioxide is absorbed


in a single absorption unit. In the double-ab-sorption modification, now routinely practiced in both new and "revamped" plants, most of the sulfur trioxide is removed from the process gas in an extra absorption step at an intermediate stage. As explained above, this may occur after either the second or the third converter bed. In the (2 + 2) configuration, ca. 85 % of the original sul-fur dioxide content of the feed gas has been con-verted to sulfur trioxide by the time the gas en-ters the intermediate absorber, and in the (3 + 1) configuration even more of the total amount of sulfur trioxide (ca. 93 —95 %) is removed by the intermediate absorber. Sulfur trioxide formed in the last bed or beds from the small quantity of residual sulfur dioxide is absorbed in the final absorber.

Most absorbers are packed towers, usually operating in countercurrent. The gases pass from bottom to top through a bed of packing, which is uniformly irrigated from the top with concen-trated sulfuric acid. Process gas leaving the con-verter system is cooled by a gas — gas heat ex-changer or a steam generator, preferably in con-junction with a feed-water preheater, to a tem-perature of ca. 180 — 220 °C before entering the absorber. It is essential that the wall temperature in the gas coolers never drops below the acid dewpoint (ca. 110 —160 °C, depending on the gas composition); otherwise, there is an acute danger of corrosion due to condensing acid, as well as mist formation. Gas entering the absorber is therefore not completely cold, and it releases heat to the absorber acid as it passes through the absorber; by the time it reaches the outlet it is at virtually the same temperature as the incoming absorber acid.

A substantial amount of heat is also generated in the absorber acid from absorption of sulfur tri-oxide and formation of sulfuric acid, and the acid temperature rises in consequence by an extent that depends on the acid-circulation rate Effi-cient sulfur trioxide absorption depends not only on uniform acid and gas distribution in the ab-sorber but also on ensuring that the temperature and concentration of the absorber acid remain at the optimum values. The acid concentration is held constant by adding process water or dryer acid to acid leaving the absorber, at a rate con-trolled by a device that measures the electrical conductivity or density. The optimum acid-inlet temperature depends on design conditions, but

it is ca. 60 — 80 °C in most plants, maintained at that level by indirect cooling. The attainable SO 3 absorption efficiency is generally > 99.9 %.

Irrigated packed-tower absorbers are not effi-cient at removing sulfuric acid mist [1]. In spite of efficient gas drying and optimum conditions for sulfur trioxide absorption, it is often impos-sible to prevent mist formation completely, es-pecially when processing high-bitumen sulfur or metallurgical feed gases with elevated hydrocar-bon contents [61], or when starting up or shutting down the plant. Furthermore, in plants with an oleum tower upstream from the intermediate ab-sorber, sulfuric acid mists may form with aero sol particle sizes in the submicron range owing to overcooling of the process gas. In such cases the mists may already have formed in the in-termediate absorber. Ordinary wire-mesh spray separators have no effect on these mists, so spe-cial mist eliminators must be installed to avoid corrosion in the downstream heat exchangers and further mist formation in the final absorber. There are various designs for mist eliminators, and not all operate on the same principle. The most appropriate type depends on the nature of the mist, especially its particle size. For exam-ple, impingement separators are best for trap-ping particles above 1— 3 gm, while submicron particles are more efficiently trapped by diffu-sion on Brownian-motion separators [61-65]. Glass fibers of varying degrees of fineness are preferentially used as the filter material.

The design principle underlying conventional absorption towers is similar to that for drying towers. Such towers have welded cylindrical steel shells lined on the inside with acid-proof bricks and silica-based, acid-proof mortar. Of-ten the lower part of the tower around the acid sump is further protected by polytetrafluoroeth-ylene (PTFE) sheeting sandwiched between the steel shell and the multilayer brick lining As with the dryer, it is essential to ensure that gas and acid distributions are uniform over the en-tire tower cross-section, and that the acid flow rate is sufficiently high to wet the entire packing layer completely. Depending on gas conditions, the packing layer may have a height of 4 — 6 m.

The development of special stainless steels over the past few years has permitted the in-stallation of steel absorption towers without any brick lining. This is a particularly great advan-tage when towers must be replaced during main-


tenance shut-downs. The absence of a brick lin-ing means that a new tower can be completely prefabricated and set onto the existing founda-tions within a very short period of time

Sulfur trioxide absorption systems based on venturi scrubbers were introduced by Lurgi in the early 1970s. These were developed specif-ically for the purpose of reducing heat losses during intermediate absorption so as to permit low-grade smelter off-gases to be processed au-tothermally in a double-absorption plant [40]. Because the gas and absorber acid flow in cocur-rent in a venturi system, the gas temperature at the absorber outlet is higher than with a con-ventional countercurrent absorber based on acid at the same inlet temperature. Thus, less heat is transferred from the gas to the absorber acid for dissipation through the acid-cooling system. Running the absorber at a higher acid temper-ature reduces gas-heat losses still further. This principle of hot absorption in cocurrent with acid at a temperature of about 120 —140 °C is useful not only for the purpose of processing low-grade feed gases but also for maximizing heat recovery generally in a double-absorption plant.

The venturi absorption system consists of vertical and horizontal venturi scrubbers ar-ranged in series. In each case, acid is injected through a nozzle into the gas inlet of the ven-turi unit. Upon leaving the venturi units the gas passes through a packed tower where droplets entrained by the gases are separated.

The intensive mixing of injected acid with the turbulent gas stream in a venturi unit provides a large liquid — gas interface and, consequently, fa-vorable mass- and heat-transfer conditions. This type of venturi system has given satisfactory ser-vice in many sulfuric acid plants, including units with H2 SO4 capacities as great as 2000 t/d.

4.1.3.4. Acid Cooling

The acid cooling system plays a vital role in de-termining the efficiency and operating safety of an entire sulfuric acid plant. Appropriate choice of a particular cooling system depends not only on the acid temperature but even more on the availability and quality of the cooling water, as well as the cost of water in relation to the cost of energy.

The introduction of special stainless steels led to significant changes in the field of sulfuric acid cooling beginning in about 1980. Acid cooling in modern plants is dominated by two types of coolers:

1) Shell-and-tube coolers, either equipped with anodic protection or fabricated from special stainless steel that requires no such protec-tion

2) Plate-type coolers

Intermediate closed-loop water circuits are also being installed more frequently both for heat recovery purposes and as a way of com-plying with environmental regulations (espe-cially with respect to sea water). A further step has been the development of systems generat-ing steam from heat released in the absorption system. A brief description of such systems is provided in Section 4.1.4.

Cast-iron cascade coolers, which were stan-dard equipment for many decades, have been al-most completely eliminated from modern acid plants; even in existing installations these cool-ers are in most cases being replaced. Decreas-ing reliability, poor quality, environmental reg-ulations, and of course better alternatives have forced this development. The use of spiral cool-ers, tank coil coolers, etc., fabricated from "stan-dard stainless" is limited to special applications. Air coolers occupy a share of the market in situ-ations where the consumption of cooling water is restricted.

Shell -and-Tube Coolers. Shell -and-tube coolers offer many advantages, includ-ing ease of installation, compact design with correspondingly low specific space re-quirements, and good heat-transfer coef-ficients (ca. 800 —1400 Win-2 K- 1 , 700 —1200 kcal111-2 h-1 K-1, depending on the de-sign conditions and the mode of construction). Shell-and-tube (s + t) coolers made of stainless steel have been used in sulfuric acid plants since the mid-1960s when the introduction of anodic protection made this practical. At present, two types of s + t coolers are preferred:

1) Coolers fabricated from "standard" stainless steels and equipped with anodic protection systems


2) Coolers fabricated from special stainless steels (Sandvik SX, 1.4575, etc.), which do not require anodic protection

These coolers make it possible to cover the com-plete range of operating parameters normally en-countered in drying and absorption systems. Sea water, brackish water, cooling-tower water, and closed-loop water can all be used as the cooling medium, provided appropriate materials are se-lected. In special cases, shell-and-tube coolers are made from other materials — Teflon or glass, for example — especially when aggressive media are being heated by the hot absorber acid, such as dilute or chemically contaminated spent sulfuric acid.

Plate Coolers. Ever since the development of adequate acid-proof gaskets, plate coolers have been increasingly called upon for sul-furic acid cooling duty. The sealing materi-als used today are elastomers such as Vi-ton (a copolymer of vinyl chloride and hexa-fluoropropene; —> Fluorine Polymers, Organic, Chap. 3.2.), which can handle acid at tem-peratures up to a maximum of ca. 110 °C. The special advantages of plate heat exchang-ers are extremely compact design, good ac-cessibility, and easy maintenance and clean-ing [66,67]. They can be built with very thin walls, so the specific material requirement for the heat exchange area is very low, and the heat-transfer coefficient is very high (up to ca. 2300 W m-2 K-1 ). Therefore, even expensive special materials can be used economically. Hastelloy C276 (DIN 2.4819) (-+ Construction Materials in Chemical Industry, Chap. 5.3.3.), for example, has proved to be very satisfactory in plate heat exchangers designed for sulfuric acid cooling with brackish or sea water, although the water must be filtered carefully to avoid de-posits in the plate cooler. For cases in which higher temperatures are required (e.g., for heat recovery purposes), an alternative type of plate cooler has been developed based on semiwelded plates. Pairs of welded plates form the sulfuric acid channels in these coolers, while plates with normal gaskets serve as the water channels.

Air Coolers. Air coolers are preferentially used in sulfuric acid plants when cooling water is not available in adequate quantity or at reason-able cost. The acid flows through tubes arranged

horizontally in flat bundles, which are equipped with external fins to improve heat transfer. The cooling air is forced or sucked past the tubes by means of a fan. The tubes are usually made of stainless steel. The acid temperature is limited to ca. 80 °C unless the cooler is equipped with anodic protection, in which case higher temper-atures can be tolerated [68]. Alternatively, if the cooler is equipped with acid-proof centrifugally cast tubes, acid temperatures up to about 110 °C can be handled.

Air coolers are relatively expensive to install, and a considerable amount of power is also re-quired to drive the fans. However, in any cost comparison with water-cooled types the capi-tal and operating costs of not only the acid-heat exchanger itself but also the requisite water-cooling system and water-treatment unit must be taken into account.

4.1.4. Practical Versions of the Contact Process

Contact sulfuric acid plants vary in a number of respects depending upon the raw material used to produce the sulfur dioxide-containing contact gases. The broadest division is into hot-gas and cold-gas plants. In hot-gas plants, which are usu-ally based on the combustion of elemental sul-fur, the hot feed gas from the sulfur furnace is cooled just to the required converter-inlet tem-perature. Cold-gas plants, based on metallurgi-cal or decomposition gases, are so designated because the crude sulfur dioxide-containing gas must be cooled to low temperatures in the de-dusting and cleaning systems before being intro-duced into the sulfuric acid plant. It is therefore necessary to reheat the cold feed gas to the req-uisite converter-inlet temperature with reaction heat from the converter system.

Cold-gas plants are sometimes confusingly described as "wet-gas" plants because cold feed gas leaving the gas-cleaning system is saturated with water vapor. However, this water vapor is removed by a dryer before the feed gas is pre-heated, so in all the remaining process steps the gas is effectively dry. By contrast, in a true wet-contact or wet-catalysis process, moist feed gas (produced, for example, by the combustion of hydrogen sulfide) is processed directly in the

feed warer


converter without preliminary drying. There-fore, the term "wet" is best avoided when re-ferring to a cold-gas plant.

These traditional contact-plant classes can be further subdivided into the double-absorption process (also referred to as the double-contact or double-catalysis process), which is the type of process now most commonly used in new plants, and the older, ordinary contact process, without intermediate absorption (also referred to as the single-catalysis or single-contact process) [69-76]. The ordinary contact process is still used where the feed gases are very low in SO 2 and where permitted by local pollution regulations. Within a given overall process concept, many variations in detailed design are possible [75-79].

4.1.4.1. Double-Absorption Process Based on Sulfur Combustion

The sulfur-burning double-absorption process is considered to be the standard sulfuric acid pro-duction process for conforming with sulfur diox-ide emission limits now in force in most coun-tries [52]. Heat is released in ail steps. Whereas

it has long been standard practice to utilize reac-tion heat from the converter for producing high-pressure steam, more recently it has become of economic interest also to recover low-level heat from the absorber and dryer-acid systems.

Process Description. Figure 18 is a flow dia-gram of a typical sulfur-burning double-absorp-tion plant with a four-bed converter in the (3 + 1) configuration (intermediate absorption after the third bed), producing ca. 98.5 % sulfuric acid. The feed gas is produced by combustion of liquid elemental sulfur with dried air (—> Sulfur Diox-ide, Chap. 4.). Air is drawn from the atmosphere through a filter and an air dryer (irrigated with concentrated sulfuric acid, see Section 4.1.3.1) into the main blower, which compresses it to a pressure of ca. 1.4 bar.

The combustion air flow is controlled in pro-portion to a predetermined sulfur feed rate so that gases of about 11 — 12 vol % SO2 are produced by combustion. Combustion gas leaves the sul-fur combustion furnace at about 1100 °C and is cooled in a waste-heat boiler to the converter-inlet temperature of ca. 420 — 450 °C. Under the adiabatic conditions prevalent in the first bed of the converter (Section 4.1.3.2), heat generated

ProLess

Figure 18. Sulfur-buming double-absorption sulfuric acid process (Lurgi) a) Steam drum; b) Sulfur furnace; c) Waste heat boiler; d) Main blower; e) Mist eliminator; f) Drying tower; g) Air filter; h) Cooler; i) Acid pump tank; j) Intermediate absorber; k) Final absorber; 1) Candie filters; m) Steam superheater; n) Boiler; o) Economizer; p) Converter; q) Intermediate heat exchanger


in the reaction raises the temperature of the gas mixture until it reaches a level at which the reac-tion is essentially in equilibrium. In practice, this corresponds to ca. 600 — 620 °C, representing a sulfur dioxide conversion of about 60 %. Passing out of the first bed of the converter, the gases are cooled in a steam superheater to the second-bed inlet temperature of 430 — 440 °C. The gas leav-ing the second bed is cooled in the evaporator of the steam system to the third-bed inlet tempera-ture, ca. 420 — 440 °C. Gas leaving the third bed is cooled in two intermediate heat exchangers and passes at ca. 200 — 220 °C into the interme-diate absorber, where sulfur trioxide is removed by absorption in concentrated sulfuric acid (Sec-tion 4.1.3.3). The remaining gas, which is cooled in the absorber to about 70 — 80 °C, is returned to the intermediate heat exchangers through a heated double-jacket tube to avoid the conden-sation of sulfuric acid. The cold gas is heated in-directly by hot gas from the third bed to the inlet temperature of the fourth bed (ca. 400 °C). The hot gases coming from the fourth bed are cooled to ca. 160 °C in an economizer (feed-water pre-heater) and passed into the final absorber. After final absorption, so-called tail gas is discharged to the atmosphere through a stack at ca. 70 —80 °C. An overall sulfur dioxide conversion ef-ficiency of ca. 99.7 % is attainable in this type of plant. Assuming the feed gas contains 10 vol % SO 2 , this corresponds to a sulfur dioxide con-centration in the tail gas of about 400 ppm SO2 .

Figure 18 also shows the acid circuits for the air dryer and the intermediate and final ab-sorbers. In this example, the end product is 98 —98.5 % sulfuric acid, which is discharged from the final absorber circuit. Water required for maintaining the correct acid concentration, apart from that absorbed in the air dryer, is supplied as process water to the absorber acid system. Water addition is controlled automatically by an acid-concentration measuring system operat-ing within narrow limits (Section 4.1.3.3). The dryer-acid circuit is designed to discharge sulfu-ric acid of 93 — 97 % concentration back to the absorber circuit. If part of the production is to be discharged as oleum (see Section 4.4), and addi-tional oleum tower is added upstream from the intermediate absorber. Plants designed for max-imum oleum production preferentially reflect the (3 + 2) configuration. Five-bed converter sys-tems with (3 + 1) configurations are used espe-

cially for achieving maximum final conversions [44, 72-74, 80].

Sulfur-burning double-absorption plants are designed for production capacities < 3000 t/d of 100 % sulfuric acid in a single stream [81].

Energy Balance. All steps in the produc-tion of sulfuric acid from elemental sulfur are exothermic. An amount of heat corresponding to about 5.4 GJ is generated overall per ton of 100 % sulfuric acid. The thermal capacity of a sulfuric acid plant with an output of 1000 t/d H2SO4 is, accordingly, ca. 63 MW. The liber-ated heat must be dissipated under controlled conditions in such a way as to maintain opti-mum gas temperatures in the converter system and optimum acid temperatures in the dryer and absorber circuits, thereby ensuring that sulfur dioxide concentrations in the tail gas are mini-mized and sulfuric acid mist and sulfur trioxide at the stack outlet are avoided insofar as pos-sible. The systems for gas and acid cooling are therefore essential components of a sulfuric acid plant.

Figure 19 shows the energy flow in the form of a Sankey diagram for a double-absorption plant of the type represented by Figure 18 [75]. Of the total energy input (100 %), 97 % is ac-counted for as energy released in the conversion of sulfur to sulfuric acid, and 3 % of the energy is consumed in driving the gas through the plant. Up to about 70 % of the total energy is normally utilized for the generation of ca. 1.35 t of high-pressure steam (40 bar, 400 °C) per tonne of sul-furic acid; the remaining 30 % is usually lost as waste heat.

The high-pressure steam is generated with high-temperature heat recovered by indirect ex-change with gases from the converter system and the sulfur furnace. The waste-heat system is completely integrated into the double-absorp-tion sulfuric acid plant (Fig. 18). In the econo-mizer that cools the gases from the fourth bed, about 13 % of the overall process heat is uti-lized for feed-water preheating. The waste-heat boiler downstream from the sulfur furnace and the evaporator after the second bed use respec-tively 38 % and 16 % of the total heat to produce saturated steam. Another 10 % of the total en-ergy is recovered in the steam superheater after the first bed for the generation of high-pressure steam.


Sut f uF c/o

Tait gas 1 5 '1/4

Produc acid 0.5 *A

Die Waste hea# Tram acid cooting 31 %

Air

Figure 19. Sankey energy-flow diagram for a 1000-t/d sulfur-burning double-absorption sulfuric acid plant (feed gas: 11 % S02) a) Blower; b) Sulfur furnace; c) Waste heat boiler; d) Catalyst bed 1; e) Steam superheater; f) Catalyst bed 2; g) Boiler; h) Cat-alyst bed 3; i) Intermediate heat exchangers; j) Intermediate absorber; k) Converter bed 4; 1) Economizer; m) Final absorber; n) Air dryer; o) Acid coolers

Of the generally unrecovered waste heat, the majority (31 % of the total energy) is removed in the acid-cooling systems: 1 % in the dryer circuit, 24 % in the intermediate absorber cir-cuit, 5 % in the final absorber circuit, and 1 % in the product-acid cooler. Heat in the dryer results from the condensation of water vapor and the di-lution of sulfuric acid. Heat from the absorbers is made up of sensible heat transferred from the moderately hot gas to the colder acid in the ab-sorber, the latent heat of condensation of sulfur trioxide, and the heats of formation and dilution of sulfuric acid. A small amount of the total en-ergy is discharged from the system as sensible heat in the tail gas (2.5 %) and in the product acid (0.5 %).

Maximizing Energy Recovery. The value of energy has risen so much that systems for maximum energy recovery are becoming in-creasingly economical despite the associated considerable increase in capital cost. This is es-pecially true for the extremely large-capacity double-absorption sulfuric acid plants now in use.

In the steam system, it is possible to im-prove energy utilization not only by increas-

ing the production of high-pressure steam but also by optimizing the steam quality [76-78]. Use of a specially designed waste-heat system makes it possible because of the high excess temperature of the gas to produce steam at up to 80 bar in a sulfuric acid plant instead of the normal 40 — 50 bar. The higher the quality of the steam the more efficiently it can be con-verted into mechanical work or electrical en-ergy. Production of high-grade steam is advan-tageous even if low-pressure steam for heating purpo ses is all that is required in neighboring production plants. In this case the high-pres-sure steam can be transformed into low-pres-sure steam in a back-pressure turbine driving a generator or other equipment [82, 83]. The Lurgi substoichiometric sulfur-combustion sys-tem (-+ Sulfur Dioxide, Chap. 4.4. is well suited to the generation of highgrade steam, and it has been used successfully in a number of modern sulfuric acid plants [77].

Steam production can be increased by rais-ing the acid and gas-outlet temperatures in the intermediate absorber to reduce gas-heat losses. As a general rule, hot converter gas is cooled be-fore entering the intermediate absorber by indi-rect exchange with cold gases returning from the


Disitfled water produclion

Figure 20. Absorber acid heat-recovery system with intermediate closed-loop hot water circuit a) Absorber acid circuit; b) Venturi S03 absorption system; c) Boiler feed-water heater; d) Feed heater; e) Filter wash-water heater; f) Wash heater; g) Sea water desalinator; h) Closed-loop hot water circuit; i) Hot water circulation pump; j) Trim cooler

same absorber. The gas exit-temperature from the intermediate heat exchanger is usually too close to the acid dewpoint for the recovery of any more heat for steam generation. Otherwise there would be a danger of sulfuric acid condens-ing, which would give rise to corrosion on the heat-exchange surfaces as well as mist forma-tion. However, use of venturi-type cocurrent ab-sorbers and the hot absorption principle permits the gas temperature after the intermediate ab-sorption stage to be increased substantially (see Section 4.1.3.3). Thus, the amount of heat re-quired for heating the cold gases from the inter-mediate absorption stage is reduced. At the same time, an adequate amount of heat at sufficiently high temperature becomes available in the sul-fur trioxide-rich gas upstream from the interme-diate absorber to run an evaporator or econo-mizer (feed-water preheater) integrated within the steam system. This can be located either be-fore or after the intermediate heat exchangers. It is possible to boost steam production in the sul-furic acid plant by an additional 10 by preheat-ing the boiler feed water to about 90 °C, utilizing waste heat in acid from the absorption systems.

If recovery of acid waste heat is to be opti-mized, the temperatures in the absorption sys-tems become very important. Raising the tem-perature creates more favorable conditions and a wide range of possibilities for heat utiliza-

tion. The extent to which the acid temperature can be increased is restricted largely by the al-lowable corrosion rate in the heat exchangers. Considerable advances have been made in re-cent years in the development of special mate-rials and techniques for controlling corrosion at higher acid temperatures. A number of double-absorption plants now incorporate acid-cooling systems modified for heat recovery and utiliza-tion. The required investment can be amortized within a short time out of energy cost savings.

Figure 20 shows one possible combination of various approaches to utilizing recovered acid waste heat [77, 84]. The heat recovery system de-picted has been installed in a preexisting sulfur-burning double-absorption plant. The heat is used to produce boiler feed water by distilla-tion of seawater, for boiler feed-water preheat-ing, and for preheating filter wash water for a neighboring phosphoric acid plant. Heat is trans-ferred from the anodically protected stainless-steel acid cooler to the various consumers by means of a closed-loop hot water circuit.

In other sulfuric acid plants, waste heat reco-vered from the acid is utilized in hot water cir-cuits for remote heating [85]. In plants oper-ating with absorber acid at elevated tempera-ture (130 —140 °C), waste heat from the acid is also used for concentrating phosphoric acid from 28 — 54 % P205 (Fig. 21). Other plants use acid

52-54 7. P 20 5 Phosphoric .rid product 25-2S Vo P 2O s Phosphoric acid


Couting weer

Figure 21. Phosphoric acid concentration using sulfuric acid waste heat a) Sulfuric acid plant intermediate absorber; b) Sulfuric acid cooler; c) Hot water circuit; d) Trim cooler; e) Evaporator; f) Heat exchanger; g) Fluorine scrubber; h) Condenser; i) Vacuum pump

waste heat for reconcentration of spent sulfuric acid in a venturi system (see Section 4.3.2) [75, 84].

A significant step toward more highly energy-efficient plants was taken with the introduction of heat recovery systems generating steam from the acid circuit. Thus, part of the heat formed in the absorption circuits (and normally dissipated with the cooling water) is transferred to a higher level and used for the production of low-pressure steam at 6 —10 bar. Depending on the nature of the plant, up to 0.5 t of steam per tonne of acid can be produced in this way in addition to the high- or medium-pressure steam obtained from a sulfur-burning plant. The overall energy effi-ciency is thus increased to > 90 % based on a "standard sulfur-burning" plant [178,179].

Such heat recovery systems are normally in-stalled in the intermediate absorption circuits. Intermediate absorption is divided into two stages: a first stage operated at ca. 180 — 200 °C and a second stage at a "normal" absorption tem-perature of ca. 80 °C. About 95 % of the incom-ing S03 is absorbed in the first stage, where no sensible heat is removed from the gas. Acid is circulated at ca. 180 — 200 °C; the heat, which is a result of absorption/acid formation, is trans-formed into steam in a special boiler, usually of

the kettle type. In the second stage, remaining S03 is absorbed and the gas is cooled to ca. 80 °C in order to provide appropriate gas conditions for downstream plant elements. Heat formed in the second stage absorber is normally used for boiler feed-water preheating or other heating purposes.

At present, two such systems are available on the market, representing different approaches to realizing the desired end:

1) The Monsanto HRS system 2) The Lurgi low-pressure steam recovery sys-

tem

The Monsanto HRS system is based on inves-tigations showing that "normal" stainless steels are subject to low corrosion rates when the acid concentration is > 99 %. That "window of con-centration" permits the design and construction of a system consisting of 300-grade stainless steel and incorporating a first-stage absorption tower (with a countercurrent flow absorber, pip-ing, heat exchanger, etc.). No brick lining is re-quired. It is self-evident that the acid concen-tration must be carefully controlled, but this is quite feasible with modern instrumentation. The second-stage absorption circuit corresponds in principle to a standard absorption system. A typ-


a

913,6 04 Acid -

113.1:1 °F (BZ frem rime twear.

fias

b

r.

Acid ta final tower pump tank

Figure 22. Monsanto heat recovery system a) Mist eliminators; b) Pump boot; c) Heat recovery tower; d) Diluter; e) Heat recovery system boiler; f) Heat recovery system water heaters

ical flowsheet for a Monsanto HRS system is shown in Figure 22.

Lurgi has taken a different approach for its low-pressure steam recovery system. In cooper-ation with Thyssen and Krupp-VDM, a stain-less steel was qualified for sulfuric acid ser-vice over a relatively wide range of concentra-tions and temperatures (see Fig. 23). The mate-rial is a ferritic stainless steel with the German material code 1.4575 and known as Superferrit (Thyssen) or Cronifer 2803 (Krupp-VDM). This material provides excellent corrosion resistance, but handling and equipment fabrication require special precautions because of brittleness. Lurgi has chosen here to again make use of its ven-turi absorption systems, already familiar from other sulfuric acid plant applications (see Sec-tion 4.3.2). The venturi unit constitutes the first-stage absorber, to which a separate acid circula-tion system is connected, including a pump tank, pumps, piping, etc. Heat generated in this system is transformed into low-pressure steam in a spe-cially designed boiler. Second-stage absorption takes place in a small packed tower connected to the venturi unit. The venturi unit, the pump tank, the heat-exchanger tube bundle, piping, etc., are

all fabricated from 1.4575 steel. Alternatively, the venturi unit and pump tank can be of a stan-dard design• mild steel with a brick lining. A simplified flow diagram is provided in Figure 24.

Concentration, %

Figure 23. Range of application for Superferrit 1.4575 in concentrated sulfuric acid

With respect to the pumps, valves, etc., spe-cial cast alloys have been developed for the re-quired high-temperature service. Additional fa-cilities, such as continuous corrosion measure-ment and noise monitoring systems, are avail-able for such heat-recovery systems as acces-sories to facilitate monitoring of all the relevant pro ces s data.


r Jr

Dernineralued walcr

Figure 24. Low-pressure steam recovery system

Pressure-Contact Process. A specific vari-ant of the sulfur-burning double-absorption pro-cess is the pressure-contact process, in which sulfur combustion, sulfur dioxide conversion, and sulfur triacide absorption are all effected at elevated pressure. The idea of conducting a contact process under pressure is not new — it was first suggested over 50 years ago [86] and the notion has since been reiterated from time to time [87-91]. So far, however, is has only been embodied in one industrial double-absorp-tion plant with a capacity of 550 — 575 t/d H2 SO4 operated by PCUK in France [92-94]. The max-imum pressure in this plant is 5 bar. It first as-sumed operation in 1972.

After drying in a conventional tower under normal pressure, combustion air is compressed to 5 bar for introduction into the sulfur fur-nace. Liquid sulfur is supplied through a special burner, and it burns at a maximum temperature of 1800 °C. The combustion gas, containing about 12 vol % SO 2 and 9 vol % 02, is introduced at ca. 1200 °C into a fire-tube waste-heat boiler, after which it passes through a gas filter before being subjected to three stages of conversion in separate converter vessels. The process gas is then cooled and introduced into the intermedi-ate absorption stage. Downstream from the inter-mediate absorber there is one further converter, from which gas passes to the final absorber. Tail gas leaving the final absorber is reheated and expanded in a turbine coupled with the air com-pressor. It is asserted that about two-thirds of the compression energy is recovered in this way [93].

As in an atmospheric-pressure double-ab-sorption plant, the process gas is cooled in gas —gas heat exchangers before entering the inter-mediate absorber. In the pressure-contact pro-cess, however, only one exchanger is used for reheating cold gas returning from the interme-diate absorber to the fourth converter stage. An-other one serves to reheat the tail gas from the final absorber before it enters the turboexpander. Heat is recovered for steam generation in econ-omizers, positioned downstream from the fourth converter and the expansion turbine, as well as in a waste-heat boiler after the sulfur furnace.

Compared with a conventional double-ab-sorption process, two special advantages have been claimed for the pressure-contact process:

1) The position of chemical equilibrium in the sulfur dioxide oxidation reaction (Eq. 10) is more favorable, permitting a higher conver-sion efficiency with less catalyst. The PCUK plant is reported to have achieved 99.8 —99.85 % conversion with a catalyst charge of only about 76 L per tonne of daily sul-furic acid production capacity [93]. The tail-gas sulfur dioxide content is reported to be reduced to about 200 — 250 ppm by volume. However, the high temperature in the sulfur furnace increases the rate of nitrogen acide formation (-+ Sulfur Dioxide, Chap. 4.4.).

2) On account of the lower operating volume with respect to the converter gases, smaller equipment can be used. This reduces the ma-terial and site-area requirements, and it raises the capacity limit on shop-fabricated equip-

4... .....

AL

dor o I ve e À

einli

Tall gas

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■ çà

egm


ment. The resulting capital and construction cost savings are said to be ca. 10 —17 % rela-tive to a typical double-absorption plant [93]. It should be mentioned, however, that these savings would be nullified in many countries by the cost of conforming with requirements for extra wall thickness and higher-grade construction materials arising from safety regulations regarding pressure vessles.

The principal disadvantage of the pressure-contact process in comparison with a conven-tional double-absorption process is that it con-sumes more power and produces less steam. The magnitude of this negative effect on operating costs depends on the value of energy.

4.1.4.2. Double-Absorption Processes Based on Metallurgical Gases

Double-absorption plants based on metallurgi-cal gases differ from hot-gas plants based on sulfur combustion in that cold feed gases must first be heated to the converter-inlet temperature using energy liberated in the oxidation of sulfur dioxide.

Process Description. Figure 25 is a flow dia-gram of a typical double-absorption plant based on metallurgical gases, with a four-bed converter in a (2 + 2) configuration for the production of 98 — 98.5 % sulfuric acid [40]. Sulfur dioxide-containing gases enter the sulfuric acid plant from the gas-cleaning system free of dust and mist but still saturated with moisture. To adjust the 02/S02 volume ratio to the required value of ca. 1.0 — 1.1, air may need to be added to the gas before it enters the dryer. After drying (Sec-tion 4.1.3.1), the sulfur dioxide-containing feed gas, at about 50 °C, is compressed by the main blower to a pressure of ca. 1.35 — 1.45 bar, de-pending on the gas-pressure drop through the plant. The liberated heat of compression raises the gas temperature to ca. 90 — 100 °C.

The sulfur dioxide-containing gases are heated to the converter-inlet temperature of ca. 420 — 450 °C by indirect heat exchange with hot gases from, successively, the fourth, third, and first catalyst beds. Converted gases are simulta-neously cooled to the appropriate inlet temper-ature for the succeeding bed and for the final absorber. In some cases, a boiler element may be installed ahead of the intermediate or final absorber. The converter system and the inter-

r

Pur

Feed gas from gas clearling plant

Figure 25. Double-absorption sulfuric acid process for metallurgical feed gas (Lurgi) a) Gas—gas heat exchangers; b) Air filter; c) Drying tower; d) Cooler; e) Acid pump tanks; f) Intermediate absorber; g) Final absorber; h) Intermediate heat exchangers; i) Converter

9 e H 20

'rail gas

I

F4S 0,


mediate heat exchangers, as well as the inter-mediate and final absorption stages, are similar to those in the sulfur-based double-absorption process described previously, except that in this case the intermediate absorption stage is placed after the second catalyst bed rather than the third. Product sulfuric acid of 98 — 98.5 % concentra-tion is withdrawn from the final absorber circuit. If 93 — 97 % sulfuric acid is to be produced, acid discharged from the dryer circuit is preferably stripped with air, which is subsequently used to adjust the 02/S02 ratio of the feed gases. This removes any dissolved sulfur dioxide, reintro-ducing it into the plant with the feed gas. Most of the water requirement for sulfuric acid pro-duction can be satisfied by water vapor absorbed in the gas dryer.

Double-absorption plants of this type are normally designed for contact gases containing 5 —10 vol % S0 2 . Attainable final 502 conver-sions are > 99.5 %. In contrast to cold-gas plants based on pyrite roasting gases, the primary pur-pose of which is to produce sulfuric acid, sulfu-ric acid plants associated with nonferrous metal smelters and acid decomposition plants [95-97] have the primary function of cleaning the off-

gases. They must therefore be adaptable to fre-quent large fluctuations in the output and com-position of off-gases from the preceding plant. The lower limit for the sulfur dioxide concen-tration of off-gases subject to processing in a double-absorption plant is determined mainly by two criteria: it must be possible to maintain the water balance, and autothermal operation must be assured. The water balance can generally be maintained with a sulfur dioxide concentration of ca. 3.5 — 4 vol % S02 when cooling water at ca. 25 °C is available for the gas cleaning system (see Section 4.1.3.1 and Figure 13).

At the other extreme, modem nonferrous metal smelters generate gases with high S02 lev-els; S02 concentrations of up to 10 or 11 vol % may be achieved after mixing/dilution. In such a case the reaction heat developed in the converter system exceeds the requirements for heating the cold gas. The result is an excessively high gas temperature at the entrance to the absorption tower. Boiler elements (economizers, evapora-tors, superheaters, etc.) are therefore often in-corporated into the converter scheme to remove excess heat. A typical flow diagram for such a plant is shown in Figure 26 [180].

Figure 26. Contact acid plant for metallurgical gases with heat recovery a) Drying tower; b) Heat exchanger 4; c) Heat exchanger 3; d) Heat exchanger 2; e) Converter; f) Superheater; g) Intermediate absorber; h) Heat exchanger 5; i) Final absorber

I

ELertrical energy 6 % III,

RceaCer gas 94 Vo

Product ami 1 %

Waee !l'ut from a<id cooling 7i'4/

1 ait gas 5 Û.A.


Figure 27. Sankey energy-flow diagram for a 1000-t/d metallurgical double-absorption sulfuric acid plant (feed gas: 8.5 % S02) a) Drying tower; b) Blower; c) Heat exchanger; d) Catalyst bed 1; e) Catalyst bed 2; f) Intermediate heat exchangers; g) Catalyst bed 3; h) Catalyst bed 4; i) Intermediate absorber; j) Final absorber; k) Acid coolers; 1) Acid tank; m) Product acid cooler

Energy Balance. Figure 27 is a Sankey dia-gram showing the energy balance for a double-absorption plant analogous to Figure 25 based on metallurgical gases. An essential characteristic of a conventional cold-gas plant is that almost all the energy is discharged as waste heat at low temperature. Of the total energy input, 94 % is accounted for as chemical energy released in the conversion of sulfur dioxide to sulfuric acid and in the condensation of water in the dryer. At a feed-gas concentration of 8.5 vol % SO 2 and a dryer inlet temperature of 30 — 40 °C, a total of about 2.7 GJ of thermal energy is liberated per ton of sulfuric acid. This corresponds to a ther-mal output of ca. 31 MW for a 1000 t/d plant.

About 45 % of the total energy is discharged through the intermediate absorber acid-cooling system, about 23 % through the final absorber acid-cooling system, and about 22 % through the dryer-acid cooling system. In order to main-tain autothermal operating conditions, gas-heat losses in the absorption system must be covered by the heat of reaction released in the converter system [40]. At thermal equilibrium, the heat content initially supplied to the converter sys-tem during preheating is thus maintained.

For a particular gas throughput, the gas-heat losses remain constant, but the lower the con-centration of sulfur dioxide in the feed gas, the less heat is released in the converter system. The lower lirait for maintaining thermal equilibrium is about 5 vol % SO2 for a conventional plant. Minimizing the gas-heat losses in the interme-diate absorption stage through the use of venturi-type cocurrent absorbers and the hot absorp-tion principle makes it possible to maintain au-tothermal operation even when processing con-tact gases containing as Little as 4 vol % SO2 .

Proper design of the gas — gas heat exchang-ers in relation to the sulfur dioxide concentration is a prerequisite for maintaining thermal equi-librium. The lower the sulfur dioxide content, the smaller will be the difference between the mean inlet and outlet temperatures at each of the catalyst beds. With lower temperature differ-ences, the specific heat-exchanger surfaces must be larger. For example, processing gases with 4 vol % SO 2 rather than 8 vol % SO 2 means that the required specific heat-exchanger surface is more than three times as large [98].

At somewhat higher sulfur dioxide concen-trations (8 — 9 vol %), the required specific hea-


Tait gas

cf

Producl acid

Feed gas tram gas

de?ning plant

Figure 28. Ordinary (single-absorption) contact sulfuric acid plant for metallurgical roaster gas a) Gas — gas heat exchangers; b) Main blower; c) Drying tower; d) Cooler; e) Acid pump tanks; f) Absorber; g) Converter

texchanger surface reaches a minimum. At still higher sulfur dioxide concentrations it rises again, and in this case it is advisable to install a waste-heat boiler element to convert the surplus reaction heat into steam.

Heat Recovery. In a conventional cold-gas double-absorption plant for processing rela-tively low-grade sulfur dioxide-containing feed gases, there is no excess high-temperature heat that can be used for generating high-pressure steam (see Fig. 27). However, if the sulfuric acid unit is associated with a fluidized-bed roaster for sulfide ores, high-pressure steam is gener-ally available from a waste-heat boiler system utilizing at least part of the heat of the roasting reaction. In pyrite or zinc blende roasters, for in-stance, steam is produced at about the same rate per tonne of sulfuric acid as in a double-absorp-tion plant based on sulfur combustion. Where the sulfuric acid plant is linked to a modern smelter (Outokumpu, QSL, Kivcet, etc.), high-level S02 gases are available for conversion, so the re-moval of surplus reaction heat in a waste-heat boiler element becomes almost obligatory. To increase the output of high-pressure steam, low-temperature heat from the absorber-acid circuits

can also be utilized for preheating the boiler feed water. Recovery of surplus heat from the acid system, which accounts for ca. 92 % of the en-tire waste heat produced in a cold-gas plant (see Fig. 27), is of the same economic importance as in a plant based on sulfur combustion. Sys-tems for acid waste-heat utilization have been in-stalled in a number of roaster gas-based double-absorption plants, generating hot water for in-dustrial purposes (e.g., P205 concentration), or for district heating schemes [99, 100].

4.1.4.3. Ordinary Single-Absorption Processes

Before the introduction of the double-absorp-tion process, the ordinary contact or single-absorption process was the standard sulfuric acid production process [19-23]. Figure 28 is an example of a flow diagram for an ordi-nary contact plant based on metallurgical gases equipped with a four-bed converter. Before en-tering the dryer, the cleaned sulfur dioxide-con-taining gases are generally diluted extensively with air to an 02/S02 volume ratio of at least 1.7. The dried gases are then compressed to


the pressure required to overcome the pressure drop of the overall plant. Although this is about 0.1 bar less than the pressure drop in a compa-rable double-absorption plant, almost the same specific blower capacity is needed on account of the substantially higher specific gas volume that results from the greater dilution of the feed gas.

The feed gas is heated stepwise in a series of indirect gas — gas heat exchangers cooling first the hot gases from the last catalyst bed and then the gases between the individual beds. The cooled sulfur trioxide-rich gases from the last bed are passed into the SO3 absorber.

In principle, the ordinary contact process re-tains more gas heat than the double-absorption processes, because there is no intermediate ab-sorption stage with its associated heat losses. In a roaster-based plant the surplus heat is simply discharged, for instance by indirect air cooling in an air cooler after the third bed (Fig. 28). The re-action heat in a plant based on sulfur combustion is recovered by steam generation in economiz-ers and evaporators. About 10 % more steam is therefore produced per tonne of sulfuric acid in a sulfur-burning ordinary contact plant than in a double-absorption plant of the same capacity [73].

Because of the favorable energy balance, an ordinary contact plant can also process gases with a lower sulfur dioxide content [101]. The lowest sulfur dioxide concentration at which thermal equilibrium can still be maintained in continuous operation without recourse to an ex-ternal heat source is ca. 2 vol %. The water bal-ance in a sulfuric acid plant based on such a low SO2 content can be maintained with the aid of a predryer — reconcentrator system developed by Lurgi [51] (see Section 4.1.3.1).

The maximum overall SO 2 conversion attain-able in an ordinary contact plant without inter-mediate sulfur trioxide absorption is 98 —99 % in continuous operation under favorable con-ditions. Assuming a feed-gas concentration of 7 vol % SO2, this corresponds to a sulfur diox-ide content in the tail gas of about 1600 ppm by volume. Stringent environmental protection regulations introduced since the early 1970s in many industrialized countries demand a higher standard of performance than this. Most new sulfuric acid plants are therefore designed ac-cording to the double-absorption principle. Or-dinary contact plants are used today only for pro-

cessing off-gases of extremely low sulfur diox-ide content, such as highly dilute smelter off-gases, which cannot sustain autothermal opera-tion in a double-absorption plant. At low feed-gas concentrations (e.g., 3 vol % SO2) and with the help of highly active (cesium-doped) cata-lysts, tail-gas SO2 concentrations are attained that are equivalent to the levels associated with "standard" double absorption plants based on high SO2 concentrations. In such cases, the or-dinary contact process, which produces concen-trated sulfuric acid as a marketable end product, is preferable to other off-gas purification meth-ods such as alkaline washing. The SO2 separa-tion efficiency with these alternative approaches is seldom higher, and they usually lead to un-marketable end products or polluting waste ma-terials.

When necessary, an existing ordinary contact plant can be brought into conformity with strict antipollution legislation either by retrofitting with an intermediate absorption stage, trans-forming it into a double-absorption plant, or by addition of a tail-gas purification system. Acid waste-heat recovery systems have also been in-stalled in ordinary contact plants [79, 102].

4.1.4.4. Wet-Catalysis Processes

Wet-catalysis processes differ from other con-tact sulfuric acid processes in that the feed gas still contains moisture when it cornes into con-tact with the catalyst. Sulfur trioxide formed by catalytic oxidation of the sulfur dioxide reacts instantly with the moisture to produce sulfuric acid in the vapor phase to an extent determined by the temperature. Liquid acid is subsequently formed by condensation of the H2SO4 vapor and not by the absorption of sulfur trioxide in concentrated sulfuric acid, as in a contact pro-cess based on dry gases. The concentration of the product acid depends on the H2 0/503 ra-tio in the catalytically converted gases as well as on the condensation temperature [103]. The wet-catalysis process is especially suitable for processing the wet, dust-free gases obtained in the combustion of hydrogen sulfide-containing off-gases, which need only be cooled to the con-verter inlet temperature of ca. 440 °C. Process-ing these moisture-laden gases in a conventional cold-gas plant would necessitate cooling to an


economically unacceptable extent in order to re-move the large excess of moisture, an expedient that would only be justified for a gas with a high dust content and a relatively high sulfur dioxide concentration.

Classical Process. The classical wet-catalysis process was developed by Lurgi as early as the 1930s. It is illustrated schematically in Figure 29. Dust-free hydrogen sulfide-con-taining gases are burned at ca. 500-1000°C (depending on their composition) with a suffi-cient excess of air to produce a contact gas with an 02/S02 volume ratio of ca. 1.1 — 1.3. The hot, wet, sulfur dioxide-containing gas is cooled to ca. 440 °C in a waste-heat boiler (or, where the volume of gas involved is very small, in a radiation cooler) and passed directly into the converter. The converter contains three or four catalyst beds, between which cold quench air is added to cool the hot reaction gases.

Gas from the last bed enters a condensation tower at ca. 420 — 430 °C. Here it is cooled to the condensation temperature of sulfuric acid by irrigation with circulated, cooled sulfuric acid. Good contact between the gas and the liquid, as well as a high irrigation density, are essen-tial to ensure that the gas is cooled insofar as possible at the gas — liquid interface of the irri-gation acid, causing the condensing sulfuric acid to coalesce with the irrigation acid. This largely avoids formation of the acid mist that would re-sult if sulfuric acid were to condense from the gas phase due to the mixing of cold and hot gas streams. Any mist that does form is subsequently removed from the tail gases by passing them through glass-fiber candle filters. The sulfuric

acid thus obtained is fed to the acid circuit of the condensation tower. The gas temperature at the condensation-tower outlet is maintained at ca. 60 °C to restrict H2SO4 and SO3 emission.

Classical wet-catalysis plants are usually de-signed for the production of 78 % sulfuric acid. Even if the H2 0/S03 ratio of the gases leaving the converter were high enough for the product concentration to exceed 78 % (assuming com-plete condensation of the sulfur trioxide and water), it would not be advisable to operate at higher acid strength because the 80 — 90 % H2SO4 concentration range is quite critical with respect to the corrosion-resistance of most of the usual construction materials. The water con-tent of moist hydrogen sulfide-containing gases is normally too high to sustain a sulfuric acid concentration > 90 %.

Acid leaving the condensation tower at ca. 80 — 85 °C is cooled to 50 — 60 °C for recircula-tion. If necessary, the concentration is adjusted to 78 % H2SO4 by addition of process water.

A sulfur dioxide conversion of about 97.5 % is achieved in the wet-catalysis process with a four-bed converter. Since wet-catalysis plants are almost exclusively used for desulfurization of hydrogen sulfide-containing off-gases, the corresponding acid production capacities are comparatively small. The largest wet-catalysis plant built by Lurgi in the United States has a capacity of 250 t/d of H2SO4.

The classical wet-catalysis process cannot normally be used for processing wet hydrogen sulfide-containing gases that are too dilute to sustain combustion without supplemental fuel, since the water content of the resulting sulfur dioxide-containing gases would be too high to

ti

H 2 S-Containing gas

Prorpss wRI - nr p-

L

Figure 29. Wet-catalysis sulfuric acid plant processing il2 S -containing gases a) Blower; b) Air filter; c) Combustion furnace; d) Waste heat boiler; e) Quench-cooled converter; f) Condensation tower; g) Coolers; h) Acid pump tank; i) Candle filter


permit production of 78 % sulfuric acid. In any case, the opportunities for marketing 78 % sulfu-ric acid are somewhat limited, and for that rea-son the wet-catalysis method has been devel-oped further to permit the conversion of gases with extremely low hydrogen sulfide content to sulfuric acid more concentrated than 90 % [104— 108]. Examples of such processes are the Concat process (Lurgi) and the WSA process (Haldor Topsoe). Another special development in this area is the wet — dry contact process with inter-mediate condensation.

Concat Process. Lurgi's Concat process, il-lustrated schematically in Figure 30, differs from the classical wet-catalysis process principally in the way in which the condensation is accom-plished. This occurs in two stages, combining a venturi unit with an irrigated packed tower [103]. In the first stage, the so-called hot condensation, the major part of the sulfur trioxide and/or sulfu-ric acid condenses at ca. 180 — 230 °C from the hot gases, which leave the converter. At this high condensation temperature only a very small por-tion of the free water passes from the gas stream into the condensed sulfuric acid on account of the relatively high partial pressure of water vapor over sulfuric acid at this temperature. Therefore, acid with a 93 % H2 SO4 content can be obtained in the first condensation stage. This section uti-lizes a venturi scrubber into which the circulat-ing acid is injected in a cocurrent sense relative to the gas to ensure intimate mixing of the gas phase with the liquid phase and the most effi-cient removal possible of the heat of formation and latent heat of condensation of the sulfuric acid.

As the gas passes between the first and sec-ond condensation stages, additional air is in-jected both for cooling purposes and to reduce the water-vapor partial pressure through dilu-tion. The second stage is a packed tower irri-gated with relatively dilute sulfuric acid from which water evaporates as a result of contact with the still moderately hot gas. In the process, the gas is cooled further, and this promotes con-densation of the residual sulfuric acid vapor. Gas leaving the condensation tower is freed from re-maining acid mists in a candie filter before being discharged to the stack.

Sulfuric acid separated in the candle filter drains into the acid circuit of the condensation tower, from the sump of which excess acid fiows over to the hot-condensation stage. Hot acid in the venturi circuit is cooled indirectly. To ensure that it does not exceed the maximum permissi-ble inlet temperature for the particular material of which the cooler is made, part of the cooled acid is recycled to the pump tank. Excess acid that accumulates in the circuit of the first stage is discharged as product after additional cooling.

On account of this two-stage condensation and the cocurrent hot-condensation principle, the Concat process is able to process gas with a sulfur dioxide content of less than 1 vol % and a water content in excess of the stoichiometric requirement, resulting in 78 % or 93 % sulfuric acid. A sulfur dioxide content < 200 ppm can be attained in the tail gas with a two- or three-bed converter depending on the sulfur dioxide concentration of the incoming gases.

Topsoe Wet Sulfuric Acid (WSA) Process [181]. The Topsoe WSA process is used for

h a 1— --111-1—

Ir H 2 S - C-untaining gas

5-1

Fuel Proc.ms water

P

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Figure 30. Lurgi Concat (hot condensation) wet-catalysis process a) Blower; b) Air filter; c) Combustion furnace; d) Waste heat boiler; e) Converter; f) Coolers; g) Acid pump tank;h) Hot-condensation venturi; i) Condensation tower; j) Candie filter

St earri.

Support fuel


Dean off-gas

Hot air

Off-gas

Cooling

Su Efu riz azi d

Loo[ing water

Figure 31. Haldor Topsoe WSA process a) Fan; b) WSA condenser; c) Acid tank; d) Pump; e) Gas cooler; f) SO2 converter; g) Burner; h) Salt system; i) Gas preheater; j) Gas heater

treating wet S0 2 - containing gases — either "cold gases" (e.g., smelter off-gas) or "hot gases" (from catalytic or thermal incineration units) — without the need for prior drying. The process ef-ficiently recovers sulfur in the form of 93 — 98 % sulfuric acid despite off-gas water contents that may exceed 20 %. Recoveries > 99 % have been achieved. The main components of the process are an SO2 converter, a heat-exchanger system, and a WSA condenser (see Fig. 31). Upstream from the WSA process, dust and trace materials such as arsenic and hydrogen fluoride are re-moved by scrubbing if these are present in sig-nificant concentrations. Cold gas is heated in two steps. The first heat exchanger uses hot air from the WSA condenser, whereas the second uses the hot converted gases. An in-line burner supplies whatever additional heat is required to reach the appropriate inlet temperature for the SO2 con-verter (400 — 420 °C). In cases where the SO2 concentration exceeds ca. 2.5 %, no supplemen-

tary heat is necessary. Surplus heat is used for steam production or dispersed in a cooler.

In the converter, SO2 is oxidized to S03 by means of a sulfuric acid catalyst. Conversion is accomplished either in a single-stage adiabatic converter or in a two-stage converter with inter-mediate cooling, depending upon the SO2 con-centration and the required level of sulfur re-moval.

In the next heat exchanger the temperature of the gas is reduced. Cooling causes the S03 to react with water in the off-gas to form sulfuric acid vapor. The heat released by this gas-phase hydration is transferred to incoming cold S02 - containing gases.

Finally, the sulfuric acid vapor is con-densed as concentrated acid in the WSA con-denser/concentrator, which is an air-cooled glass-tube device. The sulfuric acid-containing gas passes on the tube side and is cooled to ca. 100 °C by ambient air passing on the shell side.

e

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Sulfuric acid condenses selectively and flows downward on the walls of the glass tubes. The acid is concentrated in the lower parts of the tubes by countercurrent contact with hot off-gas; the acid is subsequently collected at the bottom of the condenser. Hot acid leaving the condenser is cooled in a water-cooled plate cooler before being pumped to a storage tank.

Careful control of the temperature and of sul-furic acid droplet formation makes it possible to keep the level of sulfuric acid in the cleaned off-gas low without the use of any subsequent wet electrostatic precipitator or low-velocity mechanical filter. Cleaned off-gas leaves the top of the condenser and goes to a stack. Cooling air leaves the condenser at ca. 200 °C, and is used for preheating the cold feed gas.

In power-plant boiler applications the WSA process is combined with a selective catalytic reduction (SCR) NO„ removal system. The first heat-exchanger step is omitted, because the off-gas is already hot as received after dust filtration. A rotary gas — gas heat exchanger is used to heat the off-gas ahead of the SO2 converter. The SO 2 converter is an adiabatic dust-tolerant type with a number of vertical catalyst beds placed in par-

allel array. The catalyst beds can be dedusted one by one while the SO2 converter remains online without disrupting plant operation. Hot cooling air from the WSA condenser is used as boiler combustion air.

The main construction material employed in the WSA system is either carbon steel or low-alloy steel.

Wet—Dry Contact Process. The wet — dry contact process with intermediate condensation, a joint development by Lurgi and Süd-Chemie, is a modified double-absorption process based on sulfur combustion. Instead of dried air, in-dustrial waste air contaminated with such sulfur compounds as hydrogen sulfide, carbon disul-fide, or carbonyl sulfide is used to burn the sul-fur [109]. The impurities are oxidized simulta-neously with combustion of the sulfur. The re-sultant gases, containing 9.5 — 10 vol % SO 2 and less than the stoichiometric amount of water nec-essary for the formation of 98.5 % sulfuric acid, are passed into the converter after prior cooling in a waste-heat boiler (Fig. 32).

In contrast to the normal double-absorption process based on sulfur combustion, the moist

Ta il p S

Figure 32. Wet/dry catalysis sulfuric acid plant with intermediate condensation/absorption a) Air filter; b) Blower; c) Sulfur furnace; d) Waste-heat boiler; e) Steam drum; f) Steam superheater; g) Boiler; h) Economizer; i) Converter; j) Intermediate heat exchangers; k) Intermediate condenser/absorber; 1) Coolers; m) Acid pump tanks; n) Final absorber


sulfur trioxide-containing gases formed in the first beds of a wet-catalysis system are passed on to a two-stage intermediate condenser. In the first stage, which is a vertical venturi unit, the entire water content of the gas and most of the sulfur trioxide are removed into hot sulfuric acid injected cocurrently. Remaining sulfur trioxide is absorbed in the second stage, a packed tower irrigated with cold acid. The concentration of the circulating acid is adjusted to 98.5 % by the addition of process water. Dry S02- containing residual gases from the intermediate condenser are processed in the same way as in a normal double-absorption plant (Section 4.1.4.1).

A plant of this type for the production of 120 t/d of sulfuric acid has been in operation at the Kelheim works of Süd-Chemie since 1980. This plant processes off-gases from a rayon sta-ple factory containing ca. 0.5 vol % H2 S + CS2 [109,110].

4.13. Tail-Gas Treatment

The tail gases from a sulfuric acid contact plant consist chiefly of nitrogen and residual oxygen. They also contain sulfur dioxide at a low con-centration that depends on the conversion ef-ficiency attained in the conversion stages. The content of gaseous sulfur trioxide and sulfuric acid is essentially a function of the temperature and concentration of irrigation acid in the final absorber. Under unfavorable operating condi-tions (e.g., when the sulfur dioxide-containing converter feed gases are inadequately dried, or contain hydrocarbons) sulfuric acid mists may form which are not removed in the absorption system. This is true even when the concentra-tion and temperature of the absorber acid are at their optimum values. The safest way to remove these acid mists is with a candie filter, although this is not very effective for removing excess sulfur trioxide, which may result from poor acid distribution in the absorber.

The sulfur dioxide level in the tail gases is dependent not only on the sulfur dioxide con-version achieved but also on the S02 content of the converter feed gases (Fig. 33). However, the specific sulfur dioxide emission rate (kilograms of S02 per tonne of 100 % H2SO4) is a function only of the sulfur dioxide conversion efficiency.

An essential prerequisite for attaining maxi-mum sulfur dioxide conversion and correspond-ingly low sulfur dioxide emission is maintain-ing the gas conditions at the converter inlet (gas volume, gas temperature, sulfur dioxide con-centration, and 02/S02 ratio) as constant as possible. This is in general reasonably easy in a sulfuric acid plant based on sulfur combus-tion or metallurgical gases from a fluidized-bed furnace. But in sulfuric acid plants as-sociated with metal smelters, especially cop-per smelters incorporating converter operations (-+ Sulfur Dioxide, Chap. 5.5.1.), wide fluctua-tions in the gas-flow rate as well as in the off-gas sulfur dioxide concentration are quite nor-mal [111, 112]. Consequently, the sulfur diox-ide content of tail gases from the corresponding sulfuric acid plants themselves may vary quite noticeably from the optimum value. It is also almost impossible in most cases to avoid fluctu-ating emissions during start-up [113].

5cI Conversion rate, fb — 99.9 99 B 99 7 9 9 99 9b 91 96 9L

2 3 4 5 7 10 15 20 30+0 'SIZI 7 Emission, kg/1-1 2S0,,

Figure 33. Sulfur dioxide emission as a function of S02 concentration at the converter inlet and the S02 conversion rate

Replacing the normal contact process by a double-absorption process results in a marked reduction in sulfur dioxide emissions [114-116]. The specific sulfur dioxide emission of a single-

Aqueous armai -li a ...

Process waher

SuEfuric acid

- 1 ■e,


absorption plant with a sulfur dioxide conver-sion of 97.5 % is about 17 kg of SO2 per ton of H2SO4. Corresponding values for double-ab-sorption plants with sulfur dioxide conversion efficiencies of 99.5 % and 99.7 % are ca. 3.3 kg and 2 kg of SO 2 per tonne of H2 SO4 respec-tively (Fig. 33). For a plant with an H 2 SO4 ca-pacity of 1000 tid this translates into a reduction from 710 kg/h of SO 2 to ca. 138 kg/h or 83 kg/h. Theoretically, it would be possible to improve the conversion efficiency, and thus reduce emis-sions still further, by introducing an additional intermediate absorption stage, corresponding to a triple-absorption process. In practice, however, this method is not feasible, for both technical and economic reasons [80].

Environmental regulations — in Germany, for example — call for further improvement in the conversion efficiency normally achieved with a double-catalysis plant. At locations where the maximum permissible ground-level SO2 con-centration is exceeded because of emissions from other nearby sources, it may even become necessary to reduce SO2 emissions below the value attainable with a conventional double-ab-sorption process. TWo approaches to a further reduction of the SO2 emissions might be con-sidered in such cases: installation of a fifth cata-lyst bed, or addition of a tail-gas treatment unit. Whereas the SO2 reduction in a 5-bed converter system is achieved through an improvement in conversion efficiency, reduction of SO2 emis-sions in a tail-gas treatment plant is effected by

reacting the SO2 with some type of scrubbing liquor. The 5-bed converter system will proba-bly continue to be the most convenient solution for new plants so long as the target conversion efficiency remains of approximately 99.8 %. In the case of an existing plant, "add-on" tail-gas treatment may be feasible, although some pro-cesses of this type lead to byproducts that also require treatment. Of the known tail-gas clean-ing processes, ammonia washing, the Wellman —Lord process, and the Sulfacid process are most suitable for application in normal contact plants; the Peracidox process is used only in conjunc-tion with double-absorption plants. For a single catalysis plant, upgrading to a double-absorp-tion operation remains an attractive alternative to a tail-gas scrubbing system.

Ammonia Washing. Sulfuric acid tail-gas purification processes based on ammonia wash-ing are distinguished mainly by the method selected for reprocessing the sulfur dioxide-loden scrubbing liquor [111, 115-120]. Fig-ure 34 shows a Lurgi-designed system capable of removing up to 95 % of the incoming sul-fur dioxide. This high level of sulfur dioxide removal is made possible by a two-stage gas-scrubbing system.

Tail-gases from the absorber of the sulfuric acid plant are washed with circulated ammo-nia solution in two venturi scrubbers arranged in series. Unconverted sulfur dioxide in the tail gas dissolves as ammonium sulfite or ammoni-

5B Recycle to d ryinig tower

Figure 34. Ammonia scrubbing system for an ordinary (single-absorption) contact sulfuric acid plant a) Ammonia scrubbers; b) Gas filter; c) Decomposer; d) Fan; e) Stripper


um hydrogensulfite. Any sulfur trioxide or sulfu-ric acid carried over from the absorber is trans-formed into ammonium sulfate or ammonium hydrogensulfate. Before being discharged to the stack, the cleaned gas is passed through a mist eliminator. Liquor drawn from the mist elimi-nator and fresh ammonia solution are then in-troduced into the second venturi srubber. Scrub-bing solution from the second stage overflows to the first. Thus, the solution in the second stage, where the sulfur dioxide content of the tai]. gas is lower than in the first, has a higher pH and a lower sulfite concentration than the solution in the first stage, causing the partial pressure of sulfur dioxide to be reduced even further in this stage.

A bleed stream of sulfite-rich scrubbing so-lution is withdrawn from the first stage and in-troduced into a decomposer along with concen-trated sulfuric acid. The acid decomposes the sulfites, liberating sulfur dioxide and forming an ammonium sulfate solution. This solution re-tains a certain amount of dissolved sulfur diox-ide, which is subsequently stripped out with air. Moist sulfur dioxide gas from the decomposer is combined with the stripping air and recycled to the dryer of the sulfuric acid plant.

The economics of an ammonia tail-gas scrub-bing process depend chiefly on the fate of ammo-nium sulfate solution discharged from the strip-per [118-120]. The most favorable situation is one in which the sulfuric acid plant is part of a fertilizer complex, because in that case the so-lution can probably be used directly in the pro-duction of granular fertilizers, and ammonia is likely to be available at a reasonable cost. Oth-erwise, assuming there is an adequate market, the ammonium sulfate may be sold, so that at least part of the ammonia cost can be recovered through a credit for the ammonium sulfate. Nor-mally, however, there is little prospect for selling the solution directly, and working it up into solid products would involve additional capital invest-ment as well as energy, leading to high operating costs.

In a sulfuric acid plant designed to regener-ate spent acid (-+ Sulfur Dioxide, Chap. 5.6.1.), sulfite solution formed in an ammonia tail-gas scrubbing system can be disposed of in the spent-acid decomposition furnace. This circumvents the problem of finding an outlet for byproducts, but the resulting convenience must be paid for

in lost ammonia values, since ammonia is ox-idized to nitrogen. Also, decomposition of the sulfite solution and especially the evaporation of water require additional energy, which is not matched by the heating value of the ammonia that is burned.

Wellman — Lord Process. In the Wellman —Lord process, sulfur dioxide is washed from the acid-plant tail gas by a saturated aque-ous solution of sodium sulfite and sodium hy-drogensulfite (-+ Sulfur Dioxide). Wet sulfur dioxide-containing gas expelled during thermal regeneration of the scrubbing solution is recy-cled to the contact plant. About 10 — 20 % of the recovered sulfur ends up as sodium sulfate ow-ing to oxidation to sulfite in solution. This dis-solved sodium sulfate cannot be decomposed, so it must be removed as a byproduct. The eco-nomics of the process are determined by the pu-rity and marketability of the byproduct.

Sulfacid Process. The Sulfacid process is a tail-gas purification method (developed by Lurgi) that produces dilute sulfuric acid as a byproduct [111, 115, 121, 122]. Dry tail gas nor-mally leaves the sulfur trioxide absorber of an acid plant at 75 — 80 °C. This is first humidified with steam and then passed through a reactor filled with activated carbon. Sulfur dioxide ad-sorbed by the wet activated carbon is oxidized to sulfuric acid at 50 — 80 °C by residual oxygen in the tail gas. This acid, together with traces of sulfur trioxide and sulfuric acid removed from the tail gas, is retained in the pores of the carbon. The purified gas is discharged to the stack at 40-80°C; the attainable sulfur dioxide separation efficiency is normally > 90 % [121]. Adsorbed acid is washed out by spraying the activated car-bon bed with water at regular intervals without interrupting the gas flow.

The concentration of acid discharged from the reactor is kept below 20 — 25 % in an indus-trial plant by the dilution effect of the flushing water. Increasing the acid concentration by de-creasing the water flow-rate would reduce the solubility of sulfur dioxide and, consequently, the sulfur dioxide-removal efficiency. The re-sulting dilute acid is recycled to the contact plant, where it takes the place of process water.

If the amount of water introduced with the dilute acid exceeds the water requirements of


the contact plant, the dilute acid may be concen-trated by means of the well-established venturi method (see also page 53). In this case acid leav-ing the Sulfacid reactor is passed on to a single-or double-stage venturi reconcentrator located upstream from the Sulfacid reactor. In the course of raising the acid concentration above 30 % the tail gas itself will absorb enough water vapor from the acid to render the humidification stage of the basic process unnecessary, with a conse-quent saving in steam. Furthermore, the tail-gas heat is utilized in concentrating the acid. Any ad-ditional heat required can be supplied by indirect heat exchange between the acid-plant absorber and the concentrator [122].

Peracidox Process. The Peracidox process is an oxidative sulfur dioxide-removal process developed by Lurgi and Süd-Chemie [121, 123] for purifying the already very dilute tai' gas from a double-absorption sulfuric acid plant. It does not lead to any byproducts or waste, producing only sulfuric acid, which is recycled to the main sulfuric acid plant.

Hydrogen peroxide or electrolytically pro-duced monoperoxysulfuric acid (H2S05) is used to oxidize sulfur dioxide to sulfuric acid [55, 122]:

SO2 +H202 —> H2SO4 (14)

SO2 +H2S05 +H20 —+ 2 H2SO4 (15)

Figure 35 is a flow diagram of the electrolytic version of the Peracidox process. Tail gas leav-ing the final absorber of the sulfuric acid plant is washed in a two-stage Peracidox scrubber with ca. 30 — 40 % sulfuric acid in which the oxidant is dissolved. The attainable sulfur dioxide-removal

efficiency is > 90 %. Entrained sulfuric acid va-por and residual sulfur trioxide are washed out simultaneously.

The oxidant is generated in an electrolytic unit by oxidizing H2SO4 anodically in a di-lute sulfuric acid stream produced from concen-trated acid from the product line. Diperoxysul-furic acid (H2S208) initially formed in the elec-trolytic cell disproportionates in the dilute acid to form sulfuric acid and monoperoxysulfuric acid (Caro's acid) as the final oxidizing agent. When hydrogen peroxide is used as the oxidant, the electrolysis unit and pump tank are not, of course, required — only a storage tank for the hy-drogen peroxide [55].

Oxidizing solution is introduced into the acid circuit of the second scrubbing stage at a stoi-chiometric rate proportional to the amount of sulfur dioxide to be removed. Any residual oxi-dant in solution overflowing to the first scrub-bing stage reacts there with the entering tail gases. Overflow withdrawn from the circuit of the first stage thus consists only of dilute sul-furic acid, and it is introduced into the final or intermediate absorber of the main acid plant in place of process water.

Depending on the location of the plant and prevailing prices for electric power or hydrogen peroxide, the operating costs for a Peracidox process with electrolytically produced oxidant may be lower than costs incurred with the use of hydrogen peroxide [55]. If tail-gas purification is required only intermittently during short pe-riods when sulfur dioxide emissions temporar-ily exceed permitted limits — during start-up, for example — it is more economical to use hydro-gen peroxide on account of the lower investment costs [122].

Figure 35. Lurgi Peracidox tail-gas purification system (electrolytic version) a) H2SO4 plant final absorber; b) Cooler; c) Pump tank; d) Peracidox scrubber; e) Electrolytic oxidation unit; f) Electrolyte tank


f

A Peracidox tail-gas scrubbing unit with elec-trolysis has been in operation since 1972 at a 120-t/d double-absorption sulfuric acid plant in Germany [123]. Another scrubbing unit, using hydrogen peroxide and attached to a 330-t/d double-absorption sulfuric acid plant (also in Germany) has been operating since 1982 [124].

4.1.6. Economic Factors

Economies of Scale. Due to a steep increase in demand for sulfuric acid since the early 1950s, the unit size of sulfuric acid plants has been constantly increasing. For example, in 1954 the largest sulfuric acid plants in the Federal Re-public of Germany were designed for produc-tion of ca. 250 t/d, and in the United States for about 450 t/d. Today, the largest double-absorp-tion plants have capacities of ca. 1000 t/d in the Federal Republic of Germany and ca. 3000 t/d in the United States.

The incentive for steadily increasing the size of sulfuric acid production units is "economy of scale" — savings achieved in both specific capital and production costs. This is particularly impor-tant in view of the opposing effect of increas-ingly stringent environmental protection regu-lations. The measures required for conforming with these restrictions inevitably increase the complexity of the plant, introducing energy and capital cost penalties.

Economics of Voluntary Production. The relationship between specific capital cost and unit production capacity can be seen from Fig-ure 36 for a sulfur-burning double-absorption sulfuric acid plant. The production of sulfuric acid from elemental sulfur represents the most economical and flexible option with respect not only to capital costs but also operating expendi-ture (utilities, manpower, absence of byproducts apart from steam). As a general rule, the sulfuric acid production costs in a large plant are deter-mined chiefly by the raw materials, which means the price of sulfur, since the credit for steam pro-duction usually covers all other costs (such as electric power, cooling water, labor, deprecia-tion, interest, administration, and maintenance).

In contrast, a sulfuric acid plant based on the roasting of pyrite is not only about twice as ex-pensive to build (if the cost of the roasting and

gas-cleaning systems are included) but also uses about twice as much electric power and about twice as much manpower as a sulfur-burning plant. Nonetheless, in cases involving domes-tic pyrite deposits that are economical to exploit for their nonferrous metals content, pyrite often represents a cheaper raw material source for sul-furic acid manufacture than expensive imported sulfur.

... 8 LA_ 6 [ z s c 15 éi, c Q 0

[1...

3 _

...,:,..- 0. 2 — u n

171 i__. ° 1 i . LLI11 , I I . .1

100 200 300 500 700100 20.00 Plard Eapacity, t/d

Figure 36. Relationship between specific capital cost and production capacity for a sulfur-burning double-absorption sulfuric acid plant

Without assigning a credit value for the cin-der, the cost of producing sulfuric acid from pyrite is roughly comparable to that of produc-ing it from elemental sulfur if the sulfur content of the pyrite costs ca. 25 % less than elemental sulfur. Depending on its chemical composition and prevailing market conditions, the byproduct cinder may exert either a positive or a negative influence on total costs. If the cinder can be uti-lized in a nearby cernent fatory, for example, the credit may cover up to 15 % of the pyrite costs. In the least favorable case, the cost of remov-ing and disposing of the cinder might constitute an additional burden on sulfuric acid production costs.

Economics of Involuntary ("Fatal Acid") Production. In any economic feasibility study, sulfuric acid obtained in the processing of metal-lurgical off-gases as a compulsory byproduct of a nonferrous-metal smelter operation (e.g., cop-per, nickel, zinc, lead, or molybdenum) must be valued differently from sulfuric acid of the same quality produced as a primary product in a plant based on sulfur combustion or pyrite roasting [125].


Sulfuric acid plants based on metallurgical off-gases are essentially pollution control plants, and they must operate in accordance with the production requirements of the accompanying smelter complex. Under such conditions it is not appropriate to expect the sulfuric acid plant to operate profitably, merely that proceeds from the sale of sulfuric acid should, if possible, cover the off-gas desulfurization costs. The capital costs of such a plant (including the wet gas-cleaning system as well as the sulfuric acid plant proper) depend heavily on the sulfur dioxide concentra-tion of the gases to be processed (Fig. 37 A) [40].

[I O 70

60 r 50

40 3fi

20

10 -

op 4 12 16

5.0 2 tonterdration of feed gas, vcd% •

a ___ b

A 4. '")

cl.

S Concerirrahun of feed gas,

Figure 37. Capital cost (A) and operating cost (B) for a double-absorption sulfuric acid plant and wet gas - cleaning system as a function of feed gas sulfur dioxide concentration a) Total; b) Sulfuric acid plant; c) Wet gas - cleaning system

If the sulfur dioxide concentration and the gas volume of the off-gas stream are subject to fre-quent wide fluctuations, the capital cost for the gas-cleaning system and the sulfuric acid plant

may approach that for a complete pyrite plant with a comparable acid capacity.

The current price disadvantage in the mar-keting of large quantities of sulfuric acid as a compulsory product (which frequently reflects a substantial freight cost owing to unfavorable location of the production facility, in relation to the market) is compensated by the fact that the sulfur dioxide content of the smelter off-gases is charged at zero cost in the cost calculation. How-ever, in contrast to sulfur-burning and pyrite-based plants, there is no steam credit in a met-allurgical sulfuric acid plant unless high-grade sulfur dioxide-containing smelter off-gases can be anticipated in the future [57].

The gas-processing costs and specific oper-ating costs show a dependence on the feed-gas sulfur dioxide concentration similar to the capi-tal costs, since capital-dependent costs account for most of the processing costs (Fig. 37 B).

4.2. Production by Nitrogen Oxide Processes

In a nitrogen oxide process, sulfur dioxide is oxidized to sulfuric acid in an aqueous phase, with nitrogen oxides serving as oxygen carriers. Processes of this type include the lead cham-ber processes (e.g., Mills — Packard and Gail-lard) and the tower processes (e.g., Curtius, Pe-tersen, Kachkaroff, Matignon, Salsasserra, and Moritz). These processes formerly constituted the basis for sulfuric acid production [19-23], but they have now been replaced almost com-pletely by the modern contact processes (see Chap. 3). Nevertheless, nitrogen oxide processes are of interest not only from the point of view of the history of chemical theory, technology, and industrial practice, but also because they have certain characteristics that may qualify them for use even today (in an improved design adapted to specific cases), despite certain general disad-vantages relative to contact processes.

Chemistry. Nitrogen oxide processes are based on oxidation of sulfur dioxide in the liq-uid phase. The detailed reaction mechanism is not fully understood even today, though this is one of the oldest industrially applied processes. In a simplified form, the course of reaction can be described by the following equations:

90

• 70 . E

▪

-

60

50 0 L' 40

cri 20 10

12 16 VO UV. —

Tan gas SO 2 Feed gas

SO, + 2NOFISO4 + 2 H,0 31-1,SO4 + 2N0 (16) (17) NO+1}20 2 NO2

NO2 + NO — N201 (18) (19) N,03 + 2 H,S0, — 2 NOHSO, H + 20

(20) SO, + 1/2 0. + H2 O H 2504

Sulfur dioxide absorbed in the aqueous phase is actually oxidized by nitrosyl hydrogensulfate, NOHSO4 ("nitrosylsulfuric acid"), dissolved in ca. 70 % sulfuric acid ("nitroso acid"), yielding sulfuric acid and releasing gaseous nitric oxide, NO. The colorless nitric oxide, which is virtu-ally insoluble in aqueous systems, is oxidized by free oxygen in the gas stream to form the yellow nitrogen dioxide, NO2 . Nitric oxide and nitro-gen dioxide are in equilibrium with the relatively unstable nitrogen sesquioxide, N203, which to some extent dissolves in sulfuric acid.

In sulfuric acid of greater than 70 % strength, nitrogen sesquioxide reacts to form nitrosyl hy-drogensulfate, which is a relatively stable com-pound at higher sulfuric acid concentrations. In ca. 70 % sulfuric acid, however, nitrosyl hydro-gensulfate is metastable and behaves as a very reactive oxidant toward sulfur dioxide. For this reason it is not possible to produce sulfuric acid of significantly higher concentration without ac-cepting a higher residual nitrogen oxide content.

Application as a Pollution Control Process. The low concentration of the product acid, which


(for reasons explained above) cannot econom-ically exceed ca. 78 %, constitutes the overrid-ing disadvantage of any nitrogen oxides process. Sulfuric acid of this concentration has relatively few applications today. It is not economical to concentrate it because of the high energy costs involved. Viewed strictly as a method of manu-facturing sulfuric acid, therefore, the nitrogen oxides process is no longer competitive with contact processes. It has the further disadvantage that none of the heat released is at a sufficiently high temperature to permit its economical recov-ery.

On the other hand, the low working tempera-tures of the nitrogen oxide process offer certain advantages in specific cases. Thus, it is suitable for processing gases with extremely low sulfur dioxide contents, which cannot be converted in a contact plant without supplementary heating. The nitrogen oxide process has also been pro-posed by Tyco as a method for removing both sulfur and nitrogen oxides from flue gases [126].

Ciba — Geigy Process. Since 1974, Ciba — Geigy has been developing a nitrogen oxide sul-furic acid process specifically designed for the processing of gases containing ca. 0.5 — 3 vol % SO2 and offering special advantages when the gases are laden with nitrogen oxides [127].

A flow diagram of the Ciba — Geigy nitro-gen oxide process is shown in Figure 38. The

Nitric acid

Produch acid

Figure 38. Ciba— Geigy nitrogen oxide sulfuric acid process a) Venturi scrubber; b) Star cooler; c) Product acid denitrator; d) Gas dryer; e) Cooler; f) Blower; g) Denitration tower; h) SO2 oxidation reactors; i) Nitrogen oxide absorbers; j) Concentrator


sulfur dioxide-containing gases are first passed through a wet gas-cleaning and cooling system to remove dust and reduce the water-vapor con-tent. In a subsequent gas dryer, surplus mois-ture is absorbed in 50 — 60 % sulfuric acid and transferred via the outer acid loop to a concen-trator, where it is evaporated into the tail gases. In principle, this is quite comparable to the Lurgi predryer — reconcentrator system used to main-tain the water balance in a contact-process acid plant (see page 15).

The product denitrator serves as a final nitro-gen oxides stripping stage for the product-acid bleed stream. Nitrogen oxides are removed from the irrigation acid partly by reaction with sulfur dioxide (reaction 16) and partly by hydrolysis (the reverse of reaction 19). Sulfuric acid leav-ing the tower is free of nitrogen oxides, and has a concentration of 70 — 80 %, preferably 78 % (60° Bé). Part of the sulfur dioxide feed already has been removed when the gases pass on to the denitrator.

In the denitrator up to ca. 50 % of the sul-fur dioxide feed is oxidized by nitrosyl hydro-gensulfate present in recycled acid from the first nitrogen oxide absorber. Nitric oxide is released into the gas stream. The incoming acid is heated in a recuperative acid heat exchanger, while acid leaving the denitrator is cooled before being re-cycled to the second nitrogen oxide absorber.

The two sulfur dioxide reactors serve to re-oxidize nitric oxide and remove remaining sul-fur dioxide. The transformation of sulfur diox-ide into sulfuric acid takes place at the extensive gas — liquid interface presented in the irrigated reactor towers by packings of a special design. Reaction heat from the oxidation of sulfur diox-ide is removed by cooling the circulating acid.

In the two nitrogen oxide absorbers, nitrosyl hydrogensulfate is formed by the reaction of ni-trogen oxides present in the gases with sulfuric acid recycled from the denitrator via the inner acid loop. The gas and the irrigation acid flow cocurrently in the first absorption tower and in countercurrent fashion in the second. Acid dis-charged from the second tower is pumped to the preceding nitrogen oxide absorption tower after cooling.

Efficient absorption of the nitrogen oxides de-pends not only on adequate cooling of the irri-gation acids but also on precise control of the NO/NO2 ratio, which is monitored continuously

at the outlet of the nitrogen oxide absorption stage by automatic analyzers. To ensure max-imum nitrogen oxide absorption, the plant is de-signed and operated such that gases entering the absorption towers contain nitric oxide and ni-trogen dioxide in almost equimolar quantities, which favors the formation of nitrogen sesquiox-ide. Nitric oxide in excess is scarcely absorbed at all. On the other hand, nitrogen dioxide in ex-cess does dissolve to some extent, but it is partly converted back into insoluble nitric oxide. Some autoxidation of nitric oxide to nitrogen dioxide takes place in the absorption towers, though this is a slow reaction.

Sulfur dioxide remaining in the gas at the ni-trogen oxide absorption stage will again liber-ate nitric oxide. A high sulfur dioxide content in the entering gases therefore causes a dispro-portionate rise in nitric oxide losses to the tail gases. Consequently, a high sulfur dioxide con-version efficiency in the sulfur dioxide reaction step is prerequisite to minimizing nitrogen oxide concentrations in the tail gases. The final sulfur dioxide content is actually lower than 100 ppm by volume.

Up to ca. 1 g/m3 of nitric oxide and nitro-gen dioxide may remain in the tail gas, but this is not the only cause of nitrogen oxide loss in the overall plant. Side reactions also take place in which nitrogen oxides are converted to such species as nitrous oxide, N20, and elemental ni-trogen, while a certain amount of nitrosyl hydro-gensulfate is discharged with the product acid.

Nitrogen losses must be compensated by the addition of nitric acid to the sulfur dioxide re-actors in order to maintain the nitrogen oxide balance in the entire system. Another possibil-ity is the addition of nitrogen oxides generated by burning ammonia. If the sulfur dioxide-con-taining feed gases already contain nitrogen ox-ides, the stoichiometric equivalent of nitric acid or ammonia can be saved.

The first commercial plant based on the Ciba — Geigy nitrogen oxides process has been in operation near Goslar (Federal Republic of Germany) since 1979. It is designed to desulfur-ize 10 000 m3/h of off-gases containing ca. 0.8 —1.5 % SO2 from a molybdenum sulfide roasting unit. A characteristic of this plant is the almost exclusive use of plastics as construction materi-als. In comparison with earlier tower plants, the specific reaction-space requirements (in relation


to gas throughput) have been substantially re-duced. An elaborate automatic measuring and control system restricts nitrogen oxide emis-sions to a minimum.

4.3. Regeneration of Spent Sulfuric Acid

4.3.1. Introduction

Most of the sulfuric acid produced enters not into marketable end products, but instead ends up as spent acid, impure gypsum, or other metal sul-fates, ail of which are waste products. Disposai of these chemical wastes represents an unsolved environmental problem that is becoming more serious as production capacities expand and an-tipollution regulations become more stringent.

The bulk of the spent acid is generated by the organic chemical and petrochemical industries. The nature of the acid varies widely from highly dilute to relatively concentrated, and with both low and high levels of organic and inorganic im-purities. Large volumes of uniform spent acid arise only in the course of plastics production (e.g., methacrylate, caprolactam); typical data related to the quantity and composition of re-presentative spent acids are as follows:

Composition figures for alkylation acid: Sulfuric acid content ca. 88-90 % Water content ca. 5 % Organics ca. 3 —4 % Organic sulfates ca. 2 %

Production and consumption figures for caprolactam manufacture: Caprolactam production rate 120 000 t/a H2 SO4 consumption (100 % acid) 636 kg/t of caprolactam Oleum consumption 1 300 kg/t of caprolactam Ammonium sulfate production 312 000 t/a (including

spent acid) Spent acid composition: Sulfuric acid ca. 50 % Salt content ca. 0.7 % Water balance

Production and consumption figures for methyl methacrylate (MMA) manufacture: Methyl methacrylate production 50 000 t/a H2 SO4 consumption (100 % acid) 1.8 t per tonne of MMA Spent acid production 2.7 t per tonne of MMA Spent acid composition: Sulfuric acid ca. 20 % Ammonium hydrogensulfate ca. 45 % Water ca. 30 % Organics ca. 5 % Solids ca. 2 %

An important source of spent acids con-taining inorganic impurities is the inorganic chemicals industry, especially the TiO 2 pigment industry, which generates spent acid of a rela-tively uniform quality containing ca. 20 — 23 % H2 SO4 and 7 —15 % metal sulfates. Similar in quality is the spent acid obtained from pickling operations in the metallurgical industry. Another source of waste acid is the gas-cleaning section of a roasting or other "metallurgical" acid plant.

A typical analysis of spent acid obtained from Ti02-pigment manufacture is as follows:

Sulfuric acid 22% Iron sulfate 4.5% Calcium sulfate 0.1 % Other metal sulfates 5.4 % Water 68%

In the past, typical methods of disposal or re-cycling for spent sulfuric acids included

1) Neutralization with lime 2) "Ocean dumping" 3) Reconcentration/blending with oleum

Stricter pollution-control regulations together with more limited disposai areas have now vir-tually eliminated the use of such neutralization and dumping methods in the industrialized coun-tries. Reconcentration of weak acids by blending with oleum also has its practical limits. Indeed, the method is suitable only for small acid vol-umes involving high H2 SO4 concentrations, and even then only if the specific impurity content is such as to permit recycling or use of the acid in some other process.

Regeneration of spent sulfuric acids is there-fore the only method available for obtaining a product of original quality and at the same time complying with environmental regulations, al-though the cost of acid derived from such a pro-cess greatly exceeds the production cost in a "normal" sulfuric acid production plant.

The regeneration of spent acid normally com-prises two major process steps:

1) Concentration of the spent acid to the high-est feasible level, together with separation of salts to the extent necessary

2) Decomposition of the spent acid and/or sul-fates

No single procedure can be considered uni-versally suitable for concentrating ail types of


spent sulfuric acid. The choice of an evapora-tion method also depends on the extent to which the specific application for which the regener-ated sulfuric acid is ultimately destined is able to tolerate residual impurities.

The process of concentrating of dilute sulfu-ric acid consumes substantial amounts of heat energy, not only to evaporate water but also for dehydrating the acid (see Section 2.1). For con-centration at atmospheric pressure at or near the boiling point it must also be borne in mind that the boiling point of sulfuric acid rises steeply as the H2 SO4 concentration increases.

The combination of high acid temperature, intermediate H2SO4 concentrations, and certain types of impurities can give rise to extremely corrosive conditions, necessitating the use of special construction materials for the concen-tration equipment.

Spent sulfuric acid can be concentrated by direct or indirect heating, and in one or several stages. The methods can be classified into two large groups [128]: 1) Evaporation to an H2SO4 concentration

< ca. 75 wt % 2) Evaporation to high concentration (up to the

azeotropic point, 98.3 % H 2 SO4 at atmo-spheric pressure) In concentrating to less than 75 % H2SO4,

water is almost the only substance evaporated. The H2 SO4 content of the gas phase is negli-gible on account of the low H 2 SO4 vapor pres-sure (see Section 2.1), so a scrubbing system for removing sulfuric acid is seldom necessary. Evaporating to a higher concentration (up to the azeotropic point) requires a relatively high tem-perature. However, the specific evaporation-heat requirement in the higher concentration range is relatively low because little water remains. For example, if spent acid containing 20 % H2 SO4 is concentrated to 98 % in two stages — from 20 % to 75 % and then from 75 % to 98 % — ca. 92 % of the total amount of water removed evaporates in the first step, leaving only 8 % to be evaporated in the second.

A more detailed description of the various ap-proaches to reconcentration is presented in Sec-tion 4.3.2.

The second and more fundamental regener-ation step is decomposition of the spent acid and/or sulfates. Such sulfates may be present

in the acid itself (e.g., as ammonium hydro-gensulfate) or they might be isolated in the form of free metal sulfates (e.g., FeSO 4 • H2 O from TiO2-pigment spent acid). The selection of an appropriate decomposition procedure is crucial for satisfactory operation, where the composi-tion of the acid and the solids/salts content are the determining factors. Three process concepts that have proven their worth in industrial-scale applications are:

1) Gas-phase decomposition in an empty decomposition furnace

2) Decomposition above a red-hot coke bed in a rotary kiln (Grillo process)

3) Decomposition in a fluidized bed

For more detailed descriptions of the various decomposition processes, see -+ Sulfur Dioxide, Chap. 5.6.2.

A decomposition plant of this type is nor-mally found in conjunction with a gas-cleaning plant and a sulfuric acid contact plant, such that the final product is pure sulfuric acid.

The spent-acid regeneration plant realized in Sachtleben Chemie is outlined in Fig. 39 [182]. It represents an integrated regeneration process for spent acid from a TiO2-pigment fabrication process.

Spent acid as received from the TiO2 produc-tion facility is first concentrated from ca. 22 % H2SO4 to ca. 70 % in a vacuum concentration plant. The 70 % H 2 SO4/FeSO4 suspension is then passed through the "maturation station," representing the 2nd stage of the process. Pa-ter maturation the suspension passes through a filter plant separating the acid from the sulfates (stage 3). The sulfates recovered from the fil-ter are prepared for decomposition, mixed with pyrite and carbon, and passed to an intermedi-ate storage facility (stage 4). The filtrate (sulfuric acid with 70 % H2 SO4) is further concentrated to ca. 80 % H2SO4 in another vacuum concentra-tion plant (stage 5). The acid that results is ready for immediate use in the TiO2 attack section. In stage 6, the mixture of sulfate/carbon/pyrite from stage 4 is fed into a fluidized-bed roaster for decomposition. The SO2 gas is cooled, cleaned, and transformed into sulfuric acid in a contact acid plant.

rProcess sine 1 Vacuum concentration 22 % to 70 % H 2SO 4

ProLess stage 2 Maturation

Spent from T i0 2.- maniff ac t Lee , pprox 22 % H 2SO 4

Prote55 stage Fil 'ration

Pr ocess stage 4 Mixing/ storage

Process stage 6 Dec omposition/ H 2 S0, -manufacture

Process stage S Vacuum concentration 70 fo .8 0€1/0 H 2SO 4

H 2SCI„ 80 % te TiO 2 ttack sechon


Figure 39. Integrated spent-acid regeneration process

4.3.2. Reconcentration to 70-75 % H2 SO4

Vacuum Evaporation. In this process, dilute sulfuric acid is maintained at the boiling tem-perature under vacuum by indirect heating with steam. The vapors evolved are recondensed al-most completely in a downstream cooling sys-tem so that, apart from insignificant amounts of inert gases desorbed from the acid, there is no gaseous effluent. Other dissolved volatile impurities that may be present, especially in organic-contaminated acids, can generally be condensed as well. Heavily contaminated acids containing inorganic substances, which tend to deposit solids during evaporation, are effec-

tively concentrated in forced-circulation evap-orators, whereas for comparatively pure spent acids a natural circulation system may be ade-quate [128]. Evaporators of this type are satis-factory for concentrating dilute sulfuric acid to 70 — 80 % [129].

Forced- Circulation Evaporator. In the forced-circulation evaporator shown in Fig-ure 40, the heater is separate from the flash vessel. With this arrangement it is possible to prevent evaporation in the heater itself, so solids deposition on the heating surfaces is kept to a minimum. An additional advantage of forced circulation over natural circulation is that the liquid-flow velocity in the heat exchanger is

C onde ri sa t e

Rec clin[ irated

t. 1.0 I

E

•

0.8 e -U 0 Éu

â

▪

0.6

ou

▪

4_ CL

▪ 0.2

Figure 41. Theoretical specific steam consumption as a func-tion of the number of effects

E n3'


higher and more easily controlled; as a result, the heat-transfer conditions are more stable.

Figure 40. Forced-circulation evaporator with separate heater a) Flash vessel; b) Heater; c) Circulating pump

A constriction of the inlet to the flash vessel keeps the acid under a slight pressure so that, although it is heated to 3 – 4 °C ab ove the out-let temperature of the flash vessel, water cannot evaporate until the acid is free to expand into the evaporator. Delayed boiling in the evapora-tor itself can be largely prevented by efficient distribution of the acid within the evaporator space. Owing to the relatively high rate of circu-lation, precipitated solids remain in suspension and are discharged from the system along with the reconcentrated acid.

Multiple -Effect Evaporators. The energy re-quired for water evaporation is normally sup-plied in the form of low-pressure steam. The steam requirement can be significantly reduced, especially in large-capacity evaporation sys-tems, by coupling several evaporator units to-gether in such a way as to reuse the heat content of the vapors. In the first evaporator unit, heat is supplied in the normal way by indirect steam heating of the acid circuit, but in each subse-quent evaporator unit the acid circuit is heated indirectly by vapors produced in the preceding stage. Evaporation systems arranged in this man-

ner are known as multiple -effect (or multistage) evaporators (—> 3. Evaporation, Chap. 2.2.).

The relationship between specific steam con-sumption and the number of thermally coupled stages in a multiple-effect evaporator is illus-trated graphically in Figure 41. A single evap-orator unit (i.e., a single-effect system) requires ca. 1.2 t of low-pressure steam per tonne of evap-orated water. In a double-effect system this re-quirement is reduced to only about 0.6 t. Adding a third effect leads to a further reduction, to ca. 0.4 t of low-pres sure steam per tonne of water evaporated – in other words, about one-third of the steam consumption of a single-effect sys-tem. The alternative of dividing the evaporation system into several concentration steps without such coupling (i.e., use of live-steam heating in each unit) does not reduce the specific steam re-quirement, though it does reduce the required to-tal heating-surface area relative to a single-effect system.

Vacuum and Condensation Systems. Multi-ple-effect evaporator systems are normally op-erated under vacuum. Depending on the relative cost and availability of electric power, steam, and cooling water, the vacuum may be produced by means of electric vacuum pumps or steam ejectors. When steam ejectors are used, several units are arranged in tandem to generate a rela-tively low pressure. The effect is enhanced by the installation of preceding and intermediate water-cooled vapor condensers.

Live steam con de nnate


Cooling water

Figure 42. Forced-circulation vacuum evaporation system for concentrating sulfuric acid from 20 % to 70-80 % H2 SO4 a) Evaporator heaters; b) Plate heat exchangers; c) Evaporators; d) Acid degasser; e) Ejectors; f) Contact condenser

If spent sulfuric acid containing volatile con-taminants is to be concentrated it is best to use surface condensers (tube-bundle heat ex-changers with cooling water flowing through the tubes), because with this type of system it is easiest to separate impurities from the vapor condensate. Where the contact (quench) type of condenser with direct water injection is used, the volume of wastewater can be reduced by recir-culating part of the condensate.

Construction Materials. Evaporators de-signed to concentrate sulfuric acid up to 75 % can be fabricated from fiberglass-reinforced polyester or plastic-coated, rubber-lined, enam-elled, or lead-coated steel. At higher acid con-centrations and temperatures it is possible to use steel lined with FEP sheeting (—> Films, Chap. 6.3.3.) and protected with two layers of carbon bricks. The heat exchangers are usu-ally made of graphite impregnated with phenol resin, in which case they can be operated at wall temperatures < 165 °C and acid concentra-tions < 84 % H2SO4. Cast ferrosilicon pumps in common use today are capable of withstand-ing temperatures < 140 °C and a sulfuric acid concentration of ca. 75 %. Special construction material is required for high-stress conditions.

Complete Plant. Figure 42 is a flow diagram of a typical Lurgi-designed plant for vacuum concentration of spent sulfuric acid containing 20 % H2SO4 to an H2SO4 concentration of 78 % [84]. It comprises a total of four evaporator units. The acid is preconcentrated from 20 % to 54 % H2 SO4 in the first three units, which are ar-ranged as a triple-effect system. Final concen-tration to 78 % takes place in a fourth, single-effect stage with live-steam heating. A pressure gradient is maintained across the first three units (a triple-effect stage) to ensure that water con-tinues to evaporate despite a decreasing temper-ature. Thus, the first stage operates virtually at atmospheric pressure, whereas the third runs at a much lower pressure —70 mbar, for example, at which point the vapor-condensation temper-ature is ca. 40 °C. Live steam at 5.5 bar is used for heating the first and fourth stages. This might be derived, for instance, from 3-bar and 28-bar steam, using a steam ejector followed by a steam saturator.

Dilute feed acid is preheated in plate ex-changers with the vapor condensate and con-centrated production acid. It is then degassed in a bubble-tray column by stripping with va-por from the final concentration stage; this also


serves to recover any sulfuric acid that may still be present in vapors from the final concentration stage. The degassed dilute acid is preheated fur-ther in plate exchangers with vapor and steam condensate from the evaporator heaters before being introduced into the acid circuit of the first evaporator.

In the first stage, heated by live steam, water is evaporated from the acid at about 120 °C. The steam supply is adjusted to maintain an acid con-centration in the first stage of ca. 25 % H2 SO4, where the acid boiling point is used as the con-trol parameter. Indeed, the boiling-point eleva-tion above the boiling point of pure water at a predetermined pressure can be used as an exact measure of the H2SO4 concentration provided the influence of dissolved contaminants is taken into account. In the second stage, where the acid is heated indirectly by vapors from the first stage, more water is evaporated at about 95 °C despite the higher acid concentration; this is possible because of the lower pressure. Vapors from the second stage provide heat for acid circulating in the third stage, where it is concentrated to about 55 % at ca. 65 °C under an appropriate vacuum. As mentioned previously, live steam is again used for final concentration to 78 % in the fourth stage.

Because of the pressure gradient, the acid need not be pumped from one stage to the next. After the fourth stage, the concentrated acid is cooled indirectly with incoming dilute acid in a plate heat exchanger and than pumped to storage tanks. Vapors from the third-stage evaporator are mixed with scrubbed vapors from the final evap-orator as they leave the degasifier column, and then passed through a surface condenser. Uncon-densable gases emerging from this condenser are exhausted through a multistage vacuum system consisting of steam ejectors and auxiliary sur-face condensers, and are finally discharged to the atmosphere.

The steam condensate from the heating ele-ments of the first and fourth evaporator stages is collected, expanded, and cooled through indi-rect heat exchange with the dilute acid before be-ing discharged. Vapor condensate from the sec-ond and third stages is similarly utilized for pre-heating the spent acid in plate exchangers, after which it is mixed in a collector with condensate from the condensers of the vacuum system. The residual heat of the accumulated vapor conden-

sate is utilized for initial preheating of the spent acid.

The process thus takes maximum advantage of the heat content of the vapors by means of a multiple-effect arrangement of the evapora-tors, and it also recovers the maximum possible amount of waste heat from the condensate for preheating purposes. This is of vital importance in view of the continuously rising cost of energy.

The maximum water-evaporating capacity of a single-train multistage unit with forced-circulation evaporators is presently in the range of 8 —10 t of water per hour.

Venturi Reconcentration Process. The venturi sulfuric acid reconcentration system de-veloped by Lurgi is preferably integrated ther-mally into a sulfuric acid production plant in such a way as to make use of waste heat and thereby avoid the cost of heating with steam or fossil fuels. The system is designed to utilize sensible heat from the tail gases as well as the relatively low-level heat content of acid circu-lating in the sulfur timide absorption system of an acid production plant [122, 130].

The most significant feature of the venturi reconcentration process is that—at atmospheric pressure — water is evaporated from dilute sul-furic acid at a temperature far below the boil-ing point of the acid solution. Gases passing through the venturi system become laden with water vapor in proportion to the water-vapor par-tial pressure of the acid (which is a function of the H2 SO4 concentration and the temperature), carrying this water out of the reconcentrator sys-tem for subsequent discharge to the atmosphere. Venturi systems have long been used in gas-cleaning plants as gas scrubbers, but it was only in the early 1970s that Lurgi developed venturi systems specifically designed for acid concen-tration. These have meanwhile rendered satis-factory service in various industrial plants.

The concentration unit consists of a venturi scrubber into which the acid is injected cocurrent with the gas flow, followed by a tower in which dilute spent acid and gas interact in countercur-rent. Despite a short gas retention time, water becomes equilibrated between the acid and the vapor because the two mix so efficiently and in-timately in the venturi unit. The carrier gases acquire a higher moisture load in such a two-stage evaporation system with two separate acid

Dry tais gas

8.5 % Sulfuric acid from final absorber

Ti final àbsorber

From intermediee absorber 98.5 % sulfuric acid

To intermediate absorber

FliSt tait gas ta stack


Figure 43. Two-stage venturi reconcentration system utilizing acid waste heat from a double-absorption sulfuric acid plant a) Acid heat exchangers; b) Dilute acid tank

circuits running at different H 2 SO4 concentra-tions than they would in any single-stage system, so the water-evaporation capacity is greater for the same gas throughput. At the same time, the H2 SO4 content of discharge gases remains ex-tremely low, even when the sulfuric acid concen-tration is raised to more than 70 %, on account of the lower concentration of acid circulating in the second evaporation stage. It has been demon-strated in practice that venturi reconcentrators also function well as tail-gas scrubbers, substan-tially reducing the natural H2SO4 content of tail gases from a sulfuric acid plant [75]. The heat required for dehydration and water evaporation during concentration of the acid is supplied di-rectly to the system as the sensible heat content of the carrier gases as well as by indirect heat-ing of the concentrator acid circuits with product acid from the main acid plant. The concentrat-ing capacity can be increased by introducing air heated with excess heat from the contact plant along with the hot, dry tail gas from the sulfuric acid plant.

The temperature demand associated with heat suitable for a venturi reconcentrator system, which has a comparatively low operating tem-perature, is quite compatible with the substantial amount of waste heat available in acid circulat-ing in the sulfur triœdde absorption system of a sulfuric acid plant, which must in any case be dissipated. The venturi reconcentrator system is thus a very good way of making economical use of low-level waste heat from a sulfuric acid plant.

If the concentrator unit is thoroughly integrated into the waste-heat system of the acid plant, ca. 22 — 26 t/h of water can be evaporated with the waste heat from a 1000-t/d sulfuric acid plant.

Figure 43 is a diagram of a two-stage venturi reconcentrator designed for evaporating waste sulfuric acid originally containing 30 % H2SO4 until it reaches a concentration of 70 %. Tail gases leaving the final sulfur triacide absorber of the sulfuric acid plant at ca. 80 °C are passed directly into the venturi reconcentrator. Circu-lated acid, at a concentration of about 70 % and also heated to ca. 80 °C, is injected at the top of the venturi unit in cocurrent with the gases. Adiabatic evaporation of water cools the gases and acid to about 74 °C. A special nozzle grate installed in the outlet of the venturi section pro-vides additional turbulence and assists in subse-quent separation of gas from the liquid phase.

Gases leaving the venturi section typically have a water content of ca. 20 g/m 3 . They are in-troduced into the bottom of a subsequent tower through which they rise in countercurrent with a spray of 55 % sulfuric acid from the second acid circuit, again heated to about 80 °C. Within the tower the water content of the gases increases to about 80 g/m3 . A droplet separator at the tower outlet removes any remaining droplets from the gases, which are subsequently discharged to the atmosphere via a stack. Acid accumulating in the tower sump becomes cooled to ca. 70 °C as a result of the water evaporation. Acid in both concentrator circuits is heated in Teflon tube-


bundle heat exchangers by 98.5 % sulfuric acid at ca. 100 °C derived from the intermediate and final absorbers of a double-absorption sulfuric acid plant.

The dilute acid subject to concentration, which is about 30 % in 112SO4, is fed to the tower sump of the second stage. Preconcentrated acid from the second stage simply overflows from the sump of the tower into the sump of the first-stage venturi unit. The 70 % reconcentrated acid like-wise overflows from the sump of the venturi unit to product storage.

Operation of a venturi reconcentrator system is quite straightforward. The feed rate of the spent acid is controlled relative to the acid con-centration in the venturi circuit so as to maintain the desired product-acid concentration; no fur-ther control measures are required.

Venturi reconcentrators are designed for a gas-pressure drop of 15 — 20 mbar. Apart from power consumed by the supplementary blower capacity and the electric drives of the acid circu-lation pumps, no additional energy is required.

Since they are relatively compact, venturi reconcentrators are well-suited to retrofit instal-lation in existing sulfuric acid plants.

The first of the venturi acid-reconcentration systems has been in operation at Süd-Chemie AG at Kelheim, in the Federal Republic of Ger-many, since 1972. It concentrates dilute spent sulfuric acid to a level of about 70 %, and is as-sociated with a double-absorption sulfuric acid plant designed for a capacity of 220 t/d. The largest venturi reconcentrator unit, designed for a gas throughput of 200 000 m3/h, is thermally integrated into a 2000-t/d double-absorption sul-furic acid plant at the Inspiration Copper works in Arizona (United States). Venturi reconcentra-tors thermally integrated into sulfuric acid plants are also used for processing spent acids saturated with ferrous sulfate, encountered, for example, as scrubbing acids in the gas-cleaning systems of pyrite roasting plants.

Submerged -Combustion Process. Sub-merged-combustion processes are sometimes used for concentrating spent acids with high salt contents to 60 —70 % H2SO4 [128,131-134]. Liquid or gaseous fuels and a burner are used to generate hot flue gases at 1500 — 1600 °C that are forced through concentrated acid via an im-mersion tube. The flue gases are cooled by the

adiabatic evaporation of water, simultaneously absorbing water vapor at a rate that depends on the water-vapor partial pressure above the acid. Depending on the acid concentration, the bath temperature remains in the range of ca. 120 —150 °C. The flue gases, which leave at about the same temperature, contain up to ca. 60 vol % wa-ter vapor. Before being discharged to the atmo-sphere, these gases are cleaned and cooled in a liquid separator, a scrubbing and cooling tower, a wet precipitator and, if necessary, a special sul-fur dioxide scrubbing tower.

Figure 44. Submerged-combustion evaporator for gas or oil firing a) Oil inlet; b) Igniter; c) Air inlet; d) Acid inlet; e) Acid outlet; f) Acid circulating tube; g) Immersion tube; h) Vapor outlet; i) Combustion chamber; j) Gas inlet

Figure 44 is a schematic diagram of a sub-merged-combustion evaporator. The evaporator vessel is constructed from lead-coated steel, and in most cases it is brick-lined as well. The com-bustion chamber is located above the vessel, from which the burner projects downwards be-neath the surface of the acid. A circulating tube is positioned concentrically around the immer-sion tube, through which liquid is forced by the ascending flue gases. Apart from causing vig-orous circulation of the acid, this also promotes favorable mass and heat transfer between gas and liquid.

Fuel

Air


Spent acid is fed continuously either through a separate nozzle, as shown in Figure 44, or through a dephlegmator placed on top of the gas outlet [131,132]. A mixture of concentrated acid and precipitated salts leaves the evaporator through an overflow pipe.

This system of direct heat transfer from the flue gases to the acid largely avoids the incrusta-tions that would normally be deposited on heat-ing surfaces, especially during the evaporation of sait-laden spent acid. Nevertheless, the high flue-gas temperature and the consequent high wall temperature of the immersion tube impose very heavy demands on the construction mate-rials.

The jacketed combustion chamber is cooled by incoming combustion air passing through the jacket. The immersion tube, made of cast iron, is protected in its uncooled upper region by an internai insulating lining. The lower part of the immersion tube is cooled by the acid. The crit-ical zone with respect to corrosion and erosion is the outside of the immersion tube above the top of the circulating tube. With time, this part of the immersion tube gradually becomes eaten away to the point where the lower, submerged part actually falls off, limiting the service life of the tube to a maximum of ca. 2000 operating hours [133].

A submerged-combustion evaporator is use-ful, for example, for concentrating the contam-inated sait-bearing sulfuric acid arising from TiO2 pigment manufacture from 21 % to a final concentration of 65 —68 % H2SO4 [133]. About 27 m3/h of spent acid with a specific gravity of 1.3 kg/L is fed to an evaporator vessel with a capacity of ca. 3 m3 . At the designed water-evaporation rate of ca. 1.95 t/h, 3.1— 3.2 GJ of

energy in the form of fossil fuel is consumed per ton of evaporated water. However, the specific heat requirement of the submerged-combustion evaporation process can be reduced (in this case, for example, to about 1.9 GJ per tonne of water evaporated) [133] by utilizing the heat of con-densation of the vapors for concentrating spent acid in a preceding stage.

4.3.3. Concentration to 93-98 % H2 SO4

Vapors evolved during the concentration of spent sulfuric acid to a more highly concentrated state (93 — 98 % H2SO4) contain significant quanti-ties of gaseous sulfuric acid. It is therefore ab-solutely obligatory that the vapors be cleaned before discharge to the atmosphere. Concentra-tion may involve either direct or indirect heating, with the latter approach applicable both at nor-mal pressure and under vacuum.

Chemico Direct-Fired Drum Concentra-tor. The Chemico drum-concentrator process [135] was developed in the 1920s. It is used es-pecially for concentrating spent acids from the explosives industry. The principle is similar to that of the submerged-combustion method. Flue gases resulting from the combustion of gas or fuel oui are passed through an immersion tube into the acid to be concentrated, which flows countercurrent with the gases through several vessels arranged in cascade fashion (Fig. 45). The vessels are made of lead-lined steel pro-tected with up to three layers of acid-proof bricks [136, 137]. The immersion tubes are fabricated from ferrosilicon.

Ti de3r111:u

Coricentratecl acid E-51'3 % H25041 Spent acid % H 7S13,:

Figure 45. Chemico drum concentrator a) Combustion chamber; b) Concentrator; c) Gas cooler/acid preheater


Flue gases enter the concentrator at ca. 600 °C [128]. Because of the relatively low flue-gas temperature, there is practically no thermal decomposition of the sulfuric acid, so almost no sulfur dioxide is formed [131].

The concentrator is subdivided into two chambers containing acid at different concen-trations. In the first chamber, into which the hot flue gas enters and from which concentrated acid is discharged, the acid concentration is main-tained at 93 — 94 %. The gas then passes to the second chamber at 250 — 260 °C. A portion of the hot flue-gas stream may be introduced di-rectly into the second chamber for the purpose of controlling the acid concentration. Flue gases leaving the second chamber at 180 — 200 °C are cooled to 130 — 150 °C in a second vessel. The feed acid (usually at a concentration of 70 %) is introduced into this vessel. Discharged gases must be demisted and cleaned.

The Chemico process shares with the submerged-combustion method the advantage that it avoids heating surfaces prone to acquir-ing deposits during the concentration of salt-contaminated spent sulfuric acid; all the requi-site heat is supplied by the flue gases. More-over, water evaporation takes place not in the boiling range of sulfuric acid but about 50 °C below the atmospheric pressure boiling point [136]. The process is suited to the concentration of relatively large volumes of acid. Single-train plants have been designed with capacities of up to 18 t/h of 93 — 94 % acid.

A significant drawback of the Chemico pro-cess is the large amount of contaminated waste gas that is produced, the cleaning of which is very costly. A scrubbing tower and an electri-cal gas-cleaning system were formerly used for this purpose, but modern plants employ venturi scrubbers almost exclusively. These are com-bined with cyclones for separating droplets. The sulfuric acid content of the waste gas from such a plant is reported to be 75 — 200 mg/m 3 [135].

Spent acid may also contain organic impuri-ties, in which case these and their decomposi-tion products would be entrained by the waste gases and subject to only partial elimination in the waste-gas cleaning system. When processing spent acid containing nitrates or nitrogen oxides, NO/NO2 liberated during concentration would similarly not be removed by the gas-cleaning system, resulting in a visible plume at the stack.

In view of the increasing stringency of environ-mental regulations in many locations, the utiliza-tion of this process is becoming more and more problematic [137].

Concentration Processes Involving Indi-rect Heating.

Pauling—Plinke Process. In this process, spent sulfuric acid is concentrated to 95 — 98 % H2 SO4 under normal pressure by indirect heat-ing with flue gases. The process was developed by PAULING beginning in 1915, and the neces-sary equipment is built by Plinke [138].

Actual concentration takes place in a cast iron vessel (a "Pauling pot") suspended in a furnace. The furnace is surmounted by a distillation col-umn (in most cases a bubble-tray column) made of silicon iron (see Fig. 46). Vessels of this type are built with a useful volume of up to 12 m3 ; above about 3 e an anchor mixer (also made of silicon iron) is included [138].

Ta gas cleaning

d

Spent acid

Air

Figure 46. Pauling — Plinke pot concentrator a) Condenser; b) Stripping column

Spent acid (ca. 70 % H2SO4) is fed to the column from above, and it flows downward into the vessel in countercurrent with the exhaust va-pors. Incoming acid is concentrated to 82 — 85 % H2SO4 by heat and mass transfer with the va-pors; simultaneously, sulfuric acid present in the vapors is condensed almost completely. If the spent acid has an H2SO4 content greater than 70 % the acid is admitted at a tray lower in the column.

The sulfuric acid concentration in the vessel itself must be kept above 80 % to avoid corro-sion [136]. In most cases, the concentration of the boiling acid is maintained above 95 % (corre-sponding to a boiling temperature of ca. 330 °C). Under these conditions, the average corrosion


loss-rate is about 8 —10 mm per year [138]. A service life of ca. 3 — 5 years is achieved by de- signing the vessel with a sufficiently thick wall.

The acid is heated indirectly by flue gases passing around the vessel. These flue gases, generated by burning fuel oil or gas, enter the furnace cavity at 800 —1100 °C and leave it at ca. 450 °C. The mean heating-surface load of the vessel is ca. 144 000 khn-2 11-1 = 40 kW/m2 [138]. The high flue-gas temperature means that the wall material of the vessel on the furnace side is subject to heavy stress.

Flue gases leaving the furnace can be utilized for preheating the combustion air to ca. 200 °C. Part of the flue-gas stream is cooled to 200 —250 °C and recycled to the furnace cavity. Such combustion-air preheating and flue-gas recircu-lation raises the furnace efficiency (ratio of heat supplied to the vessel to applied fuel energy) to 80 — 83 %.

Hot, concentrated sulfuric acid is extracted continuously from the vessel. It is cooled indi-rectly against water in special coolers (pot cool-ers) [138], where most of the dissolved inorganic impurities are precipitated. Organic impurities in the feed acid are decomposed by oxidation within the evaporator: either by the high tem-perature alone or — if necessary — by the addition of nitric acid.

Vapors leaving the top of the discharge col-umn at ca. 130 °C still contain as much as ca. 1 % sulfuric acid vapor, together in some cases with sulfur dioxide, nitric oxide, and organic substances, as well as inert gases [128,138]. The water vapor is separated in a condenser, whereas noncondensable components are passed on to a waste-gas cleaning system.

All concentration processes involving indi-rect heating, including the Pauling — Plinke pro-cess, have the advantage over directly heated processes that the volume of resulting waste gases to be cleaned is low. However, they suffer the drawback of heating surfaces that are prone to scaling with precipitated impurities from the spent acid. This necessitates frequent checking and entails high maintenance costs [136].

In the Pauling — Plinke process, the ratio of heating surface to vessel volume, and thus the ratio of the attainable heat-transfer rate to the rate of acid production, is relatively low. Be-cause of the limited evaporation capacity of a single unit, high throughput requires several par-

allel units [136]. Typically, 0.6 — 1.2 t/h of water can be evaporated per vessel [137].

Falling-Film Evaporation Process. In this process, spent sulfuric acid descends as a thin film on the inner surface of an externally heated tube. This arrangement provides a favorable heat-transfer rate and a large surface from which water can evaporate. Evaporated vapors ascend in countercurrent with the acid, and are substan-tially freed from entrained sulfuric acid vapor and mist by the scrubbing action of entering spent acid in a column surmounting the evapo-rator, similar to that in the Pauling — Plinke pro-cess.

A falling-film evaporator process was for-merly operated by Du Pont under vacuum con-ditions (27 — 40 mbar), with indirect heating by steam at 25 — 30 bar [139].

A more recent process —the BOSAC process, developed by Bofors — operates under normal pressure with indirect heating by flue gases [137, 140-142]. When the emerging flue gases are used for preheating the combustion air, a thermal efficiency of 82 — 85 % is achieved [141].

Bayer — Bertrams Process. This process is applicable from atmospheric pressure up to about 5 bar; indirect heating is supplied by a heat-transfer medium [137, 143, 144]. The heat carrier is a salt melt (nitrate/nitrite) that is heated indirectly by flue gases in a separate furnace.

The modem version of the process uses tubes of borosilicate or pure quartz glass with a di-ameter of 150 — 400 mm. Spent acid is fed to the stripping column at a concentration of 65 —85 % H2 SO4 and is concentrated to ca. 96 — 98 at 300 — 330 °C. Organic impurities that may be present are decomposed almost completely, with assistance as necessary from an added oxidant.

A special advantage of the falling-film con-centration process is the favorable ratio of heat-ing surface to acid volume. By connecting a large number of tubes in parallel, an evaporation ca-pacity comparable to that of a Pauling — Plinke unit can be achieved. A falling-film evaporator plant composed of 12 tubes is reported to have a capacity of 1 t/h of 96 — 97 % sulfuric acid start-ing from 65 — 70 % acid, which corresponds to a water-evaporation rate per tube of ca. 50 kg/h.

Vacuum Processes. Processes in this cate-gory transform spent sulfuric acid directly into highly concentrated material. The principle is


comparable to that used in a vacuum process for evaporation to intermediate concentration. Prob-lems are encountered with the materials of con-struction, however, and it is essential that special cleaning systems be provided to remove evapo-rated sulfuric acid from the emitted vapors. It is not possible to make use of the multiple-effect principle because of the high temperatures re-quired. The acid is heated indirectly with high-pressure steam at every stage.

Simonson — Mantius Process. Developed in 1921, this is one of the most well-known vacuum methods for attaining high acid concentrations. Most of the corresponding plants are multistage units operating with either natural or forced cir-culation, depending on their design, and with separate heat exchangers or heating tubes in-stalled directly in the evaporator vessels [145]. The evaporator vessels themselves are made of lead-lined steel further protected by acid-proof bricks or a layer of graphite or enamel. The heat exchangers or heating tubes are made of lead, graphite, silicon iron, or tantalum.

The Simonson — Mantius process has been developed further by Chemetics International Ltd. [137, 146]. In the Chemetics version of the process, the evaporators are made of enamelled steel or, in the case of very small units, glass-lined steel, and they incorporate bayonet heaters made of tantalum. In its standard form the pro-cess leads to 93 % sulfuric acid, but it can be designed to provide concentrations of 95 — 96 % if desired [137].

The largest existing unit concentrates approx-imately 28 t/d of sulfuric acid (calculated as 100 % H2SO4) from 70 % H2SO4 to 93 %; this corresponds to a water evaporation capacity of ca. 10 t/h [137]. Only acids that do not contain solids, or which will not produce accretions dur-ing concentration, can be treated by this process.

Corning/ENS Process. The Corning/EIVS process is based on horizontal evaporators [147]. Spent acid flows continuously through a hor-izontal cylindrical vessel made of borosilicate glass. PTFE deflectors are used to prevent back-mixing. The acid is heated indirectly with 15-bar steam, which is passed through an internal tan-talum tube bundle. The operating temperature of the evaporators is limited to a maximum of 180 °C to restrict corrosion.

Spent acid can be concentrated from 65 % H2 SO4 to 88 % in a single stage operating at

an absolute pressure of 93 mbar. For triple-stage concentration to 95 % H 2 SO4 , pressures in the individual stages are 133, 40, and 13 mbar [147]. The steam consumption is reported to be ca. 2 t of steam per tonne of evaporated water, corre-sponding to ca. 4 GJ per tonne of H2 O.

4.4. Production of Oleum

Oleum ("fuming sulfuric acid") [8014-95- 7] consists of a solution of sulfur trioxide in 100 % sulfuric acid. Normally, the oleum concentra-tion is expressed in terms of the mass content of so-called "free" sulfur trioxide. Part of the dis-solved sulfur trioxide reacts with sulfuric acid to form higher sulfuric acids such as disulfu-ric acid, H2S207 (see Section 2.2). Pure sulfur trioxide (Chap. 8) corresponds theoretically to oleum with 100 wt % "free" sulfur trioxide.

Oleum is produced industrially in contact plants, where sulfur trioxide-containing gases are passed through a special oleum tower. In a double-absorption plant this is located upstream from the intermediate absorber; otherwise it is upstream from the final absorber.

Oleum should contain at least 20 — 22 % "free" sulfur trioxide to avoid problems with construction materials. The upper concentration limit of about 36 % is determined by the partial pressure of sulfur trioxide in the gases as well as the oleum temperature in the oleum tower, which in turn establishes the sulfur trioxide va-por pressure of the oleum (see Section 2.1 and Fig. 6). The dependence of oleum vapor pressure on temperature and concentration in the range of interest to industry is shown in Figure 47 [148-150]. Oleum formation by sulfur trioxide ab-sorption can proceed only so long as the oleum vapor pressure remains below the partial pres-sure of sulfur trioxide in the gas phase. It follows from the vapor-pressure curves that, if the oleum temperature is 60 °C and the partial pressure of sulfur trioxide in the contact gases is 115 mbar, for example, the theoretical maximum attainable oleum concentration is about 37 %. This corre-sponds to the sulfur trioxide partial pressure in gas at a total pressure of 1.12 bar produced in the first three beds of a (3 + 1) double-absorp-tion sulfuric acid plant operating on a feed gas initially containing 10 vol % SO2, assuming the sulfur dioxide conversion at this point is 95 %.

150

100

50

15 20 25 30 35 Oleum cuntentration, >d'Y. free SO


Figure 47. Vapor pressure of oleum as a function of free S03 content and temperature a) 80°C; b) 70°C; c) 60°C; d) 50°C; e) 40°C

The potential rate of oleum production is lim-ited by the water balance in the contact plant. This is determined by the rate at which water vapor enters the plant — with the combustion air in a sulfur-burning plant, or in roaster or decom-position gases in a "cold-gas" plant — and is sub-sequently absorbed into the acid circuit in the dryer, since sufficient sulfur trioxide must re-main in the process gas alter the oleum tower to maintain an acid concentration of 98.5 % in the absorbers. Therefore, the lower the rate at which water enters the acid system via the dryer, the higher the proportion of sulfur trioxide available for oleum production.

Oleum towers are similar in design to modern acid-plant main absorbers. However, because of the specific corrosion behavior of oleum, these towers can be fabricated from stainless steel without a brick lining Sulfur trioxide-ontaining gas enters the oleum-irrigated packed tower at ca. 200 °C. Depending on the desired product concentration, the tower is irrigated with 22 % or 35 % oleum at a temperature of ca. 40 — 50 °C. Oleum of higher concentration leaves the tower

at ca. 60 — 80 °C. It is adjusted to the desired final oleum concentration by addition of concentrated sulfuric acid from the intermediate or final ab-sorber of the contact plant, and is then cooled in special coolers. Product is withdrawn from the oleum circuit downstream from the cooler. Oleum of higher concentrations (e.g., 60 — 65 %) is obtained by mixing 20 — 36 % oleum with pure sulfur trioxide.

5. Construction Materials (-+ Construction Materials in Chemical Industry)

The traditional construction material for use in a sulfuric acid plant based on the lead cham-ber or tower process was, of course, lead. By contrast, modern contact-process plants are con-structed mainly from iron and steel. This is a striking example of the extent to which the se-lection of construction materials depends on the nature of a specific process and its associated op-erating conditions [151-153]. Convers ely, new processes frequently become practical only with the development of appropriate materials. This is particularly true of acid-heat recovery systems, which are of major importance in the further de-velopment of sulfuric acid technology.

Elevated sulfuric acid temperatures are essen-tial for the economical recovery of acid waste heat. This mandates the use in acid coolers and heat exchangers of corrosion-protected materi-als that are able to withstand attack not only by the acid itself (with the aid of anodic protection as appropriate) but also, given the elevated wall temperature, attack by the cooling fluid.

One general problem relates to the extent to which a particular material will display in ser-vice the same behavior that has been observed in a laboratory with respect to corrosion rates under specific conditions (as a function of oxygen con-centration and temperature, for example). Many additional parameters that influence corrosion behavior are relevant in industrial practice. For example, dissolved impurities — even in minute concentrations, which may be difficult to deter-mine analytically — as well as suspended solids and sulfuric acid flow velocities are all signif-icant, as is the treatment and processing of a construction material prior to its installation.


Data acquired through systematic laboratory tests based on pure media, and even empiri-cal data derived from practical experience, are therefore of only limited validity for predicting the behavior of a given material in a particu-lar application. General statements regarding the suitability of a construction material with respect to sulfuric acid over a defined range of concen-tration and temperature should never be taken too literally.

A comprehensive list of important empirical values is provided in the Dechema materials ta-bles [154].

For concentrations < 90 %, sulfuric acid is classified as a nonoxidizing acid, but concen-trated sulfuric acid and oleum both exhibit oxi-dizing properties. This is the reason, for exam-ple, why hot, concentrated sulfuric acid reacts with copper and silver to form sulfur dioxide rather than hydrogen:

2H2SO4 +M° -> MIISO4 +SO2 +2H20 (21)

Carbon is also oxidized by hot, concentrated sul-furic acid and oleum, again generating sulfur dioxide. Corrosion rates for metallic materials in sulfuric acid are determined largely by the solubilities of the respective metal sulfates as a function of acid concentration and temperature.

5.1. Metallic Materials

Lead. The high corrosion resistance of lead over a wide range of sulfuric acid concentrations and temperatures is due exclusively to the for-mation of a passive layer of lead sulfate (PbSO4) on the metal surface. The solubility of lead sulfate in sulfuric acid increases with increas-ing acid concentration, and the corrosion rate increases accordingly. Lead can be used with H2 SO4 concentrations up to 78 % and tempera-tures < 110 °C. Chloride ions strongly enhance the corrosion, because they penetrate the pas-sive layer. Copper-containing lead is the most corrosion-resistant form apart from high-purity lead itself. Lead plays essentially no role what-soever in a dry-gas contact sulfuric acid plant. Only in wet-catalysis plants (< 78 % H2SO4) and especially in wet gas-cleaning systems is lead still relatively widely used in such equip-ment as scrubbers, wet-gas precipitators, tube-

bundle coolers, piping, valves, and pumps. Even here it is increasingly being replaced by plastics.

The main drawbacks of lead are high density, low mechanical strength, and high installation co st.

Mild Steel. Sulfuric acid with a concentra-tion > 60 % produces an iron sulfate layer on mild steel surfaces. The solubility of iron sulfate decreases markedly at elevated acid concentra-tions, and the corrosion rate displays a similar steep decline as the H2 SO4 concentration is in-creased above 78 %. At H2SO4 concentrations > 100 % (i.e., in the presence of free sulfur tri-oxide) corrosion increases rapidly up to a sul-fur trioxide content of 30 %, declining steeply at even higher sulfur trioxide concentrations. The passive layer is quickly eroded at elevated acid-flow velocities, and the corrosion rate rises ac-cordingly. Chloride ions contribute substantially to corrosion, because they penetrate the passive layer (pitting corrosion).

Steel is used in conjunction with 68 —100 % sulfuric acid at low temperatures for both storage and handling vessels, and it is the preferred con-struction material for an reaction vessels in con-tact plants producing 93 — 98.5 % sulfuric acid. The corrosion resistance of steel declines as the temperature rises from 50 °C to 80 °C, which is why steel surfaces in reaction vessels with turbulent flow areas are protected by a lining of acid-proof bricks. Dryers and absorbers fre-quently have a layer of plastic sheeting between the steel shell and the brick lining. Steel is unsuit-able for acid temperatures above 80 °C. When acid piping is made of steel itis advisable to limit the flow velocity to about 0.5 m/s. Steel is also used for oleum, but predominantly in situations where the "free" sulfur trioxide content exceeds 22 — 23 %. If the "free" sulfur trioxide content is low, enamelled steel is suitable, even at elevated temperature.

Austenitic Stainless Steel [183]. Under ox-idizing conditions, sulfuric acid forms a passive layer on austenitic stainless steel. This layer con-sists of iron oxide and chromium oxide together with incorporated sulfates, which improves its stability. At higher flow velocities, however, and under reducing conditions, the protective layer is eroded, causing the corrosion rate to increase. Reducing conditions result, for example, if sul-


furic acid is left standing in a shut-down plant for extended periods without being moved or aer-ated. Chloride ions also contribute to depassiva-tion. In a modem sulfuric acid plant, austenitic stainless steel (e.g., material no. 1.4571) is used in conjunction with concentrated acids up to 85 °C, and it can withstand acid velocities up to ca. 1.0 m/s (in pipelines, for example). At higher temperatures, the corrosion resistance de-creases sharply. However, the temperature resis-tance can be increased to about 160 °C by apply-ing anodic protection (passivation of a metal-lic construction material by anodic polarization with direct current) [68]. Anodic protection not only reduces the corrosion rate for certain grades of stainless steel by two orders of magnitude, but it also prevents stress-corrosion cracking. It is increasingly being used to safeguard stainless-steel acid coolers, especially in acid waste-heat recovery systems.

Because of the limitations of standard stain-less steels (types 304, 316, etc.) at higher tem-peratures and high H2SO4 concentrations, con-siderable effort has been directed toward the de-velopment of special forms of stainless steel. Several stainless steels with austenitic and fer-ritic structures are now available for sulfuric acid service at temperatures up to 140 °C (in certain cases up to 200 — 220 °C) and flow velocities up to 2.5 — 3 m/s. Anodic protection is not required for such stainless steels. The most familiar alloys for this purpose include:

Sandvik SX S aramet Hastelloy C 276 (DIN 2.4819) Cronifer 2803 (DIN 1.4575) Alloy 20 (Carpenter 20) DIN 1.4529, 1.4539 DIN 1.4562, 1.4563

The above-mentioned materials are used mainly for the fabrication of heat exchangers, piping, irrigation systems, and to a certain ex-tent drying/absorption towers. A similar devel-opment has occurred with cast materials. Pumps, valves, etc. are now available for sulfuric acid ab-sorption systems operating at temperatures up to 200 °C. The best-known alloys of this type are the Lewmet family and Rheinhütte's alloys 1.4136 S/HRS. Cast iron was for decades the dominant construction material for sulfuric acid

services but itis now being replaced in almost all fields of applications by these modern alloys.

Acid-Proof Cast Iron. Certain grades of cast iron alloyed with chromium or copper are very stable at H2 504 concentrations of 90 — 99 % and temperatures < 100 —120 °C. Cast iron is used mainly for acid piping, irrigation systems in acid towers, and irrigation coolers. It is not suitable for use with oleum; although the surface-wear corrosion rate is extremely low, salts tend to form alongside the graphite lamellae, giving rise to cracks and material failure.

Silicon Iron. Silicon iron contains 14 —18 % silicon. Its high resistance to acids is due to the formation of a protective layer of silica on the surface of the casting. It is therefore somewhat sensitive to hydrogen fluoride. Silicon iron is highly resistant to sulfuric acid at all concentra-tions, from extremely dilute up to highly concen-trated, and over a wide temperature range. Sili-con iron with at least 16 % Si is useful for pumps and agitators under the high thermal stress en-countered in a sulfuric acid concentration plant. Unfortunately, however, it is highly sensitive to both thermal and mechanical shock. It also can-not be used for oleum, because its silicide con-tent is subject to oxidation, and the correspond-ing sait formation gives rise to cracks and rup-tures along the grain boundaries.

5.2. Nonmetallic Materials

Plastics. Plastic materials [for example, polypropylene, poly(vinyl chloride), or polyeth-ylene] are preferred for use in systems with low acid concentrations because of their high re-sistance to corrosion. Their usefulness is lim-ited, however, by high coefficients of thermal expansion and a loss of mechanical strength with increasing temperature, resulting in creep. Poly(vinyl chloride) (PVC) can be used with acid concentrations < 78 % H2SO4 at tempera-tures < 50 °C, polypropylene (PP) at concentra-tions < 90 % and temperatures < 80 °C. On ac-count of its excellent corrosion resistance, how-ever, polytetrafluoroethylene (PTFE) is applica-ble even at high acid concentrations (< 100 %) and temperatures > 100 °C for equipment and


pipe linings as well as in special acid heat ex-changers.

Graphite. Nonporous carbon and graphite are regarded as resistant construction materials under nonoxidizing conditions (i.e., in sulfuric acid with a concentration up to 90 %), and they are used in condensation towers, coolers, storage tanks, and pumps at temperatures < 170 °C.

Glass. Because of its excellent corrosion re-sistance at all acid concentrations, glass is used in special acid heat exchangers — for heating di-lute waste acid with hot absorber acid, for exam-ple. A significant disadvantage of this material is its fragility.

Enamel. Acid-proof enamel offers the same resistance as glass. Newly developed enamel grades are largely insensitive to thermal shock. Enamelled steel and cast iron are used for acid piping at elevated temperature.

6. Uses of Sulfuric Acid and Economic Aspects

Sulfuric acid is one of the most widely used of all industrial chemicals. But, as has already been noted, most of its uses can be considered as in-direct, because it functions as a reagent rather than as an ingredient: surprisingly little of it ap-pears in end products, and most ends up as spent acid or some type of sulfate waste. A number of products do incorporate the sulfur of sulfu-ric acid, but nearly all of them are low-volume, specialty items.

6.1. Indirect Uses

The largest single consumer of sulfuric acid by far is the fertilizer industry. Most goes into the production of phosphoric acid, which in turn is used to manufacture such fertilizer materials as triple superphosphate and mono- and diammo-nium phosphates. Lesser amounts are used for producing superphosphate and ammonium sul-fate. About 60 % of the sulfuric acid produced is utilized in fertilizer manufacture [35, 155]. In the Federal Republic of Germany, however, the fer-

tilizer industry accounts for a much less signif-icant proportion of total sulfuric acid consump-tion: only about 14 % [33]. This is because of the historical predominance in that country of thermally treated phosphates and basic slag (a byproduct of steel making) relative to chemi-cally produced phosphate fertilizers. Although thermal phosphates are generally in decline in relation to total consumption owing to their high cost, the balance is accounted for by increased use of fertilizers based on imported phosphate intermediates rather than an expansion in the do-mestic basic phosphate processing industry.

Elsewhere in industry, substantial quantities of sulfuric acid are used as an acidic dehydrating reaction medium in organic chemical andpetro-chemical processes involving such reactions as nitration, condensation, and dehydration, as well as in oil refining, in which it is used for refining, alkylation, and purification of crude-oil distil-lates.

In the inorganic chemical industry sulfuric acid is used notably in the production of TiO2 pigments, hydrochloric acid, and hydrofluoric acid.

In the metal processing industry, sulfuric acid is used for pickling and descaling steel, for leaching copper, uranium, and vanadium ores in hydrometallurgical ore processing, and in the preparation of electrolytic baths for nonferrous-metal purification and plating.

Certain wood pulping processes in the paper industry require sulfuric acid, as do some textile and chemical fiber processes and leather tan-ning.

6.2. Direct Uses

Under certain conditions, sulfuric acid is occa-sionally used directly in agriculture for reha-bilitating extremely alkaline soils, such as are found in the desert regions of the western United States. This is not a very important use in volume terms, however.

Probably the largest use of sulfuric acid in which the sulfur becomes incorporated in the fi-nal product is organic sulfonation, particularly for the production of detergents [156]. Other mi-nor organic chemicals and pharmaceuticals are also made by sulfonation.


One of the most familiar consumer products containing sulfuric acid — the lead-acid battery —accounts for only at tiny fraction of total sulfuric acid consumption.

6.3. Economic Aspects

Table 1 provides a summary of world production and consumption of sulfuric acid for the years 1982 and 1992, during which world production increased by 9.9 % (despite a drop of 10.4 % bet-ween 1989 and 1992).

Table 1. Sulfuric acid production and consumption for 1982 and 1992 (103 t of 100% H2 SO4) [source: British Sulphur]

Production Consumption

1982 1992 1982 1992

World total 130 263 144 646 132 189 145 505 Western Europe 24 094 20 295 23 806 20 649 Belgium/Luxembourg 1 798 1 917 2 092 2 697 France 3 927 3 132 4 000 3 456 Germany * 4 670 3 800 4 290 2 785 Italy 2 360 1 800 2 320 1 740 Netherlands 1 692 1 125 1 600 1 530 Spain 2 690 2 490 2 623 2 190 United Kingdom 2 587 1 716 2 559 1 946 Central Europe ** 8 216 3 390 8 510 3 310 former Czechoslovakia 1 252 510 1 415 510 Poland 2 682 931 2 600 831 Romania 1 669 500 1 700 540 former Yugoslavia 1 120 900 1 220 900 Former USSR 23 801 17 400 23 600 17 400 Africa 10 141 15 702 10 265 15 670 Morocco 3 072 7 279 3 072 7 279 South Africa 3 382 2 589 3 380 2 589 Tunisia 2 525 3 347 2 525 3 347 North America 32 490 44 030 33 458 44 740 Canada 3 131 3 605 3 060 2 350 United States 29 359 40 425 30 398 42 390 Latin America 6 308 8 602 6 960 8 824 Brazil 2 229 3 077 2 331 3 241 Chile 460 1 194 460 1 325 Mexico 2 920 2 825 3 370 2 849 Asia 20 872 33 815 20 980 33 446 China 8 174 14 060 8 270 14 060 India 2 232 3 970 2 323 3 970 Indonesia 1 330 1 370 Japan 6 531 7 120 5 763 5 948 South Korea 1 596 2 493 1 780 2 663 Taiwan 680 600 700 949 Oceania 2 491 1 412 2 490 1 466 Australia 1 971 950 1 920 1 030

* 1982 data include both the Federal Republic of Germany and the German Democratic Republic. ** German Democratic Republic omitted from the 1982 totals.

7. Analytical Techniques

Concentration and impurities content are the most frequently measured parameters for sul-furic acid quality-control purposes. In addition, quantitative determinations of sulfur dioxide and triacide are essential in the production plant for monitoring the performance of the process and ensuring that the tail-gas composition conforms with environmental regulations. More detailed accounts of analytical methods are available in [157-159].

7.1. Concentration Measurement

The simplest manual method for determining sulfuric acid concentration is density measure-ment with a densitometer. This method is reli-able for sulfuric acid with a strength of up to 95 % provided the influence of temperature is taken into account. The method can also be ex-tended to acids of higher concentration and also to oleum by first diluting a measured sample un-der controlled conditions and then measuring its density.

Continuous monitoring of acid concentration in an industrial sulfuric acid plant is generally ac-complished by means of automatic recording de-vices based either on conductivity or ultrasonic measurement. The common chemical laboratory approach to measuring sulfuric acid content is volumetric titration with sodium hydroxide so-lution, but in certain cases it is necessary to re-sort to gravimetric determination of precipitated barium sulfate.

7.2. Measurement of Impurities

Iron, arsenic, sulfur dioxide, nitrogen com-pounds, chloride, and fluoride in sulfuric acid can be determined by conventional means as fol-lows:

Iron. Iron content can be established either photometrically after addition of thiocyanate or 1,10-phenanthroline, or else by iodometry. If the sample is first reduced chemically, iron can also be determined by titration with permanganate.


Arsenic. The sample is first subjected to re-ducing conditions to convert arsenic to arsine. The latter is then driven off by heating and ab-sorbed in pyridine containing silver diethyldi-thiocarbamate. The concentration of the result-ing red complex can be determined by measur-ing its absorbance. Other techniques include the coulometric method of GUTZEIT and a gravi-metric method in which all the arsenic present is reduced to arsenic trichloride, which can be distilled off and determined as sulfide.

Sulfur Dioxide. Dissolved sulfur dioxide is determined iodometrically (-+ Sulfur Dioxide, Chap. 10.).

Nitrogen Oxides and Nitric Acid. Dis-solved nitrogen oxides can be measured by titra-tion with permanganate. Alternatively, the sam-ple may be treated with 1-naphthylamine and sulfanilic acid to convert the nitrogen oxides to an azo dye, which is then determined photomet-rically. Nitric acid and nitrosylsulfuric acid form a colored species upon addition of brucine, and this can also be determined photometrically.

Chlorides and Fluorides. Chloride is deter-mined gravimetrically by precipitation with sil-ver nitrate. Fluoride can be measured by etching according to the Spielhaczek method.

A wide range of sulfuric acid impurities is now determined on the basis of physical meth-ods; e.g.: — Mercury, selenium, and arsenic by flameless

atomic absorption spectrometry Iron, nickel, chromium, zinc, cadmium, lead, copper, vanadium, aluminum, sodium, and potassium by atomic emission spectrography (with inductively coupled plasma) or atomic absorption spectrometry

7.3. Analysis of Acid -Plant Gas Streams

The sulfur dioxide content of contact and tail gases is established by one of the methods de-scribed in -+ Sulfur Dioxide, Chap. 10..

When determining the sulfur trioxide content of sulfuric acid-plant tail gases it is important to bear in mind that residual sulfur dioxide is easily oxidized, which can lead to unrealistically high

readings. The following methods give satisfac-tory results.

The gas sample in question is passed through an aqueous solution of hydrogen peroxide and the resulting sulfuric acid is determined by titra-tion or gravimetric analysis (see above). The fraction of this acid derived from sulfur diox-ide is determined by measuring the sulfur diox-ide concentration of a second sample and calcu-lating the sulfuric acid equivalent. This is sub-tracted from the result of the first determination to leave the amount of sulfuric acid due to sulfur trioxide alone.

Alternatively, the gas can be passed through an isopropyl alcohol solution, which absorbs sulfur trioxide without oxidizing sulfur dioxide. The resulting sulfuric acid then is titrated against barium chloride solution using thorin as an in-dicator [160,161].

S03/H2 SO4 mists entrained in tail gas are separated in special filters and determined by measurement of the trapped acid. Particle spec-tra of S03/H2SO4 mists are determined by means of cascade impactors [162].

8. Sulfur Trioxide

8.1. Properties

Physical Properties. Sulfur trioxide [7446-11-9], S03, MT 80.06, is the anhydride of sulfuric acid. It is known in all three states of malter. In the gaseous and liquid phases, an equilibrium exists between the monomer, S03 , and the annular trimer, S 309 . This equilibrium is shifted towards 5309 at low temperature. In the presence of slight traces of moisture (ca. 100 ppm 112 0), liquid sulfur trioxide (below ca. 27 °C) and solid sulfur trioxide are transformed into solid polymers, which form asbestos-like, felted crystalline needles. These consist of 6-S03 and a-S03 , which are thought to corre-spond to long SO3 chains with water saturation at the chain ends — the so-called polysulfuric acids, HO(S0 2 0)x H [163].

Information regarding the physical properties of sulfur trioxide is tabulated in various works [16,164-166]. The available data originate from different measurements made over the course of the past 60 years, and they are not always mu-tually consistent, in part because it is extremely


difficult to prepare pure, absolutely anhydrous sulfur trioxide.

Table 2 provides a selection of important physical data. Pure, solid sulfur trioxide, re-ferred to as ry-S03, forms silky orthorhombic (ice-like) crystals. The melting points cited for the polymeric forms a-S03 and e-S03 indicate temperatures at which these solids depolymerize to liquid sulfur trioxide. From an industrial oper-ation standpoint it is important to recognize that the liquid sulfur trioxide originating from a-S03 and /3-S03 exerts a significant vapor pressure: 2.4 bar at 62.2 °C [163].

Table 2. Physical properties of sulfur trioxide

Property Value Reference

Gaseous SO3 Nominal density, g/L (0° C, 1013 mbar) 3.57 [164] Specific heat Cp ,k.1111 -3 K-1

100 ° C 2.543 [16] 500 ° C 3.191 [16]

Liquid SO3 Density, g/cm3 (25 ° C) 1.9 [163] bp, ° C (1013 mbar) 44.8 [164] Heat of evaporation (boiling point), J/g 538 [163] Vapor pressure, bar

20 ° C 0.26 [166] 30 ° C 0.47 50 ° C 1.32 100 ° C 8 [163]

Critical temperature, ° C 217.7 [163] Critical pressure, bar 81.9 [163] Solid SO3 ry-S03 mp, ° C 16.86 [163] Heat of fusion, J/g f3-S03

119 [163]

"Melting point", ° C a-S03

30.4 [163]

"Melting point", ° C 62.2 [163]

The temperature dependence of various phys-ical properties of sulfur trioxide has been re-presented graphically [165]. Formulae have also been worked out from which it is possible to cal-culate approximate values at any temperature for a number of properties, including vapor pres-sure, specific heat, and density [163, 167].

Chemical Properties [163, 168]. Because of sulfur trioxide's importance as an industrial in-termediate in the production of sulfuric acid, the thermodynamics and kinetics of its gener-ation by the oxidation of sulfur dioxide (Eq. 10)

have been extensively studied, as has the re-verse reaction. Pure sulfur trioxide is in fact ex-tremely resistant to thermal decomposition be-cause of kinetic inhibition, even at elevated tem-peratures where thermodynamic equilibrium is shifted heavily toward S0 2 + 02 . However, cer-tain catalytically active substances are able to increase the rate of equilibration substantially. In the presence of metals such as platinum [24] or of metal oxides and sulfates (e.g., of iron, cop-per [169], and, of course, vanadium) the decom-position approaches equilibrium at temperatures above ca. 700 °C.

The sulfur in sulfur trioxide is present in its maximum (hexavalent) oxidation state. Sulfur trioxide is therefore a strong oxidizing agent, and at the same time it is one of the stongest known Lewis acids. This bifunctional reactivity is apparent, for example, in the reaction with hy-drogen halides. With hydrogen fluoride, sulfur trioxide reacts as a Lewis acid to form fluoro-sulfuric acid. Hydrogen bromide and hydrogen iodide, on the other hand, are oxidized under the same conditions to the respective free halogens, while the sulfur trioxide is reduced to give sulfur dioxide or even hydrogen sulfide [163].

Sulfur trioxide reacts at 50 — 150 °C with elemental sulfur to give sulfur dioxide. This reaction is used industrially for the produc-tion of pure sulfur dioxide (-+ Sulfur Dioxide, Chap. 7.1.). Coke or pulverized coal at elevated temperature will reduce sulfur trioxide to sul-fur dioxide, which is the basis for the regen-eration step of the Bergbau-Forschung process for flue-gas desulfurization (—> Sulfur Dioxide, Chap. 5.7.2.).

Under absolutely dry conditions, sulfur triox-ide is unreactive toward most metals. With metal oxides it reacts at moderately high temperature to form the corresponding metal sulfates. Sulfur trioxide undergoes an extremely violent reac-tion with water to form H2SO4 or higher sulfuric acids. Sulfur trioxide gas reacts instantaneously with wet gases to form sulfuric acid mists, which are extremely difficult to precipitate.

As a general rule, sulfur trioxide reacts very briskly with all organic compounds. In the course of the reaction, the organic compound may be sulfonated, oxidized (decomposed) with the release of water, or dehydrated.

Gas return to sulfuric am/ plant

tint gas from sulfuric acid pLant

30-36 % Oleurn

24-M ConUng water


SO 3 Vapor

Figure 48. Production of 100 % sulfur trioxide by distillation from oleum a) Evaporator; b) Heat exchanger; c) S03 condenser

8.2. Manufacture

Sulfur trioxide is of course always produced in dilute gaseous form in a contact-process sulfu-ric acid plant through the oxidation of sulfur dioxide-containing gases. So far, however, the only reported preparation of pure sulfur triox-ide from dilute S03- containing gases has been a pilot-plant scale process involving cryoscopic condensation [170]. The usual procedure instead involves distillation of oleum. The heat required for oleum distillation is most conveniently sup-plied by hot contact gas from the associated sulfuric acid plant [171]. Figure 48 shows the method use at Hoechst.

In this approach, 30 — 36 % oleum is pre-heated indirectly by liquid — liquid heat ex-changers and subsequently fed to the top of a falling-film evaporator for distillation. In the evaporator the oleum flows downward as a film on the inner walls of steel tubes that are heated on the outside with hot contact gas in counter-current. The weak oleum discharged from the evaporator sump, with a residual "free" S03 con-tent of 24 — 25 %, is cooled through indirect heat exchange with the entering concentrated oleum, after which it is recycled to the sulfuric acid con-tact plant for concentration in the oleum tower once again to 30 — 36 %. It is important that the S03 content of the oleum not drop below 20 % in the interest of avoiding severe corrosion of the steel tubes at the temperatures prevailing in the evaporator. Sulfur triacide gas leaving the top of the falling-film evaporator is condensed in a

tubular condenser made of steel or aluminum. The cooling water is maintained at ca. 30 °C to ensure that the sulfur trioxide temperature does not fall below 27 °C, since sulfur trioxide would otherwise solidify. The evaporation capacity of the falling-film evaporator is controlled fora par-ticular oleum input rate by adjusting the flow rate of the hot contact gas.

The largest units of this type built so far are designed to produce ca. 3 t/h of S03. Apart from falling-film evaporators, forced-circulation evaporators are also used for oleum distillation. In most cases, stainless steel is utilized as the construction material.

8.3. Handling and Uses

When storing and handling liquid sulfur triacide it is extremely important to prevent polymeriza-tion, because any solid formed is very incon-venient to remove. Sulfur triacide reacts explo-sively with water, so the vessels must never be cleaned with water. Concentrated sulfuric acid is required for dissolving solid deposits of sulfur triacide.

In order to maintain sulfur trioxide in the liq-uid state during storage, tanks and piping must be kept at ca. 30 °C by supplementary heating. Addition of a stabilizer also effectively inhibits the formation of solid polymers at temperatures above the natural solidification point of sulfur trioxide. A number of inorganic and organic compounds are used or have been recommended


as stabilizers, but the extensive patent literature on the subject provides no indication of their mode of action [163].

Liquid sulfur trioxide is among the most dangerous of ail industrial materials, and strict safety regulations govern its handling [172, 173]. In the Federal Republic of Germany, for example, every act of transporting liquid sulfur tridioxide over public roads requires a special permit, which can be withdrawn at any time.

Pure sulfur trioxide is used in organic syn-thesis for sulfonation reactions, including the manufacture of chlorosulfonic acid, thionyl chloride, aminosulfonic acid, dimethyl sulfate, and sulfamide. The reaction is difficult to control if the sulfur trioxide is introduced in liquid form, so it is preferentially supplied as a gas (pure or diluted with an inert gas), as a solution in liquid sulfur dioxide or some other solvent that is inert under the reaction conditions, or as an addition compound with an organic base.

9. Toxicology

Given their high chemical reactivities, it is not surprising that both sulfuric acid and sulfur tri-oxide have marked toxic and ecological effects [174-176].

permitted workplace concentration of vapors and mists under German regulations (MAK) is 1 mg/m3 (ca. 0.245 ppm by volume).

Environmental Effects. Sulfuric acid re-leases are highly deleterious with respect to ground and surface waters. The substance is toxic to both fish and algae, both directly and as a result of reaction with other materials in the water. Any concentration > 1.2 mg/L is consid-ered lethal to fish; 6.3 mg/L or more causes death within 24h [176].

9.2. Sulfur Trioxide

In principle, sulfur trioxide has the same toxic effects as sulfuric acid [173]. Inhalation of the gas itself or the sulfuric acid mist formed when it cornes into contact with humid air causes irri-tation, burning, and degeneration of the tissue of moist skin, eyes, and mucous membranes, espe-cially in the respiratory tract. This happens rela-tively rapidly at concentrations > 10 ppm by vol-ume. Laryngal edema, chronic bronchitis, and dental damage are possible. In contrast, symp-toms of resorptive poisoning have not been ob-served.

9.1. Sulfuric Acid

Effects on Animais. Sulfuric acid has a highly corrosive effect on the eyes, the mucous membranes, and the skin, even in low concen-trations. Because it completely destroys living tissue, concentrated sulfuric acid causes burns that penetrate deeply and heal only slowly. Swal-lowing sulfuric acid produces extreme pain in the digestive tract, vomiting, and shock, and there is a danger of perforation. Sulfuric acid va-pors or mists irritate the eyes and the muscous membranes of the nose, pharynx, and respira-tory tract, causing heavy coughing and breath-lessness. Chronic inflammation of the upper res-piratory tract may result, and eventually impair-ment of the lungs (chronic bronchitis). Even con-centrations well below 0.1 vol % (corresponding to ca. 4 g/m3 ) will render breathing impossible. Pain and damage to the teeth have also been ob-served, especially the incisors. The maximum

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Superabsorbents 1

Superabsorbents MARKUS FRANK, Stockhausen GmbH & Co KG, Krefeld, Germany

1. Introduction ............................... 1 5.5. Permeability ............................... 12 2. Definition .................................... 1 5.6. Test Methods ............................... 13 3. Theory of Polyelectrolyte Networks 2 5.7. Stability of Superabsorbents ......... 13 3.1. Elasticity ...................................... 2 6. Uses ............................................. 13 3.2. Swelling ...................................... 3 6.1. Personal Hygiene Products ........... 13 33. Phase Transition ........................... 4 6.1.1. Diapers ................................................ 14 4. Chemistry and Production of Super- 6.1.2. Feminine Hygiene Products ............. 14

absorbents .................................... 5 6.2. Packaging Applications ................ 15 4.1. Raw Materials ............................. 5 6.3. Son Conditioners ......................... 15 4.2. Radical Solution Polymerization . . 6 6.4. Cable Applications ...................... 15 4.2.1. Principles ............................................. 6 6.5. Civil Engineering and Construction 16 4.2.2. Production Processes ........................ 7 6.6. Other Applications ...................... 16 4.3. Inverse Suspension Polymerization 7 7. Other Forms of Superabsorbent 4.4. Surface Treatment ...................... 8 Polymers ...................................... 16 5. Properties, Characterization, and 7.1. Fibers .......................................... 17

Analysis ...................................... 8 7.2. Foams and Films ......................... 17 5.1. Absorption Capacity .................... 9 8. Environmental Aspects and Toxicol- 5.2. Shear Modulus ............................. 10 ogy ............................................... 17 53. Kinetics of Absorption .................. 11 9. Economic Aspects ......................... 18 5.4. Swelling under Load .................... 11 10. References .................................... 18

1. Introduction Superabsorbents are materials that can absorb up to 1000 times their mass of water. They are usu-ally white, crystalline powders. When they come into contact with aqueous liquids, within a short period of time the powder changes into a gel-like substance and the liquid is absorbed. In this context the absorbency, however, depends very strongly on the type of aqueous liquid. In water dissolved ions in particular do strongly reduce the absorbency. Absorbents for non aqueous liq-uids are not part of the concept superabsorbents. The interested reader is referred to the relevant literature [1-4]

The development and commercial success of superabsorbents is closely connected with disposable diapers. Although the first superab-sorbents were developed as water reservoirs for soils and as carriers for active ingredients, it was their use in diapers in the 1980s that helped in their breakthrough.

The first superabsorbents were polyacryloni-triles, grafted on to starch and subsequently saponified. Later, superabsorbents based on

© 2005 Wiley-VCH Verlag GmbH & Co. KGaA, Weinheim 10.1002/14356007125101

acrylic acid quickly gained recognition for eco-nomic reasons and due to their superior prop-erties. A detailed overview on superabsorbents can be found in [5,6].

2. Definition

In the literature superabsorbents appear under different names. Hydrogel, polyelectrolyte gel, water-swellable polymer, water-absorbent poly-mer, superabsorbent material (SAM), and super-absorbent polymer (SAP) are the most frequent designations. GROSS used the designation xero-gellant, which describes both a dry gel and the behavior of the material [7].

Superabsorbents are capable of absorbing large quantities of aqueous fluids spontaneously and rapidly. The aqueous fluid is strongly re-tained and is not released mechanically. While swelling they essentially keep their original shape (Fig. 1). They only change in their dimen-sions and rheological behavior: a brittle solid material becomes a gel.

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