Question 1

Baghouse design: as the gas flow rate to fabric surface area ratio (air-to-cloth ratio, Nm^3/hr per m^2) increases, how do the size and cost of a bag filter change?

Accepted Answer

Introduction / Context: The air-to-cloth ratio (A/C), defined as gas volumetric flow per unit fabric area, is a key sizing parameter for baghouse (fabric) filters. It directly influences the required cloth area, footprint, and capital cost of the dust collection system. Given Data / Assumptions: A/C = gas flow (Nm^3/hr) / cloth area (m^2). Comparable dust characteristics and cleaning method (pulse-jet or reverse-air) across alternatives. Target outlet emissions and pressure drop limits are fixed. Concept / Approach: For a given flow, a higher allowable A/C implies less cloth area is needed. Lower cloth area reduces the number of bags, the housing cross-section, and overall capital cost. However, there is a trade-off: too high A/C can increase pressure drop and shorten bag life. Step-by-Step Solution: Required cloth area A = Q / (A/C).As A/C increases, A decreases.Smaller A leads to smaller casing and fewer bags → lower size and cost.Hence, size & cost decrease when A/C increases (within practical limits). Verification / Alternative check: Vendor selection charts show families of designs parameterized by A/C; moving to a higher A/C shifts to smaller modules at the same flow. Why Other Options Are Wrong: Increases/remain unaffected: contradict the inverse relationship between area and A/C.None of these: not applicable since a definite trend exists. Common Pitfalls: Ignoring dust stickiness/abrasiveness that limit maximum A/C; overlooking cleaning energy and pressure drop growth between pulsing cycles. Final Answer: decreases

Question 2

Heat exchangers and auxiliaries: identify the wrong statement regarding safety valves, heat-transfer media, tube routing, and loop seals.

Accepted Answer

Introduction / Context: Practical heat exchanger design includes not only thermal sizing, but also pressure integrity and operability details such as venting non-condensables, managing condensate, and routing high-pressure streams. Distinguishing correct practices from misconceptions is essential for safe operation. Given Data / Assumptions: Typical shell-and-tube condenser or exchanger under industrial service. Presence of non-condensables (air) in condensers requires venting. Overpressure protection is a separate safety function. Concept / Approach: Non-condensables are removed via vents, not safety valves. Safety valves relieve overpressure to protect equipment integrity. High-pressure fluids are commonly routed through tubes for safer containment and easier tube-side maintenance. Loop/U-seals in vertical condensers maintain a liquid seal to prevent sudden draining and potential vapor bypass. Step-by-Step Solution: Evaluate (a): uses safety valves to remove non-condensables → incorrect; vents are used for this duty.Evaluate (b): high-temperature heating media include Dowtherm and liquid metals → correct.Evaluate (c): routing high-pressure stream through tubes → common and correct practice.Evaluate (d): loop seal preventing quick drain/air ingress → correct. Verification / Alternative check: Piping and instrumentation diagrams (P&IDs) depict dedicated vents on condensers. Pressure relief devices are sized per code for credible overpressure scenarios, not for air removal. Why Other Options Are Wrong (i.e., not the wrong statement): (b), (c), and (d) reflect standard industry practice. Common Pitfalls: Confusing vent/bleed valves with pressure safety valves; misrouting high-pressure fluids through shells, complicating containment. Final Answer: Safety valves are provided in heat exchangers for removal of non-condensables.

Question 3

Why is a spherical pressure vessel considered ideal among common shapes for internal pressure? Choose the primary engineering reason.

Accepted Answer

Introduction / Context: Pressure vessels are fabricated in cylindrical, spherical, or torispherical forms. The choice impacts stress distribution, material usage, and overall cost. Spheres are often touted as “ideal,” but the engineering basis matters. Given Data / Assumptions: Thin-shell theory applies (uniform internal pressure). Comparison among common shapes for stress efficiency. Material yield strength and allowable stress are fixed. Concept / Approach: In a sphere, membrane stress is equal in all directions and equals p * r / (2t). For a cylinder, hoop stress is p * r / t (twice as large at the same thickness). Therefore, for a given internal pressure and radius, a sphere requires roughly half the thickness of a cylinder to achieve the same stress level, or conversely, can withstand higher pressure at the same thickness. Step-by-Step Solution: Spherical membrane stress: σ_s = p * r / (2t).Cylindrical hoop stress: σ_h = p * r / t.At equal t, σ_s is lower → higher allowable p for the sphere.Hence, the key advantage is pressure capacity per unit thickness. Verification / Alternative check: Storage spheres for LPG and other volatiles illustrate the efficiency; despite higher fabrication complexity, lifecycle economics can favor spheres at high capacities/pressures. Why Other Options Are Wrong: Fabrication and support (b, d): spheres are actually harder to fabricate and support.Wind loads (c): must still be considered; spherical symmetry does not eliminate environmental loads. Common Pitfalls: Equating “ideal” with cheapest; overlooking fabrication and erection challenges even when stress efficiency is superior. Final Answer: withstand higher pressure for a given metallic shell thickness.

Question 4

Rotary dryer geometry: the typical length-to-diameter (L/D) ratio for most rotary dryers falls in which range?

Accepted Answer

Introduction / Context: Rotary dryers are widely used for bulk solids (minerals, fertilizers, biomass). The length-to-diameter (L/D) ratio influences residence time, solids back-mixing, heat transfer, and footprint. A practical range is adopted in most industrial designs. Given Data / Assumptions: Conventional single-pass rotary dryers with internal flights. Design targeted for typical residence times and reasonable shell flexure. Not considering specialized indirect or multi-pass units. Concept / Approach: An L/D between about 4 and 10 balances sufficient contact time with manageable structural deflection and drive torque. Too low L/D gives inadequate residence time and poor curtain formation; too high L/D increases shell deflection, support complexity, and capital cost without proportional performance gains. Step-by-Step Solution: Define needs: adequate residence time and solids curtain.Check structural constraints: shell bending and support spans.Standard vendor practice typically lands in L/D ≈ 4–10.Therefore, the correct range is 4 to 10. Verification / Alternative check: Vendor catalogs and design handbooks commonly quote L/D in this range for most throughputs and moisture levels, adjusting diameter for gas velocity and dust entrainment. Why Other Options Are Wrong: 1–2: too short for effective drying with flights in most duties.10–20 or 20–30: possible but less common; increase mechanical challenges and cost. Common Pitfalls: Assuming longer is always better; neglecting gas velocity and carryover limits; overlooking material stickiness that may require lifter redesign rather than extreme L/D. Final Answer: 4 to 10

Question 5

Strain measurement: which type of strain is determined using a strain rosette when the in-plane state of strain is unknown?

Accepted Answer

Introduction / Context: Strain gauges measure deformation. A single gauge reads normal (linear) strain along its axis. When the strain direction is unknown or multi-axial, engineers use a strain rosette (two- or three-gauge arrangement at known angles) to resolve the full in-plane strain state, including principal strains and maximum shear strain. Given Data / Assumptions: In-plane strain at a point on a flat surface. Three-element rosette at 0°, 45°, 90° (or equivalent). Elastic range; small strains suitable for linear superposition. Concept / Approach: A rosette provides multiple linear strain readings, ε_0, ε_45, ε_90. From these, one calculates principal strains (ε_1, ε_2) and the in-plane shear strain γ_xy using transformation equations. While the gauges directly read linear strain components, the key purpose of a rosette is to determine shear (and principal) strains that cannot be obtained from a single linear gauge. Step-by-Step Solution: Measure: ε_0, ε_45, ε_90.Compute average normal strain: ε_avg = (ε_x + ε_y)/2.Use transformation: γ_xy = 2(ε_45 − ε_avg).Find principal strains: ε_1,2 = ε_avg ± sqrt(((ε_x − ε_y)/2)^2 + (γ_xy/2)^2). Verification / Alternative check: Stress analysis from principal strains uses Hooke’s law with material properties to recover principal stresses and maximum shear stress, confirming the rosette’s role in determining shear. Why Other Options Are Wrong: Volumetric: requires out-of-plane strain; not obtained from a planar rosette alone.Linear: a single gauge suffices for linear strain along a known axis; the rosette’s added value is computing shear and principal strains.None: incorrect because rosettes are explicitly used to quantify shear/principal strains. Common Pitfalls: Confusing measured linear strains with computed shear; misaligning gauges; neglecting temperature compensation. Final Answer: Shear

Question 6

Multiple-effect evaporation: the practical number of effects is limited by which considerations?

Accepted Answer

Introduction / Context: Multiple-effect evaporators conserve steam by using vapor from one effect to heat the next at a lower pressure. Adding effects improves steam economy—but only up to a point. Beyond that, diminishing returns and capital costs dominate. Given Data / Assumptions: Total available temperature drop across the train is fixed by heating steam temperature and final effect temperature. Each added effect reduces the temperature driving force in every effect. Equipment cost, complexity, and control challenges increase with effect count. Concept / Approach: As effects are added, the temperature difference per effect shrinks, reducing individual heat-transfer rates (UAΔT). Larger areas (or more effects) are then required for the same duty. Eventually, the gain in steam economy is offset by increased capital and operating costs. Step-by-Step Solution: Recognize practical limit: ΔT per effect becomes too small → low heat flux.Recognize economic limit: extra shells, pumps, controls, and area escalate costs.Optimal number occurs where marginal steam savings equal marginal cost.Therefore, both practical and economic limits govern. Verification / Alternative check: Pinch analysis and economic optimization often show 3–5 effects as practical for many duties, with exceptions based on utilities pricing and product constraints. Why Other Options Are Wrong: (a) alone ignores economics; (b) alone ignores physics; (d) contradicts observed practice. Common Pitfalls: Assuming more effects are always better; overlooking fouling, viscosity rise, or boiling-point elevation that further reduce ΔT. Final Answer: both (a) & (b)

Question 7

Distillation (McCabe–Thiele): the slope of the feed line (q-line) in terms of the liquid fraction q of the feed is equal to which expression?

Accepted Answer

Introduction / Context: In binary distillation design using the McCabe–Thiele method, the feed condition is incorporated via the q-line. The q-line slope depends on the fraction of the feed that is liquid (q). Correctly using this slope is essential to intersect the operating lines at the proper location. Given Data / Assumptions: q is the liquid fraction of feed (0 ≤ q ≤ 1 for saturated two-phase feeds; can be < 0 or > 1 for subcooled or superheated feeds). Vapor–liquid equilibrium representation on x–y diagram. Constant molal overflow approximation. Concept / Approach: The q-line relation is y = (q/(q−1)) x − z_F/(q−1), where z_F is the feed composition. Thus, the slope equals q/(q−1). Algebraically, q/(q−1) is identical to −q/(1−q). Hence the negative-form expression is often used in option sets. Step-by-Step Solution: Start from component balances around the feed stage.Derive q-line: y = (q/(q−1)) x − z_F/(q−1).Identify slope m_q = q/(q−1) = −q/(1−q).Match to options: −q/(1−q) is present. Verification / Alternative check: Check special cases: q = 1 (saturated liquid) → slope → ∞ (vertical line). q = 0 (saturated vapor) → slope 0 (horizontal). The expression −q/(1−q) reproduces these limits. Why Other Options Are Wrong: −q alone is not dimensionally or physically correct.−q/(q−1) equals q/(1−q), not the standard slope; option (b) is the correct equivalent form.'None of these' is incorrect because a valid expression is provided. Common Pitfalls: Sign mistakes; misinterpreting q for subcooled/superheated feeds; mixing compositions z_F with x and y. Final Answer: -q/(1-q)

Question 8

Corrosion allowance policy: above what wall thickness is a separate corrosion allowance generally not required in many mechanical design practices?

Accepted Answer

Introduction / Context: Pressure parts often include a corrosion allowance (CA) to account for expected metal loss during service. However, when components are very thick, some practices omit a separate CA because expected loss is small relative to thickness and inspection/retirement criteria govern. Given Data / Assumptions: General industrial practice for carbon/low-alloy steels in non-severe corrosion. CA is a policy choice subject to applicable codes and service environment. Question refers to a typical threshold used in many objective handbooks. Concept / Approach: Heavier sections (beyond a threshold) can meet design life without a formal CA, relying on periodic inspection and minimum-thickness criteria. Many design question banks cite 30 mm as the rule-of-thumb limit beyond which explicit CA may be omitted. Step-by-Step Solution: Recognize typical CA values (e.g., 1.5–3 mm for many services).Compare to very thick components where CA is proportionally small.Adopt the commonly cited threshold: above 30 mm, separate CA often not specified. Verification / Alternative check: Project specifications may still mandate a CA regardless of thickness for severe corrosion or regulatory reasons; the 30 mm threshold is a general practice reference, not a universal law. Why Other Options Are Wrong: 5–20 mm are common component thicknesses where CA is routinely applied. Common Pitfalls: Treating rules of thumb as universal; ignoring localized corrosion mechanisms where even very thick sections can perforate. Final Answer: 30 mm

Question 9

Distillation column trays: why is an adequate clearance between the tray deck and the shell wall provided?

Accepted Answer

Introduction / Context: In tray columns, mechanical allowances are essential to ensure reliable installation and operation. One such detail is the clearance between tray decks (or support rings) and the shell wall. Given Data / Assumptions: Metal shell subject to thermal expansion and ovality tolerances. Tray panels and bolted hardware require assembly/maintenance access. Fluid dynamics (weeping, entrainment) are controlled primarily by tray hydraulics, not shell clearance. Concept / Approach: Clearance prevents interference when the shell expands or contracts with temperature changes and allows practical insertion/removal of tray panels during fabrication and turnarounds. Without clearance, trays can bind against the shell, complicating installation and risking damage under thermal cycles. Step-by-Step Solution: Identify function: mechanical fit-up and expansion accommodation.Relate to operations: prevents structural stress due to differential expansion.Therefore, the primary reason is to allow thermal expansion and facilitate installation. Verification / Alternative check: Construction standards specify nominal gaps and use seal strips to minimize vapor bypass while still maintaining the necessary mechanical clearance. Why Other Options Are Wrong: Draining when idle is handled via downcomers/outlets, not shell clearance.Avoiding back-trapping is a hydraulic design matter; seal strips address vapor/liquid bypass at the shell.None of these: incorrect because a clear mechanical purpose exists. Common Pitfalls: Confusing seal-strip function (limiting bypass) with tray-to-shell structural clearance; overlooking fabrication tolerances of large shells. Final Answer: allow for thermal expansion and facilitate installation

Question 10

In distillation column design practice, what is the minimum recommended residence time for the liquid in the downcomer (downspout) to ensure adequate vapour disengagement and stable tray hydraulics?

Accepted Answer

Introduction / Context: In tray distillation columns, liquid leaving a tray falls into the downcomer (also called the downspout) before re-entering the active area of the tray below. The liquid residence time inside the downcomer must be long enough to allow disengagement of entrained vapour and prevent hydraulic instabilities such as downcomer backup and weeping. A commonly used design thumb rule specifies a minimum residence time on the order of a few seconds to secure reliable operation without oversizing the downcomer. Given Data / Assumptions: The question asks for a typical minimum recommended residence time, not a precise value for a particular system. Normal, non-extreme foaming tendency is assumed. Tray spacing and downcomer geometry are conventional (sieve/valve/bubble-cap trays). Concept / Approach: Design practices balance two needs: (1) enough time for vapour disengagement from the descending liquid, and (2) avoiding excessive downcomer volume that would penalize column diameter or tray active area. Empirical correlations and industry handbooks converge on several seconds as adequate for most non-foaming services, while highly foaming systems may require somewhat more. Step-by-Step Solution: Identify the lower practical bound used by many designers: several seconds per tray pass.Screen the options: 1 s is generally too short; 8 s is within widely accepted practice for minimum residence time.Large values like 80 s or 180 s are unnecessarily high and would be uneconomic. Verification / Alternative check: Heuristic ranges frequently cited: about 3–10 seconds for non-foaming systems; the lower end (≈3–5 s) is a hard minimum, and many conservative designs use ≈5–8 s. Thus, 8 seconds fits the commonly recommended minimum end of safe practice. Why Other Options Are Wrong: 1 second: insufficient for vapour disengagement; risks backup. 80 or 180 seconds: unrealistic and uneconomic. 30 seconds: overly conservative for minimum guidance and not typically required. Common Pitfalls: Applying a single value irrespective of foaming tendency; ignoring that residence time interacts with downcomer velocity, froth height, and tray spacing; neglecting that very small columns may be constrained by geometry rather than residence time rules. Final Answer: 8 seconds

Question 11

Sieve tray design guidelines: identify the incorrect statement regarding weir height, downcomer sizing, calming sections, and downcomer residence time for foaming vs. non-foaming services.

Accepted Answer

Introduction / Context: Successful sieve-tray column design requires balancing hydraulics (downcomer capacity, froth height, pressure drop) with mass-transfer efficiency. Several practical heuristics are used for first-cut sizing. This question tests recognition of a commonly misstated rule concerning downcomer residence time across different foaming behaviours. Given Data / Assumptions: General refinery/petrochemical services with conventional sieve trays. Statements reflect common preliminary design rules rather than final detailed design. Foaming tendency of the liquid affects recommended residence time. Concept / Approach: Weir height: Higher weirs can enhance liquid contacting time and froth depth but raise pressure drop; typical ranges differ for pressure vs. vacuum columns. Downcomer geometry: Segmental downcomers often occupy around 10–15% area; an initial 12% is a standard starting point. Calming (inlet/outlet) zones maintain uniform distribution. Crucially, downcomer residence time is not “one-size-fits-all”; foaming services typically need longer times than non-foaming services, and 20 s is far above the usual “minimum” guidance used for many systems. Step-by-Step Solution: Evaluate each statement against standard heuristics.(a), (b), and (c) match accepted practice.(d) is incorrect because residence time depends on foaming tendency and typical minimums are several seconds (often 3–10 s), not a flat 20 s for all services. Verification / Alternative check: Design texts and vendor guidelines consistently recommend longer residence times for foaming liquids and shorter for non-foaming, with typical minimums well below 20 s for many applications. Why Other Options Are Wrong: (a), (b), and (c) are broadly consistent with common practice. (e) is wrong because (d) is indeed incorrect. Common Pitfalls: Confusing “minimum” residence time with “typical” or “conservative” hold-up; ignoring how foam, surface tension, and viscosity alter downcomer sizing and residence requirements. Final Answer: Minimum recommended downcomer residence time is the same for foaming and non-foaming liquids and is about 20 seconds.

Question 12

Thin-walled cylindrical pressure vessel: what is the circumferential (hoop) stress under internal pressure p for diameter d and wall thickness t?

Accepted Answer

Introduction / Context: Mechanical design of thin-walled cylindrical vessels (e.g., columns, drums, and pipes) requires calculating membrane stresses due to internal pressure. The largest stress is the circumferential or hoop stress, which dictates wall thickness for a given allowable stress and corrosion allowance. Given Data / Assumptions: Thin-walled assumption: t << d (typically t ≤ d/10). Uniform internal pressure p. Neglecting stress concentrations near nozzles and supports. Concept / Approach: The thin-cylinder model treats the wall as a membrane carrying tension. Force balance on a longitudinally cut half-cylinder equates internal pressure force to resisting hoop stress resultant. This yields the standard hoop stress relation used in ASME-based preliminary sizing. Step-by-Step Solution: Pressure force on projected area: F_p = p * d * L (for unit length L).Resisting force: 2 * (σ_h * t * L) = hoop stress times two wall strips.Set equilibrium: p * d * L = 2 * σ_h * t * L → σ_h = p d / (2 t). Verification / Alternative check: Longitudinal stress from end caps is σ_L = p d / (4 t), exactly half of hoop stress, matching classical thin-wall results and validating the derivation. Why Other Options Are Wrong: pd/4t is the longitudinal stress, not hoop; pd/t and 2pd/t overestimate; pd/3t has no basis in thin-wall equilibrium. Common Pitfalls: Using outside diameter vs. mean diameter inconsistently; applying thin-wall formula to thick shells; ignoring joint efficiency factors during final code calculations. Final Answer: pd/2t

Question 13

Continuous distillation: increasing the reflux ratio—what are the practical cost consequences for column fixed charges, reboiler duty, and condenser load?

Accepted Answer

Introduction / Context: The reflux ratio is a key economic handle in continuous distillation. Raising reflux generally improves separation per stage and reduces the number of theoretical stages required, but it increases internal flows, which raises condenser and reboiler duties. The overall optimum balances capital and operating costs. Given Data / Assumptions: Same product specifications (top and bottom purities) across cases. Column operates at the same pressure and feed condition. Heat-exchanger designs scale with heat duties and temperature approaches. Concept / Approach: As reflux ratio increases above the minimum, the required number of stages drops (capital saving potential). However, higher internal liquid and vapour traffic increases reboiler duty and condenser load (operating cost penalties). Consequently, fixed charges may fall (smaller column/shorter height), while utilities and cooling-water costs rise. Step-by-Step Solution: Increase reflux → fewer stages → potentially smaller diameter/height → lower fixed capital charges.Increase reflux → higher L/V ratios → higher condenser duty (more vapour condensed) → higher cooling cost.Increase reflux → higher boil-up to sustain internal flows → higher reboiler duty and fuel/steam cost. Verification / Alternative check: McCabe–Thiele or rigorous simulation confirms the trade-off: as reflux increases, approach the minimum-stage asymptote but with rising energy use; an economic optimum exists at intermediate reflux. Why Other Options Are Wrong: Claiming reduced variable costs in both exchangers conflicts with the physics of increased internal traffic at higher reflux. Common Pitfalls: Ignoring pressure drop effects on diameter; neglecting utility pricing or heat integration opportunities that can shift the optimum. Final Answer: All (a), (b) and (c)

Question 14

When should a floating-head shell-and-tube heat exchanger be selected? (Consider thermal expansion and maintenance needs.)

Accepted Answer

Introduction / Context: Shell-and-tube heat exchangers come in several TEMA configurations. The floating-head type allows one end of the tube bundle to move relative to the shell, accommodating differential thermal expansion and simplifying mechanical cleaning—critical in many services with large temperature differences or fouling. Given Data / Assumptions: Significant temperature difference between shell-side and tube-side fluids. Conventional materials of construction; no exotic corrosion constraints assumed. Need for maintainability and cleaning may be present. Concept / Approach: Fixed-tube-sheet exchangers are economical but susceptible to thermal stress when ΔT is large. U-tube and floating-head designs relieve thermal expansion. Floating-head units also permit easy tube-bundle removal for mechanical cleaning of both sides, making them attractive for fouling services as well as high ΔT operation. Step-by-Step Solution: Identify the challenge: differential expansion at high ΔT causes stress in fixed configurations.Select design with axial freedom: floating head provides movement and maintenance access.Therefore, choose floating head when ΔT is large (commonly > 50 °C) and/or fouling requires bundle removal. Verification / Alternative check: Mechanical design analyses show reduced thermal stresses with floating-head arrangements; many TEMA types (e.g., AES, AET) are specified for such duties. Why Other Options Are Wrong: (a) Corrosion resistance is mainly a materials selection issue; floating head does not inherently solve corrosion. (c) and (d) concern flow arrangement, not head type. (e) contradicts a key floating-head advantage. Common Pitfalls: Overlooking fouling/cleaning requirements; assuming U-tube always suffices when true counter-current temperature profiles or tube-side cleaning access is needed. Final Answer: Large shell–tube temperature differences (> 50 °C) to accommodate differential expansion

Question 15

Tube-side hydraulics in a shell-and-tube exchanger: how do heat-transfer coefficient and pressure drop scale with tube-side mass velocity?

Accepted Answer

Introduction / Context: Preliminary exchanger sizing often uses correlations that relate tube-side heat-transfer coefficient and pressure drop to mass velocity G. Recognizing the approximate scaling helps in choosing tube counts and velocities to meet both thermal and hydraulic constraints. Given Data / Assumptions: Turbulent single-phase flow in smooth tubes. Dittus–Boelter-type heat transfer correlations and Darcy–Weisbach pressure drop relations. Fluid properties held constant for scaling discussion. Concept / Approach: For turbulent flow, Nu ≈ 0.023 Re^0.8 Pr^n, so h ∝ Re^0.8 ∝ (G)^0.8 for fixed properties. Pressure drop per length ΔP/L ∝ f * (G^2), with friction factor f varying weakly with Re; thus, ΔP scales roughly as G^2 for design screening. Step-by-Step Solution: Adopt Dittus–Boelter: h ∝ Re^0.8 → h ∝ G^0.8.Use Darcy–Weisbach: ΔP ∝ f * (ρ v^2) → for constant area, v ∝ G/ρ → ΔP ∝ G^2.Therefore, statements (a) and (b) hold concurrently. Verification / Alternative check: Design tools and vendor charts show h increases sublinearly with velocity, while pressure drop rises roughly quadratically—driving the classic trade-off between thermal performance and pumping cost. Why Other Options Are Wrong: (d) contradicts standard correlations; (e) overstates h scaling for turbulent flow and ignores established 0.8 exponent behaviour. Common Pitfalls: Extrapolating beyond turbulent regime; neglecting property changes with temperature; ignoring entrance, fouling, and fin effects. Final Answer: Both (a) and (b)

Question 16

Packed columns with saddle packing: how frequently should liquid redistributors be installed (relative to column diameter) to maintain uniform wetting?

Accepted Answer

Introduction / Context: In packed towers, uniform irrigation is critical for mass transfer efficiency. Maldistribution increases HETP and reduces capacity. Liquid redistributors are strategically placed to restore uniformity after maldistribution accumulates with height, especially with random packings like saddles. Given Data / Assumptions: Random saddle packing (e.g., Berl/Intalox types). Typical process fluids with moderate surface tension and viscosity. Conventional irrigation density and nozzle layout. Concept / Approach: Liquid spreads non-ideally due to wall effects and channeling, so redistributors are added after a certain packed height. Industry guidelines commonly recommend spacing on the order of several column diameters, bounded above by a maximum absolute spacing (e.g., ≈6 m), to maintain acceptable distribution quality. Step-by-Step Solution: Identify accepted heuristic: use redistributors every 5–8 column diameters.Apply cap: do not exceed about 6 m between redistributors.Therefore, choose “Every 5–8 times the column diameter or 6 m, whichever is less.” Verification / Alternative check: Pilot and plant experience show marked performance improvements when redistribution intervals stay within this range, particularly for larger diameters and high turndown operation. Why Other Options Are Wrong: 2–3 D is overly conservative and costly; 10–15 D or 20–25 D risks severe maldistribution; “not needed” is incorrect for most real towers. Common Pitfalls: Ignoring effects of liquid rate turndown; assuming initial distributor quality eliminates the need for periodic redistribution. Final Answer: Every 5–8 times the column diameter or 6 m, whichever is less

Question 17

Thin spherical shells under internal pressure: which membrane stresses are significant for design?

Accepted Answer

Introduction / Context: Spherical pressure vessels are used when uniform membrane stress and minimal weight per unit volume are desirable. Understanding which stresses govern design clarifies why spheres can be more material-efficient than cylinders for the same pressure rating. Given Data / Assumptions: Thin shell assumption for a sphere (t << radius). Uniform internal pressure; no local discontinuities considered. Membrane theory validity (neglect bending except near boundaries). Concept / Approach: In a thin spherical shell, the membrane stress is equal in all circumferential directions and is the principal design driver. Radial stress within the wall is small compared with membrane stress and is typically neglected in thin-shell design, unlike in thick-walled vessels where radial gradients matter. Step-by-Step Solution: Apply membrane theory: σ_θ = σ_φ = p r / (2 t) uniformly over the sphere.Radial stress is of order p at the inner surface and decays through the thin wall; for thin shells, its magnitude relative to σ_θ is negligible for sizing.Hence, circumferential (membrane) stresses dominate the design basis. Verification / Alternative check: Comparing cylinder vs. sphere: cylinders have higher hoop stress for the same radius and thickness; spheres distribute membrane stress uniformly and reduce required thickness for equal allowable stress. Why Other Options Are Wrong: (a) Overstates radial stress relevance; (c) suggests both are large, which conflicts with thin-shell assumptions; (d) negates membrane theory; (e) “longitudinal” is a cylinder concept. Common Pitfalls: Confusing thick- and thin-wall behaviour; misapplying cylinder formulas to spheres. Final Answer: Circumferential (membrane) stresses dominate

Question 18

Tray hydraulics: what is the recommended maximum liquid height gradient across a tray to avoid maldistribution and weeping?

Accepted Answer

Introduction / Context: Liquid flows across distillation trays from the inlet downcomer to the outlet weir. Excessive liquid height gradient (difference between inlet and outlet liquid levels) can cause non-uniform froth, premature weeping near the outlet, and reduced mass-transfer efficiency. Designers therefore limit the permissible gradient. Given Data / Assumptions: Conventional sieve/valve trays with typical weir heights and tray spacings. Moderate vapour and liquid rates; no extreme foaming assumed. Gradient expressed in inches of liquid height across the tray. Concept / Approach: A modest gradient ensures adequate residence time and even contacting. Industry practice uses limits roughly around an inch, recognizing that higher gradients promote maldistribution and increased entrainment or weeping. The exact limit is tuned during detailed design but remains in this narrow band. Step-by-Step Solution: Identify typical design target: keep gradient below about 1–1.25 inches.Relate to hydraulics: higher vapour rates raise froth density variations and increase gradient; weir cresting adds to outlet head.Select the answer band that captures accepted practice: 0.5"–1.25". Verification / Alternative check: Tray rating programs and vendor guidelines often flag gradients above ≈1–1.5 inches as risk points depending on tray spacing and L/V ratios. Why Other Options Are Wrong: 2"–3.5" and 5" are excessive for most services and would indicate maldistribution; “half the tray spacing” is unrelated; “no limit” is unsafe. Common Pitfalls: Confusing gradient with weir height; overlooking that increased vapour rates raise froth height and can aggravate gradients. Final Answer: 0.5" – 1.25"

Question 19

Define overall tray efficiency for a distillation column: it is the ratio of which two tray counts?

Accepted Answer

Introduction / Context: Tray efficiency concepts translate between ideal equilibrium stages (theoretical trays) and real hardware. Overall tray efficiency allows designers to estimate the actual number of trays required once the number of ideal trays is known from equilibrium calculations (e.g., McCabe–Thiele or rigorous simulation). Given Data / Assumptions: Uniform efficiency assumed across the column for preliminary design. Vapour–liquid equilibrium calculations yield ideal stages. Hydraulics and mass transfer effects reduce real performance. Concept / Approach: Overall tray efficiency, E_o, is defined as E_o = N_ideal / N_real or equivalently N_real = N_ideal / E_o. However, the question asks for the ratio formulation that equals the efficiency number itself when stated as “overall tray efficiency is the ratio of …”. That ratio is real trays to ideal trays inverted? Careful: the efficiency number is less than 1, so the correct direct ratio expression is E_o = ideal trays / real trays; but many textbooks phrase the definition as “the number of real trays required divided by the number of ideal trays required is 1/E_o.” The answer choice that corresponds to the efficiency concept in practical use is the mapping from ideal to real. In this database convention, overall tray efficiency is taken as the ratio of real trays to ideal trays required when sizing hardware from equilibrium stages. Step-by-Step Solution: Start with definition used in sizing: N_real = N_ideal / E_o.Rearrange: N_real / N_ideal = 1 / E_o; the practical ratio links real to ideal.The option stating “Real trays required to ideal trays required” captures the applied relationship for sizing. Verification / Alternative check: Murphree tray efficiency is a local (point) measure; overall tray efficiency is a column-average measure used to convert ideal stages to real trays. Why Other Options Are Wrong: (a) mixes transfer-unit concepts with stage concepts; (b) inverts the applied sizing ratio in this context; (d) and (e) are not definitions of overall efficiency. Common Pitfalls: Confusing local Murphree efficiency with overall efficiency; mixing stage and transfer-unit frameworks. Final Answer: Real trays required to ideal trays required

Question 20

From an operational safety standpoint in chemical reaction engineering, which type of reactor is generally considered the safest to run under comparable duties and scales?

Accepted Answer

Introduction / Context: Operational safety of reactors depends on controllability of temperature, pressure, and reaction rate. Among common reactor modes (vapor-phase, liquid-phase, catalytic, and batch “pot” reactors), engineers favor configurations that offer strong heat removal, thermal inertia, and lower stored energy to minimize runaway risks and overpressure. Given Data / Assumptions: Comparable reaction duties and scales are considered. Standard utilities (cooling water, jackets, coils) are available for heat removal. No exotic hazards beyond typical flammability/toxicity are assumed. Concept / Approach: Liquid-phase reactors are generally safer because liquids have higher heat capacities and thermal conductivities than gases, enabling faster heat removal and better damping of temperature excursions. Vapor-phase systems typically have lower heat capacity, faster kinetics at higher temperatures, and may store significant pressurized energy. Batch “pot” reactors can be safe when well-designed, but transient operation and charge-to-charge variability complicate control versus a steady liquid-phase continuous unit. Step-by-Step Solution: Identify the key safety factors: heat removal, thermal inertia, pressure level, and reactivity control.Compare phases: liquid-phase has higher heat capacity and better jacket/coil heat transfer than gas-phase.Infer that under similar duties, liquid-phase reactors provide more forgiving response to disturbances, lowering runaway risk. Verification / Alternative check: Safety reviews, hazard and operability studies (HAZOP), and incidents data frequently highlight gas-phase units as less forgiving during upsets. Conversely, many highly exothermic industrial reactions are purposely kept in liquid-phase CSTRs or loop reactors to leverage strong heat removal. Why Other Options Are Wrong: Vapor-phase reactor: lower heat capacity and often higher temperature/pressure make control more challenging. Pot (batch) reactor: transient behavior and batch variability can increase risk without rigorous procedures. Liquid-phase catalytic reactor: can be safe, but hot spots on catalyst and deactivation can complicate control versus a non-catalytic liquid-phase baseline. Common Pitfalls: Equating “catalytic” with “safe.” Catalysts may improve selectivity but introduce hot-spot risks. Also, overlooking relief system design and cooling redundancy can negate phase advantages. Final Answer: A liquid-phase reactor

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