To give an example of one of the most successful applications of chemical exchange to separation of isotopes of an element heavier than hydrogen that may have industrial application, a brief description will be given of the process and equipment used by Taylor and Spindel [T2] to product 15 N 99.8 percent pure. This separation depends on the exchange reaction

15NO + H14N03 *= 14N0 + H15N03

which takes place in the gas phase because of the presence there of the species NO, N02, N203, N204, H20, HN02i and HN03. These interact at acceptably high rates at temperatures of 25°C or higher. The separation factor for this process, defined as the ratio of ISN/14N in the liquid phase to 1SN/14N in the gas phase, was found by Taylor and Spindel to be 1.055 at 25°C in 10 M HN03, and to decrease with increasing acid concentration and increasing temperature. Because the value of the equilibrium constant for the foregoing reaction calculated from spectroscopic data is 1.096, it appears that isotopic exchange reactions between species other than HN03 and NO enter into the observed overall exchange equilibrium. This reaction, however, may be used to characterize the process.

Taylor and Spindel found that the optimum conditions for operating this process on the laboratory scale were 8 to 10M HN03, 25 to 50°C, and atmospheric pressure. Although a higher temperature speeds up attainment of exchange equilibrium, a is lower, and more N02 is present with a lower exchange equilibrium constant.

The process used by Taylor and Spindel is illustrated in Fig. 13.41. liquid aqueous HN03 flows downward through a packed column countercurrent to an upflowing gas stream consisting largely of NO with lesser amounts of other nitrogen compounds. Nitric acid containing the
normal abundance of 1SN, 0.365 a/о, is fed at the top of the larger column, no. 1, and is enriched in 1SN by the foregoing exchange reaction as it flows down this column. At the foot of the column, where its 15N content is around 7 percent, the reflux ratio of NO vapor to product may be substantially reduced. This is done by diverting 4 percent of the acid downflow to the smaller column, no. 2. The remaining 96 percent of the acid downflow is sent to NO reflux generator no. 1, where it is reduced to NO by reaction with S02:

НгО + HN03 + f S02 — f H2S04 + NO

The NO is returned to column no. 1 as reboil vapor.

The HN03 flowing down through the smaller column, no. 2, countercurrent to NO is enriched further in 1SN to 99.8 percent at the foot of the column. At this point some of the downflowing HNO3 is withdrawn as plant product, and the remainder of the HN03 is reduced to NO with S02 in reflux generator no. 2. This NO is used to reboil column no. 2.

NO vapor depleted in 1SN leaving column no. 1 at the top of the plant is converted to HNO3 depleted in 15 N by mixing it with air and passing the mixture counter to downflowing water in a packed column, where the reaction

——- H20

——- S02

-► H2S04

Figure 13.41 Plant used for production of 1SN by Taylor and Spindel.

NO + fOj +|HjO-*-HN03

takes place.

The net result of the process, then, is to separate HN03 containing the natural abundance of, SN into product HN03 highly enriched in 1SN and waste HN03 slightly depleted in 15 N, while converting S02 and air to H2S04. The minimum ratio of H2S04 to 1SN is | times the minimum molal reboil vapor ratio, which is given by Eq. (12.80), or

3Xp-xF g _ 3 0.998 — 0.00365 1.055 2 xF a— 1 2 0.00365 0.055

This high reflux requirement is not a complete economic drain because H2S04 is a more valuable material than S02. In this respect, Taylor and Spindel’s process is in a more favorable economic position than the chemical exchange system of Fig. 13.24 to concentrate deuterium, which consumes aluminum to make less valuable A1203.

In their engineering analysis of the HN03-NO process, Garrett and Schacter [G2] considered a plant to produce 30.2 g-mol lsN/day while simultaneously producing 239,670 g-mol H2S04/day. They recommended use of substantially the same conditions employed by Taylor and Spindel and estimated that 1SN could be produced at a cost of $4/g. This relatively low cost is due to the credit for converting S02 to H2S04.

It is important to note that the use of a cascade of columns of decreasing size, such as in Fig. 13.41, does not affect the consumption of chemicals for reflux, because this depends on the interstage flow required at the feed point. The cascade of columns of decreasing size does, however, reduce the total volume and the holdup of desired isotope. If the cascade of columns were not used for the 15 N separation example, with its low feed concentration and separation factor close to unity, the holdup would be so great that product concentration would not reach

99.8 percent in any practical time.

NOMENCLATURE

a defined by Eq. (13.34)

A annual cost, $/year A tower cross-sectional area b defined by Eq. (13.36) c defined by Eq. (13.35) c unit cost D separative capacity E efficiency F molar feed rate

F’ molar flow rate of supplementary feed to hot tower g ratio of steam rate to minimum rate G vapor molar flow rate h height of transfer unit H moles of hydrogen

H humidity, mol water/mol noncondensible gas I inventory, mol / annual charge against investment к equilibrium constant for gas-phase exchange reaction К total tails flow rate

К equilibrium constant for gas-liquid exchange reaction

L liquid molar flow rate M molecular weight n number of stages p pressure

P mol product (or molar product flow rate)

P kg D2 O/year Q rate of loss of availability r fractional recovery

R gas constant

s entropy per mole

S entropy

5 solubility, mol dissolved gas/mol water t time

T absolute temperature

T0 absolute temperature at which heat is rejected v vapor velocity, cm/s

V tower volume

V vapor molar flow rate

W moles of tails (or molar tails flow rate)

W power

x atom fraction or mole fraction in liquid

у atom fraction or mole fraction in vapor

z distance from top of tower

Z height of tower

a stage separation factor

a* relative volatility, separation factor in distillation /3 heads separation factor 7 relative volatility of H2 S to HDS r relative abundance in vapor

£ relative abundance in liquid

?r vapor pressure

p density

со overall separation factor

Subscripts

a stream produced in heating liquid or cooling gas b bottom of tower c cold tower

F feed stream

h hot tower

і stage number

m stage number

P product stream Q turbine work

S supplementary feed point t top of tower

V tower volume W tails stream

0 gas stream from cold tower to hot tower, Fig. 13.34

(a) Find the number of theoretical plates in the cold stripping section ncS, the cold enriching section n^, and the hot tower nh.

(b) Compare this process with the methylamine process of Fig. 13.40 and the ammonia process of Fig. 13.37 with respect to:

(1) Number of cold-tower plates

(2) Number of hot-tower plates

(3) Flow ratio, hydrogen to D2 0

(4) Flow ratio, liquid to D2 0

CHAPTER