Chemical & Process Technology

Earlier post “Process Critical Line” has presented a checklist of process critical line. During design phase, these process shall be checked in detail to minimize or avoid problem such as vibration, hammering, capacity reduction, cavitation, etc to occur. This post will further present good engineering practice for process critical line.

Gravity Flow
Any line subject to gravity flow e.g. drain, flare, vent, etc, low pocket shall be avoided. Liquid or solid accumulate in low pocket potentially result corrosion and blockage. Line should be sloped (and/or free draining) from sources to receiver.

Pump Suction
Line to pump suction should be designed to allow self floating as far as possible where lowest liquid level is above the pump highest point.

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Ensure minimum submergence of tank to avoid potential vapor being sucked in to pump suction line due to vortex. If positive submergence is not achievable, installation of vortex breaker is another option. Read more in “Vortex Breaker to Avoid Vapor Entrainment”.

Any high pocket shall be avoided and provision of eccentric reducer at the pump suction to avoid potentially vapor lock prior to pump start-up.

Ensure NPSHa is always higher than NPSHr with a positive margin e.g. 1m for entire operation range (turndown to design capacity) and operation conditions (highest operating temperature). There are 17 Ways to Reduce Likelihood of Pump Cavitation.

Minimize suction length and fitting as much as possible to minimize potential of pump cavitation.

Compressor Suction
Compressor suction knock out drum (KOD) may be equipped with mist eliminator e.g. wiremesh to promote droplet coalescing and separation.

KOD vapor exit nozzle should be designed large enough to minimize exit momentum (rho V2 less than 6000 Pa) in order to minimize reentrainment of coalesced liquid droplet into vapor.

Compressor in general can tolerate small amount of liquid. If absolute no liquid is allowed enter compressor as imposed by compressor manufacturer, one may consider provision of insulation to minimize ambient and JT cooling and heat tracing to compensate heat loss due to above mentioned cooling.

Absolute no low pocket shall present in the compressor suction line as low pocket can accumulate liquid and slug of liquid can cause severe damage to compressor.

May consider a compressor suction strainer for start-up and commissioning. As compressor is sensitive to suction line pressure drop, any additional fitting and device at compressor suction can lead to capacity reduction, installation of suction strainer shall be analyzed in detail during design phase.




Flashing / Two phase Gas-Liquid Flow
Slugging and plugging flow in vertical and horizontal potentially results significant vibration to piping. During process design phase, slugging and plugging flow shall be avoided for entire operating range (turndown to design capacity) and operating conditions.

May consider provision of vapor liquid separation and run separator separate header for vapor and liquid line if slugging / plugging flow is unavoidable. For steam header, provide sufficient steam traps to drain-off condensate and minimize potential of slugging flow.

Extra and strengthen support may be provided to avoid severe vibration and failure on pipe crack.

Liquid-Liquid Coalescer
Vapor generation in liquid-liquid coalescer may accumulate and result under-performed liquid-liquid separation. May consider to provide sufficient static head to suppress vapor generation in liquid-liquid coalescer. It is always recommended to provide a vapor equalization line back to separator to release any vapor form in liquid-liquid coalescer.

Low Pressure Line
Minimizing pressure drop in low pressure line is the key factor to ensure proper performance of system. Minimize line length, fittings, elbow, etc and use of smooth surface pipe e.g. stainless steel may be considered.

Potential Surge Line
Steam supply line experience heat loss and condensation due to partially damaged insulation and extreme low ambient temperature. Flashing condensate with steam return to collection header mix with cold condensate. Both condition would results sudden steam collapse and lead to implosion. Steam implosion would generate severe movement of condensate in the collection header and severe vibration of header. Therefore proper maintenance of insulation is extremely important in keep steam line from transient surge. Besides, provide sufficient steam trap to eliminate condensate from steam line.

Long pipeline transferring incompressible fluid e.g. LNG rundown line, produce water injection line, etc potentially experience transient surge (water hammer) in the event of closure of shutdown valve. Transient surge analysis shall be conducted during design phase to ensure surge is avoided. Slower closure of shutdown valve is one of the key component in minimizing surge in long pipe line. Non-slam check valve on the pump discharge may also assist in minimizing surge in long pipeline with pump. Surge suppression system may be considered in the event surge is unavoidable. One shall take note that provision of pressure relief valve may not help to eliminating surge due to slow response time of PRV.



Pressure Relief Valve Inlet & Outlet
May consider discussion and recommendation in :

• Concerns & Recommendations on PSV INLET line 
• Few Concerns & Recommendations of PSV Discharge Tail pipe

Control Valve & Restriction Orifice
Flow Induced Vibration (FIV) and Acoustic Induced Vibration (AIV) may be studied to identify location of piping which potentially experience high risk of low frequency and high frequency vibration. Minimizing small bore connection may be considered e.g. provision connection with more than 2 inches, avoid using connection smaller than 2 inches. For small bore connection, may consider brazing and extra support to strengthen the connection and avoid pipe cracking.

Anti-cavitation trim could be considered for control valve potentially experience cavitation. Similarly provision of multiple restriction orifice (RO) in series or multi-ported RO may be considered if cavitation occurs in RO.

Pressure and measurement can be extremely complex and complicated. However, for most systems it is relatively easy to obtain accurate pressure measurements if the proper techniques are used. 

What is fluid pressure ? Fluid pressure can be understood as the measure of force per-unit-area exerted by a fluid, acting perpendicularly to any surface it contacts (a fluid can be either a gas or liquid, fluid and liquid are not synonymous). The standard SI unit for pressure measurement is the Pascal (Pa) which is equivalent to one newton per square meter (N/m2) or the KiloPascal (kPa) where 1 kPa = 1000 Pa.......

Process design involve line sizing and pressure profile definition. All line size will be presented in Piping & Instrumentation Diagram (P&ID). Nevertheless, there is no line length, elbow  and elevation define in P&ID. Upon receipt of P&ID, Piping engineer will begin the piping routing activities and assign necessary length, elbow and elevation to the line. This piping routing may not consistent with assumption taken by process engineer during earlier process design. Significant increase in pressure drop, wrong routing of pipe , incorrect sloping, etc could lead to severe vibration, valve chartering, reduced capacity, under-perform equipment, etc. Therefore it is important for a process engineer to identify Process Critical Line for detail isometric checking. Following will tabulate typical line may experience potential problem and required  detail process checking.

In general most fluid is transferred by pressure from source to destination during normal operation. Pressure head available at source will overcome frictional loss, velocity head and static head. This allow fluid transfer from low point to elevated point. Typical example is transfer liquid from closed drain drum to production separator with the pressure head develop by a reciprocating pump. This kind of pipe is typically know as pressurized pipe. Nevertheless, there are some fluid is transferred by gravity force (or static head). Typical system is closed drain network, process line designed for gravity transfer, etc. Improper design of gravity flow would lead to reduce or no flow.

Cavitation is phenomenon cause by bubble generation follow by bubble collapsed. More thorough discussion on cavitation phenomenon, cavitation damages and the way to minimize / avoid cavitation can be found in following post :

Typically to minimize / avoid cavitation damage is to ensure Net Positive Suction Head required (NPSHr) by the pump is lower than the NPSH available (NPSHa) by the system itself. Pump suction line size and routing is a dominant factor affects NPSHa. Improper design of pump suction line would lead to severe cavitation, vibration and pump damage.

Centrifugal Compressor capacity is subject to designated flow and compressor inlet condition. Any changes in suction condition (e.g. decrease in density) would seriously affect compressor capacity (e.g. decrease in capacity). Improper design of line between Compressor suction Knock-out drum (KOD) and compressor inlet nozzle would lead to high pressure drop, subsequently lower density and capacity decrease. 

Long compressor line (from KOD to compressor) would increase potential of heat loss to ambient (severe during winter time) and results condensation. Present of condensate in vapor to compressor and impinge on compressor impeller when vapor is accelerated potentially damage compressor impeller and severe vibration in compressor.

Flare / vent collection header has significant impact on built-up backpressure to pressure relief valve. (PRV) Severe pressure drop can lead built-up back pressure exceed it allowable limit e.g. 10% for conventional type PRV. Warm fluid mix with cold fluid in flare header may results two phase gas liquid flow in flare header. Similarly, severe flare header vibration can occur with the present of slugging / plugging flow.  Low point in flare line potentially results liquid accumulation in flare line and corrosion may occur. In the major relief event, high velocity vapor pushing accumulated liquid would results slugging flow in the flare line. Liquid column flowing at vapor velocity knocking of elbow/bend may generate severe vibration.

Liquid at saturation point coming from separator potentially flash and two phase gas liquid flow. Typical flow regime is Bubbly flow. Similarly saturated vapor experience ambient cooling and line frictional loss results condensation and two phase gas liquid flow. Typical flow regime is Mist flow. Both Bubbly and Mist flow are not destructive in nature and properly a normal support would be sufficient. Nevertheless, slugging and plugging flow in vertical and horizontal potential results significant vibration to piping. Extra and strengthen support is required to avoid severe vibration and failure on pipe crack. More discussion on Problems Caused by Two Phase Gas-Liquid Flow.

Saturated liquid from separator feeding liquid-liquid separator, any pressure drop increase potentially lead to vapor accumulation and  under-performed liquid-liquid separation.

Low pressure stream e.g. overhead from amine regeneration column, end flash gas from end flash column, etc is very sensitive to frictional loss.Low pressure here is pressure very close to atmospheric pressure. Any increase in frictional loss will seriously reduce flow through the pipe.

Steam supply line experience heat loss and condensation due to partially damaged insulation and extreme low ambient temperature. Flashing condensate with steam return to collection header mix with cold condensate. Both condition would results sudden steam collapse and lead to implosion. Steam implosion would generate severe movement of condensate in the collection header and severe vibration of header. Long pipeline transferring incompressible fluid e.g. LNG rundown line, produce water injection line, etc potentially experience transient surge (water hammer) in the event of closure of shutdown valve. Piping surge is severe in nature and potentially lead to pipe crack and support failure.

Some line is normally flow with vapor contains CO2 & H2S and sulfide stress corrosion cracking (SSCC) and general CO2 corrosion is not expected as only vapor flow. During winter low ambient temperature and under turndown operation, ambient cooling potentially lead to vapor condensation and induced SSCC and general corrosion on under-designed piping. Typical example is Condensate stabilizer overhead. Similarly warm wet flare header is normally dry due to continuous dry gas purging. In the event, PRV passing leaks wet vapor into warm wet header or any PRV open follow by closure of PRVs, wet vapor potentially condensed and accumulate in low point and results general corrosion.

Line normally design for low pressure drop, any increase in pressure drop could to capacity reduction and/or under-perform downstream unit. Typical example is high pressure gas feeding liquefaction Main Cryogenic Heat Exchanger. Any reduction in Feed pressure to MCHE would lead to higher heat of vaporization and reduce LNG production.

Under normal design condition, PRV inlet line non-recoverable pressure loss shall be limited to 3% of PRV set pressure  Any significant increase in line length and elbow (due to piping routing) will results non-recoverable pressure loss increase and lead to PRV chattering. 

Upon opening of PRV, instantaneous large gas or vapor passing PRV. High frequency noise is generated results acoustic induced vibration (AIV)  which potentially cause discharge pipe cracking. Instantaneous large gas/vapor flow accelerated from zero velocity to maximum velocity will induced high reaction force to downstream piping. Under-designed pipe may crack on high reaction force.

Saturated liquid passing a control valve or restriction orifice, pressure will began to decrease and lowest pressure closed to vena contracta, follow by pressure recovery once is passed the vena contracta. Lowest pressure point could be lower than vapor pressure of fluid. Vapor bubble will begin to form and once fluid passed through the vena contracta, vapor will start to collapse and results jet wave impacting control valve or restriction or piping. Above phenomenon generally known as cavitation which generate severe vibration to the piping.

Any scenario is normally ignore or miss by engineer where control valve downstream piping may not design for occurrence sonic flow downstream piping. This typically occur in line with control valve discharge to flare/vent header. Sonic flow potentially reduce flowing capacity and how reaction force to piping.

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Carbon Dioxide (CO2) is known as one of industrial waste gas causing global warming...there are number of measures are implemented to reduce global warming. CO2 reduction measures are CO2 capturing, reinjection into aquifer, injection into reservior for well maintenance, gas injection for enhanced oil recovery (EOR), etc. CO2 being captured will be injected in the reservior. The injection can be as high as 300-400 atm which is higher than the critical pressure of CO2. Carbon dioxide is in its supercritical fluid state when both the temperature and pressure equal or exceed the critical point of 31°C and 73 atm (see diagram). In its supercritical state, CO2 has both gas-like and liquid-like qualities, and it is this dual characteristic of supercritical fluids that provides the ideal conditions for extracting compounds with a high degree of recovery in a short period of time. Read more in below...

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Pressurized liquid or vapor-liquid in equilibrium or vapor only system during normal operation, when it is expose to external fire attack, heat inputs into vessel (or system) may possibly increase internal pressure and temperature. For system with high design pressure (or Maximum allowable working pressure, MAWP), the pressure relief valve (PRV) protecting the vessel (or system) may have same set pressure as the design pressure (or MAWP). Subject to design code of the vessel, overpressure allowed by code is different from code to code. Typically for ASME unfired vessel section VIII, the maximum allowable overpressure is 10% of set pressure. This will results maximum allowable accumulation pressure (or relieving pressure) reach at 110% of set pressure.


In some event, the relieving pressure is higher than the fluid critical pressure. For example, a CO2 injection compressor, the injection pressure at is 65 barg. The design pressure may be 72 barg and relieving pressure is approximately 79.2 barg, is higher than CO2 critical pressure of 72.9 barg. The PRV is relieving at supercritical condition.

Conventional method in determining PRV orifice size is presented in API RP-521. Part 1. It considered all “un-wetted” vessels are same regardless the fluid is supercritical, a vapor or a gas. Nevertheless, one shall take note that  this method based on the physical properties of air and the perfect gas laws with no change in fluid temperature. 

API method may be conservative, there are chattering and oversized PRV and discharge problem. A  rigorous method has been discussed by R.C. DOANE in "Designing for pressure safety valves in supercritical service" published in Hydrocarbon Processing Jan 2010. This method assuming all thermodynamic paths are well defined by a Process Simulator. Thermodynamic path of fluid from operating condition to relieving and drop to back pressure to PRV may be defined by four typical steps.

For Step 1, earlier post "Constant Density To Obtain Relieving Condition" has discussed similar subject previously. 

For the step 3, another definition of PRV relieving path can be Isentalpic from relieving to throat follow by isentropic from throat to PRV backpressure as discussed in "Discussion on ISENTROPIC and ISENTHALPIC process via Relief Valve".

In the "Designing for pressure safety valves in supercritical service" article, table 1 "Supercritical relief valve sizing example problem—normal butane" has tabulated step-by-step calculation. This table has incorporated equation 1 to equation 10 in this article. The calculation consist of segment 1-to-2, 2-to-3, 3-to-4 and 4-to-5. In recent work in establishing similar task carried out by this example, one suspicious discrepancy is identified. Details as follow.

From segment 1-to-2 to segment 4-to-5, volumetric flow is increased from 1304 ft3/hr to 1315 ft3/hr (segment 2-to-3) and decreased in subsequent segments to 1250 ft3/hr. The volumetric flow is in the range of 1250 to 1315 ft3/hr. HOWEVER, mass flow is increased almost double from 6,374 lb/hr to 13,405 lb/hr (segment 2-to-3) and decreased in similar range (11620 to 12,371 lb/hr). Why there is significant different in segment 1-to-2 compare to other segment ?

Detail checking found that Mass Flux (G in lb/ft2s) is calculated by dividing Orifice Velocity (v) by Specific Volume (V) at orifice condition for segment 1-to-2. On the other hand, Mass Flux (G in lb/ft2s) is calculated by dividing Orifice Velocity (V) by Specific Volume (V) at outlet condition for other segments. Different method in calculating Mass Flux has created the discrepancy.

Opinion

As the flow is choked (or highest velocity for subsonic) at the PRV nozzle,  Mass Flux (G in lb/ft2s) should be calculated by dividing Orifice Velocity (v) by Specific Volume (V) at orifice condition.

Well... Above query will be raised and response will be posted once we have received it.

Related Topics

• Tedious & Simple Method In Determining Specific volume For Isentropic Nozzle Flow Mass Flux
• Constant Density To Obtain Relieving Condition
• Discussion on ISENTROPIC and ISENTHALPIC process via Relief Valve
• API Std 520 Part 1 Dec 2008 is Released
• API Std 521 ADDENDUM, MAY 2008 - Check Out Revised Section
• Requirement of overpressure protection devices on system design to PIPING code