1 1. Long tube vertical evaporators with either natural or forced circulation are most popular. Tubes are 19–63 mm (0.75–24.8 in.) in diameter and 3.66–9.14 m (12–30 ft) long.

      2 2. In forced circulation, linear velocities in the tubes are in the range of 4.57–6.09 m/s (15–20 ft/s).

      3 3. Elevation of boiling point by dissolved solids results in temperature differences of 3–10°F between solution and saturated vapor.

      4 4. When the boiling point rise is appreciable, the economic number of effects in series with forward feed is 4–6.

      5 5. When the boiling point rise is small, minimum cost is obtained with 8–10 effects in series.

      6 6. In backward feed the more concentrated solution is heated with the highest temperature steam so that heating surface is lessened, but the solution must be pumped between stages.

      7 7. The steam economy of an N-stage battery is approximately 0.8 N-lb evaporation/lb of outside steam.

      8 8. Interstage steam pressures can be boosted with steam jet compressors of 20–30% efficiency or with mechanical compressors of 70–75% efficiency.

      1 1. The dispersed phase should be the one that has the higher volumetric rate, except in equipment subject to back-mixing where it should be the one with the smaller volumetric rate. It should be the phase that wets the material of construction less well. Since the holdup of continuous phase is greater, that phase should be made up of the less expensive or less hazardous material.

      2 2. There are no known commercial applications of reflux to extraction processes, although the theory is favorable.

      3 3. Mixer–settler arrangements are limited to at most five stages. Mixing is accomplished with rotating impellers or circulating pumps. Settlers are designed on the assumption that droplet sizes are about 150 µm in diameter. In open vessels, residence times of 30–60 min or superficial velocities of 0.15–0.46 m/min (0.5–1.5 ft/min) are provided in settlers. Extraction-stage efficiencies commonly are taken as 80%.

      4 4. Spray towers as tall as 6–12 m (20–40 ft) cannot be depended on to function as more than a single stage.

      5 5. Packed towers are employed when 5–10 stages suffice. Pall rings 25–38 mm (1–1.5 in.) in size are best. Dispersed-phase loadings should not exceed 10.2m3/min-m2 (25 gal./min-ft2), and HETS of 1.5–3.0 m (5–10 ft) may be realized. The dispersed phase must be redistributed every 1.5–2.1 m (5–7 ft). Packed towers are not satisfactory when the surface tension is more than 10 dyne/cm.

      6 6. Sieve tray towers have holes of only 3–8 mm diameter. Velocities through the holes are kept below 0.24 m/s (0.8 ft/s) to avoid formation of small drops. Re-dispersion of either phase at each tray can be designed for. Tray spacings are 152–600 mm (6–24 in). Tray efficiencies are in the range of 20–30%.

      7 7. Pulse packed and sieve tray towers may operate at frequencies of 90 cycles/min and amplitudes of 6–25 mm. In large-diameter tower, HETS of about 1 m has been observed. Surface tensions as high as 30–40 dyn/cm have no adverse effect.

      8 8. Reciprocating tray towers can have holes of 150 mm (9/16 in.) diameter, 50–60% open area, stroke length 190 mm (0.75 in.), 100–150 strokes/min, and plate spacing normally 50 mm (2 in.) but in the range of 25.0–150 mm (1–6 in.). In a 760-mm (30-in.) diameter tower, HETS is 500–650 mm (20–25 in.) and throughput is 13.7 m3/min-m2 (2000 gal./h-ft2). Power requirements are much less than those of pulsed towers.

      9 9. Rotating disk contractors or other rotary agitated towers realize HETS in the range of 0.1–0.5 m (0.33–1.64 ft). The especially efficient Kuhni with perforated disks of 40% free cross section has HETS of 0.2 m (0.66 ft) and a capacity of 50 m3 /m2-h (164 ft3/ft2-h).

      1 1. Process are classified by their rate of cake buildup in a laboratory vacuum leaf filter: rapid, 0.1–10.0 cm/s; medium, 0.1–10.0 cm/min; and slow, 0.1–10.0 cm/h.

      2 2. Continuous filtration should not be attempted if 1/8 in. cake thickness cannot be formed in less than 5 min.

      3 3. Rapid filtering is accomplished with belts, top feed drums, or pusher centrifuges.

      4 4. Medium rate filtering is accomplished with vacuum drums or disks or peeler centrifuges.

      5 5. Slow-filtering slurries are handled in pressure filters or sedimenting centrifuges.

      6 6. Clarification with negligible cake buildup is accomplished with cartridges, precoat drums, or sand filters.

      7 7. Laboratory tests are advisable when the filtering surface is expected to be more than a few square meters, when cake washing is critical, when cake drying may be a problem, and when precoating may be needed.

      8 8. For finely ground ores and minerals, rotary drum filtration rates may be 15,000 lb/day-ft2 at 20 rev/h and 18–25 in. Hg vacuum.

      9 9. Coarse solids and crystals may be filtered at rates of 6000 lb/day-ft2 at 20 rev/h and 2–6 in. Hg vacuum.

      1 1. Properties of particles that are conducive to smooth fluidization include rounded or smooth shape, enough toughness to resist attrition, sizes in the range of 50–500 µm diameter, and a spectrum of sizes with ratio of largest to smallest in the range of 10–25.

      2 2. Cracking catalysts are members of a broad class characterized by diameters of 30–150 µm, density of 1.5 g/ml or so, and appreciable expansion of the bed before fluidization sets in, minimum bubbling velocity greater than minimum fluidizing velocity, and rapid disengagement of bubbles.

      3 3. The other extreme of smoothly fluidizing particles are typified by coarse sand and glass beads, both of which have been the subject of much laboratory investigation. Their sizes are in the range of 150–500 µm, densities 1.5–4.0 g/ml, have small bed expansion and about the same magnitudes of minimum bubbling and minimum fluidizing velocities, and they also have rapidly disengaging bubbles.

      4 4. Cohesive particles and large particles of 1 mm or more do not fluidize well and usually are processed in other ways.

      5 5. Rough correlations have been made of minimum fluidization velocity, minimum bubbling velocity, bed expansion, bed level fluctuation, and disengaging height. Experts recommend, however, that any real design be based on pilot-plant work.

      6 6. Practical operations are conducted at two or more multiples of the minimum fluidizing velocity. In reactors, the entrained material is recovered with cyclones and returned to process. In driers, the fine particles dry most quickly so the entrained material need not be recycled.

      1 1. For conservative estimate set F = 0.9 for shell and tube exchangers with no phase changes, q = UAF∆Tlm. When ∆T at exchanger ends differ greatly then check F, reconfigure if F is less than 0.85.

      2 2. Take true countercurrent flow in a shell-and-tube exchanger as a basis.

      3 3. Standard tubes are 19.0 mm (3/4 in.) outer diameter (OD), 25.4 mm (1 in.) triangular spacing, 4.9 m (16 ft) long.A shell of 300 mm (1 ft) diameter accommodates 9.3 m2 (100 ft2);600 mm (2 ft) diameter accommodates 37.2 m2 (400 ft2);900 mm (3 ft) diameter accommodates 102 m2 (1100 ft2).

      4 4. Tube side is for corrosive, fouling, scaling, and high-pressure fluids.

      5 5. Shell side is for viscous and condensing fluids.

      6 6. Pressure drops are 0.1 bar (1.5 psi) for boiling and 0.2–0.62 bar (3–9 psi) for other services.

      7 7. Minimum temperature approach is 10°C (20°F) for fluids and 5°C (10°F) for refrigerants.

      8 8.