Design Tips 10

Piping

• Provide SDVs at either end of HC piping on bridges interconnecting offshore platforms, to minimize loss of inventory

• Pressure Equalization: For gas SDV ≥ 6” and dP>2 bar. 2” pressure equalization with manual or SDV. To slowly pressurise downstream. To avoid damage to SDV seat seals by gas at high velocity with mill scale impacting TSO. See API 14B and 14C, allowable leakage rates. Gas rush can damage downstream piping, vibration, if not properly supported. Pressurization duration depends on downstream volume. Equalized slowly via a RO or globe valve that takes the ΔP. Cheaper to replace 2” globe/ RO

• Pressure Equalization across ESDVs: Some prefer an upstream manual valve of same size equalized instead of 2” equalization across ESDV to protect ESDVs integrity, should the 2” valve is accidentally left open. Others opt for 2” equalization valve across ESDV, provided it is managed in SIS. May allow testing the ESDVs in service

• Pressure Equalization: Mass flow via 2” low. Avoids JT effect cooling of downstream piping below MDMT. P&ID Caution Box: “Pressurize slowly watching for external icing/ cold. Wait for downstream piping to warm up”. Pressurization can be done with diesel via a portable pump to avoid a 2” (small bore) line subject to rupture in HP lines. Larger vessels are left with hydrotest water or filled with diesel to NLL to minimise duration. Liquid piping may not require pressurisation line as a few drops of liquid can pressurize downstream piping

• Pressure Equalization: Reduces opening torque of large SDVs. Critical ESDVs actuators sized for full torque even with equalization. Smaller actuators for other ROVs. For a large manual valve, 2” line may have manual (LC/LO) valves to save money. When a downstream equipment cannot handle sudden flow, say a contactor or filter coalescer, provide pressure equalization regardless of size and pressure rating

• Pressure Equalization: 2” SDV in automatic start-up sequence as in compressor trains. It opens first and the main SDV opens after pre-set dP is reached. Then 2” SDV is auto-closed. Whenever the main SDV is closed, signal is also sent to the 2” SDV also to close it, though it remains closed

• Control Valve Bypass - Gate Vs Globe: Both throttle flow. Globe better in smaller sizes. Gate for large sizes. Butterfly in large gas and water lines. Globe valves were used in showers; now ¼ turn ball valves

• Control Valves: Quick opening type for shut-off service; not for modulating. Equal % type when valve ΔP < 30% of total ΔP. Linear type when bulk of the ΔP is at the valve

• Avoid bends downstream of LCV/ Steam/ Water Injection points. Potential source of rupture. Min 20D to bends. See Safety Alerts

• Check Valves: Avoid orphan check valves, i.e., without a downstream block valve. Maintenance of such check valves require downstream system isolation/ shutdown. Take spec break upstream of check valve

• Check Valves: Dissimilar or double or any number of check valves in series is not a LP/HP protection. Check valves reduce bulk backflow but leak and equalize pressure across. With incompressible liquids, all it takes is a few drops. Regardless of tightness specified, foulants - sand, coke, mill scale etc may not allow proper check valve seating. Even with 0.1% leak, pressure is equalized over time with upstream LP system. Upstream LP system requires relief protection if the downstream pressure is high

• Check Valves: API 521 abandoned 2 dissimilar check valves in series. “Complete check valve failure is assumed for all check valves in series that are not inspected and maintained and for a single check valve regardless of, if it is inspected and maintained”. Take spec break correctly

• Choked Flow: Choked flow issues continue to baffle process engineers. Too many queries on the subject and see wrong calculations/ conclusions regularly

• Choked flow at sonic velocity occurs at-the-throat when outlet pressure Po is <≈ 0.5*inlet pressure, Pi. (1) Further reduction in Po doesn’t increase mass flow. But mass flow increases proportionally with Pi. See API 520 PSV formula, W α APi (2) While throat velocity remains constant at sonic velocity, increasing Pi increases throat pressure, Pt and throat density, ρt. Wrong to say at sonic velocity, flow cannot increase further (3) Sonic velocity is at the throat and not in outlet piping. Velocity in downstream piping is decided by piping ID and Po. Wrong calculations are regularly made assuming sonic velocity in downstream piping, multiplying it by downstream piping area* downstream density

• Choked Flow: Occurs at (1) PSV or RO or CV throat (2) tail pipe to header and (3) flare tip

• Choked Flow: Orifice/ Control Valves. Flow α sqrt(Pi-Po). For choked flow, backpressure felt is not Po but throat pressure, Pt. Instrument Engineers use Pi-Pt in critical flow calculations, call it as critical drop, erroneously concluding that a RO/CV can drop Pi only to half of it downstream. This confusion has percolated to process engineers further confused by vendors. This confusion is only in RO or Control Valve sizing. Not in PSV sizing as API 520 has 2 separate formulae for critical and sub-critical flows. Next time you hear critical drop story, please point out - if you open LPG cylinder at 14 bar or N2 cylinder at 200 bar, atmospheric pressure doesn’t become 7 or 100 bar

• Choked Flow: In flow application, upstream can NOT decide downstream pressure. It is decided by datum pressure further down. Simple explanation. In a river flowing from Town A to Town B, Town A is at a higher elevation. Its water level is decided by flow and level in Town B. There is a level continuity. In a waterfall, upstream level does NOT decide downstream level for any flow. Downstream level has no influence on flow. So also, with choked flow with pressure discontinuity at throat

• Choked Flow: If velocity in a 6” is 70% sonic, in a 4” velocity will NOT be 70*(6/4)^2 = 160% sonic. Pressure in 4” will go up to increase density and maintain sonic velocity

• Choked Flow: PSV backpressure = flare tip ΔP + header ΔP + tail pipe ΔP or choked pressure at tail pipe / header junction. As long as backpressure is less than allowed, PSV capacity is not affected - irrespective of tail pipe/ header/tip velocity and ΔP. Keep velocity < 70% Mach (Some clients may want 50% in new installations) to avoid vibration and noise. You can’t avoid sonic velocity and choked flow at PSV nozzle or tail pipe exit or HP flare tip. Sonic at tip increases backpressure in upstream header, reduces volume flow and velocity; allows smaller headers and laterals. Before HP Flares in upstream Oil & Gas, API and industry practice was governed by refineries where highest backpressure was decided by Crude column top at 0.5 bar. API 521 past editions used to mention flare tip velocity to 20% sonic and never talked about sonic flares