In addition to needing an exchange equilibrium constant different from unity, exchange processes for concentrating deuterium require a reflux-making step in which part of the deuterium in the liquid phase leaving an enriching column is transferred to the vapor phase returned to the column. This can be done either by a chemical-phase conversion operation in the monothermal exchange processes to be described in this section or by another exchange column at a higher temperature in the dual-temperature exchange processes to be described in Secs. 11 through 14.

The exchange towers of Fig. 13.21 are an example of a monothermal exchange process for concentrating deuterium, with the electrolytic cells providing reflux-making phase conversion. Because the equilibrium constant for the reaction

HD(?) + H2 0(0 — H2 (g) + HDO(/)

is 3.2 at 60°C, a flow sheet similar to this figure would permit concentration of deuterium even if no separation occurred in electrolysis.

Chemical reflux-making steps, such as the electrolysis of water in Fig. 13.21, cost more than thermal reflux-making steps such as the reboiler or condenser in distillation. Consequently, there are only a few examples of economical monothermal exchange processes for concentrating deuterium. Hydrogen-water exchange refluxed by water electrolysis as in Fig. 13.21 is one example that can be economical where electricity is cheap enough and electrolytic hydrogen valuable. The only other commercial example of production of deuterium by monothermal exchange is use of the ammonia-hydrogen exchange reaction

HDfe) + NH3(0 — H2 fe) + NH2 Щ)

This was used m the exchange plant at Mazingarbe (item 12, Table 13.2) and is planned in plants in India (items 16 and 18). The economic attractiveness of this process comes from two factors:

1. The relative ease with which hydrogen reflux can be obtained by thermal cracking of ammonia

2. The relatively low cost of providing liquid ammonia reflux as incremental production of an existing synthetic ammonia plant

10.1 Monothermal Ammonia-Hydrogen Exchange

Figure 13.23 illustrates the principle of enriching deuterium by monothermal ammonia — hydrogen exchange in conjunction with a synthetic ammonia plant. The flow quantities have been developed from partial information reported for the Mazingarbe plant [LI, El, N2]. Feed, at point (1), consists of ammonia synthesis gas, 3H2/1N2, which has been purified to reduce its content of water, CO, and C02 to less than 1 ppm and compressed to the pressure of ammonia

Figure 13.23 Monothermal ammonia-hydrogen exchange process. Flow quantities, kg-mol/h.

synthesis, 350 atiri^ in this example, as at Mazingarbe. At (2) feed joins additional synthesis gas circulating countercurrent to liquid ammonia in the stripping exchange column В and the enriching exchange column C, both operated at —25°C. At this temperature, the equilibrium constant for the foregoing deuterium exchange reaction is 5.2. At the countercurrent flow conditions of Fig. 13.23, deuterium is transferred from upflowing synthesis gas to downflowing ammonia. For the reported [El] 85 percent recovery of deuterium, the atom percent deuterium in synthesis gas is reduced from 0.0132 percent in feed [LI] to 0.00198 percent in the gas leaving the stripping column (4). For the reported [LI] 100-fold enrichment, the ammonia leaving the enriching column (11) is enriched to 1.32 percent deuterium. Final concentration to 99.8 percent deuterium [N2] is by distillation of ammonia, G.

At Mazingarbe synthesis gas leaving the stripping column at (4) was converted to ammonia in the synthetic ammonia plant A of Houffleres du Bassin du Nord et du Pas-de-Calais. In Fig. 13.23 about 5 percent of the synthesis gas is purged at (7) to remove inert impurities, mostly methane and argon, present in feed. About 65 percent of the ammonia synthesized in A is withdrawn as stripped ammonia product at (6) and the remainder is returned (8) to the exchange column as liquid reflux.

At the bottom of the enriching column C, a small portion of the partially enriched ammonia (13) is withdrawn for final concentration, and the remainder (19) is dissociated in D into partially enriched synthesis gas (12) to serve as reboil vapor for the enriching column. To obtain nearly complete dissociation of ammonia without using excessively high temperatures, the Mazingarbe plant is reported [LI] to have reduced the ammonia pressure in the cracking section to 55 atm. At this pressure and the assumed cracking temperature of 740°C, the ammonia content at equilibrium would be under 1 percent. For simplicity, the small amounts of ammonia in recycle gas (12) and hydrogen dissolved in recycle ammonia (3) have been neglected in Fig. 13.23.

To cause the deuterium exchange reaction between hydrogen and ammonia to take place at a useful rate, it is necessary to have 1 to 2 m/o (mole percent) of potassium amide, KNH2, dissolved in the liquid phase to serve as a homogenous catalyst [C2]. Presence of potassium amide complicates the process for three reasons:

1. Potassium amide is expensive and must be recovered and recycled.

2. It remains in deuterium exchange equilibrium with ammonia:

KNHD + NH3 ^ KNH2 + NH2D

3. It reacts, sometimes violently, with oxygen-containing impurities in synthesis gas.

To recover and recycle potassium amide, the partially enriched ammonia containing dissolved partially enriched potassium amide leaving the exchange column at (11) is depressured to 55 atm, heated to 125°C, and separated at E into ammonia vapor and a concentrated liquid solution of potassium amide (16). Because this solution contains 1.32 percent deuterium, its deuterium content must be reduced before it is recycled to the top of the stripping exchange column. In Fig. 13.23 this is done by countercurrent exchange with stripped ammonia vapor (9) in the catalyst deuterium stripper F. This is a conventional sieve-plate column in which deuterium is transferred from dissolved potassium amide to ammonia vapor in the overall reaction

KNHD(0 + NH3(?) — KNH2(/) + NHj D(g)

which proceeds rapidly with an equilibrium constant near unity. The deuterium-depleted catalyst solution leaving F is repressured to 350 atm and returned (10) to the top of the

^l atm = 1.01325 bar.

stripping exchange column B. At Mazingarbe, catalyst stripping was done with synthesis gas feed.

Potassium amide reacts with water, C02, CO, and oxygen, forming solid impurities that would plug the columns. With oxygen it forms potentially explosive potassium azide, KN3. To prevent loss of exchange catalyst and formation of undesirable reaction products in the exchange system, synthesis gas feed is purified ahead of (1) by two systems not shown. It is dried by molecular sieves and then passed through a guard chamber containing sacrificial KNH2 dissolved in liquid ammonia to remove oxygen-containing impurities [N2].

The flow quantities of synthesis gas feed (I), stripped ammonia product (6), and enriched ammonia product (14) were calculated from the net ammonia production rate of 330 MT/day [El] reported at Mazingarbe, the reported [N2] heavy-water production rate of 26 MT of 99.8 percent D20 per year at 85 percent recovery [El], and the 0.0132 percent deuterium in feed [LI]. The reflux rates of ammonia (3) to column В and synthesis gas (12) to column C and the number of theoretical stages were evaluated in unpublished design studies at Massachusetts Institute of Technology which led to a requirement of n$ = 2.6 theoretical stages in stripping column В and ri£ = 3.1 theoretical stages in enriching column C. These values were obtained from the Kremser equation (13.92). The number of stages in the stripping section, п$, is

In [(ay2 — x3)/(oy4 — x5)]

In (oL3/Kj)

The number of stages in the enriching section, ng, is

In [(oy12 -*n)/(oy2 ~x3)] "£ hi (oLu/KI2)

In these equations у is the atom fraction deuterium in hydrogen, and x is the atom fraction deuterium in the solution of potassium amide in ammonia at the numbered points in Fig. 13.23. К is the molar flow rate of hydrogen in the vapor phase, and L is the molar flow rate of hydrogen in the liquid solution of KNH2 in ammonia. For example, L3 = f(438.80) + 6.20 = 664.4.

The small number of theoretical stages is a consequence of the high value of the separation factor, 5.2, and is the principal advantage of the monothermal ammonia-hydrogen exchange process. There are, however, a number of offsetting disadvantages. Even in the presence of KNH2 catalyst, the rate of the exchange reaction is low, primarily because of the low solubility of hydrogen in liquid ammonia. Even at the high pressure of 350 atm used in Fig. 13.23, the hydrogen content of the liquid is only 0.85 m/o. With conventional sieve-plate or bubble-plate columns the plate efficiency would be only 1 or 2 percent, necessitating use of hundreds of plates. The Mazingarbe plant is reported [LI] to have used special ejectors for upflow of gas to entrain liquid and increase interphase transfer area. Because of the pressure drop taken by the upflowing gas, it was necessary to pump the liquid from one plate to the plate next below. Even with this enhanced contacting, it was necessary to use towers 36 m high [El].

The Mazingarbe exchange plant produced its first heavy water in January 1968. It was taken out of service in 1972 because of an explosion in the ammonia synthesis plant, which has not been repaired. Operation of the exchange plant itself was satisfactory; the availability factor was 92 percent in 1970. Lefrancois stated [LI] that an operating temperature of —10°C in the exchange towers would have been preferable to the design temperature of —25°C.