Distillation

In short

Distillation separates a liquid mixture into products of different composition by exploiting the difference in volatility of its components: the vapour formed is richer in the more volatile component than the liquid it came from. It is the single largest separation operation in the process industries — refineries, petrochemical crackers, distilleries, air separation plants, solvent recovery units — and also the highest consumer of thermal energy in them. For exams it is the heaviest-weight sub-topic of Mass Transfer in GATE (CH), ESE and university papers: McCabe–Thiele construction, q-line, minimum reflux, Fenske–Underwood–Gilliland (FUG) short-cut design and Rayleigh batch distillation appear almost every year. In interviews at IOCL, BPCL, HPCL, ONGC, Reliance, GAIL, EIL, L&T, Linde, Praj Industries and Air Liquide, questions move quickly from relative volatility to why a real column floods, weeps or breaks an azeotrope.

Explained

Vapour–liquid equilibrium: the foundation

• Distillation is possible only because at equilibrium y\* ≠ x. The separation driving force vanishes when y\* = x (an azeotrope or a pure component).
• Raoult's law (ideal liquid): partial pressure p_A = x_A·P_A^sat. Total pressure P = x_A P_A^sat + (1−x_A) P_B^sat, hence the bubble-point line; y_A = x_A P_A^sat/P gives the dew-point line.
• Antoine equation supplies P^sat(T): log₁₀P^sat = A − B/(T + C). Constants are specific to units (mmHg or kPa, °C or K) — always state them.
• Non-ideal liquids need activity coefficients: y_i P = γ_i x_i P_i^sat (modified Raoult). γ from Van Laar, Wilson, NRTL or UNIQUAC; γ > 1 (positive deviation) → minimum-boiling azeotrope, γ < 1 → maximum-boiling azeotrope.
• Relative volatility α_AB = (y_A/x_A)/(y_B/x_B) = K_A/K_B. For a binary, the constant-α equation y = αx/[1 + (α−1)x]. α = 1 means no separation; the larger α, the fewer stages.
• Diagrams examined: T–x–y (boiling point diagram, with bubble/dew curves and tie lines), y–x (equilibrium curve), and H–x–y (enthalpy–composition, used in Ponchon–Savarit).
• Lever rule on a tie line fixes the amounts of the two phases; Gibbs phase rule F = C − P + 2 shows a binary two-phase system has 2 degrees of freedom (fix P and x → T and y fixed).

Types of distillation

1. Differential (simple/Rayleigh) distillation: liquid charged once, vapour removed and condensed continuously; no reflux, poorest separation, basis of the Rayleigh equation.
2. Flash (equilibrium) distillation: feed heated and flashed through a valve into a drum; one equilibrium stage only. Used ahead of crude columns and in topping units.
3. Continuous fractionation: multistage column with rectifying (above feed) and stripping (below feed) sections, reflux at top and boil-up at bottom. The industrial workhorse.
4. Batch (with reflux): charge in the reboiler, distillate withdrawn over time — at constant reflux (distillate purity falls) or constant distillate composition (reflux increased with time). Suited to small volumes, many products, pharma and specialty chemicals.
5. Steam distillation: live steam carries over a heat-sensitive, water-immiscible compound below its normal boiling point — essential oils (mint, lemongrass, sandalwood, eucalyptus) at Indian aroma-chemical units, and fatty-acid stripping.
6. Vacuum distillation: reduced pressure lowers boiling temperature, avoiding thermal cracking — refinery VDU, mono/di-glyceride and vitamin distillation.
7. Azeotropic distillation: an entrainer forms a new, lower-boiling azeotrope removed overhead (benzene/cyclohexane formerly used for ethanol; now largely replaced).
8. Extractive distillation: a high-boiling solvent alters relative volatility without forming an azeotrope — NMP/sulpholane for aromatics extraction, ethylene glycol for ethanol dehydration.
9. Reactive distillation: reaction and separation in one shell, shifting equilibrium (methyl acetate, MTBE, biodiesel transesterification).
10. Molecular (short-path) distillation: very high vacuum, vapour path shorter than the mean free path; for high-molecular-weight, heat-labile materials.
11. Azeotrope-breaking alternatives: pressure-swing distillation (azeotrope composition shifts with pressure — THF/water), and hybrid distillation + pervaporation or molecular sieve (fuel-ethanol dehydration at Indian distilleries under the Ethanol Blended Petrol programme, which reached the E20 target).

McCabe–Thiele method (binary, constant molal overflow)

• Assumptions: equal and constant molar latent heats, negligible heat of mixing and sensible-heat effects, adiabatic column, constant pressure — so L and V are constant within each section (constant molal overflow, CMO).
• Rectifying-section operating line: y = [R/(R+1)]x + x_D/(R+1); slope L/V = R/(R+1), intercept x_D/(R+1); passes through (x_D, x_D).
• Stripping-section operating line: y = (L̄/V̄)x − (W/V̄)x_W; passes through (x_W, x_W).
• Feed (q-) line: y = [q/(q−1)]x − x_F/(q−1), where q = moles of saturated liquid produced on the feed plate per mole of feed = (H_V − H_F)/(H_V − H_L). It passes through (x_F, x_F).
  • q > 1 cold (subcooled) liquid, slope positive steep; q = 1 saturated liquid, vertical; 0 < q < 1 partially vaporised, negative slope; q = 0 saturated vapour, horizontal; q < 0 superheated vapour.
• Construction: plot equilibrium curve and 45° line → mark x_D, x_F, x_W → draw q-line → draw rectifying line from (x_D, x_D) with slope R/(R+1) → join its intersection with the q-line to (x_W, x_W) → step off triangles between equilibrium curve and operating lines. Number of triangles = ideal stages; the one that straddles the q-line intersection is the optimum feed plate; the reboiler counts as one stage, so trays = N − 1 (a total condenser is not a stage; a partial condenser acting as one equilibrium stage is).
• Minimum reflux R_min: rectifying line rotated until it touches the equilibrium curve — normally at the q-line intersection (pinch point); for a curve with an inflexion, at the point of tangency. Requires infinite stages.
• Total reflux (R = ∞): operating lines coincide with the 45° line → minimum number of stages N_min, zero product.
• Optimum reflux: industrially R ≈ 1.2–1.5 R_min — a balance between column diameter/energy (rises with R) and column height/stage count (falls with R); total annual cost shows a flat minimum.
• Murphree plate efficiency (vapour basis) η_MV = (y_n − y_{n−1})/(y_n\* − y_{n−1}); overall efficiency η_O = N_ideal/N_actual. Typical tray efficiencies 50–80%; O'Connell correlation relates η_O to μ·α of the feed.
• Open steam in the stripping section (used when the bottom product is water): the stripping line then passes through (x_W, 0) instead of (x_W, x_W) — no reboiler stage.
• Multiple feeds and side draws add extra operating lines, one per section.
• Ponchon–Savarit (enthalpy–composition) method removes the CMO assumption; uses difference points Δ_D and Δ_W on an H–x–y diagram, giving exact L/V from enthalpy balances. Needed when latent heats differ widely or the feed is strongly subcooled.

Multicomponent and short-cut design (FUG)

• Specify two key components: light key (LK) — lightest appreciably present in the bottoms; heavy key (HK) — heaviest appreciably present in the distillate. Lighter-than-LK = non-distributed light, heavier-than-HK = non-distributed heavy.
• Fenske: N_min from the split of the keys at total reflux, using geometric-mean α.
• Underwood: two equations give R_min via the root θ lying between α_HK and α_LK.
• Gilliland correlation: empirical link between (N − N_min)/(N + 1) and (R − R_min)/(R + 1) → actual N at the chosen R.
• Kirkbride equation: ratio of rectifying to stripping stages → feed-plate location.
• Rigorous solution uses stage-by-stage MESH equations (Lewis–Matheson, Thiele–Geddes, Wang–Henke bubble-point, Naphtali–Sandholm) in Aspen Plus, HYSYS, DWSIM or ProSim.
• Sequencing heuristics: remove the most plentiful component first; do the easiest (largest α) split first; leave the most difficult/highest-purity split last; direct sequence (lightest overhead one at a time) is the common default.

Column internals, hydraulics and operating problems

• Trays: sieve (cheap, simplest, narrow turndown), valve (good turndown ~4–5:1), bubble-cap (best turndown, costliest), dual-flow. Parameters: tray spacing (typically 450–600 mm), weir height (40–50 mm), downcomer area, hole area 5–15% of active area.
• Packing: random (Pall rings, Raschig rings, Intalox saddles) and structured (corrugated wire gauze/sheet) — low pressure drop, used in vacuum and heat-sensitive duty and for revamps; characterised by HETP rather than stage count.
• Malfunctions: flooding (liquid cannot descend — jet/entrainment flooding at high vapour rate, downcomer flooding at high liquid rate; signalled by rising ΔP and loss of bottom level control), weeping/dumping (vapour too low to hold liquid on the tray), entrainment (liquid droplets carried up, reducing efficiency, Souders–Brown/Fair correlation sets the allowable vapour velocity), coning, channelling and maldistribution in packing, foaming (derate by a system factor), fouling and coking.
• Auxiliaries: kettle/thermosyphon/forced-circulation reboiler, total or partial condenser, reflux drum, feed preheater, pump-arounds, side strippers, demisters.
• Energy integration: feed preheat against bottoms, vapour recompression (heat pump), multi-effect and divided-wall columns (one shell, two splits — 20–30% savings), pinch analysis. Distillation is often ~40–50% of a refinery's heat duty, so energy performance is audited under the Perform, Achieve and Trade (PAT) scheme of the Bureau of Energy Efficiency under the Energy Conservation Act, 2001 (as amended in 2022).
• Indian codes and practice: pressure-containing shells to ASME BPVC Section VIII Div. 1 and IS 2825 (unfired pressure vessels) — IS 2825 is an older standard still widely referenced in Indian tender documents, with ASME used for most new designs; flanges to ASME B16.5; relief sizing to API 520/521 and API 2000 for low-pressure storage. Hazardous process units fall under the Manufacture, Storage and Import of Hazardous Chemical Rules, 1989 and the Factories-successor provisions of the Occupational Safety, Health and Working Conditions Code, 2020, in force from 21 November 2025 along with the other three Labour Codes. Potable and industrial alcohol distilleries are licensed by State Excise departments; distillery effluent (spent wash) must meet CPCB norms, with molasses distilleries required to operate on zero liquid discharge.

Key points

• Volatility basis — vapour is richer in the component with higher vapour pressure; separation is impossible when α = 1.
• Relative volatility α — α = K_LK/K_HK; it alone decides how hard the separation is (α ≈ 1.05 for propylene/propane needs 150+ trays).
• Raoult's law — ideal-solution VLE, y = xP^sat/P; real mixtures need activity coefficients.
• Azeotrope — a mixture that boils with y = x; ethanol–water at ~95.6 wt% ethanol, 78.1 °C, 1 atm (minimum boiling); cannot be passed by ordinary distillation at that pressure.
• q-line — feed condition line through (x_F, x_F) with slope q/(q−1); vertical for saturated liquid, horizontal for saturated vapour.
• Reflux ratio R — R = L/D at the top; higher R gives fewer stages but more reboiler and condenser duty and a wider column.
• R_min and N_min — R_min needs infinite stages (pinch); N_min occurs at total reflux (Fenske); real design at R ≈ 1.2–1.5 R_min.
• Constant molal overflow — the McCabe–Thiele assumption making L and V constant per section; Ponchon–Savarit avoids it.
• Optimum feed plate — the tray where the step crosses the intersection of the two operating lines.
• Reboiler is a stage — a total condenser is not; so actual trays = ideal stages − 1 (÷ efficiency).
• Murphree efficiency — actual vapour enrichment ÷ equilibrium enrichment on one tray; 50–80% typical.
• Rayleigh equation — ln(F/W) = ∫dx/(y\*−x), the batch/differential distillation material balance.
• Flooding vs weeping — flooding is excess vapour or liquid loading (high ΔP); weeping is too little vapour (loss of efficiency, liquid drains through holes).
• Steam and vacuum distillation — both lower the operating temperature to protect heat-sensitive products.
• Extractive vs azeotropic distillation — extractive uses a heavy solvent that leaves with the bottoms; azeotropic uses a light entrainer that leaves overhead.
• Pressure effect — raising pressure lowers α (components converge towards the critical region) and raises bottom temperature; vacuum raises α but needs larger diameter.
• Energy intensity — distillation dominates plant heat duty; divided-wall columns, heat pumps and feed-bottoms exchange are the standard savings routes.

Formulas and examples

Example 1 — Overall and component balance. Feed 100 kmol/h of benzene–toluene, z_F = 0.45 benzene; x_D = 0.95, x_W = 0.05.

• Benzene balance: 100(0.45) = D(0.95) + (100−D)(0.05) → 45 = 0.95D + 5 − 0.05D → 40 = 0.90D → D = 44.44 kmol/h, W = 55.56 kmol/h.
• Check: 44.44(0.95) + 55.56(0.05) = 42.22 + 2.78 = 45.00 ✓
• Benzene recovery in distillate = 42.22/45 = 93.8%.

Example 2 — Relative volatility, q-line and minimum reflux. At 1 atm take α = 2.5 (constant) for the above system; feed is 40% vaporised.

• Equilibrium: y = 2.5x/(1 + 1.5x). At x = 0.45 → y\* = 1.125/1.675 = 0.672.
• q = fraction liquid in feed = 0.60. q-line slope = q/(q−1) = 0.6/(−0.4) = −1.5; line: y = −1.5x + 0.45/0.4 = −1.5x + 1.125.
• Pinch = intersection of q-line with equilibrium curve: solve −1.5x + 1.125 = 2.5x/(1+1.5x).

→ (−1.5x + 1.125)(1 + 1.5x) = 2.5x → −1.5x − 2.25x² + 1.125 + 1.6875x = 2.5x → −2.25x² − 2.3125x + 1.125 = 0 → 2.25x² + 2.3125x − 1.125 = 0. x = [−2.3125 + √(5.3477 + 10.125)]/4.5 = (−2.3125 + 3.9335)/4.5 = x_q = 0.3602; y_q = −1.5(0.3602) + 1.125 = 0.5847.

• R_min = (x_D − y_q)/(y_q − x_q) = (0.95 − 0.5847)/(0.5847 − 0.3602) = 0.3653/0.2245 = 1.627.
• Design at R = 1.4 R_min = 2.28. Rectifying slope = 2.28/3.28 = 0.695, intercept = 0.95/3.28 = 0.290.
• Internal flows: L = R·D = 2.28 × 44.44 = 101.3 kmol/h; V = L + D = 145.7 kmol/h.
• Condenser duty with λ = 32,000 kJ/kmol: Q_c = 145.7 × 32,000 = 4.66 × 10⁶ kJ/h = 1,295 kW.

Example 3 — Fenske, Gilliland and feed plate. Same split, α_avg = 2.5.

• N_min = ln[(0.95/0.05)(0.95/0.05)]/ln 2.5 = ln(19 × 19)/0.9163 = ln(361)/0.9163 = 5.8889/0.9163 = 6.43 stages.
• Gilliland at R = 2.28: X = (2.28 − 1.627)/(2.28 + 1) = 0.653/3.28 = 0.199.

1 + 54.4X = 11.83; 11 + 117.2X = 34.32; ratio = 0.3447. (X−1)/√X = (−0.801)/0.4461 = −1.7956. Exponent = 0.3447 × (−1.7956) = −0.6191 → exp = 0.5383 → Y = 0.4617. N = (N_min + Y)/(1 − Y) = (6.43 + 0.4617)/0.5383 = 12.8 ≈ 13 ideal stages (including the reboiler) → 12 ideal trays.

• At 65% overall tray efficiency: actual trays = 12/0.65 = 18.5 ≈ 19 trays. At 500 mm spacing, tangent-to-tangent height ≈ 19 × 0.5 + 3 m (top and bottom space) ≈ 12.5 m.

Example 4 — Rayleigh batch distillation. Charge 100 kmol with x_F = 0.50, α = 2.5 constant; distil until x_W = 0.25. Find residue and average distillate composition.

• ln(F/W) = [1/1.5]·ln[(0.50/0.25)(0.75/0.50)] + ln(0.75/0.50) = (0.6667)·ln(3.0) + ln(1.5) = 0.6667(1.0986) + 0.4055 = 0.7324 + 0.4055 = 1.1379.
• F/W = e^1.1379 = 3.120 → W = 32.05 kmol, distillate D = 67.95 kmol.
• Average y_D by balance: (100 × 0.50 − 32.05 × 0.25)/67.95 = (50 − 8.01)/67.95 = 0.618.

Example 5 — Flash (single-stage) separation. Feed 100 kmol/h, z_F = 0.60, flashed to 50% vaporisation, α = 2.5.

• V = L = 50 kmol/h. Balance: 60 = 50y + 50x → y + x = 1.2, with y = 2.5x/(1+1.5x).
• 2.5x/(1+1.5x) = 1.2 − x → 2.5x = (1.2 − x)(1 + 1.5x) = 1.2 + 1.8x − x − 1.5x² → 2.5x = 1.2 + 0.8x − 1.5x² → 1.5x² + 1.7x − 1.2 = 0.
• x = [−1.7 + √(2.89 + 7.2)]/3 = (−1.7 + 3.1765)/3 = 0.4922; y = 1.2 − 0.4922 = 0.7078.
• Check with equilibrium: 2.5(0.4922)/(1+0.7383) = 1.2305/1.7383 = 0.7078 ✓ — one stage lifts 0.60 only to 0.708, which is why multistage columns with reflux are used.

Where it is used

• Refineries (IOCL, BPCL, HPCL, Reliance Jamnagar, Nayara, HMEL): atmospheric crude distillation units with side strippers and pump-arounds, vacuum distillation of reduced crude, naphtha splitters, debutanisers; CDU/VDU operation and optimisation is a core panel-engineer role.
• Petrochemicals and gas processing (GAIL, OPaL, HMEL, Haldia Petrochemicals): C2/C3 splitters, de-ethanisers, de-propanisers, benzene–toluene–xylene fractionation using extractive distillation with sulpholane or NMP.
• Distilleries and ethanol plants (Praj Industries, Thermax, Shree Renuka, Balrampur Chini): multi-pressure distillation with molecular-sieve dehydration to produce fuel-grade anhydrous ethanol for the Ethanol Blended Petrol programme (E20 achieved; higher blends under evaluation), with spent-wash concentration and zero liquid discharge.
• Cryogenic air separation (INOX Air Products, Linde India, Air Liquide India): double-column (Linde double column) distillation of liquid air into N₂, O₂ and argon, feeding steel plants and hospital oxygen.
• Pharmaceuticals, agrochemicals and solvent recovery (Divi's, Aarti, Deepak Nitrite, SRF): batch and thin-film/short-path distillation for solvent recovery and heat-sensitive intermediates; recovery economics directly reduce raw-material and hazardous-waste cost.
• Aroma chemicals, edible oils and specialities: steam distillation of mint and lemongrass oil (Uttar Pradesh mint belt), deodorisation and fatty-acid distillation in vegetable-oil refineries.
• Job roles: process design engineer (EIL, L&T, Toyo, Technip, Tata Consulting), simulation engineer (Aspen/HYSYS), commissioning and panel operations, technical services and energy audit (BEE-certified energy auditor under PAT).

Interview questions

• Q: What is relative volatility and why does it matter more than boiling-point difference? A: It is the ratio of the K-values (equilibrium distribution ratios) of two components; it directly measures how far equilibrium vapour departs from the liquid, so it sets the number of stages — and it accounts for non-ideality, which a boiling-point difference alone does not.
• Q: Why is a column operated at a reflux ratio above the minimum but well below total reflux? A: At R_min stages are infinite (height and cost unbounded); at total reflux there is no product. Operating at about 1.2–1.5 R_min sits near the flat minimum of total annual cost, trading energy and diameter against height.
• Q: What does the q-line slope tell you about the feed? A: Slope q/(q−1): vertical means saturated liquid, horizontal saturated vapour, negative slope a partially vaporised feed, positive and steep a subcooled liquid, and positive but between 0 and 1 a superheated vapour.
• Q: Why is the reboiler counted as a theoretical stage but a total condenser is not? A: The reboiler generates vapour in equilibrium with the bottoms liquid, so it performs a separation; a total condenser merely condenses all vapour without changing composition. A partial condenser does count as a stage.
• Q: How do you distil past an azeotrope such as ethanol–water? A: Change the thermodynamics — pressure-swing distillation, extractive distillation with a heavy solvent such as ethylene glycol, azeotropic distillation with an entrainer, or a hybrid with molecular-sieve adsorption or pervaporation, which is the standard route in Indian fuel-ethanol plants.
• Q: What happens to the separation when you raise column pressure? A: Relative volatility generally falls (components converge towards the critical region), so more stages or more reflux are needed; but vapour density rises, so the column diameter shrinks, and condensing at higher temperature may allow cooling water instead of refrigeration.
• Q: How do you recognise flooding on a running column? A: Pressure drop across the trays rises sharply and erratically, bottom level climbs or loses control, top and bottom temperatures converge and product purities collapse together; the cure is to cut reboiler duty, feed rate or reflux.
• Q: What is the difference between weeping and flooding? A: Weeping is too little vapour — liquid drains through the tray holes, cutting efficiency; flooding is too much vapour or liquid — liquid cannot descend and accumulates up the column.
• Q: Why does vacuum distillation follow atmospheric distillation in a refinery? A: Heavier fractions would need temperatures above about 370 °C to vaporise at atmospheric pressure, at which they thermally crack and coke; vacuum lowers the boiling point enough to recover vacuum gas oil without cracking.
• Q: What is the Rayleigh equation used for? A: Differential or simple batch distillation with no reflux — it integrates the unsteady material balance to relate the amount and composition of residue to the amount distilled.
• Q: When would you choose packing over trays? A: Vacuum or low pressure-drop service, heat-sensitive products needing low holdup, corrosive duty favouring ceramic or plastic packing, small-diameter columns, and revamps needing extra capacity in the existing shell; trays stay preferred for fouling, high liquid rates and very large diameters.
• Q: How does a divided-wall column save energy? A: A vertical partition inside one shell performs a three-product separation that would otherwise take two columns, eliminating the duplicated remixing and one reboiler/condenser pair — typically 20–30% energy and capital savings.
• Q: Why is Murphree efficiency sometimes above 100%? A: On a large tray the liquid is not perfectly mixed; vapour contacting liquid near the inlet weir meets liquid richer than the average leaving composition, so local enrichment can exceed that predicted from the exit liquid.

Common mistakes

• Mixing mass and mole fractions — all distillation equations are on a molar basis; ethanol–water data quoted in wt% must be converted first.
• Using R = L/V instead of R = L/D, which makes the rectifying-line slope and all internal flows wrong.
• Drawing the q-line through (x_D, x_D) or (x_W, x_W) rather than (x_F, x_F), or getting the sign of q/(q−1) wrong for a partially vaporised feed.
• Counting the reboiler as a tray, or forgetting to divide ideal stages by tray efficiency before quoting a tray count.
• Finding R_min from the q-line intersection on an equilibrium curve that has an inflexion, where the true pinch is the point of tangency.
• Assuming constant α across the whole column when the feed is wide-boiling — use a geometric mean of top and bottom values, or stage-wise α.
• Claiming any purity is reachable with enough stages when an azeotrope exists at the operating pressure, or assuming a higher reflux will break it.
• In the lab, reading a distillation temperature from the liquid instead of the vapour at the still head, and running without a packed/fractionating column while expecting sharp separation.

Related topics

Mass transfer: diffusion · Absorption and stripping · Design of trays and packed columns · Extraction and leaching · Membrane separations · Chemical engineering thermodynamics (VLE and activity coefficient models)