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Lpg Export Pump Controls

lpg pump

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#1 shekhar dhuri

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Posted 11 December 2014 - 05:51 AM

Hi,

Refer Attachment for the Process flow scheme of the LPG storage and export system.
LPG produced from Gas Processing trains is stored in the horizontal mounded LPG bullet. Operating conditions of the LPG at the bullet is 14.3 barg @ 65 degC. These conditions are very close to its bubble point of the LPG. Inlet flow rate to the bullet depends on the number of train/s in operation since LPG produced from each train is 4 m3/h (600 bpd) which is 50% of the rated capacity of the pump (i.e. 8 m3/h). Dimensions of the bullet are 4.6 m ID X 26 m L and it is sized to store 2 days of LPG production from both trains (i.e. 1200 bpd).
On-off valve is provided in the pump suction line to enable batch operation during single train operation. Pump start is manual (either local or remote) and permitted once on-off valve open position is ensured.
In order to satisfy NPSHR of 2.5 m as specified by pump supplier in the performance curve, trim cooler has been provided in the pump suction. This option was found to be cost effective instead of the instead of elevating the LPG bullet to meet NPSH requirements. Cooling medium for the trim cooler is cooling water with inlet temperature of 35 degC. The trim cooler is designed with sufficient margin on surface area to ensure that LPG is remains completely sub-cooled in the pump suction piping. Cooling water inlet flow is controlled based on LPG temperature in the pump suction. LPG operating conditions at the pump suction are considered as 14.06 bag @ 55 degC (max temp). Two pumps are provided (1 operating + 1 stand by) each having rated capacity of the 8 m3/h.
During pump flow is maintain using discharge flow control valve (set at 8 m3/h). Protection of the pump assured by following:
• Low low pressure trip at the pump suction (set @ 13.5 barg)
• Low low liquid level trip in the LPG bullet
• High high temperature at the pump suction (set @ 60 degC)
• Automatic Recirculation Valve (ARV) is provided by pump supplier which is set @ 5.5 m3/h
• Pump discharge pressure over-ride (over flow) @ 58 barg is provided avoid pump running at the end of curve operation. However this over-ride can be deleted if pump motor has been sized for the end of curve operation. Please suggest.
LPG pressure at the battery limit of the CPF (Central Processing Facility) is taken as 70 barg and LPG is exported to the metering station via 45 km, 4” underground pipeline. Pressure control is provided downstream of the custody metering skids to maintain tie-in pressure of 60 barg at the transfer point.

Attached Files


Edited by shekhar dhuri, 13 December 2014 - 11:23 PM.


#2 Bobby Strain

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Posted 11 December 2014 - 10:15 AM

Is there a question?

 

Bobby



#3 Aneken

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Posted 12 December 2014 - 06:45 AM

In between,Can anyone please how can we change NPSH by trim cooler.?


Edited by Aneken, 12 December 2014 - 06:47 AM.


#4 shekhar dhuri

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Posted 13 December 2014 - 11:18 PM

Bobby,
I described Process Controls understood by me and I seek suggestions / confirmation from experts on this. Being, export pumps, I don't want make any design errors which might lead to the loss of product export.

Aneken,
LPG product coming from Gas Processing train is at saturated conditions therefore by using trim cooler, it is sub-cooled. As described in the note, due to sub-cooling by water, suction pressure became higher the (Vapour pressure - Line losses). Therefore, NPSH available has increased.
Using a trim cooler may not be the best design approach and some literature recommends, 'LPG Pumps to be installed in the Pit'. However, Pit is not acceptable by our Client. Therefore, we had to use trim cooler instead of increasing height of the LPG bullet. Please refer my previous posts on this Topic for info.

Regards,
Shekhar

#5 chemsac2

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Posted 14 December 2014 - 07:47 AM

Shekhar,

 

Not sure if low-low suction pressure trip would be of any help in detecting NPSHA issues for the pump as NPSHA is ensured by cooling the suction and there-by reducing vapour pressure. If cooling water fails, suction pressure would still be at its normal value, but pump would cavitate due to increased vapour pressure.

 

Also, TAHH of 60 deg C is also set at high a value for a mounded bullet and may not detect cooling water failure. Also, TAHH for winter and summer would have to be different even for a mounded bullet depending on site location.

 

PIC giving set-point to FIC is also not necessary for a pump with automatic re-circulation valve, FIC and PIC at tie-in. PIC would have no role as long as FIC is operational and would not help if FIC has failed. I do not think end-of-curve operation is possible. This can be further confirmed based on pressure drop of pipeline and full open control valves.  

 

PIC/PCV at tie-in point can be acceptable, but pressure drop of flow control valve shall be higher than that of PCV at tie-in point. 

 

These are few points I could think of from the scheme.

 

Regards,

 

Sachin



#6 Aneken

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Posted 15 December 2014 - 07:11 AM

Thanks Shekhar.



#7 shekhar dhuri

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Posted 16 December 2014 - 07:53 AM

chemsac2,

Thank you for your suggestions.

I agree with you that low-low suction pressure trip may not be required. So I will trip pump on Low low liquid level.

I would like to clarify that max temperature of the LPG coming from the Gas Plant is 65 degC and I have assumed no heat transfer from mound. So I think, TAHH set at 60 degC will ensure cooling water failure. During winter, temperature of the LPG coming to bullet may be lower but we have asked Gas Plant Supplier that to maintain constant temperature of 14.5 barg in summer and winter. In other words, LPG will be sub-cooled during winter and trim cooler will be redundant. I am not sure if changing set points during summer and winter is a good idea since I am not sure if operators will actually change set points.

I think ARV will ensure the pump protection during minimum flow and may not prevent the motor damage during end of curve operation (low head and high flow). During end of curve operation PCV will over-ride a flow control.

At battery limit, I have only considered, PIC and no FIC. Do you think it is required ?? Here line pressure drop is negligible and required pressure at tie in is 60 barg. Therefore, pressure drop of PCV at transfer point has been considered at 8-9 barg and Pressure drop via FCV is only 0.7 - 1 barg.

Does it make sense ??

Regards,
Shekhar Dhuri

Edited by shekhar dhuri, 16 December 2014 - 07:54 AM.


#8 SARTHAK

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Posted 19 December 2014 - 10:47 AM

Thank you all for Valuable Information. 

 

I work at LPG Bottling Plant where LPG is filled from Mounded Bullet to Cylinders by using Centrifugal Pump . Rated Flow is 50m3/hr for our Vertical Canned Pump and we are operating at 40-48 m3/hr flow rane ,mean our maximum requirement is increased from 43m3/hr to 48 m3/hr recently. But currently we are facing problem of discharge pressure fluctuation at Filling Point downstream discharge Piping to Carosal filling Cylinders . Rated Discharge is Max 17kg/cm2 . But there is pressure relief valve at discharge piping which is return piping Line connected directly to Mounded Bullet.

 

How one can maintain flow rate when there are shocks and Vibrations at Discharge of LPG Pump? What happens in our case as soon as flow rate reduces the Pressue at discharge increases from 15 to 17 kg and this actuates the PRV so for next round when flow is actually required it will be less and this hampers productivity of Bottled Cylinder?/

 

 

What should be done for maintaing constant Pressure at Pump discharge irespective of whether actually filling of LPG takes place or LPG Pump is idle , but still we want no fluctuation in pressue at all. Pressue Gauge installed at the end of discharge Pipe should always show constant pressure...

 

Please suggest any modifications??



#9 shekhar dhuri

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Posted 21 December 2014 - 01:13 AM

Sarthak,

If you provide some more information and sketch, providing suggestions will be easier for everyone.

What I could say after reading information provided by you is, discharge PSV setting is very close to the rated point. Generally, I prefer to design discharge piping for the Shut-off pressure of the pump (i.e. maximum pressure @ zero flow) if possible.

One way to avoid lifting of discharge PSV is increase in the set pressure (spring design). However, ensure that, all components in the pump circuit are designed for the new set pressure. For piping, check maximum allowable pressure as per ASME31.3. Once PSV set pressure is changed, then forward flow will be possible even at higher differential pressure conditions.




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