Troubleshooting LPG Pump Flashing and End-Curve Duty

On the pump performance curve, the fault is visible before startup: the specified duty point of 166 m³/h at 400 ft head sits near the right-hand end and far from the best-efficiency point. That selection leaves little capacity margin. A second, more immediate problem appears in the pressure data: 7 barg suction and 14.57 bara vapor pressure are on different pressure bases and, after conversion, indicate flashing if both apply at 83°C.

Start here. Reconcile temperature, composition, density units, pressure basis, and measurement location. Do not approve the pump by estimating pressure at the impeller eye or by reviewing motor power alone.

Check the property data first

The stated fluid data contain one unit error and several conditions that must be tied together:

• Temperature: 83°C
• Reported relative density: 0.557 kg/cm³
• Viscosity: 0.137 cP
• Vapor pressure: 14.57 bara
• Flow: 166 m³/h
• Differential head: 400 ft
• Suction pressure: 7 barg
• Discharge pressure: 13.87 barg

Relative density is dimensionless. If 0.557 is specific gravity, use an approximate density of 557 kg/m³. The written value 0.557 kg/cm³ would equal 557,000 kg/m³ and cannot represent LPG. Correct the process datasheet before comparing bids.

The supplied pressure rise supports the specific-gravity interpretation. The pressure difference is 13.87 - 7 = 6.87 bar. If density is 557 kg/m³, a 400 ft head, or 121.92 m, produces:

ΔP = ρgH = 557 × 9.81 × 121.92 ≈ 666,000 Pa ≈ 6.66 bar

The small difference between 6.66 bar and 6.87 bar can come from rounding, the exact LPG density at the rated condition, or pressure-tap locations. It does not support the literal density unit.

LPG composition also matters. A guessed 50% propane and 50% butane mixture is not a valid design property when seasonal composition changes. Obtain the specified composition range and a vapor-pressure or bubble-point calculation at the maximum operating temperature. Tie every property to the same composition, temperature, and pressure case.

Put every pressure on an absolute basis

Vapor pressure is stated as absolute pressure, while suction and discharge are gauge pressures. Convert before making the flashing decision:

Pabs = Pg + Patm

Using nominal atmospheric pressure, 7 barg is approximately 8 bara. That is roughly 6.6 bar below the stated 14.57 bara vapor pressure. If both values apply at the pump suction and at 83°C, the stream cannot remain single-phase liquid there.

Nitrogen pressure over the storage liquid can raise total tank pressure and suppress bulk boiling. It does not prevent local flashing when liquid pressure falls below the mixture bubble pressure at the local temperature. The deciding reading is the absolute liquid pressure at the suction reference point, not the nitrogen pressure by itself.

Also correct the pressure-gradient logic. The pump suction pressure does not have to exceed tank pressure for liquid to enter the pump. Flow from the vessel to the pump requires vessel pressure plus static liquid head to exceed suction-line losses and the pressure at the pump nozzle. With flow present, the pump-nozzle pressure will normally be below the equivalent vessel driving pressure.

Symptom or reading	Likely cause
7 barg suction compared directly with 14.57 bara vapor pressure	Gauge and absolute pressure bases have been mixed
Calculated eye pressure exceeds suction-nozzle pressure	Elevation, reference location, or internal acceleration losses have been handled incorrectly
Unstable flow, pressure, noise, or motor current	Flashing, cavitation, vapor ingestion, or operation too far right on the curve
Duty point near the curve endpoint	Pump is undersized for additional flow or the selected impeller/speed leaves no useful uprate path
High motor load at increased flow	Required power rises along the selected pump curve and exceeds the rated operating case

Calculate NPSHA from the vessel to the suction nozzle

Do not call a system calculation NPSHR. The system provides net positive suction head available, or NPSHA. The pump vendor supplies NPSHR from the pump test curve.

For a pressurized vessel, calculate NPSHA in metres of pumped liquid:

NPSHA = (Pvessel,abs - Pvapor,abs)/(ρg) + zsurface-to-pump - hfriction

Use the minimum vessel pressure, minimum liquid level, maximum liquid temperature, limiting LPG composition, and suction losses at 166 m³/h. Include entrance, pipe, valves, strainers, fittings, reducers, and any equipment between the vessel and pump. Use the actual elevation difference between minimum liquid level and the pump datum.

If a reliable pressure measurement is available at the suction nozzle during rated flow, NPSHA can instead be determined from suction absolute pressure, velocity head, and vapor pressure using one consistent datum. Do not mix a vessel-pressure calculation with a nozzle-pressure measurement or count static head twice.

The stated 3 m distance from the first-stage impeller to the discharge does not establish suction static head and does not calculate eye pressure. Measure the elevation from the minimum vessel liquid level to the pump suction datum and document the suction piping geometry.

Compare NPSHA with the vendor NPSHR curve

The impeller eye is normally a local pressure minimum because the inlet accelerates the liquid. An estimate of 8.7 barg or 9 barg at the eye from a 7 barg suction reading is therefore not credible without a defined elevation change and a validated hydraulic model. Internal eye pressure is not the normal tender acceptance variable.

Use the vendor's NPSHR curve at the proposed speed, impeller diameter, flow, and liquid condition. Then apply the NPSH margin required by the project specification and the applicable edition of API 610. Do not substitute an unspecified generic margin.

• If NPSHA is less than NPSHR, reject or redesign the suction system and pump selection.
• If NPSHA exceeds NPSHR but misses the project margin, treat the bid as noncompliant.
• If vapor is already present at the suction nozzle, an NPSH comparison for a single-phase liquid pump does not cure the inlet condition. Raise suction absolute pressure, lower temperature, reduce losses, increase static head, or select equipment rated for the resulting two-phase service.
• If the comparison passes only at normal vessel level or normal temperature, it is not a valid limiting-case selection.

Moving the pump closer to the vessel reduces friction loss. Lowering it relative to minimum liquid level adds static head. Increasing suction-line diameter reduces velocity and friction loss. Check each change with a new hydraulic calculation rather than assuming one modification is sufficient.

Validate the rated head and pressure cases

The rated pressure rise is broadly compatible with 400 ft head when 0.557 is treated as specific gravity. That confirms the intended duty calculation, but not the suction condition or curve position.

Ask the bidder to mark every required operating case on the certified performance curve:

• Rated flow and head: 166 m³/h at 400 ft
• Minimum and maximum required flow
• Minimum and maximum system differential pressure
• Maximum-temperature and limiting-composition case
• Minimum vessel pressure and minimum liquid level
• Proposed speed and impeller diameter

Confirm whether 7 barg suction and 14.57 bara vapor pressure actually belong to the same operating case. If they do not, replace them with a case matrix. Mixing a suction pressure from one temperature with vapor pressure from another creates a false flashing conclusion and an unusable NPSHA calculation.

Reject the end-of-curve selection unless it is corrected

A duty point near the curve endpoint is a tender problem even when the pump can touch the specified head and flow. The pump has little room for higher flow, changing LPG properties, system resistance error, or future duty. Operation far from the best-efficiency point also raises hydraulic loading, vibration risk, recirculation risk, and sensitivity to suction disturbances.

Check the rated point against both the preferred and allowable operating regions required by the specified API 610 edition and the purchaser's datasheet. If it lies outside a required region, record a deviation and reject the selection unless the bidder changes the hydraulic selection. Do not accept a verbal statement that the pump can operate at the curve endpoint.

Request an alternative impeller, casing, speed, or pump size that places all continuous operating points in the required region. Avoid selecting the largest impeller available in the casing when the project needs a practical future uprate path. A variable-frequency drive can move the operating point by changing speed, but it cannot repair inadequate NPSHA or make two-phase inlet flow suitable for a liquid-only pump.

Check the driver across the full curve

The motor check comes after the hydraulic and flashing checks. Increasing motor rating does not fix vapor at the inlet, inadequate NPSH margin, or an unacceptable operating region.

Review absorbed power across the complete permitted operating range, not only at the rated point. Use the bidder's power curve for the proposed impeller, speed, density, and viscosity. Compare the maximum absorbed power with the motor rating and the project service-factor requirement.

For variable-speed operation, require curves or calculated limits for each permitted speed. Check that control logic cannot command a speed and system condition that drives the pump beyond its allowable flow, power, pressure, or NPSH limits. The specified 0.137 cP viscosity is low, so the selection remains particularly sensitive to suction pressure and vapor formation rather than viscous-flow effects.

Resolve the tender and verify the corrected selection

1. Correct the fluid entry to either specific gravity 0.557 or a properly stated density at the rated condition.
2. Obtain the LPG composition range, including the limiting seasonal case, and calculate vapor pressure at 83°C for that case.
3. Convert vessel, suction, and discharge pressures to absolute pressure for phase and NPSH work. Retain gauge pressure only for differential-pressure checks where the atmospheric term cancels.
4. Calculate NPSHA at minimum vessel pressure, minimum level, maximum temperature, rated flow, and maximum credible suction-line resistance.
5. Obtain the vendor NPSHR, head, efficiency, and absorbed-power curves for the exact proposed speed and impeller.
6. Apply the NPSH margin and operating-region requirements from the project specification and specified API 610 edition.
7. Require a revised selection if the rated point remains near the curve endpoint, outside the required region, or without an uprate path.
8. Confirm motor capacity at the maximum absorbed-power point across every permitted operating case.

Verify the resolving branch by checking that the corrected suction absolute pressure remains above vapor pressure and that NPSHA exceeds NPSHR by the project-required margin. Confirm all continuous-duty points lie in the required curve region, the driver covers maximum absorbed power, and the selected impeller and speed match the quoted curves.

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